Carbon Dioxide Utilization for Global Sustainability

Studies in Surface Science and Catalysis Advisory Editors: B. Delmon and J.T. Yates Series Editor: G. Centi Vol. 153 C...

Author: Sang-Eon Park | Jong-San Chang | Kyu-Wan Lee

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Studies in Surface Science and Catalysis Advisory Editors: B. Delmon and J.T. Yates Series Editor: G. Centi

Vol. 153

CARBON DIOXIDE UTILIZATION FOR GLOBAL SUSTAINABILITY Proceedings of the 7 th International Conference on Carbon Dioxide Utilization, Seoul, Korea, 12-16 October, 2003

Edited by Sang-Eon Park Department of Chemistry Inha University Incheon, Korea Jong-San Chang Research Center for Nanocatalysis KRICT, Taejon, Korea Kyu-Wan Lee Advanced Chemical Technology Division KRICT, Taejon, Korea

2004

ELSEVIER Amsterdam - Boston - Heidelberg - London - New York - Oxford - Paris - San Diego San Francisco - Singapore - Sydney - Tokyo

Studies in Surface Science and Catalysis 153 CARBON DIOXIDE UTILIZATION FOR GLOBAL SUSTAINABILITY

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V

CONTENTS

Preface Chairman's Remark

xiii xv

Organizing committee

xvii

International advisory board

xvii

International scientific committee Local advisory board

xvii xviii

Local organizing committee

xviii

Supporting organizations

xviii

Supporting organizations for technical excursion

xix

I. Heterogeneous catalysis of CO 2 Heterogeneous catalytic reactions with CO 2 : Status and perspectives G. Centi and S. Perathoner

1

Chemicals from CO 2 via heterogeneous catalysis at moderate conditions Y. Sun

9

Synthetic hydrocarbon fuels and CO2 utilization G. Schaub, D. Unruh and M. Rohde

17

Design of copper based hybrid catalysts for CO2 hydrogenation Y.-K. Park, J.-K. Jeon andS.-K. lhm

25

Plasma enhanced preparation of highly dispersed Ni/ALO^ catalyst for CO 2 reforming of Methane D. Cheng, Q. Xia. and C. Liu

33

Synthesis of dimethyl carbonate by transesterification over CaO/carbon solid base catalysts T. Wei, M. Wang, W. Wei, Y. Sun and B. Zhong

41

The synthesis of clean fuels by F-T reaction from CO 2 rich biosyngas J.-S. Ryu, K.-W. Lee, M.-J. Choi and H.-S. Yoo

47

Reduction of carbon dioxide using metal powders H. Katsumata, K. Matsushita, S. Kaneco, T. Suzuki and K. Ohta

55

Catalytic hydrogenation of CO 2 to methanol over Pd/ZnO: Metal support interaction C.-H. Kim, J.S. Lee andD.L. Trimm

61

CAMERE Process for methanol synthesis from CO 2 hydrogenation O.-5. Joo, K.-D. Jung and Y. Jung

67

vi

Product distribution analysis for catalytic reduction of CO 2 in a bench scale fixed bed reactor S.-B. Lee, J.-S. Kim, W.-Y. Lee, K.-W. Lee, and M.-J. Choi

73

Process evaluation of biomass to liquid fuel production system with gasification and liquid fuel synthesis T. Minowa, T. Hanaoka and S. Yokoyama

79

Novel catalysts for gasification of biomass with high energy efficiency T. Miyazawa, M. Asadullah, M. Nurunnabi, B. Li, K. Kunimori and K. Tomishige

85

Improving carbon utilization in biomass conversion to synthetic hydrocarbons via Fischer-Tropsch synthesis D. Unruh, M. Rohde and G Schaub

91

Fischer-Tropsch synthesis with CO 2 -containing syngas from biomass - Kinetic analysis of fixed bed reactor model experiments M. Rohde, D. Unruh, P. Pias, K.-W. Lee and G. Schaub

97

Hydrogen production from woody biomass by novel gasification using CO 2 sorbent T. Hanaoka, S. Fujimoto, T. Yoshida, K. Kamei, M. Harada, Y. Suzuki, S. Yokoyama and T. Minowa

103

A feasibility study of synthesis of oxygenates directly from methane and carbon dioxide using dielectric-barrier discharges J. Wang, C. Liu and B. Eliassion

109

Synthesis gas production from CO 2 and H2O with nonthermal plasma S. Futamura and H. Kabashima

119

CO 2 reduction by blast furnace top gas recycling combined with waste hydrocarbon gasification J.-O. Choi and J.-G. Kim

125

Microwave—assisted reactions of oxiranes with carbon dioxide in ionic liquids M. Nuchter, B. Ondruschka, J. Hoffmann and A. Tied

131

CO 2 reforming of n-heptane on a Ni/Al2O^ catalyst K.J. Puolakka and A.O.I. Krause

137

Hexaaluminate catalysts of the novel process of syngas production through catalytic oxidation and steam-CO 2 reforming of methane O. Yokota, T. Tanaka, Z. Hou and T. Yashima

141

Redox behavior of Cu-ferrite for CO 2 decomposition H.C. Shin, S.C. Choi, C. S. Kim, O.-5. Joo and K.-D. Jung

145

CO 2 reforming by CH 4 over Ni-YSZ modified catalysts D.J. Moon, J.W. Ryu, D.M. Kang, B.G. Lee andB.S. Ahn

149

The interaction between CO 2 and CH 4 on Ru-Co-catalysts L.B. Shapovalova, CD. Zakumbaeva, I.A. Shlygina andA.A.

153 Shyrtbaeva

Selective formation of light olefins by the cracking of heavy naphtha over acid catalysts S.X Han, C.W. Lee, J.R. Kim, N.S. Han, W.C. Choi, C.-H. Shin and Y.-K. Park

157

Effect of additives and a preparation method on catalytic activity of Cu/ZnO/ZrO 2 system in the carbon dioxide hydrogenation to methanol J. Stoczynski, R. Grabowski, A. Koztowska, M. Lachowska and J. Skrzypek

161

vii

Selective formation of ethylene carbonate from ethylene glycol and carbon dioxide over CeO 2 -ZrO 2 solid solution catalysts K. Tomishige, H. Yasuda, M. Nurunnabi, B. Li and K. Kunimori

165

Effects of palladium particle size in hydrogenation of carbon dioxide to methanol over Pd/ZnO catalysts D. Lee, J.-Y. Lee and J.S. Lee

169

Ga, Mn and Mg promoted copper/zinc/zirconia - catalysts for hydrogenation of carbon dioxide to methanol M. Lachowska and J. Skrzypek

173

Catalytic hydrogenation of carbon dioxide to light olefins in a fluidized bed reactor J.S. Kim, S.-B. Lee, M.-J. Choi, Y.K. and K.-W. Lee

177

Synthesis of dimethyl carbonate from CH3OH and CO 2 with Cei_xZrxO2 catalysts G.B. Han, N.K. Park, J.H. Jun, W.C.Chang, B.G. Lee, B.S. Ahn, S.O. Ryu and T.J. Lee

181

The effect of catalyst pore structure change into the selectivity and conversion of CO 2 hydrogenation over Fe-K/y-AI2O3 S.-C. Lee, J.-H. Jang, B.-Y. Lee, T. Jeong andS.-J. Choung

185

Tri-reforming of CH 4 using CO 2 for production of synthesis gas to dimethyl ether S.-H. Lee, W. Cho, W.-S. Ju, Y.-S. Baek, J.S. Chang and S.-E. Park

189

Carbon dioxide reduction technology with SOFC system D.J. Moon, J.W. Ryu and S.D. Lee

193

Synthesis of dimethyl carbonate from urea and methanol over metal oxides M. Wang, N. Zhao, W. Wei and Y. Sun

197

Development of heterogeneous catalyst system for esterification of free fatty acid contained in used vegetable oil H.-J. Kim, B.S. Kang, M.-J. Kim, D.-K. Kim, J.S. Lee and K.-Y. Lee

201

Investigation of synthesis gas production from natural gas and CO 2 by microwave plasma technology W. Cho, W.S. Ju, S.-H. Lee, Y.-S. Baek and Y.C. Kim

205

CO 2 hydrate kinetics in electrolyte solutions containing clay minerals J. Lee, J. Park, M.-K. Chun, K. Lee, J.-H. Lee, J.-M. Kang and H. Lee

209

II. Homogemeous catalysis of CO2 New catalysts for the conversion of urea into carbamates and carbonates with C1 and C2 alcohols M. Aresta, A. Dibenedetto, C. Devita, O.A. Bourova and O.N. Chupakhin

213

Group 5 (V, Nb and Ta) clcmcnt-alkoxidcs as catalysts in the trans-cstcrification of ethylenecarbonate with methanol, ethanol and allyl alcohol M. Aresta, A. Dibenedetto and C. Pastore

221

Hydroformylation with carbon dioxide using ionic liquid media K. Tominaga and Y. Sasaki

227

Ionic liquid-derived imidazolium metal halides for the coupling reaction of ethylene oxide and CO 2 O.S. Kwon, J. Palgunadi, H. Lee, J.Y. Bae,B.S. Ahn, H.S. Kim and S.-H. Lee

233

viii

Aliphatic polycarbonate synthesis by alternating copolymerization of carbon dioxide with cyclohexene oxide using heterogeneous zinc complex /. Kim, M.J. Yi, S.H. Pyun, D.W. Park and C.-S. Ha

239

Alternating copolymerization of carbon dioxide and epoxide [1] Aluminum Schiff base complex - quartenary ammonium salt systems as novel initiators H. Sugimoto, H. Ohtsuka and S. Inoue

243

Alternating copolymerization of carbon dioxide and epoxide [2] The first example of polycarbonate synthesis from 1-atm carbon dioxide by manganese porphyrin H. Sugimoto, H. Ohshima and S. Inoue

247

Semifluorinated block copolymer surfactants for \vater-in-CO2 microemulsions K.T. him, H.J. Kim, S.H. Jin, S.-J. Choi and Y.S. Gal

251

Aliphatic polycarbonate synthesis by alternating copolymerization of carbon dioxide with cyclohexene oxide using /2-diiminate zinc complex S.M. Kim, I. Kim, D.W. Park and C.-S. Ha

255

Synthesis of poly(DOMA-co-AN) by addition of carbon dioxide to poly(GMA-co-AN) and the miscibility behavior of its blends with PEI J.-H. Hur, D.-O. Lim, K.-H. Kim, I. Kim, H.-C. Woo andD.-W. Park

259

Half-sandwich complexes with dihydroxy polypyridine: water-soluble, highly efficient catalysts for hydrogenation of bicarbonate attributable to electron-donating ability of oxyanion on catalyst ligand Y. Himeda, N. Onozawa-Komatsuzaki, H. Sugihara, H. Arakawa and K. Kasuga

263

Synthesis of propylene carbonate from carbon dioxide and propylene oxide using ionic liquids D.-W. Park, J.-H. Hur, E.-S. Jeong, S.-W. Park and I. Kim

267

III. Electro and photocatalytic reduction of CO2 Spectroscopic characterization of intermediates in CO 2 reduction with rhenium photocatalysts E. Fujita, Y. Hayashi, S. Kita and B.S. Brunschwig

271

Electrochemical reduction of CO 2 at alloy electrode in methanol S. Kaneco, H. Yamauchi, H. Katsumata, T. Suzuki and K. Ohta

277

Photoactivation of Ti centers in mesoporous silicate sieve under visible and UV light W. Lin and H. Frei

283

Photocatalytic reduction of CO 2 with H2O on Ti-containing mesoporous silica hydrophobically modified using fluoride ions H. Yamashita, M. Okazaki, K. Ikeue and M. Anpo

289

Effect of CO? concentration on growth and photosynthesis of spirulina platensis S.-G Kim, C.-S. Park, Y.-H. Park, S.-T. Lee and H.-M. Oh

295

High performance photocatalytic reduction of CO 2 with H2O by TiSBA-15 mesoporous material J. -S. Hwang, J. -S. Chang, S. -E. Park, K. Ikeue and M. Anpo

299

ix

IV. CO2 as oxidant New CO 2 chemistry - recent advances in utilizing CO 2 as an oxidant and current understanding on its role S.-E. Park and J.S. Yoo

303

Tri-reforming of methane over Ni catalysts for CO 2 conversion to syngas with desired H2/CO ratios using flue gas of power plants without CO 2 separation C. Song, W. Pan, S.T. Srimat, J. Zheng, Y. Li, Y.-H. Wang, B.-Q. Xu and Q.-M. Zhu

315

CO 2 dehydrogenation of propane over Cr-MCM-41 catalyst K. Takehira, Y. Oishi, T. Shishido, T. Kawabata, K. Takaki, Q. Zhang and Y. Wang

323

CO 2 reforming of n-heptane on a Ni/Al 2 O, catalyst K.J. Puolakka and A.O.I. Krause

329

Dehydrogenation of ethylbenzene over promoted vanadium oxide catalyst with carbon dioxide: from laboratory to bench-scale test D.-Y. Hong, S.H. Jhung, J.-M. Lee, M.S. Park, J.S. Chang and S.-E. Park

333

Oxidative dehydrogenation of ethane with carbon dioxide over supported chromium oxide catalysts M. Ji, D.-Y. Hong, J.S. Chang, V.P. Vislovskiy, J.S. Yoo and S.-E. Park

339

Oxidehydrogenation of ethane with CO 2 over transition metal doped MCM-41 mesoporous catalysts Y.-L Bi, V.C. Corberdn, H. Zhuang and K.-J. Then

343

Selective formation of styrene via oxidative dehydrogenation of 4-vinylcyclohexene over ZrO 2 supported iron oxide catalysts J.S. Chang, D.-Y. Hong, Y.-K. Park and S.-E. Park

347

Catalytic transformation of cyclohexanol over mixed metal oxides with and without CO 2 V.R. Chumbhale, J.S. Kim, S.B. Lee andM.J. Choi

351

V. Studies in supercritical CO2 The utility of carbon dioxide in homogeneously-catalyzed organic synthesis P.G. Jessop

355

Aldol reactions of propanal using MgO catalyst in supercritical CO 2 K. Matsui, H. Kawanami and H. Hayashi

363

Critical point and phase envelope calculations - Some practical aspects based on CO 2 -mixtures H.A. Turunen, J.-P. Pokki, R. Raudaskoski, E. Turpeinen, E. Muurinen, K.I. Keskinen and R.L. Keiski

369

Evaluation of sequesterable carbon dioxide in Japanese coal samples at sub-critical and supercritical conditions M. M. Toribio, Y. Oshima and S. Shimada

375

Synthesis and characterization of poly(styrene-cobutyl methacrylate) in supercritical carbon dioxide B. Shin, W. Bae, H. Lee and H. Kim

381

acrylate) and poly(styrene-cobutyl

Phenol hydrogenation over supported metal catalysts under supercritical carbon dioxide C. V. Rode, U.D. Joshi, T. Sato, O. Sato and M. Shirai

385

Synthesis of various block copolymers containing poly( L -lactide) in supercritical carbon dioxide H.S. Hwang, E.J. Park, Y.T. Jeong, H. Heo and K.T. Lim

389

X

Vapour pressure of CO 2 as well as phase envelopes and critical points for CO 2 and CH 4 mixtures calculated by Aspen Plus and Flowbat simulation software R. Raudaskoski, E. Turpeinen, H. Turunen, K. Rahkamaa-Tolonen and R.L. Keiski

393

VI. CO2 separation and recovery CO 2 recovery pilot plant K.R. Jang, J.H. Kim, J.G. Shim, Y.M. Ahn and H.M. Earn

397

A 2-stage PSA process for the recovery of CO 2 from flue gas and its power consumption S.-H. Cho, J.-H. Park, H.-T. Beam, S.-S. Han andJ.-N. Kim

405

Novel nanoporous "Molecular basket" adsorbent for CO 2 capture C. Song, X. Xu, J.M. Andresen, B.G. Miller and A. W. Scaroni

411

Reversible adsorption of carbon dioxide on amine-modified SBA-15 from flue gas containing water vapor N. Hiyoshi, K. Yogo and T. Yashima

417

Hollow fiber membrane contactor hybrid system for CO 2 recovery K.-H. Lee, S.-H. Yeon, B. Sea and Y.-I. Park

423

Optimization of CO 2 absorption process with MEA solution S. Kim, H.-T. Kim and B.C. Choi

429

Permeability modification for CO 2 EOR G. Turn, C. Tao, R. Qu, W. Han, J. Deng and P. Behrenbruch

435

Catalytic reactivity for the formation of dimethyl ether from synthesis gas over hybrid catalysts W.-S. Ju, C.W. Choi, S.-H. Lee, W. Cho, J.-S. Hwang, S.-E. Park and Y.-S. Baek

441

VII. Biology and biochemistry of CO2 utilization Design and experimental practice of photobioreactor incorporating microalgae for efficient photosynthetic CO 2 fixation performance Y. Watanabe

445

Biomass Japan strategy S. Yokoyama

453

Carbon dioxide removal and biomass production by Anabaena cylindrica I AM Ml using reactor in series A. Wijanarko and K. Ohtaguchi

461

Viability and activity of microorganisms in pressurized and supercritical carbon dioxide B. Zhang, X. Liu and Y. Gan

469

Novel conversion of carbon dioxide to methanol catalyzed by sol-gel immobilized dehydrogenases Z. Jiang, S. Xu and H. Wu

475

Combined thermoneutral processes for CO 2 emission avoidance and fuel saving in the metallurgical and lime industries M. Halmann and A. Steinfeld

481

xi

VIII. Utilization of CO 2 hydrates Greenhouse gas hydrates in the ocean S.M. Masutani

487

Replacement of methane hydrate by carbon dioxide: I3C NMR study for studying a limit to the degree of substitution H. Lee, Y. Seo, Y.-T. Seo, D.Y. Kim, I.L. MoudrakovskiandJ.A. Ripmeester

495

Continuous CO 2 clathrate hydrate formation by a motionless mixer H. Tajima, A. Yamasaki and F. Kiyono

501

A new type separation process of condensable greenhouse gases by the formation of clathrate hydrates K. Shiojiri, T. Okano, Y. Yanagisawa, M. Fujii, A. Yamasaki, H. Tajima and F. Kiyono

507

The OACE (Ocean Abyssal Carbon Experiment) research project R. Kojima, I. Aya, K. Yamane, S. Namie, Y. Nakajima, H. Shirota, P. G. Brewer, E.T. Peltzer, P. M. Haugan, T. Johannessen and R.G.J. Bellerby

513

Addition of carbon dioxide to phenylglycidyl ether using quaternary ammonium salts grafted on modified MCM-41 support D.-W. Park, D.-H. Shin, B.-S. Yu, K.-H. Kim, S.-W. Park, I. Kim andJ.-W. Lee

519

Chemical absorption of carbon dioxide with NaOH in non-Newtonian w/o emulsion S.-W. Park, B.-S. Choi, T.-Y. KimandJ.-W. Lee

523

CO 2 absorption and regeneration using Na and K based sorbents S.C. Lee, B.Y. Choi, SJ. Lee, S.Y.Jung, C.K. Ryu andJ.C. Kim

527

Preparation of 6FDA-based polyimide membranes for CO 2 gas separation K.-J. Kim, W.-W. SoandS.-J. Moon

531

Chemical absorption of carbon dioxide into toluene solution of oxirane S.-W. Park, D.-W. Park, T.-Y. Kim and J.-W. Lee

535

A study about CO 2 absorption process design for thermal power plant flue gas S.-Y. Park, B.-M. Min, J.-S. Lee and H.-M. Eum

539

Membrane/PSA hybrid process for carbon dioxide recovery at low concentration J.-K. Jeon, S.-K. Ihm, Y.-K. Park, J.S. Kim, J.-I. Dong, S. Kim, J.M. Kim, 5.-5. Kim and K.-S. Yoo

543

Effect of isotherm selection on performance prediction of CO? PSA process J.-K. Jeon, S.-K. Ihm, Y.-K. Park, J.S. Kim, S.D. Kim, S. Kim, J.M. Kim, 5.-5. Kim and K.-S. Yoo

547

Chemical absorption of carbon dioxide with AMP in aqueous PAA solution S.-W. Park, B.-S. Choi, H.-I. Jung and J.-W. Lee

551

IX. General uses of CO 2 The role of CO 2 fixation in the strategy for mitigating global warming Y. Kara

555

xii

The preparation of metallic tungsten and flower-like structure tungsten materials with using CO/CO 2 mixtures J.-D. Lee, D.H. Lee, N.K. Park, 5.0. Ryu, W.C. Chang and T.J. Lee

561

On-line gas chromatographic analysis of CO 2 hydrogenation products M. Reinikainen and M. Niemela

565

Synthesis of ultrafine TiO2 particles from hydrolysis of Ti(O'Pr)4 with PEO-b-PFOMA reverse micelles in CO 2 K.T. Lim, H.S. Hwang, S.-S. Hong, C. Park, W. Ryoo and K.P. Johnston

569

Synthesis of propylene carbonate from urea and propylene glycol Q. Li, N. Zhao, W. Wei and Y. Sun

573

Dissolution behavior of liquid CO 2 droplets injected at the intermediate ocean depth in ClarionClipperton

577

C.-B. KimandN.-J. Kim Growth and carbon fixation of cyanobacterium spirulina platensis with different nitrogen sources A. Choi, B.-D. Yoon, H.S. Kim andH.-M. Oh

581

Author index

585

Keyword index

591

Other volumes in the series

597

xiii

PREFACE The global environmental problems, especially the global warming caused by the accelerative accumulation of carbon dioxide in the atmosphere, are nowadays the most crucial for human beings as well as global sustainability. Due to this global environmental issue, the scientists and engineers, who had a strong interest and wished to contribute to solve the carbon dioxide problem from technological aspects, gathered and discussed the countermeasures in this conference. As a result of this effort, carbon dioxide utilization has been making a progress and distinct contributions in conducting the strategies for carbon dioxide mitigation and in finding the solution of these environmental problems by adoption of a variety of technologies. In this regard, since the first conference, held in 1991 at Nagoya, Japan, the ICCDU conferences have been held every two years (alternating between North America, Europe, and Asia) with the aim of covering the latest developments in the field of carbon dioxide chemistry as it pertains to utilization. The 7th International Conference on Carbon Dioxide Utilization (7th ICCDU) was held successfully at the Hoam Convention Center in Seoul, Korea from October 12 to October 16, 2003. We would like to express sincere thank all the members of the Local Organizing Committee for their active supports for the conference, the several organizations for their financial support, and all the distinguished participants who have traveled to Korea from around the world. The attendance in this conference was very encouraging with respect to the futuristic perspective of the scientific field and the efforts for the development on carbon dioxide utilization. The 7th ICCDU Conference was organized by Korea Research Institute of Chemical Technology (KRICT) and Carbon Dioxide Reduction & Sequestration (CDRS) R&D Center in Korea. During the 4 day meeting of the 7th ICCDU, 5 plenary lectures, 14 keynote lectures, and 53 papers were presented orally in 4 sessions and 66 papers were presented as posters. Their topics of the 7th ICCDU covered the followings: CO2 separation and recovery including advanced CO2 separation system and materials for CO2 separation CO2 as oxidant including dehydrogenation and oxidative conversions -

heterogeneous catalysis of CO2 including catalytic CO2 conversions, plasma & microwave application and CO2 rich syngas from biomass & fossil fuels

-

homogeneous catalysis of CO2 including chemicals from CO2, polymers using CO2 and electro- and photocatalytic reduction of CO2

-

biology and biochemistry of CO2 utilization

xiv

studies in supercritical CO2 including spectroscopic techniques and reactions in supercritical CO2 -

utilization of CO2/CH4 hydrates

This book was conceived as the proceedings of the 7th ICCDU, which reflect the most recent research results and stimulating scientific discussions with new challenges in advancing the development of carbon dioxide utilization. These topics are important in green chemistry and technology that are considered as essential fields for the global sustainability. The organizers wish to express sincere appreciation to attendees of the 7th ICCDU and the authors for submitting their manuscripts to the proceedings. We are grateful to the outstanding scientists who accepted our invitation to overview vital research areas in plenary lectures and the keynote lectures that introduce the important topics of each session covered by the conference. We are also grateful to Dr. Choong Seob Kim (President of KRICT), Dr. Sang Do Park (Director of CDRS), Prof. Yoichi Kaya (RITE/NEDO), Prof. Eunyoung Heo (Seoul National Univ.) for their generous financial and administrative support. We wish to thank Dr. Soo Min Oh (KRICT), Dr. Myoung-Jae Choi (KRICT), and Prof. Woonsup Shin (Sogang Univ.), who have spent so much time and efforts for the success of the 7th ICCDU conference. Furthermore, we wish to thank members of the Research Center for Nanocatalysis (RCNC), KRICT, especially Dr. Sung Hwa Jhung and Mr. Do-Young Hong, who very efficiently helped in the preparation of the proceedings. Finally we wish to acknowledge the help and support by co-operating organizations and sponsors from industries.

Seoul, October 2003 Editors Sang-Eon Park Jong-San Chang Kyu-Wan Lee

XV

Chairman's Remark Thanks to the Contributors of 7th ICCDU, This proceedings contain papers presented and displayed as posters at the Seventh International Conference on Carbon Dioxide Utilization (ICCDU VII) held for four days during October 12-16, 2003 at Hoam Convention Center in Seoul, Korea. The ICCDU conferences are held every two years alternating continental. The First ICCDU was held in Nagoya, Japan in 1991, the second at Bari, Italy in 1993, the third at Oklahoma, USA in 1995, the fourth in Kyoto, Japan in 1997, the fifth at Karlsruhe, Germany in 1999, the sixth at Breckenridge, USA in 2001 and the Seventh in Seoul, Korea in 2003. The eighth will be held in Norway in 2005. This Conference was aimed to cover the latest developments in the field of carbon dioxide chemistry as it pertains to utilization in the world. In the 7th Conference included seven fields and 21 detail subjects and emphasized specially on the presentations of research results scale in pilot and demonstration operation and utilization of biomass and CO2/methane hydrates which are recently more interested as sustainable and renewable energy sources. This ICCDU VII, was organized jointly by Korea Research Institute of Chemical Technology (KRICT) and Carbon Dioxide Reduction & Sequestration R&D Center (CDRS). The conference was financially supported by the Ministry of Science and Technology of Korea, Korea Research Institute of Chemical Technology, Carbon Dioxide Reduction & Sequestration R&D Center and Research Institute of Innovative Technology for the Earth (RITE)/New Energy and Industrial Technology Development Organization (NEDO) in Japan. I would like express my sincere thank to the all attendees and organizations for their contribution to and cooperation in the 7th ICCDU. Greetings from Yanji, Jilin, China Chairman of ICCDU VII Prof. Dr. Ing. Kyu-Wan Lee Yanbian University of Science & Technology (YUST), Yanji, Jilin, China 133000


xvii

Organization

Organizing Committee Chairman Kyu-Wan Lee

KRICT, Korea

Secretary Sang-Eon Park

Inha Univ., Korea

International Advisory Board Michele Aresta

Campus Univ., Italy

David Beecy

Department of Energy, USA

Son-Ki Ihm

KAIST, Korea

Donald Darensbourg

Taxas A&M Univ., USA

Tomoyki Inui

Kyoto Univ., Japan

Eckhard Dinjus

Karlsruhe Research Centre, Germany

Yoichi Kaya

Research Institute of Innovative Technology for the Earth, Japan

Baldur Eliasson

ABB Switzerland Ltd., Switzerland

Kenneth Nicholas

Univ. of Oklahoma, USA

International Scientific Committee Alfons Baiker

Swiss Federal Institue of Technology, Switzerland

Philip Jessop

Queen's Univ., Canada

Danielle Ballivet

Universite de Bourgogne, France

Klaus S. Lackner

Los Alamos National Laboratory, USA

Dan DuBois

Univ. of North Carolina, USA

Jae-Sung Lee

POSTECH, Korea

David Dyrssen

Chalmers Univ. of Technology and Gothenburg Univ., Sweden

Kyu-Wan Lee

KRICT, Korea

Giuseppe Filardo

Univ. of Palermo, Italy

Chang-jun Liu

Tianjin Univ., China

Martin M. Halmann

Weizmann Institute of Science, Israel

Yu-han Sun

Chinese Academic of Sciences, China

xviii

Allan A. Johansson

The Technical Research Centre of Finland (VTT), Finland

Tatsuaki Yashima

Research Institute of Innovative Technology for the Earth, Japan

Shohei Inoue

Tokyo Univ. of Science, Japan

Shin-ya Yokoyama

AIST Chugoku National Institute of Advanced Industrial Science and Technology, Japan

Local Advisory Board Yong-Bok Chae

Seoul National Univ., Korea

Hyun-Ku Rhee


Ho-ln Lee


Sang-Chul Shim

Kyungpook National Univ., Korea

Local Organizing Committee Kyu-Wan Lee (Chairman)

KRICT, Korea

Sang-Eon Park (Secretary)

Inha Univ., Korea

Jong-San Chang

KRICT, Korea

Jin-Suck Lee

KIER, Korea

Soon-Haeng Cho

KIER, Korea

Kew-Ho Lee

KRICT, Korea

Jae-Ou Choi

Pohang Institute of Industrial Science and Technology, Korea

Dae-Won Park

Pusan National Univ., Korea

Sung Hwa Jhung

KRICT, Korea

Myoung-Jae Choi

KRICT, Korea

Sang-Do Park

KIER, Korea

Hee-Moon Eum

KEPRI, Korea

Woonsup Shin

Sogang Univ., Korea

Sung-Hwan Han

Hanyang Univ., Korea

Supporting Organizations The Organizing Committee gratefully acknowledges the receipt of financial support from

Korea Research Institute of Chemical Technology (KRICT) Ministry of Science and Technology (MOST) Korea Science and Engineering Foundation (KOSEF)

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Research Center for Nanocatalysis (RCNC) Carbon Dioxide Reduction & Sequestration R&D Center (CDRS) Research Institute of Innovative Technology for the Earth (RITE) New Energy and Industrial Technology Development Organization (NEDO) Korean Institute of Chemical Engineers (KIChE)

Supporting Organizations for Technical Excursion Korea Gas Corporation (KOGAS) Seoul Thermal Plant (KEPCO)


Studies in Surface Science and Catalysis 153 S.-E. Park, J.-S. Chang and K.-W. Lee (Editors) © 2004 Elsevier B.V. All rights reserved.

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Heterogeneous Catalytic Reactions with CO2: Status and Perspectives Gabriele CENTI and Siglinda PERATHONER Department of Industrial Chemistry and Engineering of Materials and ELCASS (European Laboratory for Surface Science and Catalysis), University of Messina, Italy.' The use of CO2 as feedstock for producing chemicals is not only interesting to contribute to mitigation of greenhouse gas emissions, but an interesting challenge to explore new concepts and new opportunities for catalysis and industrial chemistry. The present state and perspectives in heterogeneous catalytic reactions with CO2 are briefly surveyed both in terms of (i) use of catalysis for CO2 fixation and (ii) use of CO2 in heterogeneous catalytic reactions as mild oxidant, H-transfer agent and catalyst surface modifier. Discussion is limited to heterogeneous gas-solid catalytic reactions. Some results and opportunities for gas-phase photoelectrocatalytic conversion of CO2 to fuels (hydrocarbons and alcohols) are also shown. Keywords: carbon dioxide, catalysis, reforming of CO2, methanol, CO2 as oxidant, photoelectrocatalytic conversion of CO2, CO2 as catalyst modifier 1. INTRODUCTION The use of CO2 as feedstock for producing chemicals will not have a major impact on the mitigation of greenhouse gas emissions, at least in a short-medium term scenario, but is an integral part of carbon management [1-3], because provides a strategic path for its recycle and reuse. There are several motivations for producing chemicals from CO2: - CO2 is a cheap (especially considering possible CCVemission taxes) and nontoxic feedstock that can replace toxic chemicals such as phosgene or isocyanates. - The production of chemicals from CO2 can lead to new and valuable materials such as polycarbonate as well as to new routes in the synthesis of chemicals (for example, hydrocarboxylation of butadiene to adipic acid). - The production of chemicals from CO2 has a small, but positive impact on the global carbon balance. - The use of CO2 as feedstock or co-reactant is a challenging opportunity which stimulates new approaches in industrial chemistry, as will be briefly discussed below. CO2 is a highly oxidized and thermodynamically stable compound. Its conversion requires high free energy content substances (NH3, amines, etc.) or an external supply of energy (photochemical, electrical or thermal). The use of a catalyst to promote reaction rates and/or the selective pathway of reaction is typically required [4]. Catalysis, either homogeneous, heterogeneous or enzymatic were successfully applied in the conversion of greenhouse gases and in the chemical syntheses using CO2 as feedstock [1-11]. 1

Salita Sperone 31, 98166 Messina, Italy. Fax: +39-090-391518, e-mail: [email protected], [email protected]

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Heterogeneous catalysis shows in general advantages linked with reaction rate, stability, separation, handling and reuse of the catalyst, and reactor design. Discussion is limited here to heterogeneous gas-solid catalytic reactions. Although the range of compounds synthesized from CO2 by heterogeneous catalytic routes is comparatively narrow with respect to homogeneous catalytic routes, the latter have been discussed in detail previously [8,9]. Main products which can be produced from CO2 by heterogeneous gas-phase processes include methanol, methylamines and formic acid derivatives, syn gas (CO, H2) and derived products (dimethyl ether and synthetic hydrocarbons by Fischer-Tropsch synthesis). Parts of these reactions have been previously reviewed [7,10] or are discussed in other contributions of this volume. Therefore, discussion will be focused here on the analysis of the state-of-the-art and the perspectives of the reactions of CO2 reforming with methane and CO2 hydrogenation, but will include also a short discussion of the gas-phase photoelectrocatalytic conversion of CO2 which is a challenging option for converting CO2 to fuels [6,12]. The use of CO2 in heterogeneous catalytic reactions offers also challenging opportunities as novel mild oxidant, H-transfer agent and catalyst modifier. This emerging area of research will be also discussed. The use of supercritical CO2 as the reaction medium for catalytic heterogeneous reactions is another interesting research area [13,14], but not discussed here. 2. USE OF HETEROGENEOUS CATALYSIS FOR CO2 FIXATION The heterogeneous catalytic conversion of CO2 involves mainly gas phase hydrogenation reactions, with the exception of the CO2 reforming with CH4, and occurs at high temperatures, although should be mentioned that heterogeneous catalysts have been successful applied also in liquid phase reactions such as in the oxidative carboxylation of olefins to organic carbonates [15]. The major reactions of heterogeneous catalytic conversion of CO2 are listed below: CO2 + H2 X CO + H2O CO2 + CH4 X 2CO + 2H2 CO2 + 3H2 X CH3OH + H2O

(1) (2) (3)

The hydrogenation of CO2 to CH4 (Sabatier methanation) has been not included in this list. There are limited economic and environmental incentives to synthesize the most reduced hydrocarbon (CH4), if not using renewable energy sources (such as solar energy), although this reaction was extensively studied on noble (Rh, Rh) or transition (Ni, Fe) metals supported on oxides or zeolites,. A full life cycle energetic and economic analysis is necessary to assess the potential routes for mitigating CO2 emissions [16,17], but often data on this aspect are lacking in literature. Equations (1-3) are equilibria which occur simultaneously and therefore the product yields are below 100%. The role of the catalyst in terms of metal, support, additives, and preparation method is crucial to optimize the performances in these reactions, but equally important is the reactor design (fluidodynamics, quenching, flow rate, and reaction conditions). Optimization requires an integrated catalyst and reactor design [18], which, however, is often lacking. 2.1 CO2 reforming or reaction with methane The reaction of CO2 reforming with methane (eq. 2) is known from several years and commercial processes already exist; in particular, the SPARG® (Sulfur PAssivated ReforminG of Haldor Topsee), a process characterized by the use of a nickel based catalyst passivated by sulphur to reduce the rate of coke formation, and the CALCOR® (Caloric GmbH) processes which is characterized by a structured catalytic bed in the reactor to limit coke formation. The CO2 reforming with CH4 has received renewed interest (as alternative to steam reforming) in recent years [19-26], since it decreases the H2/CO ratio from about 3 to

approximately 1, allowing improved selectivities for the production of synthetic hydrocarbons by Fischer-Tropsch (FT) chemistry, and oxygenated chemicals (dimethylether, in particular). Due to RWGS (Reverse Water Gas Shift, eq. 1), the H2/CO ratio is less than 1 and therefore RWGS should be minimized. Other alkanes such as propane may be also used [27]. Generally, lower reaction temperatures are necessary when the nr. of C atoms increases. Typical reaction temperatures for CO2 reforming are 900-1000K and the major process cost is the supply of the heat of reaction necessary for this high temperature endothermic reaction. It was often criticized that the CO2 reforming with methane cannot be considered a reaction which allows the mitigation in the emissions of CO2, because the energetic costs necessary to supply the heat of reaction at high temperature as well as to recover and purify the CO2 from flue gases overcome the benefits (in terms of greenhouse gas emissions) of using CO2 as the feedstock. An advanced solution proposed to directly use the flue gas and avoid external supply of energy is the tri-reforming process [28]. The tri-reforming is a synergetic combination of CO2 reforming, steam reforming, and partial oxidation of natural gas. The syngas (CO/H2) produced can be used for making ultra-clean fuels or chemicals by FT or oxo synthesis, and for generation of electricity by gas turbines [28]. CO2 may be converted to hydrocarbons on FT-type catalysts (for example, Fe-Zn promoted with Cu and K [29]) in a one-step process instead that in a two-step process (first producing syngas and then converting them via FT synthesis). The adsorbed oxygen formed in CO2 dissociation reacts with adsorbed hydrogen to form H2O or with CO to reform CO2, and furthermore water reacts with CO to form again CO2. Therefore, although the addition of CO2 during FT synthesis could lead to the minimization of CO2 formation and to the preferential removal of oxygen as H2O, leading to improved selectivity to olefin content and C5+ reaction products, adsorbed CO is the intermediate in FT synthesis, differently from the case of methanol synthesis discussed later. As a consequence, a two step process for converting CO2 to FT-type products is necessary. The main problem in the CO2 reforming with methane is the formation of filamentous-type coke on the catalyst via CH4 decomposition or, most likely, CO disproportionation. Between the catalysts, supported nickel was the most studied, but interesting results have been also obtained with supported noble metals (Rh, Ru, and Pt). Sulphur passivation may increase resistance to coke deposition, but decreases activity. The role of the support is very important to limit deactivation. MgO, A12O3, or Ca aluminate are often used. Recent trends include the use of alkaline-earth or rare-earth oxides (as support for Ni) to prevent coke formatiom, and the use of ZrO2 (or doped ZrO2) as support for Pt or Ni [29,30]. One of the motivation for the lower coking rate on zirconia is that stable carbonates are formed, due to the acid character of the support. The simplified reaction mechanism of CO2 reforming with methane on Pt/ZrO2 catalysts is shown in Scheme 1. The key aspect is that CH4 and CO2 adsorb on different active sites and that the carbonate (formed by strong CO2 chemisorption at the support/noble metal interface) is reduced to formate by CHX species formed by methane dissociation on the noble metal. However, a different reaction mechanism involving the direct reaction of CO2 with surface carbonaceous species formed by CH4 dissociation may also occurs. This second mechanism is dominating for some catalysts such as RI1/AI2O3. The reaction of CO2 with surface carbonaceous species formed by methane dissociation has been applied recently for the challenging reaction of direct synthesis of acetic acid [19,32] in order to avoid the thermodynamic limitations: CH4 + CO2 -> CH3COOH

(4)

Pd/SiO2 allows the formation of acetic acid, although in low amounts, at temperatures around 450K. The investigation originates from patented results [33] reporting IR data of the formation of

4

Scheme 1 Simplified reaction mechanism of CO2 reforming with methane on Pt/ZrO2 catalysts acetic acid (as adsorbed species) when a CH4/CO2 feed is passed over a 5% Pd/charcoal catalyst. 2.2 CO2 hydrogenation The synthesis of methanol starting from CO2 (eq. 3) is a well known reaction which attracted recently new interest, because of the potential of CH3OH in a future hydrogen economy and in downstream chemicals production. Catalysts should possess high selectivity to CH3OH and low selectivity to CO from the RWGS reaction (eq. 1) and therefore should be selective at low temperature and high pressure, since the RWGS reaction is favoured at high temperature. Main catalysts used are Cu/ZnCXAbCh) mixed oxides, but Co, Pd, and Re are also active metals [34]. Differently from the FT synthesis, there are definitive proofs, for example using labelled 14 CO2 and 14CO, that CO2 is the reaction intermediate in methanol synthesis also starting from CO/H2 mixture (CO2 forms by WGSR- Water Gas Shift Reaction), although may apparently be contradictory that CO reduction to CH3OH should pass through CO2 as intermediate. The reason is related to the mechanism of CO2 adsorption on the catalyst [34,35]. In methanol synthesis catalysts the largest part of the active sites are covered with strongly adsorbed species under reaction conditions and therefore reactants chemisorption is a critical aspect of the reactivity. Still debate exists on the nature of the active sites [Cu, Cu-Zn(O)] and reaction intermediate (surface formate, carbonate or formaldehyde). Generally, it was observed that CO2 conversion is proportional to the metallic copper surface area (determined by N2O titration, for example), but different dependences were observed as a function of the nature of the support in Cu/ZnO based catalysts [36]. The reason is that the microstructure of the metallic copper is very important in determining activity. Recent results reported a correlation between TOF (TurnOver Frequency) and Cu lattice strains present in Cu/ZnO catalysts [37]. The strained Cu crystallites derives from the presence of some substitutional atoms of Zn and lattice miss-match deriving from the contact with ZnO crystallites. The highly strained Cu crystallites result more active in methanol synthesis. CO2 hydrogenation may leads to also other products such as formic acid, but the lower stability of these products leads generally to low selectivity in gas phase heterogeneous reactions. However, the presence of other reactants (such as amines) which can react with the formate surface intermediates to give rise to more stable products leads to selective syntheses, such as the production of diethylformammide and diethylamine [7]. Methanol which forms from CO2/H2 reaction may be also further converted using bi- or multi-functional catalysts (for example, Cu/Zn catalysts + Fe/ZSM-11) to dimethylether or iso-hydrocarbons. There is a rich and interesting

5

possible chemistry which has been not fully analyzed and which is expected to give rise to interesting applications in a near future [38].

3. USE OF CO 2 IN HETEROGENEOUS CATALYTIC REACTIONS The use of CO2 as mild oxidant, H-transfer agent and as modifier of the catalyst surface offers some interesting challenges for the heterogeneous catalysis research which are only started to be explored, although already indicate some interesting opportunities to improve existing processes. CO2 was studied as mild oxidant in various type of reactions, such as the acrylonitrile synthesis from acetonitrile and methanol on K/MgO catalyst [39], the ethane ODH (Oxidative DeHydrogenation) to ethylene on 5% Cr/SiO2 [40], and the ethylbenzene ODH to styrene on K/Fe2O3 [41,42]. Recently the use of CO2 as mild oxidant in the conversion of light alkanes to oxygenated products on V2O5 supported on oxidized diamond [43] was also reported. The most interesting results were obtained in the ethylbenzene to styrene conversion, where the use of CO2 as mild oxidant considerably limits the fast deactivation characterizing this industrially relevant reaction. It is likely that this can be the first example of industrial application of CO2 as mild oxidant. Minura and Saito [44], studying the reaction on a Fe2O3(10%)/Al2O3 catalyst observed that both single step mechanism and two step mechanism single step two step

Ph-CH2CH3 + CO2 -> Ph-CH=CH2 + CO + H2O Ph-CH2CH3 -> Ph-CH=CH2 + H2 H2 + CO2 -» CO + H2O

(5) (6) (7)

contribute to the overall reactivity. CO2 or the active oxygen species produced by CO2 dissociation may also react with surface carbonaceous species limiting their formation. CO2 is also applied as mild oxidant in the selective oxidation of single-walled carbon nanotubes (SWNT) to open the end of the tubes (temperatures about 850K) [45]. The use of CO2 is more selective (and simpler) than other methods such as laser ablation. The use of CO2 as H-acceptor has been also reported, although it should be noted that may be often difficult to distinguish when CO2 acts as H-acceptor (like in eq. 7) or instead as mild oxidant (eq. 5, for example). Interesting results have been reported in light alkane ODH: CnH2n+2 + CO2 -» CnH2n + CO + H2O (8) C + CO2 -> 2CO (9) using conventional dehydrogenation catalysts, but using alternating feeds of alkanes and CO2 being the 2nd reaction (eq. 9) slower with respect to the first one (eq. 8) [46]. CO2 acting as Hacceptor has been reported to promote the ethylbenzene ODH on mesoporous Mg-V-Al mixed oxides [47] and the propane aromatization on metal-loaded ZSM-5 catalysts [48]. In the latter reaction, CO2 promotes catalyst stability, but especially the selectivity to aromatics. CO2 as modifier of the catalyst performances has been also reported [49], although the investigation in this area should be considered still at the beginning. CO2 is widely used as ballast (diluent) in industrial reactions (for example, in selective oxidation reactions) [50] and therefore the understanding of its influence on the catalyst properties is industrially relevant and also offers new opportunities to develop new or improved processes. CO2 was indicated to promote the formation of V5+ in mesoporous Mg-V-Al mixed oxides for the ethylbenzene to styrene conversion [47], and to create new active sites in alkylaromatics oxidation on Fe-Mo-oxides encaged in B-Silicalite [51] (Fe-Mo/DBH) and in methane combustion on Pd/CexZri_xO2 [52]. In alkylaromatics oxidation on Fe-Mo/DBH catalysts it was observed that CO2 promotes p-ethyltoluene oxidation to p-methylstyrene, while inhibits the consecutive oxidation of p-methylstyrene to p-tolualdehyde [51]. It was suggested that new

6

peroxocarbonate species forms on the surface of the catalyst in the presence of CO2 and O2 and these species are responsible both of the promotion of the activity in H-abstraction reactions (pmethylstyrene formation) and of the inhibitions of the O-insertion reactions (p-methylstyrene oxidation). The hypothesis is suggestive, but not proven, although the role of peroxocarbonate species as modifiers of the surface activity has been suggested also in other reactions such as the nitrogen oxides conversion on Pt-BaO/Alumina NOx-storage-reduction catalysts [53]. Ruiz et al. [52] studying the methane oxidation on Pd/CexZr1.xO2 catalysts observed instead a correlation between catalyst activity and proportion of oxygen atoms bound to Cerium after catalytic test. In the presence of CO2 the amount of the latter species increases proportionally to the increase of catalyst activity, suggesting that these oxygen active species form by CO2 dissociation to CO and O. Bernal et al. [54] demonstrated the re-oxidation of ceria by CO2 in the presence of Pt due to the generation of spillover oxygen by CO2 dissociation on platinum. It is known that spillover oxygen can enhance the surface reactivity of the catalyst. Therefore, the mechanism proposed by Ruiz et al. [52] is possible, while more doubts exist on the effective stability and reactivity of peroxocarbonate species. No definite proof exists in favour of one over the other mechanisms of promotion by CO2 of the surface reactivity and therefore more studies are necessary. Probably also different mechanisms may be present depending on the catalyst nature. 4. (PHOTO)ELECTROCATALYTIC GAS-PHASE CONVERSION OF CO2 In the CO2 conversion using solar energy, direct (photochemical) or indirect (photoelectrochemical) routes are possible. The 2nd has the theoretical advantage of higher efficiency by reducing the rate of recombination of e7h+ pairs generated in the photoexcitation process. Photoelectrochemical conversion of CO2 has been studied mainly in liquid phase using either homo- or heterogeneous catalysts [2]. It should be noted that a first step of the CO2 conversion is the generation of the CO2" anion radical [55] which is easily converts to products such as formic, oxalic and glycolic acids which recovery from aqueous solution is costly. Products such as methanol and methane may also obtained, but with low selectivities. Few studies have been instead reported on the gas phase electrocatalytic conversion, although this is more interesting in terms of reaction products (higher hydrocarbons and alcohols) and reduced cost of their separation. An attractive possibility of using solar energy to produce protons and electrons (by water photodissociation) which then can be used to reduce CO2 to alcohols or hydrocarbons, is the photo-electrochemical (PEC) reactor [6,12,56] which operates according to the following scheme: -e—>

H2O + A v / ^ . 0 2 t

II

T"-CO 2 ->CH 4 , CH3OH, ....

The symbol || indicates a proton membrane. In a PEC reactor, one side is composed of a titania-based photocatalyst which oxidises water to O2 using solar light. This process produces protons and electrons that are transported to the other side of the device by means of a membrane and an electric connection, respectively. On the other side of the photo-electrochemical device the protons and electrons react with CO2 in the presence of an electrocatalyst to produce hydrocarbons such as methane or methanol. The net reaction is thus the conversion of combustion products to fuels/chemicals using solar energy, allowing a net decrease in the emissions of CO2 into the atmosphere. Furthermore, the pure oxygen produced on the photocatalytic side can be used to improve the efficiency of the combustion or can be used in other chemical processes. Several aspects are important in determining the performance of this type of PEC reactor and the efficiency in the use of the solar energy to convert CO2 back to fuel (i.e., the process

7 economics). A critical aspect is the development of the electrocatalyst which should be active at or near room temperature, resistant to poisoning, and selective in avoiding side reactions (H2 and CO formation). Furthermore, the catalyst should operate in the gas phase electroreduction of CO2, while most of the literature data refer to liquid phase behaviour only [2]. Pt or Pd nanoclusters on GDM (Gas Diffusion Membranes) show enhanced properties with respect to commercial materials for the gas phase electrocatalytic reduction of CO2 [12]. A recent extension under investigation is the use of the concept to synthesize Fischer-Tropsch hydrocarbons from CO 2 . 5. CONCLUSIONS The use of CO2 as feedstock for producing chemicals is not only interesting to contribute to mitigation of greenhouse gas emissions, but an interesting challenge to explore new concepts and new opportunities for catalysis and industrial chemistry. REFERENCES 1. C. Song, A.M. Gaffney, K. Fujimoto (edited by), CO2 conversion and utilization, ACS Symposium Series No. 809, Publisher: American Chemical Society, Washington, D. C. (2002). 2. M.M. Halmann, M. Steinberg, Greenhou se Gas Carbon Dioxide Mitigation, CRC Press: Boca Raton (FL, US) 1999. 3. H. Arakawa, M. Aresta, J.N. Armor, M.A. Barteau,E.J. Beckman, A.T. Bell, J.E. Bercaw, C. Creutz, E. Dinjus, D.A. Dixon, K. Domen, D.L. DuBois, J. Eckert, E. Fujita, D.H. Gibson, W.A. Goddard, D.W. Goodman, J. Keller, G.J. Kubas, H.H. Kung, J.E. Lyons, L.E. Manzer, T.J. Marks, K. Morokuma, K.M. Nicholas, R. Periana, L. Que, J. Rostrup-Nielson, W.M.H. Sachtler, L.D. Schmidt, A. Sen, G.A. Somorjai, P.C. Stair, B.R. Stults, W. Tumas, Chemical Reviews, 101 (2001) 953. 4. A. E. Van Diepen, F. Kapteijn, M. Makkee, J. A. Moulijn, in Environmental Catalysis, F.J.J.G. Janssen, R.A. van Santen Eds., World Scientific Publishing 1999, pp. 219. 5. G. Centi, S. Perathoner, F. Vazzana, CHEMTECH, 29(12) (1999) 48. 6. G. Centi, S. Perathoner, Z. Rak, Appl. Catal. B: Env. 41 (2003) 143. 7. A. Baiker, Appl. Organometallic Chem., 14 (2000) 751. 8. M. Aresta, E. Quaranta, CHEMTECH, (1997) 32. 9. M. Aresta, Studies in Surface Science and Catal., 114 (1998) 65. 10. T.Inui, Preprints-American Chem. Society, Div. of Petroleum Chemistry (2000), 45(1), 113. 11. R.H. Heyn, Carbon dioxide conversion, in Encyclopedia of Catalysis, J. Wiley & Sons Publisher (2002). 12. G. Centi, S. Perathoner, Z. Rak, Studies in Surface Science and Catal., 145 (2003) 283. 13. A. Baiker, Chem. Rev, 99 (1999) 453. 14. W. Leitner, Appl. Organometallic Chem., 14 (2000) 809. 15. M. Aresta, A. Dibenedetto, J. of Mol. Catal. A: Chem., 182-183 (2002)399. 16. J.H. Edwards, Catal. Today, 23 (1995) 59. 17. J.N. Armor, Studies in Surface Science and Catalysis, 114 (1998) 141. 18. G. Centi, S. Perathoner, CATTECH, 7 (2003) 78. 19. G. Chen, J. of Natural Gas Chem., 11 (2002) 109. 20. M.C.J. Bradord, M.A. Vannice, Catal. Rev.-Sci. Eng., 41 (1999) 1. 21. Y.H. Hu, E. Ruckenstein, Catal. Rev.- Sci. Eng., 44 (2002) 423. 22. K. Asami, X. Li, K. Fujimoto, Y. Koyama, A. Sakurama, N. Kometani, Y. Yonezawa, Catal. Today, 84 (2003) 27. 23. J. Munera, S. Irusta, L. Cornaglia, E. Lombardo, Appl. Catal., A: General, 245 (2003) 383. 24. T. Inui, Catal. Today, 29 (1996)329.

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Studies in Surface Science and Catalysis 153 S.-E. Park, J.-S. Chang and K.-W. Lee (Editors) © 2004 Published by Elsevier B.V.

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Chemicals from CO2 via heterogeneous catalysis at moderate conditions Yuhan Sun State Key Laboratory of Coal Conversion, Institute of Coal Chemistry Chinese Academic of Sciences, Taiyuan, 030001 PR China For the mitigation of global warming due to carbon dioxide, catalytic conversion of CO2 has been extensively investigated for last decade. Most of studies in this field focused on the utilization of carbon dioxide mainly as a carbon source or as an oxide reagent. As a result, those processes required a significant amount of energy input because CO2 is thermodynamically stable. In our laboratory, CO2 is considered as carboxylation reagent via heterogeneous catalysis under moderate conditions in order to take full advantage of carbon and oxygen source in CO2. By reacting CO2 with so-called "active hydrogen" containing organic compounds, the manufacture of valuable chemicals such as ethylene/propylene carbonate and dimethyl carbonate could be economically carried out via the development of product tree. 1. INTRODUCTION Carbon dioxide is the most abundant substance produced by human activities. Its further accumulation in the atmosphere must be avoided in order to prevent risky situations derived by the reinforcement of the greenhouse effect. In the last century, it has been paid much attention to reduce CO2 emission and utilize CO2'1'2'. CO2 is a thermodynamically very stable molecule (free energy is -394.01kl/mol). The conversion of carbon dioxide into industrially useful compounds has been a challenge for synthetic chemists and then has recently attracted more interest in view of the so-called "Sustainable Society" ^ and "Green Chemistry"'4' due to its environmentally friendly nature (nontoxic and nonflammable) as the most abundant and the cheapest carbon source in the earth. As a result, chemical approaches to CO2 conversion were developed in the last two decades. Herein, those can be categorized as: (i) CO2, only as a carbon source or a oxide reagent, is activated through catalytic reduction processes with various kinds of reductants; (ii) CO2, as a carboxylation reagent, is activated into valuable chemicals with so-called "active hydrogen" in organic compounds at mild conditions. In the former case, most of the studies were concentrated on the utilization of CO2 via catalytic reduction, such as CO2 hydrogenation to carbon and methanol'5"7', syngas synthesis by CO2 and CH^7'8-' etc. *To whom all corresponding should be addressed, E-mail: yhsun(o),sxicc.ac.cn or [email protected]; Fax: +86-351-4041153; Tel: +86-351-4053801

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Unfortunately, these processes are usually carried out at high temperature and/or pressure, and then required a significant amount of energy input. Moreover, both conversion and selectivity hardly meet the satisfactory economics at present. In the latter case, the reaction condition is very mild and economical due to the use of so-called "active hydrogen" in organic compounds for valuable chemicals. At the same time, such an approach is to take full advantage of carbon and oxygen source of CO2. The present utilization mainly includes urea synthesis from CO2 and NH3, salycilic acid from CO2 and phenol[10], long chain carboxylyic acids by CO2 with olefins or dienes[11], polycarbonates and/or cyclic carbonates by coupling CC^with epoxides[12], and dimethyl carbonate from CO2 and methanol[13], and alkylureas by amines and CC^'14^. It is worth to note that the synthesis of urea and salycilic acid does not require any catalysts: they are pure thermal reactions. Also, the synthesis of urea represents the largest single process for the utilization of CO2. Additionally, urea can be considered as an active form of carbon dioxide. As mentioned above, carbon dioxide, as an oxide reagent, is known to be much less active than molecular oxygen, even with hydrogen as a reduction reagent. However, CO2, as a carboxylation reagent, could easily be activated into valuable chemicals with so-called "active hydrogen" containing organic compounds at mild conditions. Thus, CO2 could be activated into valuable chemicals at mild condition. Of course, the organic compounds with "active hydrogen" play very important role in the process of reaction. Obviously, the reactant used to activate CO2 could be phenol, ammonia, olefins, dienes, epoxides, methanol, alkylureas, which all possess very strong electron rich atom in their molecules. The hydrogen located at the near electron rich atom of these molecules, which is so-called "active hydrogen " such as the hydrogen in a site, shows the strong tendency to positively attack the oxygen atom of carbon dioxide. As a result, these molecules can easily be converted into nucleophilic reagents and react with the carbon atom of carbon dioxide at mild conditions. Unfortunately, those processes using CO2 as a carboxylation reagent are seldom applied in industry because most of catalysts used in these reaction are organometallic complexes with other organic solvents as promoters. Furthermore, these homogeneous catalysts and promoters also lead to the difficulties for the product separation and then the high cost. Thus, the development of new heterogeneous catalysts for CO2 carboxylation reaction is the key for the industrial utilization of CO2 in large scale. In our laboratory, CO2 is considered as carboxylation reagent via heterogeneous catalysis under moderate conditions and the manufacture of valuable chemicals such as ethylene/propylene carbonate and dimethyl carbonate is developed by reacting CO2 with so-called "active hydrogen" containing organic compounds. Here we introduce some results regarding the CO2 utilization as a carboxylation reagent over high active heterogeneous catalysts. 2. RESULTS AND DISCUSSION

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2.1. CO2 Cyloaddition with epoxides to Ethylene/Propylene Carbonates One of the most attractive synthetic goals starting from carbon dioxide is the five-membered cyclic carbonates system because five-membered cyclic carbonates have many synthetic uses and have generally been synthesized from the corresponding diols and phosgene or related compounds'151. As we know, epoxides possess the active hydrogen because of the influence of the negative electron of oxygen in the molecule, and might lead to very easily react with CO2. hi fact, many inorganic and organic compounds including amines'-1 , phosphanes'161, organotin halides'171, alkali metal salts'18^, transition metal complexes'191, alkali metal halide-crown ether'20' or polyethylene glycol-400'21], and phthalocyaninatoaluminum in supercritical carbon dioxide'221 are known to catalyze the activation of CO2 molecule. For these processes, high temperatures, high pressures, and toxic polar solvents such as DMF or CH2CI2 have been thought to be necessary. However, under these reaction conditions, some inactive or polymerization-sensitive oxiranes are hardly converted to the corresponding cyclic carbonates. Moreover, the reaction was carried out in an intermission autoclave, leading to low production capacity. Herein, we report an extremely simple and ecologically safer route to cyclic carbonate from the reactions of epoxide with carbon dioxide in the presence of new heterogeneous catalysts with high catalytic performance at moderate conditions without organic solvent. In order to further improve the catalyst lifetime and the productivity capability, a new continual structured reactor was also investigated in our laboratory. KI supported on different supports was investigated for the heterogeneous process (see Table 1). Firstly, the reaction was operated at moderate condition (i.e.l20°C and 2.0MPa), and the pressure was lower than the saturate pressure of carbon dioxide at room temperature. Secondly, both propylene oxide (PO) or ethylene oxide (EO) conversion and five-cyclic carbonate selectivity were nearly 100% although neither solvents nor promoters were used in this catalytic synthesis. Thus, the product did not need to be separated, which simplified the production of five-cyclic carbonate. The mechanism of this reaction on the heterogeneous catalyst was introduced in the previous papers '23]. In order to avoid the hot spot of the catalysts and then further improve the lifetime of the catalysts and productivity capability, a new continual structured reactor, the capacity of which is 100 t/a, was designed. Figure 1 showed that both EO conversion and EC yield hardly changed in lOOOh with KI supported activated carbon in the novel continual structured reactor. This demonstrates that CO2 could very easily be activated into valuable chemicals with so-called "active hydrogen" in epoxides with heterogeneous catalysts at mild conditions. Table 1. Effect of supports on catalytic performance Catalysts

PO or EO Convention, %

Product Yield , %

KI/y-Al2O3

100.0

99.0

Kl/activated carbon

100.0

99.3

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KI/SiO2 100.0 NOTE: T=120°C, P=2.0MPa, CO2/epoxide(mole rate)=10:l.

99.0

Figure 1 The performance of catalysts atl20°C and 2.0MPa 2.2. Dimethyl carbonate from urea and methanol Urea synthesis is a traditional industrial process for CO2 utilization. Of course, in the process of urea synthesis, we can treat ammonia as a very active molecule with "active hydrogen" due to the negative electron of nitrogen in the molecule. Therefore, CO2 can easily react with ammonia even without any catalysts. As a CO2 active form, urea is cheap and widely available, so it is thought to be a potential material for the synthesis of important organic compounds. In fact, ammonia, as one of the products, could be separated and recycled. In our laboratory a kind of heterogeneous catalysts, which are solid base catalysts, are explored. These catalysts have been demonstrated to show excellent catalytic activity without any other organic solvent used in the reaction process. The urea conversions were nearly 100%, and the yield of DMC, especially for ZnO-K2O and ZnO-Carbon, exceeded 50%. At the same time, a catalytic distillation technology is also used in order to shift the equilibrium and improved the DMC yield. Figure 2 showed that the catalyst was highly stable in

Figure 2 The performance of catalyst on stream at 190 C and 0.8MPa in the catalytic distillation reactor the catalytic distillation reactor with the production capacity of 100 t/a. The detail reaction

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performance and reaction mechanism would be introduced in other paper of the meeting. 2.3. Propylene Carbonate from Urea and Propylene Glycol The synthesis of propylene carbonate from urea and propylene glycol, which was very safe and economical synthesis method, was carried out in our laboratory. The reaction process is as below:

Figure 3 Effect of different catalysts on propylene carbonate yield (Reaction conditions: 170 °C,300mmHgand2h) The performance of some heterogeneous catalysts are illustrated in Figure 3. The acidic oxides (such as AI2O3, SiC>2 and TiC^) showed lower catalyst activity than alkali metal oxides (such as MgO, CaO and BaO). But, It was interesting that ZnO and PbO showed the highest activity (propylene carbonate yield could reach over 99%). Thus, it could be concluded that the catalysts with a suitable acid/basic balance showed good performance. The detail mechanism is under investigation. 2.4. Dimethyl Carbonate from CO2 and Methanol The direct synthesis of dimethyl carbonate from CO2 and methanol is the most attractive due to the low-cost of CO2. In fact, methanol is also an active molecule with "active hydrogen". Unfortunately, the synthesis is limited by thermodynamic equilibrium, and favorable only at low temperature. It has been reported that DMC can be synthesized from methanol and CO2 in the presence of dialkoxydibutyltin, metal(IV) tera-alkoxide[24], magnesium dialkoxide^25' and other organometallics'26'. But the activity was very low even in

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the presence of dehydrates and promoters. The catalysts were also easily decomposed by water, and considered not to overcome limits of the reaction. In addition, the reaction was in most cases conducted above 423 K, which was far from the critical temperature of CO2. hi our laboratory, direct conversion of CO2 and methanol to DMC was carried out using nickel acetate as the catalyst in CO2 supercritical phase, and CO2 acted as a reactant and a reaction medium. The typical results were shown in Table 2. It was obvious that in non-supercritical (413K and 5.0MPa) only 2.1% methanol converted and 76.1% methyl acetate produced as by-product after 12h reaction. However, the yield of DMC was highly improved at near supercritical CO2 conditions (305K for 2 hrs), and DMC was the unique product. As the pressure changed from 7.4 to 10.3 MPa, the yield of DMC reached the maximum at 9.3MPa, which was 13 times higher than that at non-supercritical conditions, showing that the production of DMC was pressure sensitive. Such phenomenon was also observed in DMC synthesis from CO2 and trimethyl orthoacetate at 423K for 24h. It should be noted that the improvement could be attributed to the supercritical phase behavior itself rather than high CO2 concentration. Based on the characterization by BET, XRD and EXAFS, the reaction mechanism was proposed in the previous paper.

Pressure Methanol Conversion, °, /o DMC yieldb (Mpa) 62.0 5.0 2.1* 569 7.4 8.9 587 8.3 9.3 796 9.3 11.3 279 10.3 6.2 a Cat. / MeOH =l:25(mol), 32°C, 2h. * 12h b Calculated as mmole product per mole of cat.

Methylacetate yield 76.1 0 0 0 0

3. CONCLUSION The present work demonstrated that CO2, as a carboxylation reagent, is easily activated into valuable chemicals along with so-called "active hydrogen" containing organic compounds at mild conditions. Moreover, heterogeneous catalysts are proved to be very effective. Thus, chemicals from CO2 via heterogeneous catalysis under moderate conditions is very promising and economical route for CO2 utilization.

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Synthetic hydrocarbon fuels and CO2 utilization G. Schaub*, D. Unruh and M. Rohde Engler-Bunte-Institut, Universitat Karlsruhe (TH), Engler-Bunte-Ring 1, D-76131 Karlsruhe, Germany ""corresponding author [email protected] The potential of generating or utilizing CO2 during Fischer-Tropsch (FT) hydrocarbon synthesis affects overall efficiencies in synthetic hydrocarbon fuel production. Looking at different carbon-containing raw materials, the behavior of CO2 during FT synthesis is determined by syngas composition (ratio H2/CO, availability of H2), type of catalyst and reaction conditions. Natural gas and biomass as actual examples for potential feedstocks have different characteristics in this respect. Natural gas has an excess of hydrogen, whereas in biomass conversion, carbon can be removed as CO2 to solve the H2 deficiency problem, thus lowering overall carbon utilization efficiency. Low carbon efficiencies are also achieved with coal as feedstock, leading to large fossil CO2 emissions. Finally, pure CO2 (from combustion flue gas or from the atmosphere) can be used as carbon source for the chemical conversion of H2 into a liquid fuel, however competing with carbon in a reduced state. 1. HYDROCARBON SYNTHESIS High value hydrocarbon transportation fuels (in particular diesel fuel) can be synthesized from any carbon-containing raw material via synthesis gas production and subsequent Fischer-Tropsch (FT) synthesis (Fig. 1). This may become interesting because fuel supply today is based to a large extent on hydrocarbons and demands will increase very significantly in the future, due to the expected increase in the number of motor vehicles in the world. Car manufacturers expect a total of 1.4 x 109 of cars worldwide for 2030 (2003 : 0.7 x 109), which will need motor fuels with low pollution effects. On the other hand, petroleum as raw material for today's fuel supply may reach its production maximum in the near future with decreasing production levels afterwards. Climate change concerns may in addition force to make better use of and substitute fossil raw materials. In this context, synthetic hydrocarbon fuels may serve as a short- or mid-term substitute because (i) today's distribution infrastructures and modern engines technologies can be used, and (ii) due to the high fuel qualities (e.g. low sulfur and aromatics content, optimum distillation curve, high cetane index). The implication of CO2 in fuel synthesis includes (i) generation of CO2 due to carbon removal required for increasing H/C ratios and to energy requirements of the conversion process, met by fuel combustion, and (ii) potential chemical conversion of CO2 in the synthesis step, to increase the ultimate yield of hydrocarbon products in cases where sufficient H2 is available. This paper presents the state-of-the-art and addresses various aspects of CO2 generation and utilization in connection with Fischer-Tropsch hydrocarbon synthesis.

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CH 4 + vO 2 + wH 2 O CHj 6 O 0 7 + vO 2 + wH 2 O

— —

xCO + yH 2 + zCO 2 xCO + yH 2 + zCO 2

(1) (2)

CO + 2 H 2 — -fCH 2 ^- + H 2 O (CO 2 + H 2 * = > CO + H 2 O) -f CH2-)- + uH2 —

(3) (4) -eCH2+^-

(5)

Fig. 1. Overall flow scheme for the conversion of organic feedstocks (natural gas, biomass) to liquid hydrocarbon fuels Traditional Fischer-Tropsch synthesis using CO/H2 feedgas produces paraffinic and olefinic hydrocarbons, ranging from methane to high molecular weight waxes. They are produced in polymerization reactions on the surface of a solid catalyst (Eq. 3 in Fig. 1). With CO/H2 feedgas, reactions over iron, cobalt or ruthenium catalysts lead to hydrocarbon molecules that are mainly linear. The mechanism is dominated by a competition of hydrocarbon chain growth and chain termination (desorption from the catalyst surface). The carbon number distribution of FT-products ideally follows a simple statistical model with a chain growth probability parameter a [1]. Chain length distribution can be adjusted and controlled to a limited extent to gain yields of desired products by varying (i) the kind of catalyst (promoted Fe, Co), and (ii) operating conditions (temperature, feedgas composition). If diesel fuel with molecular carbon numbers in the range 12 to 20 is the desired product, maximum yields can be achieved by a two-step strategy. During synthesis, long chain waxy hydrocarbon molecules (C21+) are produced as intermediates (Fig. 2), that are subsequently hydrocracked and hydroisomerized to produce a high-quality diesel end product. Mechanistic aspects, kinetics and selectivity con-trol continue to be critical issues in Fischer-Tropsch chemistry even \—] 1 „ 1 j—I after decades of research and 0.8 "I applications. Fig. 2. Correlation of primary synthesis product composition and hydrocarbon chain growth parameter a, according to AndersonSchulz-Flory model.

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Due to the complex chemical conversion process and depending on the composition of the feedstock, the overall yields of the fuel product are significantly lower than the yields with petroleum as raw material (Tab. 1). Deficiency in hydrogen and presence of oxygen in the raw material lower the achievable hydrocarbon yields, and even methane with an excess of hydrogen leads to a loss in energy and carbon efficiency. Table 1 Overall yields of synthetic hydrocarbon products from various feedstocks as compared to natural hydrocarbons from petroleum, r|Hu : conversion efficiency based on lower heating value TlHu(%)

natural gas') biomass2) coal3) petroleum4) *

100 CH4 100CH16O07 100CH06O01 100CH18 l.o

0 3)

68-80 24-40 28-40 94

fCH2)-fCH 2 )-fCH 2 )-fCH 2 )v

53-63 30-50 35-50 94

Z *

Audus et al. 2001 [2], Sie et al. 1991 [3], 2) Tijmensen et al. 2002 [4] Hoogendorn 1976 [5], from: JUntgen et al., in: Elliott (Ed.) 1980 [6], 4 ) own estimate

2. BEHAVIOR OF CO2 IN FISCHER-TROPSCH SYNTHESIS Hydrogenation of CO2 and effects on Fischer-Tropsch reactions have been studied on various catalysts which can be divided according to their catalytic activities in CO/CO2 shift reaction. Some prominent effects reported in the recent literature are summarized in Figs. 3 and 4. So far, iron exhibits the most promising properties for hydrogenation of CO2 to form longchain hydrocarbons [7]. As shown in Fig. 3, the same organic products are formed independently whether the FT reaction starts from CO or from CO2. However, with cobalt as an active metal, exhibiting low or no CO/CO2 shift activity at temperatures typical of FT reactions, the product composition shifts with increasing CO2 content from a typical FT-type to almost exclusively methane. Here, the FT regime of specific inhibition of product desorption through strong reversible CO adsorption [8] cannot be obtained under low COpressure conditions. Results from experimental studies in lab-scale fixed bed reactors with Fe/AbOs/Cu/K led to a reaction network and a kinetic reaction model [9]. CO2 is reduced to CO as the primary step (equilibrium-limited with kinetic control) followed by subsequent CO hydrogenation to hydrocarbons (with kinetic control), Eqs (4) and (3) in Fig. 1. This sequence of reactions allows higher overall CO2 conversion than the CO2 shift reaction alone. Kinetic analysis of hydrogenation experiments with CO2/H2 synthesis gas led to a system of rate equations and rate parameters, leading to conversion- and yield-time curves typical of consecutive reactions (Fig. 4 left). Recent studies show that this kinetic reaction model developed for CO2/H2 syngas can also be applied to CO/CO2/H2 syngases [10]. Low-temperature FT experiments in a perfectly mixed flow (slurry) reactor with various Co catalysts showed that CO2 behaves as an inert gas component, however a negative effect on FT-reaction rate was observed in one case (Co/SiO2-Prec, Fig. 4 right). Although the

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reason for this deactivation is not known yet, it is obvious that in principle Co-catalysts can be developed with a performance not affected by the presence of CO2.

Fig. 3. Measured hydrocarbon product distributions with Fe/Al2O3/Cu/K and Co/MnO/SiO2/Pt catalyst, H2/CO and H2/CO2 synthesis gas, lab-scale fixed bed reactor (2 g catalyst), T = 190 °C (Co)/250 °C (Fe), p = 1 MPa, (Pm/pcox) in = 2 (Co), (pm/pcox) in = 2.3 (Fe), tmod = 4 g s/cm3 (mcat/V in NTP) [11].

Fig. 4 (left): CO2 conversion (X) and yields of CO and organic products (Y) versus residence time (m Fe /V in NTp ), lab-scale fixed bed reactor, catalyst: Fe/Al2O3/Cu/K, 100/13/10/10, T = 300 °C, p = 1 MPa, (PH2/PC02) in = 3, curves: calculated with kinetic model [9]. (right): Effect of CO2 partial pressure on organic product formation rates with Co- and Fe-catalysts, T = 250 °C (Fe)/190 °C (Co), p = 1 MPa, perfectly mixed (slurry) reactor (10 g catalyst), PH2/pco/pH2o/pAr = 0.23/0.14/0.03/0-0.55 MPa (Fe); 0.35/0.18/0.45/0 MPa (Co-Imp); 0.53/0.28/0.18/0 MPa (Co-Prec). xmod = 14 g s/cm3 (mcat/V in NTP), curves: calculated with rate equations, Imp : impregnation, Prec : precipitation [11]. Based on the experimental information available so far the following conclusions can be drawn for the potential of CO2 utilization in FT synthesis: (i) the catalyst used should have a high CO/CO2 shift activity, (ii) sufficient H2 must be available, (iii) chemical equilibrium limitations (CO/CO2 shift) must be lessened, e.g. by increasing reaction temperature or removing H2O [12]. 3. FEEDSTOCK NATURAL GAS ("GAS-TO-LIQUIDS" GTL) Conversion of natural gas to synthetic liquid hydrocarbon fuels is presently seen as an option to utilize large volumes of otherwise unusable natural gas (stranded gas, associated gas). Driving forces herein are (i) low price of the feedstock, (ii) clean and high-value fuel

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products, and (iii) avoidance of fossil CO2 emissions caused by flaring of associated gas. Two commercial GTL plants are presently in operation, several plants are in design and erection status [13]. In the case of CH4, practically no CO2 appears in the syngas, because enough H2 is available. CO2 emissions do result from meeting the energy requirements of the overall GTL process [2]. Carbon utilization efficiencies of 68-80 % and energy efficiencies of 53-63 % are reported (Table 1). Even the feedstock with the highest possible hydrogen content leads to extra CO2 emissions due to its chemical conversion. The excess hydrogen in CH4 could be used to convert CO2 in two cases, (i) Large natural gas reserves are known with high contents of CO2. Hydrogenation of CO2 to Fischer-Tropsch products by using the extra hydrogen in CH4 may be of interest for natural gas with low CO2content (< 25 % in a mixture CH4 and CO2). According to stoichiometry, the hydrogen available in CH4 is sufficient for hydrogenating both the carbon and the oxygen. If the natural gas contains higher hydrocarbons (ethane, propane etc.) there is less hydrogen available per amount of carbon present, (ii) When combining conversion of natural gas and a feedstock with a lower H/C ratio (like biomass), the extra hydrogen in CH4 could be used to hydrogenate part of the CO2 resulting from synthesis gas production of the H-deficient feedstock (for biomass, see section 4). 4. FEEDSTOCK BIOMASS ("BIOMASS-TO-LIQUIDS" BTL) Conversion of biomass to liquid hydrocarbon products has recently gained interest, e.g. in Europe [4], where political aims were defined regarding the introduction of biofuels in the transportation sector (5.75 % of all transportation fuels by 2010). This is because substitution of fossil fuels by biomass has a potential to contribute to a reduction of fossil CO2 emissions (Fig. 5).

Fig. 5. Significant global carbon flows (natural cycle and contribution caused by human activities) and carbon storages (based on Bolin 1970 [14], Schlesinger 1997 [15]) The most critical factors for achievable overall efficiencies are those related to the gasification step: (i) O2-consumption depending on reactor principle (temperature, pressure,

22

heat management), reactor size, biomass moisture content, and affecting the composition of the syngas produced (see Fig. 1), (ii) yields of (H2 + CO) being essential for the overall yields of liquid hydrocarbon product (representative gas compositions summarized in Table 2). Quantitative mass balance data for biomass gasification to synthesis gas production from biomass are not available so far (although fuel gas production is a proven technology). As practical challenges in biomass conversion to hydrocarbons can be seen (i) an increase of hydrocarbon yields (because the available biomass is a limiting factor). Conversion of CO2 in the syngas could offer a possibility for increasing overall yields, e.g. by addition of external hydrogen and removal of H2O during FT synthesis [12]. The potential increase of hydrocarbon yields is indicated in Fig. 6. (ii) Due to the low energy density of biomass and its decentralized production, as well as due to the complexity of the conversion process, overall cost must also be considered as a limiting factor. Preliminary studies indicate hydrocarbon production cost figures of 2.5-4 times higher than for today's petroleum-derived hydrocarbons [4], with a reduction potential based on process development progress. In this respect, the appropriate plant size is critical related to centralized/decentralized concepts of biomass processing. Table 2 Composition of synthesis gases produced from biomass (and natural gas) feedstock

model gas

O2/H2O

5) 57.4

6) (66.7)

15.8

28.4

(33.3)

27.2

35.9

14.2

~0

0.1

3.1

11.6

n.a.

n.a.

1.8

0.9

0.8

2.0

2.0

~2

0.48

0.24

0.20

0.23

0.58

"~ 1

gasifying medium pressure

H2O atm

xH2

52.4

O2 atm 2) 35.2

xCO

28.7

41.1

38.6

xCO2

16.8

22.0

xCH4

2.1

xH2/xCO

D vol% dry

xH2/(2xCo +3x C 02)

natural gas

biomass H2O/O2 press 3) 31.0

H2O/O2 press 4) 31.7

x, : volume fraction x 100 1) own estimate, based on [16], 2) UET [17], 3) Creusot Loire [18], 4) IGT with process adjustments, estimate from [4], 5) corresponding to steam gasification, max. H2-availability [19], 6) combination non-catalytic partial oxidation and steam reforming, estimate based on [3]

5. FEEDSTOCK COAL As for potential CO2 utilization, the situation with coal is similar to that of biomass. The feedstock is poor in hydrogen, so syngas H2 content must be increased by shift conversion of CO, leading to high amounts of CO2 that are ultimately emitted to the atmosphere. Resulting carbon utilization and energy efficiencies are low (Table 1), leading to high extra-fossil CO2 emissions. Large integrated coal-to-liquid plants have been operating successfully in South Africa since the 1950s [6], and very significant progress in FT technology (regarding

23

catalysts, reactors, process integration) has been achieved during this long period of operation.

Fig. 6. Effect of O2 consumption (per volume of product gas) in biomass gasification process on maximum achievable yields of Fischer-Tropsch hydrocarbon product and potential for CO2 utilization (own estimate based on stoichiometric approach)

6. FEEDSTOCK CO2 Utilization of CO2 as carbon source for the chemical conversion of H2 into a liquid synthetic fuel was suggested by various authors (e.g. [20]). Flue gases from combustion or air can serve as source for CO2 [21]. Although methanol was first proposed as final product, also hydrocarbons can be synthesized based on CO2/H2 syngas. Chemical thermodynamics and stoichiometry indicate that CO2 is less attractive as carbon source for synthesis than carbon in a reduced state (as in the feedstocks biomass or natural gas). Significantly higher amounts of H2 are consumed (see Table 3), and equilibrium limitation of CO2 shift conversion results from CO2 being a very stable molecule. The higher synthesis temperatures required would lead to lower chain growth probabilities (Fig. 2) and, as a consequence, to a different product upgrading strategy to yield the desired transportation fuels. Table 3 Stoichiometry of H2 conversion to synthetic hydrocarbons with CO2 and biomass as carbon sources

7. CONCLUSIONS Synthetic hydrocarbon fuels from natural gas and biomass may significantly contribute to meet future transportation fuel demands. Petroleum will pass a production peak in the future, and climate change concerns will enhance substitution of fossil raw materials, e.g. by biomass. Synthesis of hydrocarbon fuels from a carbon-containing raw material generally leads to generation of CO2, due to (i) the energy requirements of the conversion process, met by fuel combustion, and (ii) the necessary carbon removal if H/C ratios have to be increased, as with raw materials poor in hydrogen and rich in oxygen (like biomass, coal). Utilization of CO2,

24

present in the syngas, leading to increased product yields, can be achieved if sufficient H2 is available and catalyst/reaction conditions are adjusted to convert CO2 to CO and subsequently to FT hydrocarbons. Utilization of CO2 may be of significant interest if concerns of potential climate changes will force human activities to strictly control fossil CO2 emissions. If biomass as a non-fossil raw material available in limited amounts is to be upgraded to liquid transportation fuels, product yield increases by a factor 2 to 3 can be achieved if CO2 from gasification is successfully used as an additional carbon source in FT synthesis. ACKNOWLEDGEMENT Discussions with Prof. Hans Schulz and Prof. Kyu-Wan Lee are gratefully acknowledged. REFERENCES 1. Flory P., J. Am. Chem. Soc. 58 (1936) 1877. 2. Audus H., Choi G., Heath A., Tam S.S., Conference on Natural Gas Conversion, Alaska, June 2001. 3. Sie ST., Senden M.M.G., van Wechem H.M.H., Catalysis Today 8 (1991) 371. 4. Tijmensen M.J.A., Faaij A.P.C., Hamelinck C.N., van Hardeveld M.R.M., Exploration of the Possibilities for Production of Fischer-Tropsch-Liquids and Power via Biomass Gasification, Biomass & Bioenergy 23 (2002) 129-152 (2). 5. Hoogendorn J.C., Gas from Coal for Synthesis of Hydrocarbons, paper presented at American Institute of Mining Engineers, 23 rd Annual Meeting 1974. 6. Juntgen H., Klein J., Knoblauch K., Schroter H.J., Schulze J., in: M.A. Elliott (ed.): Chemistry of Coal Utilization, 2nd Supplementary Vol., Wiley New York 1981. 7. Lee M.D., Lee J.F., Chang C.S., Bull. Chem. Soc. Jpn. 62 (1989) 2756. 8. Schulz H., van Stehen E., Claeys M , Topics in Catalysis 2 (1995) 223. 9. Riedel T., Schaub G., Jun K.-W., Lee K.-W., Ind. Eng. Chem. Res. 40 (2001) 1355. 10. Rohde M., Unruh D., Pias P., Lee K.-W., Schaub G., 7th International Conference on Carbon Dioxide Utilization, Seoul, Korea, Oct. 2003. 11. Riedel T., PhD dissertation, Universitat Karlsruhe 2003. 12. Unruh D., Rohde M., Schaub G., 7th International Conference on Carbon Dioxide Utilization, Seoul, Korea, Oct. 2003. 13. Thackerey F., Fischer-Tropsch Gas-to-Liquids, Prospects and Implications, SMI Publ. Ltd., London, 2000. 14. Bolin B., The Carbon Cycle, Scientific American, September 1970, 130. 15. Schlesinger W.H., Biogeochemistry, An Analysis of Global Change, 2" Edition Academic Press 1997. 16. Hofbauer H., Rauch R., Foscolo P., Matera D., Proc 1st World Conference on Biomass for Energy and Industry, James & James London (2001) 1999. 17. Althapp A., UET Freiberg, Personal Communication (2003). 18. Lemasle J.M., in: H.V. Bridgewater (ed.): Thermo-chemical Processing of Biomass, Butterworths London 1984. 19. Lee K.W., Jun K.W., Nam S.S., Hwang J.S., Choi M.J., Abstracts 6* International Conference on Carbon Dioxide Utilization, Beckenridge USA, Sept. 2001, p. 18. 20. Ushikoshi K., Mori K., Watanabe T., Takeuchi M., Studies in Surface Science and Catalysis, Vol. 114 (1998) 357-362. 21. Specht M., Bandi A., Elser M., Staiss F., Studies in Surface Science and Catalysis, Vol. 114(1998)363-366.


25

Design of Copper Based Hybrid Catalysts for CO 2 Hydrogenation Young-Kwon Park1, Jong-Ki Jeon2 and Son-Ki Ihm* 'Faculty of Environmental Engineering, University of Seoul, Seoul 130-743, Korea. 2

Dept. of Chemical Engineering, DongYang University, YoungJou, Kyungbook 750-711, Korea.

'National Research Laboratory for Environmental Catalysis, Dept. of Chemical and Biomolecular Engineering, Korea Advanced Institute of Science and Technology, Taejon, 305-701, Korea.

Catalyst design requires a number of experimentation and decision steps from initiating the idea to developing the catalytic processes. CO2 hydrogenation is limited by thermodynamic constraints in the formation of methanol. This thermodynamic barrier could be mitigated by in-situ transformation of methanol inside the reaction system where the Cu-based hybrid catalysts can be applied to obtain chemical products such as hydrocarbons or oxygenates. The modification of Cu-based catalyst for methanol synthesis, the selection of proper zeolite counterpart, the consideration of different reaction mechanisms, and the medium acidity induced by the solid-solid interaction were discussed by reviewing the information reported, and some guidelines for designing Cu-based hybrid catalysts were proposed for better control of hydrocarbon yield and product distribution.

1. INTRODUCTION Catalytic hydrogenation of carbon dioxide has been recognized as one of the best methods to mitigate global warming as well as to solve carbon resource problems [1-5]. Methanol synthesis from carbon dioxide and hydrogen has been most widely investigated, even if the conversion of carbon dioxide is limited by the thermodynamic constraint [1,2]. The thermodynamic equilibrium could be shifted if methanol is converted within the reaction system. One simple approach is to apply the hybrid or composite catalysts composed of methanol synthesis catalyst and methanol conversion catalyst, which was reported to show some good yields to hydrocarbons [2-5]. This hybrid catalyst system seems to be more advantageous than the combination of Fischer-Tropsch synthesis and CO2 reforming in view of product distribution where the Schultz-Flory distribution law does not always prevail. Development of hybrid catalyst requires a lot of trial and error efforts, which can be efficiently

* To whom correspondence should be addressed. FAX: +82-42-869-5955. E-mail: [email protected] This work was partially supported by National Research Laboratory Program and Brain Korea 21 Project

26 reduced through catalyst design. Characteristics of Cu-based hybrid catalysts in CO2 hydrogenation will be discussed in terms of hydrocarbon yield and product distribution, and some main parameters for a better design of the hybrid catalysts will be summarized.

1.1. Catalyst design Catalyst design can be regarded as a logical application of available information to the selection of a catalyst for a given reaction [6]. Considerable effort is now being extended to provide a logical scientific basis for many of the processes that have been developed arbitrarily.

A general

sequence

of

operations

is

schematically shown in Fig. 1. Once the reaction to be catalyzed can be defined, it is necessary to describe the reaction in chemical terms, so as to allow some check of the thermodynamic feasibility. If this appraisal was satisfactory, catalysts may be designed, prepared and tested. At this moment this is done largely on the basis of experience and trial and error, and a logical basis of selection of catalysts to be tested is the subject. It is also important to develop a database on the activity patterns of various catalysts correlated with the type of target

Fig. 1. Schematic illustration of

molecules. Besides the catalytic activity, other factors

designing procedure

such as flow rate, temperature, pressure drop constraints and durability should be considered in the design of catalyst.

2. METHANOL SYNTHESIS FROM CO2 HYDROGENATION OVER COPPER CATALYSTS 2.1. Cu based catalysts for methanol synthesis from CO2 hydrogenation Various kinds of metal catalysts were reported to be active for methanol synthesis from CO2 hydrogenation. Sugawa et al. [7] reported that activity of metal catalyst for methanol synthesis increased with the following order, Cu » Co=Pd=Re > Ni > Fe > > Ru=Pt > Os > Ir=Ag=Rh > Au. Ihm et al. [8] investigated the effect of the Cu dispersion in Cu/ZrO2 catalyst on the catalytic activity in CO2 hydrogenation. The methanol synthesis activity of the Cu/ZrC>2 catalyst was greatly influenced by the Cu dispersion (or Cu crystallite size); the smaller the crystallite size, the higher the rate of methanol synthesis. When some components of Ce, Cr, Pd, K, V and Zn were added as promoters into Cu/ZrC>2, the crystallite size of Cu particles changed significantly. Saito et al. [9] classified Cu/ZnO

27

based catalyst into two categories in terms of Cu dispersion and specific activity. It is believed that AI2O3 or ZrO2 enhanced Cu dispersion and that Ga2O3 and Cr2O3 increased specific activity. 2.2. Reaction mechanism of methanol synthesis over Cu based catalyst While many reaction mechanisms for methanol synthesis in CO2 hydrogenation have been suggested, the commonly accepted reaction intermediate is formate(HCOO) [8,10]. Ihm et al. [8] suggested that reaction intermediates of methanol synthesis were carbonate, formate, formaldehyde and/or methoxy, and the rate determining step for methanol synthesis seemed to be the conversion of formate into formaldehyde or methoxy. Also they proposed the synergistic effect between copper and ZnO; formate migration onto ZnO and its hydrogenation to methanol on ZnO [10].

3. METHANOL CONVERSION REACTION OVER ACIDIC ZEOLITES Methanol conversion reaction can be carried out over various kinds of acidic zeolites such as MFI type and SAPO type zeolites. For MFI type catalyst like HZSM-5 or H-Ga-silicate, the main products were higher (C5+) hydrocarbon including aromatics. On the other hand, SAPO type catalyst like SAPO-34 or SAPO-5, the main product was C3 or C4 hydrocarbon. In view of main hydrocarbon from the methanol conversion, it is expected in CO2 hydrogenation that higher hydrocarbon could be obtained with HZSM-5 and C3 hydrocarbon with SAPO-34 if they are hybridized with Cu-based methanol synthesis catalyst.

4. HYDROCARBON SYNTHESIS OVER HYBRID CATALYSTS 4.1. Hybrid system to shift the thermodynamic equilibrium It can be demonstrated in Table 1 that the hybrid catalyst (Cu/ZnO/ZrO2 + SAPO-34) in a single bed showed higher activity and higher hydrocarbon selectivity than the two-bed system. This enhancement of activity should be due to the relaxation of thermodynamic limitation through the hybridization between Cu based methanol synthesis catalyst and SAPO-34.

4.2. Hydrocarbon synthesis over Cu based hybrid catalyst Table 2 shows the results of CO2 hydrogenation over hybrid catalysts. The aforementioned methanol conversion reaction over HZSM-5 and SAPO-34 might suggest that higher hydrocarbon could be obtained with HZSM-5 hybrid catalyst and that C3 hydrocarbon with SAPO-34 hybrid catalyst. This is true in the case of SAPO hybrid system where C3 hydrocarbons are major products. In the case of HZSM-5 hybrid system, however, the main hydrocarbon was ethane and neither higher hydrocarbons nor aromatics were formed. Higher yield and selectivity for C2+ hydrocarbons were also reported by Jeon et al. [4] and Fujiwara et al. [3] over hybrid catalysts ; CuZnOZrO2/SAPO-5 and

28 CuZnCr oxide/HY zeolite, respectively. They proposed the following reaction pathway; methanol synthesis from carbon dioxide and hydrogen over copper catalyst, methanol/dimethyl ether to lower alkene over the zeolite, alkene oligomerization and isomerization over the zeolite, and hydrogenation to alkane over the copper catalyst. Table 1 Activities of CO2 hydrogenation in different arrangements of catalyst beds with (CuO/ZnO/ZrO2 and/or SAPO-34 [11] Arrangement (b) (a) (c) 26.9 Conversion (%) 27.7 32.9 98.4 97.0 78.7 CO selectivity (%) Oxygenate selectivity (%) 0.0 1.5 0.9 1.6 1.5 20.4 H.C. selectivity (%) 0.44 0.40 6.7 H.C. yield 400°C, 3.0MPa, H2/CO2:3, W/F:10g-cat h/mol, time on stream:2h (a) : Reactants contacted first with 0.5 g of CuO/ZnO/ZrO2 and then with 0.5 g of SAPO-34, (b): Reactants contacted first with 0.5 g of SAPO-34 and then with 0.5 g of CuO/ZnO/ZrO2, (c) : Reaction was carried out over a hybrid catalyst of lg in a single bed. Table 2. Activities of CO2 hydrogenation over various hybrid catalysts [11] Con. Selectivity (%) H.C.s distribution (%) H.C.s yield Catalyst (%) H.C.s CO Oxy C , C2 C3 C4 C 5+ total C2-C4 Cat.l : 27.8 4.8 93.4 1.8 35.9 59.5 3.8 0.4 0.4 1.3 0.8 (A)+HZSM-51) Cat 2 : 29.2 7.5 88.2 4.3 22.6 56.7 17.0 1.5 2.2 2.2 1.7 (A)+HZSM-52) Cat 3 : 32.9 20.4 78.7 0.9 9.2 35.1 43.3 11.4 1.0 6.7 6.0 (A)+SAPO-34 400°C, 3.0MPa, H2/CO2:3, W/F:10g-cat h/mol, time on stream:2h, 1) Si/Al=22, 2) Si/Al=11.8 (A):CuO/ZnO/ZrO2 (6:3:1)

5. DESIGN FACTORS FOR HYDROCARBON SYNTHESIS OVER Cu BASED HYBRID CATALYST 5.1. Design of methanol synthesis catalyst Fujiwara et al. [12] carried out CO2 hydrogenation using hybrid catalysts comprised of Cu-ZnCr/HY zeolite (Table 3). When methanol synthesis catalyst, prepared from CuO, ZnO and CrO3, was mixed with HY zeolite, C2+ hydrocarbons were obtained in a good selectivity (Run 1). The same catalyst made into granule gave better results (Run 3). The yield of hydrocarbon was higher and selectivity to methane was lower. No olefin was observed. When methanol synthesis catalyst was prepared by coprecipitation, the yield of hydrocarbons decreased. They suggested that sodium remaining even after 5 times wash deactivated the active site of zeolite. Similar tendency was reported

29 on the hydrocarbon synthesis between two Cu-Zn/HY composite catalysts, in which one Cu-Zn catalyst was precipitated by Na2CO3, and another Cu-Zn catalyst was precipitated by oxalic acid. Since Cu-Zn-chromate catalysts have lower ability to hydrogenate carbon double bond [3], a Cu-Znchromate catalyst containing 1% of Cu was most effective for olefin synthesis among various hybrid catalysts (Run 4). It was reported that the addition of alkaline metals to Cu-Zn-chromate catalyst, especially Cs, improved the selectivities to ethene and propene. Table 3. Activities of CO2 hydrogenation over various hybrid catalysts [12] HC distribution (wt%) Conv Conv. to (%) C2" Sel. Catalyst H.C. CO Oxy (%) C c2 c3 c4 c5+ A/HY 14 0 37.3 9.2 27.3 23 0.8 8 37 18 2.1 29.2 32.4 A2)/HY 10 5 14 1.1 33 38 0 A3)/HY 24 12 35 24 0 40.9 14.4 26.0 0.5 5 B4)/HY 5.2 30.2 39 25 24 9 3 58 35.5 0.1 400°C, 50 atm, H2/CO2:3, SV=3000ml/g-cat h, A: Cu-Zn-Cr(3:3:l), B: Cu-Zn-chromate 1) C27(C2+C2=), 2)Prepared by coprecipitation, 3)Mixing of oxide granules 4)Cu/(Cu+Zn)= 0.01

(%)°

Run 1 2 3 4

5.2. Design of methanol conversion catalyst Table 2 shows the hydrogenation activities of different hybrid catalyst systems. Of the hybrid catalysts composed of CuO/ZnO/ZrO2 (A), the SAPO-34 hybrid catalyst (Cat. 3) showed a larger conversion and hydrocarbon selectivity than the HZSM-5 hybrid catalysts (Cat.l and Cat.2). Furthermore SAPO-34 hybrid catalyst and HZSM-5 hybrid catalyst resulted in different product distributions, which should be due to different mechanisms of hydrocarbons formation i.e., consecutive mechanism for HZSM-5 and carbon pool mechanism for SAPO-34 [14], respectively. Fig. 2-(a) shows the typical spectra of NH3-TPD for pure zeolites and hybrid catalysts. SAPO-34 has a larger amount of acid site than HZSM-5 catalyst. A new peak around 350 °C which was not observed in each of metal oxides and zeolites appeared when the zeolites were hybridized with CuO/ZnO/ZrO2 methanol synthesis catalysts. It is noted that CuO/ZnO/ZrO2 catalyst showed no TPD peak throughout the whole temperature region. Quadrupole mass analysis of desorbed gases showed that the high intensity peak around 350°C in hybrid catalyst appeared due to not only NH3 but also N2 and H2O (Fig. 2-(b)). Ammonia adsorbed on the acid sites acted as a reducing agent, resulting in the reduction of nearby metal oxides, mainly CuO, and NH3 itself might be decomposed into N2 and H2O. The area of this new peak was found to be in good correlation with the activity of hybrid catalyst for CO2 hydrogenation, especially hydrocarbon yield as shown in Fig. 2-(c). The hydrocarbon yield of hybrid catalyst having SAPO-34 was much higher than that of hybrid catalyst having HZSM-5 (Table 2), which could be predicted by the fact that the medium acidity peak of the former showed a larger peak area. This implied that the amount of medium acidity was well correlated with the amount of original zeolite's Bronstead acid sites [5]. HZSM-5 (Si/Al=l 1.8) showed higher amount of medium acidity than HZSM-5 (Si/Al=22). For the various MFI hybrid catalysts (HZSM-5, H-Ga-silicate, H-

30 Fe-silicate), the HZSM-5 hybrid catalyst showed the highest yield and selectivity to C2+ hydrocarbon, which had the largest amount of medium acid sites [5,13].

Fig. 2. (a): TPD profiles of pure zeoltes and corresponding hybrid catalysts (R=Si/Al), (b): quadrupole mass spectra of SAPO-34 and hybrid catalyst, (c): relationship between TPD medium peak area and catalytic activity for [CuO/ZnO/ZrO 2 ] hybrid catalysts at a same conversion [11] In an attempt to increase the hydrocarbon formation, iron was incorporated into the methanol synthesis catalyst which was hybridized with SAPO-34. The SAPO-34 hybrid catalysts with various Fe2O3/CuO/ZnO/ZrO2 ratio such as 1:5:3:1,3:3:3:1, 5:1:3:1, and 0:6:3:1 showed a great increase in the CO2 conversion as well as the hydrocarbon yield. However the product distribution is quite similar to the typical Schulz-Flory distribution, which is different from that of CuO/ZnO/ZrO2 based hybrid catalysts. Even small fraction of iron component in the hybrid catalyst seemed to favor the FischerTropsch reaction rather than the methanol synthesis reaction. As expected, however, the Fe2O3/CuO/ZnO/ZrO2+SAPO-34 hybrid catalysts did not show any medium acidity in NH3-TPD, and the activity correlation could not be established due to the difference in hydrocarbon formation mechanism (i.e., Fischer-Tropsch synthesis from carbon dioxide). Ihm et al. [11] also observed that the interactions between CuO/ZnO/ZrO2 and SAPO-34 led to mutual modifications in the acidic property of zeolite and the reducibility of metal oxides. FT-IR showed that intense hydroxyl stretching bands at 3625 and 3600 cm'1 in pure SAPO-34 were masked

31 due to hybridization. Multiple reduction peaks above 200 °C due to interactions between CuO and SAPO-34 were confirmed by TPR experiments.

5.3. Pore size effect The pore size of zeolite can affect the product distribution. While small size of SAPO-34 or 44 gave C3 as main product, the large pore size of SAPO-5 (0.8 nm) gave C4 hydrocarbon as major product. In addition, Fujiwara et al. [3] obtained C3, C4, C5+ hydrocarbon over HY zeolite. 5.4. Mechanism of hydrocarbon formation : Effect of CO and ethene addition Addition of CO increased the hydrocarbon yield and also gave higher (C3+) hydrocarbon including aromatics for HZSM-5. This seemed to be due to the presence of water. In CO2 hydrogenation, methanol conversion into hydrocarbon and methanol decomposition into CO are competing with each other[5]. For methanol conversion reaction over hybrid catalyst, the methanol decomposition was enhanced and the formation of higher hydrocarbon was suppressed with water addition. As CO was added, the water could be removed by water gas shift reaction. This might give a chance for ethene intermediate to grow into higher hydrocarbon over HZSM-5. Ihm et al. [5] performed an analysis on the hydrocarbon product distribution based on the reaction mechanism by ethene co-reaction. In the case of HZSM-5 without ethene, ethane was obtained as a main product because water might prevent ethene from adsorbing on zeolite surface. When ethene was cofed with CO2 and H2, the selectivity to higher hydrocarbons increased tremendously and even aromatics were formed. Cofed ethene would easily adsorb on zeolite surface and propagate into higher hydrocarbons. Ethene should have participated in the consecutive growth of olefmic intermediates into higher hydrocarbon. In the case of SAPO-34, however, most of the cofed ethene were rejected without participating in the chain growth. This might be due to different mechanisms of hydrocarbons formation i.e., consecutive mechanism for HZSM-5 and carbon pool mechanism for SAPO-34 [14].

6. DIMETHYLETHER (DME) SYNTHESIS DME can easily be obtained by combining the methanol synthesis and methanol conversion processes in a single step to obtain oxygenates directly from CO/H2 or CO2/H2 [15]. The in situ transformation of methanol to DME would also shift the thermodynamic equilibria and thereby improve the total oxygenates yield (methanol + DME). The hybrid catalyst of Cu/ZnO/ZrO2 and HGa-silicate showed the high yield of oxygenates [15]. Dubois et al. [16] also demonstrated that combination of Cu/ZnO/Al2O3 catalyst with a solid acid, which enhanced the DME formation, indeed provided a strong driving force for CO2 conversion. The reaction conditions should be controlled accordingly.

32

7. CONCLUSION From the review of Cu based hybrid catalysts, some guidelines for catalyst design can be suggested as follows. 1. To control hydrocarbon yield: - Na species introduced during the catalyst preparation should be thoroughly removed by sufficient washing so that the acid sites of zeolites were not poisoned. - Metal oxide catalysts prepared by physical mixing would result in higher hydrocarbon yields than those prepared by coprecipitation. - Zeolite with higher acidity should be favored as a component of hybrid catalyst. This may result in higher hydrocarbon yield due to the formation of higher amount of medium acidity. 2. To control hydrocarbon distribution: - For olefin synthesis, a catalytic species with weaker hydrogenation ability such as chromate is preferred. - A methanol conversion mechanism over zeolite should be considered. - Addition of CO and ethene in the feed could increase selectivity of higher hydrocarbons over HZSM5 or HY hybrid catalyst. - For LPG synthesis, it is desirable to use SAPO-34 or SAPO-5.

REFERENCES 1. M. Fujiwara, H. Ando, M. Tanaka and Y. Souma, Bull. Chem. Soc. Jpn, 67 (1994) 546. 2. K. Fujimoto and T. Shikada, Appl. Catal., 31 (1987) 13. 3. M. Fujiwara, H. Ando, M. Tanka and Y. Souma, Appl. Catal. A:Gen., 130 (1995) 105. 4. J.K. Jeon, K.E. Jeon, Y.K. Park and S.K. Ihm, Appl. Catal. A:Gen., 124 (1995) 91. 5. Y.K. Park, K.C. Park and S.K. Ihm, Catal. Today, 44 (1998) 165. 6. Trimm, D.L, Design of Industrial Catalysts, Elsevier, The Netherlands (1980). 7. S. Sugawa, K. Sayama, K. Okabe and H. Arakawa, Energy Convers. Manage. 36 (1995) 665. 8. S.K. Ihm, Y.K. Park, J.K. Jeon, K.C. Park, and D.K. Lee, Stud. Surf. Sci. Catal., 114 (1998) 505. 9. M. Saito, T. Fujitani, M. Takeuchi, and T. Watanabe, Appl. Catal. A:Gen., 138 (1996) 311. 10. O.S. Joo, K.D. Jung, S.H. Han, S.J. Uhm, D.K. Lee, and S.K. Ihm, Appl. Catal. A: Gen., 135 (1996)273. 11. S.K. Ihm, S.W. Baek, Y.K. Park, and K.C. Park, Stud. Surf. Sci. Catal., 135 (2001) 277. 12. Y. Souma, M. Fujiwara, R. Kieffer, H. Ando, and Q, Xu, Stud. Surf. Sci. Catal, 114 (1998) 327. 13. Y.K. Park, K.C. Park, K.E. Jeong, J.K. Jeon, S.K. Ihm, and D.K. Lee, J. Korean Ind. Eng. Chem, 8 (1997) 140. 14.1.M. Dahl, S. Kolboe, J. Catal, 149 (1994) 458. 15. S.K. Ihm, S.W. Baek, Y.K. Park, and J.K. Jeon, ACS Symposium Series, 852 (2003) 183. 16. J.L. Dubois, K. Sayama, H. Arakawa, Chem. Lett, (1992) 1115.


33

Plasma Enhanced Preparation of Highly Dispersed Ni/Al2O3 Catalyst for CO2 Reforming of Methane Dang-guo Cheng, Qing Xia and Chang-jun Liu* ABB Plasma Greenhouse Gas Chemistry Laboratory, Key Laboratory for Green Chemical Technology of Ministry of Education and School of Chemical Engineering, Tianjin University, Tianjin 300072, RR. China FAX: +86-22-27890078. E-mail: [email protected] A special plasma catalyst preparation, which includes a glow discharge plasma treatment followed by calcinations thermally, was developed for CO2 reforming of methane. Such prepared Ni/AkC^ catalyst exhibits a highly dispersion of active species over the support and shows a better stability as well. The plasma prepared catalyst exhibits unusual high conversions at low temperature (lower than 873 K) but the best selectivities for hydrogen and carbon monoxide present at 973 K. TEM, XPS and TGA analyses indicate the carbon species formed during reaction on the plasma prepared catalyst is different from those reported in the literature over the conventional N1/AI2O3 catalysts. 1. INTRODUCTION Converting methane into CO and H2 (syngas) in a safe and efficient manner has drawn much attention recently because of the increasing demand for syngas, which is an excellent feed stock for the methanol and Fischer-Tropsch syntheses and for many other reactions like carbonlyation, hydrogenation and reduction.1'2 Another incentive comes from the facts that the synthetic fuel (hydrocarbon and alcohol) from syngas can help to reduce the environmental pollution, especially, at the crisis of the energy. The "dry" reforming of methane with carbon dioxide is an excellent pathway to produce syngas. Although this process is highly endothermic requiring an intensive energy input, it is of interest because the low H2/CO ratio (about 1/1) is generated, which is suitable for the Fischer-Tropsch synthesis to higher hydrocarbons, and also because CO2 (a major greenhouse gas) is consumed in a useful manner.2'3 It has been demonstrated that most of the GROUP VIII metals are more or less catalytically active towards the CO2 reforming of methane. Due to the inertness of methane molecule, a high reaction temperature, typically 1100 K, is normally required to get a reasonable yield to syngas.4 " 6 At such a high temperature, a conventional supported metal catalyst is often subjected to sintering and coking, which result in the deactivation of catalyst. Stability, however, is a very important factor that should be seriously considered for a practical industrial process. Among the catalysts developed, alumina supported nickel catalyst was extensively used for its relatively high activity and cheap cost. We previously reported a novel catalyst preparation using a glow discharge plasma treatment following by thermal calcinations for partial oxidation of methane 7'8, methane combustion 9'10 and NO reduction by methane u . A better low-temperature activity with an enhanced-stability was achieved over the glow discharge plasma treated catalysts, compared

34

to the catalysts prepared conventionally. In this work, we attempt to prepare the alumina supported nickel catalysts using this typical glow discharge plasma preparation to achieve a higher activity at lower temperatures and improved stability for CO2 reforming of methane. 2. EXPERIMENTAL The preparation and the plasma treatment of catalyst have been previously reported. 7 " u A argon glow discharge plasma was employed to treat the catalyst. The electrode configuration of glow discharge plasma is the same as that reported previously. During the catalyst treatment, the catalyst was located in the "positive column" of glow discharge 10 that was generated at ca. 12 Pa and at room temperature using argon as the glow discharge plasmaforming gas. Glow discharge is one of cold plasma phenomena that are characterized by having very energetic ("hot") electrons while the bulk of the molecules are still near ambient temperatures. The principal procedure of the catalyst preparation in this work can be described as follow: 1) impregnation; 2) drying; 3) glow discharge plasma treatment; and 4) calcination thermally. The conventional wetness impregnation has been employed in the first step. The AI2O3 powder was first impregnated in an aqueous solution of Ni(NC>3)2 for ca. 6 hours at room temperature. After impregnation, the obtained sample was treated by glow discharge plasma at room temperature for only 1 hour. After glow discharge treatment, the sample was calcined at 600 °C for 6 hours. All catalyst samples were then pressed, crushed, and sieved to a size of 40-60 mesh for the reaction. The loading amount of nickel on the alumina support is 9 wt%. The reaction of CO2 reforming of methane was carried out at atmospheric pressure in a 6 mm i.d quartz-tube fixed-bed reactor. A thermocouple was inserted into the center of the catalyst bed to measure the reaction temperature. The reaction time is 48 h for all the temperatures. Argon was used as the dilution gas during the reaction. The ratio of the feeding gases is 1:1:2 of methane/carbon dioxide/argon. The total flow rate is 40 ml/min, which, for 50 mg of catalyst loaded, corresponds to a space velocity of 48000 ml/hgcat. The reactants and products were analyzed using an online gas chromatography (Agilent 4890D) with a Porapak Q column. An ice-cold trap was placed between the reactor exit and the GC sampling valve to remove the water formed during the reaction. Catalyst characterization was conducted to analyze the carbon deposits of used catalysts and catalyst structure using XPS, TGA, XRD and TEM. XPS analysis employed a mutipass analyzer (PHI-1600) with pass energy of 187.85 eV for survey spectra and 29.35 eV for high-resolution spectra. X-ray was generated by using a Cu Ka X-ray source. And, X-ray diffraction (XRD), performed with a D/max 2000 (Rigaku) system using a filter Cu Ka radiation, was used to identify the bulk phase of the catalyst and the support. To measure the carbon species, thermogravimetric analyses (TGAs) were also performed in a TA-50 thermal analyzer. To do so, nitrogen (99.99% in purity) and air (99.99% in purity) at flow rates of 20 ml/min and 30 ml/min were introduced into the instrument, respectively. Sample (about 10 mg) were loaded into a thermobalance and heated to 800 °C at a heating rate of 15 °C/min until a constant weight was obtained. The weight loss was recorded simultaneously. TEM analysis was conducted using a JEOL JEM -100CXIITEM system. In this work, conversions of methane and carbon dioxide and selectivities of hydrogen and carbon monoxide were calculated according to the following formulas: X (CH 4 )%=(FCH 4 > IN X (CO 2 )%=(FCO 2 ,IN -

FCH 4 ,OUT)/FCH 4J IN X FCO 2 ,OUT)/FCO 2> IN X

100% 100%

35 S ( H 2 ) % = F H 2 , O U T / [ 2 X (FCH 4 ,IN - FCH4, OUT)] X 100% S ( C O ) % = F C O , O U T / [ ( F C H 4 J I N - FCH4, OUT) + (FCO 2 ,IN F; = Ci*FtotaI

FCO2>OUT)]X

100%

where X, S and F are conversion, selectivity and gas flow rate, respectively. Ftotai is the total feed gas flow rate or reaction effluent gas rate. C; is the molar fraction of i in the feed gas or reaction effluent gas, which is detected by GC. 3. RESULTS AND DISCUSSION 3.1. Stability test Figures 1 and 2 showed the effect of reaction temperature on the catalytic activity and the H2/CO in syngas, respectively. Unusual high conversions at low temperatures have been achieved, as shown in Figure 1. Within the temperature range tested, the conversions of methane and carbon dioxide reduce first with the increasing temperature. When the temperature reaches 973 K, the conversions reach the minimum. After that, the conversions

Figure l.The effect of reaction temperature on the catalytic activity (reaction conditions: CH 4 :CO 2 :Ar=10:10:20, space velocity=48000 ml/h-gcat, reaction time=48h,Ni 9% loading.) increase with the temperature. The interesting thing is the fact that we gained the minimum of conversion but the maximum of selectivity at 973 K. The ratio of hydrogen and carbon monoxide is also close to 1/1 at the temperature of 973 K. This indicated that the reaction temperature of 973 K is the best for CO 2 reforming of methane over the plasma treated catalyst. The different catalyst activities at different temperatures indicate too a different reaction mechanism. Es

Figure 2. The ratio of hydrogen and carbon monoxide versus reaction temperature

36

pecially,for the low temperatures reactions (below973 K), the reaction mechanism ca n be very different. It was considered tha t, at low temperatures, the high activity wouldbe obtained from glow discharge pi asma enhanced acidities as discussed prev iouslyin our investigation on catalytic met hane combustion 1O'U. The enhanced catal yst acidities promote the decomposition of methane. However, we did not observe as erious carbon deposit over used catalystsw ith the enhanced catalyst acidities. The us ed catalyst surface remains "clean" as sho wn in Figure 3. Evidently, some well-defi ned facets can be even observed in this TEM image. It is well known that the ca rbon deposition is the major reason for th Figure 3. TEM image of used catalyst e deactivation of the developed conventio taken after the stability test (reaction nal Ni/AbCh catalysts for the dry reformi conditions: reaction temperature=923 K, ng of methane. Obviously, the glow disch CH 4 :CO 2 :Ar=l:l:2, space velocity=48000 arge treated catalysts exhibit an excellent ml/h-gcat.) anti-carbon deposit performance. Figure 4 shows the conversions of methane and carbon dioxide and selectivities of hydrogen and CO obtained with the stability test at 973 K. Evidently, the catalyst prepared with glow discharge plasma treatment following by calcination is effectively stable. The catalytic activity is relatively high as well at 973 K. Further investigation is being conducted to improve the selectivity with the low temperature reactions (below 973 K).

Figure 4. Catalytic activity and stability of Ni/Al2O3 catalyst (reaction conditions: CH4:CO2: Ar=10:10:20, space velocity=48000 ml/h-gcat, reaction temperature=973 K, Ni 9% loading.)

37

3.2. Catalyst characterization The XRD analysis was carried out to investigate the structural characteristics of the catalyst and the support. Figure 5 presents the XRD patterns of the support and the catalyst. It is well known that there are three main peaks at XRD pattern of Ni/Al2O3 catalyst for CO2 reforming of methane, defined as Ni, NiO and NiAl2O4.12'13 But, very interesting, we did not found Ni peaks in the catalyst treated by argon glow discharge plasma. This indicates that Ni particles were probably highly dispersed over the alumina or presented in a non-crystal phase. Figure 5. XRD patterns of catalyst and support This is probably one reason that (0#Al2O3; l#:Ni/Al2O3) the catalyst used in this work achieved a very good catalytic activity at low temperatures. The previous investigations indicated that, for CO2 reforming of methane, methane decomposition is the main reaction that would even lead to carbon deposition. According to the previous XPS characteristics, there are four kind of carbon species on the Ni/Al2O3 catalyst surface for methane reforming with carbon dioxide. They were referred as inactive carbon at 281.4 eV, carbide carbon at 282.5eV, contaminate carbon at 284.6 eV and carbonate at 289 eV in XPS spectra. It is Figure 6. XPS spectra of the support, fresh and used the inactive carbon and catalysts carbonate that will lead to (0#:Al2O3; 1#: fresh Ni/Al2O3; 2#: reacted at 873K; 3#: deactivation of catalysts. reacted at 923K; 4#: reacted at 973K; Carbide carbon is in favor of CH4/CO2/Ar=l0/10/20; space velocity=48000 ml/h-gcat.) the reforming reaction, which take part in the formation of carbon monoxide. Figure 6 exhibits Cis spectra of used plasma treated catalyst taken after CO2 reforming at 873, 923 and 973 K. Upon Figure 6, at lower reaction temperatures (873 and 923 K), only carbide carbon and contaminated carbon present

38

on the used catalysts. No carbonate and inactive carbon species are observed. For the catalyst sample taken from the reaction at 973 K, no carbonate species is observed. The Cis peak contains the contributions from carbide and inactive carbon species. The XPS characterization confirms that the glow discharge treated catalyst can inhibit the formation of carbonate species during CO2 reforming of methane. This could explain the enhanced stability achieved in this work. The TGA analyses will give us further evidence with it, as discussed below. Zhang et al. reported that three types of carbonaceous species were found on the Ni/Al2O3 catalysts in TGA analysis, designated as C Q at 423 to 493 K, Cp at 803 to 873 K and C Y at >873 K. At low reaction temperatures (773 to 873 K), the Cp species is the majo r carbon species present on the surface (its amount corresponds to several mono-layers of equivalent carbon on Ni crystallites), along with small amounts of the C a species. However, the species C p can be changed into the C x species when time of exposure is longer than 2 h. The active C a species is responsible for the formation of syngas, while most of C Y species is responsible for catalyst deactivation. The Cp species is a surface poison or spectator at low reaction temperature (873 K). Figure 7 shows the TGA analysis of plasma prepared catalyst used at reaction temperature of 973 K. The C a and Cp carbonaceous species can be identified. However, no C Y species can be observed. This indicates that the glow discharge catalyst treatment described in this work can inhibit the formation of carbon species that will lead to the deactivation of catalyst. The same conclusion has been made from XPS analyses discussed above.

Figure 7. TGA spectra of catalyst after reaction at 973K (reaction conditions: CH4:CO2:Ar= 10:10:20, space velocity=48000 ml/h-gcat, reaction temperature = 973 K, reaction time=48h, the loading Ni=9%) Table 1 presents the mass percentage of each carbonaceous species in the used catalysts at different reaction temperature, calculated from TGA analyses. Evidently, the amount of C a species reduces slightly with the increasing reaction temperature. But, the amount of C p species is relatively high at the reaction temperature of 823 K, and it reaches a minimum at 873 K, then increases rapidly. This can explain why we obtained a relatively high conversion

39

of reactant and low selectivity of products at low reaction temperature (as shown in Figure 1). In addition, the TGA data matched with the XPS analysis well. Table 1. The amount percentage of three carbonaceous species on used catalysts at different reaction temperatures Reaction temp. /K 823

Co 6.5%

Cp 4.3%

873

5.6%

923

4.4%

1.6% 2.1%

973

4.3%

4.4%

CY

4. CONCLUSION The present investigation confirms that the plasma catalyst preparation described in this work can lead to a production of catalyst with a better stability for CO2 reforming of methane. The plasma prepared catalyst shows a very different performance with different temperatures. Unusual low temperature conversions have been achieved but the reaction temperature of 973 K is the best for selectivities for hydrogen and carbon monoxide at the present design of the catalyst. The Ni/AÔs catalysts prepared in this work exhibit an unusual dispersion characteristic. From hydrogen chemisorption (will be reported in the future work), XRD characterization and TEM analysis, no metal particles can be identified. One possibility is that the metal species is ultra-highly dispersed over the support. Another possibility is that the Ni species changes to amorphous species after the plasma treatment. Further investigation is needed to study this unusual dispersion performance. In addition, XPS and TGA characterization demonstrate the carbon deposition on the plasma prepared catalysts are different from that reported in the literature. No carbonate species was detected over the plasma treated catalyst. This indicates that the plasma catalyst treatment can inhibit the formation of carbonate species. Since the carbonate species will normally induce deactivation of the catalyst, the plasma preparation will lead to a better catalyst with a very good anti-carbon deposit performance. The TEM characteristics of the used catalyst has given us further evidence with it. Even after reaction, the plasma prepare d catalyst still exhibits some well-defined facets of the support and presents a "cl ean" catalyst surface. From the present understanding, the plasma treatment will indu ce a significant change in the interaction between the active metal species and the support. Further improvement is being conducted for an excellent low temperature activity with better selectivity of syngas. ACKNOWLEDGEMENT Support from Major Research Foundation of Ministry of Education of China (under the contract number Major#0212) is very appreciated. The TEM analysis was conducted under the assistance of Mr. Yutian Shen, Analysis Center, Hebei University of Technology in Tianjin, which is also very appreciated.

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REFERENCES 1. Bradford, M.C.J., and Vannice, M.A., Appl.Catal. A, 1996, 142, 73. 2. Lunsford, J.H., Catal. Today, 2000, 63, 165. 3.Tomishige, K., Kanazawa, S., Ito, S.-I. and Kunimori, K., Appl.Catal. A, 2003, 244,71. 4. Choi, J.-S., Moon, K.-L, Kim, Y.G., Lee, J.S., Kim, C.-H. and Trimm, D.L., Catal. Lett., 1998, 52,43. 5. Chang, J.-S., Park, S.E., Yoo, Y.W. and Park, J.-N., J. Catal, 2000, 195, 1. 6. Frusteri, R, Arena, R, Calogero, G, Torre, T. and Parmaliana, A., Catal. Comm., 2001, 2,49. 7. Wang,J.-G, Liu,C.-J., Zhang,Y.-P., Yu,K.-L., Zhu,X.-L. and He,R, Catal. Today, 2004, 89, 183. 8. Wang,J.-G., Liu,C.-J., Zhang,Y.-P., Zhu,X.-L., Zou, J.J. and Eliasson, B., Chem. Lett., 2002, 1068. 9. Liu, C.-J., Yu, K.-L., Zhang, Y.-P., Zhu, X . - L . , He, R and Eliasson, B . , Catal Commun., 2003, 4, 303. 10. Liu,C.-l, Yu,K.-L., Zhu,X.-L., Zhang,Y.-P., He,R and Eliasson,B., Appl. Catal. B, 2004, 47, 95. 11. Zhang,Y.-P, Ma,P.-S., Zhu,X.-L., Liu,C.-J. and Shen,Y-T., Catal. Commun., 2004, 5, 35. 12. Seok, S.-H., Choi, S.H., Park, E.D., Han, S.H. and Lee, J.S., J. Catal, 2002, 209,6. 13. Wang, S.B. and Lu, G.Q., Appl Catal. A, 1998, 169, 271. 14. Chang, J.-S., Park, S.-E. and Chon, H., Appl.Catal. A, 1996, 145, 111. 15. Zhang, Z.L. and Verykios, X.E., Catal. Today, 1994, 21,589.

Studies in Surface Science and Catalysis 153 S.-E. Park, J.-S. Chang and K.-W. Lee (Editors) ©2004 Published by ElsevierB.V.

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Synthesis of dimethyl carbonate by transesterification over CaO/carbon solid base catalysts Tong Wei, Mouhua Wang, Wei Wei, Yuhan Sun*, Bing Zhong State Key Laboratory of Coal Conversion, Institute of Coal Chemistry Chinese Academic of Sciences, Taiyuan, 030001 PR China Carbon supported CaO solid base catalysts, CaO/C and CaO-C were prepared by loading CaO on carbon through impregnation and pugging method, respectively. The catalytic performance of CaO-C prepared by pugging method was much higher than that of CaO/C prepared by impregnation method due to its larger effective pore diameter and smaller inner diffusion resistance. When the CaO-C composite was used in catalytic distillation reactor, PC conversion and DMC yield reached 100% and 98% at 337K with 0.3 h"1 of LHSV, respectively. 1. INTRODUCTION Dimethyl carbonate (DMC) is an important intermediate for polycarbonate resins as well as a useful carbonylation and methylation agent , and it is promising as a substitute for phosgene, dimethyl sulfate, or methyl iodide due to its negligible toxicity. DMC could be prepared by oxidative carbonylation of methanol, carbonylation of methyl nitrite or transesterification of cyclic carbonate with methanol [3'41. Because of the moderate reaction conditions and the avoidance of equipment corrosion, transesterification method (see Scheme 1) has attracted much attention in recent years. The transesterification of cyclic carbonate with methanol could be catalyzed by both acid and base, but basic catalysts were more effective . In our previous work, CaO prepared from the dissociation of CaCO3 at elevated temperature was found very effective for synthesis of DMC from methanol and propylene carbonate (PC) [6 '. However, ultra fine CaO prepared from CaCO3 were difficult to be filtered from the products and reused. Therefore, CaO were loaded on carbon by impregnation and pugging method to prepare carbon supported CaO-based solid bases in the present work, respectively, which were used for the reaction of propylene carbonate (PC) with methanol. The effects of preparation method on the structure of the solid base and subsequently their catalytic performance were investigated in detail.

*To whom correspondence should be addressed: E-mail: vhsun(a>,sxicc.ac.cn, weiwei(a),sxicc.ac.cn: Fax: +86-351-4041153 Tel: +86-351-4053801

42

2. EXPERIMENTAL 2.1 Preparation of catalyst CaO was prepared by heating CaC03 at 1173K for 1 h in N2 atmosphere. CaO/C catalyst was prepared by impregnating calcium acetate on active carbon (0.15-0.18mm) from its aqueous solution, followed by drying in air at 393K for 6 h and then calcined in N2 at 1073K for 1 h. The CaO content in CaO/C was 28.3 wt%. CaO-C composite made from linear phenolic resin 217 (provided by Tianjin Resin Company), hexamethylenetetramine and CaCO3 (weight ratio=5:l:5). After grounded and mixed homogeneously, the mixture was pumped at 20 MPa and solidified at 473K for 10 h in N2, and then broken into small particles ranging from 0.18 to 0.28 mm. The solidified mixture was activated at 1173K for 1 h inN 2 to prepare CaO-C composite catalyst. The CaO content in CaO-C composites was 50 wt%. For the comparison, phenolic carbon was prepared from linear phenolic resin and hexamethylenetetramine only with the ratio and procedure same as that of CaO-C. 2.2 Characterization XRD of samples was carried out in Rigaku D/max- y A using Cu target with Ni filtration. Pore distribution and surface area of the samples were determined with the BET method through Micromeritics ASAP-2000, and the BET surface area for CaO/C, CaO-C composite and phenolic carbon were 1012.0 m2/g, 288.8 m2/g and 207.7m2/g, respectively. CO2-TPD measurement was performed at a heat rate of 12K/min under He flow (50mL/min), and CO2 desorbed was detected by a BALZA Q-Mass spectrometer. 2.3 Evaluation of the catalyst The reaction was carried out in a batch reactor with the mole ratio of methanol to propylene of 4:1, and the different amount of catalyst samples (0.90 wt% (CaO), 2.80 wt% (CaO/C), 1.80 wt% (CaO-C)) were used to keep the same CaO content in the reactor. After the reaction proceeded for a certain time at expected temperature under strong stirring, the reactor was cooled down to room temperature. The reaction condition for catalytic distillation reactor was described in the caption of Fig. 7. The products were analyzed on a gas chromatograph with a TCD after centrifugal separation from the catalyst. 3. RESULTS AND DISCUSSION

43

3.1 Characterization of the catalysts It is well known that carbon is an inert supporter, which hardly reacts with an active substance. Therefore, CaO was loaded on carbon by impregnation and pugging method to prepare solid base catalyst with high performance for DMC synthesis by transesterification, respectively. The XRD patterns of CaO, CaO/C and CaO-C are illustrated in Figure 1. It indicated that the loading of CaO on carbon hardly changes its crystal structure, although CaO could be dispersed more homogeneously on carbon when the catalyst was prepared by impregnation method. Since base strength and basicity were the main factors that influence the catalytic behavior of solid base for this reaction, the base strength and basicity of CaO, CaO/C and CaO-C was determined by CO2-TPD (see Fig. 2 and Table 1). CaO and carbon supported CaO showed the same CO2 desorption at 923K, implying that the base strength of CaO hardly changed whatever it was supported on carbon. In addition, the basicity of CaO/C was much higher than that of pure CaO and CaO-C due to the homogeneous dispersion of CaO on carbon (see Fig.l and Table 1). These indicated that carbon supported CaO solid base catalysts, which possessed the same base strength as pure CaO, could be prepared by both impregnation method and pugging method. Moreover, CaO could be dispersed homogeneously on carbon by impregnation method and subsequently the basicty of the catalyst was improved remarkably. The pore distribution of CaO/C and CaO-C catalysts is shown in Fig. 3. For CaO/C catalyst, both micropore and mesopore were present with the volume ratio of micropores (0.44cm /g ) to mesopores (0.26 cm /g) was 1.67. This indicated that most active CaO was loaded on the microporous surface. As far as CaO-C catalyst was considered, it was suggested that only mesopores were effective although both micropores and mesopors with diameter were present. Since only micropores existed in phenolic carbon, the mesopores in CaO-C composite might result from the addition of CaCO3 in the starting materials. This was reasonable if CaCO3 dissociated into CaO and CO2 at 1173K, and then the effluent of CO2 led to the formation of mesopores.

Fig. 1. XRD pattern of CaO, CaO-C and CaO/C catalyst

Fig. 2.

CO2-TPD of CaO, CaO/C and CaO-C catalyst * based on CaO

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Table 1 C0 2 uptake of CaO, CaO-C and CaO/C catalyst Catalyst

C0 2 uptake/ (mmol/g)*

CaO

0.32

CaO/C

0.95

CaO-C

0.34

Fig. 3. Pore structure of CaO/C and CaO-C catalyst 3.2 Catalytic performance of the catalysts The catalytic performance of CaO/C and CaO-C are illustrated in Fig. 4. As we reported in our previous work '7 , CaO exhibited excellent catalytic performance for the reaction. At 353K, DMC yield was 43% after 2 h with CaO as catalyst. The high performance hardly decreased when CaO was loaded on carbon by pugging method, but DMC yield was only 4.5% when CaO/C used as catalyst. As mentioned above, the crystal structure and base strength, which were the main factors that affect the catalytic behavior of CaO, hardly changed when CaO was loaded on carbon, but the basicity of CaO/C was far higher than hat of pure CaO and CaO-C. Therefore, the catalytic performance difference should come from the mass transfer. CaO used in the present work were ultra fine particles with diameter about 10~20nm (see Fig. 5) and the reaction mainly proceeded on the surface of the particles under strong mixing. Whereas, for supported catalysts CaO/C and CaO-C, the reaction mainly took place on the inner surface. The diameter of effective pores of CaO-C catalyst was about 40 nm, while the effective pores in CaO/C were far narrower than that of CaO-C, mainly micropores and a small quantity of mesopores with diameter only about 4 nm. This led to the increase of inner diffusion resistance, and consequently the total reaction rate decreased remarkably. Furthermore, the reaction was carried out in liquid phase, and the molecule movement was slow in pores due to the strong interaction between molecules via H-bonds. Therefore, the inner diffusion was the main factor with CaO/C as catalyst, which decreased the reaction rate. Detailed investigation of inner diffusion on the reaction rate was discussed in other paper [7 ]. From the results above, it can be seen that when CaO was loaded on carbon by pugging method, it still showed high performance for transesterification of propylene carbonate with

45

methanol. To illustrate the reusability of the catalyst, CaO-C was reused two times (see Fig. 6). The catalytic activity hardly changed when CaO /C catalyst was used for three times. Thus, such a type of catalyst had the good stability. Transesterification of propylene carbonate with methanol was a reversible reaction, in which DMC yield is limited by thermal equilibrium. At 101.3kPa, azeotropic temperature of DMC and methanol was 337K. When the reaction was carried out on catalytic distillation reactor at this temperature, DMC could be removed from the catalyst as soon as it was produced, so the equilibrium could be pushed and consequently PC conversion and DMC yield were improved. It can be seen from Figure 7 that PC conversion and DMC yield reached 100% and 98 %, respectively, when PC LHSV was 0.3 h"1 at 337K. With the rise of PC LHSV, PC conversion and DMC yield decreased regularly due to the increase of feed rate. As a result, PC conversion was 78% at 1.2 h"1 of PCLHSV. By contrast, Jiang[8] et al used 12-tungstophosphoric acid supported carbon molecular sieves as catalyst in catalytic distillation apparatus. PC conversion was only 45% when PC LHSV was 0.01 h'1. This indicated that the CaO-C composite was an efficient and convenient heterogeneous catalyst for synthesis of DMC from PC and methanol.

Fig. 4. Catalytic performance of CaO, CaO/C and CaO-C catalyst. Temperature: 353K Time: 2h

Fig. 5. TEM of pure CaO catalyst

Fig. 6. Reusability of CaO-C Fig. 7. Effect of PC LHSV on PC conversion and DMC yield. Catalyst: CaO-C ( 50% CaO); catalyst catalyst. Temperature: 323K size: 0.9-2.0mm; reflux ratio: 4; reaction time: 12h;

46

4. CONCLUSION In conclusion, carbon supported CaO solid base could be prepared by loading CaO on carbon via impregnation and pugging method, respectively. The catalysts had the same crystal structure and base strength as pure CaO. However, the catalytic performance of CaO-C prepared by pugging method was much more higher than that of CaO/C prepared by impregnation method due to its larger pore diameter and then smaller inner diffusion resistance. In addition, CaO-C catalyst could be reused with little deactivation. When the CaO-C composite was used in catalytic distillation reactor, PC conversion and DMC yield reached 100% and 98% at 337K with 0.3 h"1 of LHSV, respectively. REFERENCE 1. 2. 3. 4. 5. 6. 7. 8.

Y. Sato, T. Yamamoto, Y. Souma, Catal. Letts, 65 (2000) 123. M. A. Pacheco, C. L. Marshall, Energy & Fuels, 11(1991) 2. P. Tundo, Pure Appl.Chem., 73 (2001) 1117. S. Fujita, B. M. Bhanage, Y. Ikushima and M. Arai, Green Chemistry, 3(2001) 87. Y. Ikeda, T. Sakaihori, K. Tomishige, K. Fujimoto, Cata. Letters, 66(2000) 59. T. Wei, M. Wang, W. Wei, et al, ACS symposium, No. 852 (2003) 175 T. Wei, M. Wang, W. Wei, et al, Green Chemistry, 5(2003) 343. Z. Y.Jiang, W.Yong, Pertochemical Technology (China), 30(2001)173.


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The Synthesis of Clean Fuels by F-T Reaction from CO2 Rich Biosyngas Jae-Seong Ryu,a Kyu-Wan Lee,b* Myoung-Jae Choi,aHee-Soo Yooc a

Korea Research Institute of Chemical Technology, P.O.Box 107 Yusong Daejeon 305-600, Korea b Yanbian University of Science & Technology (YUST), Yanji, Jilin, China c Department of Chemistry, Chungbuk National University, 48 Gaeshin-dong Heungduk-ku Chonju Chungbuk 361-763, Korea This paper presents the results of Fischer-Tropsch (F-T) reaction in a fixed bed reaction system with CO 2 rich biosyngas to produce clean fuels which contain no sulfur, nitrogen and aromatics. Generally CO2 is eliminated to get pure synthesis gas for the F-T reaction, but in this work, the authors did not separate CO2 which is the most effective global warming gas and investigated the role of CO2 in the reactions. In this study we targeted following points (i) how can we let not only CO but also CO2 react, (ii) how to suppress the water gas shift reaction, (iii) how to treat the effluent gas after reaction, and (iv) how to sustain the catalyst activity. To investigate such phenomena, we carried out the reactions over Fe-Cu based catalyst with various potassium content. In the reaction we used mainly a standard gas which is composed of CO:CO2:H2 = 2:1:4 vol.%. We found the best composition of catalyst and gas and optimal conditions to satisfy above targets. Furthermore we elucidated the causes of catalyst deactivation and thus proposed a two step reaction system to let more CO2 react. 1. INTRODUCTION Recently the biomass is highlighted as a tool to sink CO2 and as environmental friendly renewable and sustainable clean energy source because biomass is grown by photosynthesis with CO2 and water. Besides, the Fischer-Tropsch (F-T) synthesis from biomass is the only way to synthesize liquid and solid paraffin products containing alpha olefins without sulfur and nitrogen. Generally for the F-T reaction, CO2 was eliminated to get pure syngas, but in this work, we didn't separate CO2. The behavior of CO2 was investigated during the reaction in a fixed bed reactor from the point of view of WGSR (water gas shift reaction) suppression which leads to more CO to be reacted to the products, because CO2 is the most effective global warming gas and very stable compound to react. We have for a long time done the catalytic hydrogenation of CO2 into hydrocarbons [1] and methanol/DME [2]. 2. EXPERIMENTAL 1) Catalyst preparation The Fe-Cu-Al(Si) precursors were prepared by two steps, including co-precipitation and 'Corresponding author: [email protected]

48

impregnation. The details appeared elsewhere[3]. The catalyst composition was determined to beFe:Cu:Al:K= 100:6.6:15.7:x (x: variable) by ICP-AES. 2) Characterization of Catalysts The fresh, activated, and deactivated catalysts were characterized by BET surface analysis, Mossbauer spectroscopy[3], XRD, XPS, TPR, TPDC, elemental analysis, SEM, etc. 3) Reaction The reaction was performed in a continuous fixed bed reactor. The catalyst used for laboratory scale 0.1-1.Og and 20~40g in bench scale reactions. The details were published elsewhere [3,4] 3. RESULTS AND DISCUSSION 1) Characterization of catalysts The catalysts are characterized by BET adsorption behavior of CO, CO2 and H2. The results are summarized in Table 1. The surface area and the uptakes of CO2 and hydrogen were increased by the addition of alumina while those in silica supported catalyst were decreased more remarkably. Alumina accelerated the K dispersion more than silica. In our previous paper [4] and as shown in Table 1 and 2, the alumina supported catalysts converted CO at least more 90% although the CO2 conversion is very low. Therefore we used Fe/Cu/Al/K catalyst for the reaction. Table 1. Characterization of prepared catalysts BET area Catalyst CO CO2 2 (y mol/g) (m /g) Fe/Cu/Al/K(100/5/17/4) 95.5 209 4.4 Fe/Cu/Al/K(100/6/16/4) 89.5 248 Fe/Cu/Si/K(100/6/16/4) 209.0 8.9 H2S exposed* 47.0 * 20atm, 275°C for 30h and 1800ml/gcat. 30ppm H2S in standard gas

H2

11.6 9.8

2) F-T Reactions It is impossible to get hydrogen sufficient gas by the biomass gasification to react coproduced CO and CO2 completely, because of the high oxygen content of its composition, C:H:O = 6:12:5. And the composition of biosyngas is variable with the reaction conditions. Therefore, we tested several gas mixtures composed of different CO:CO2:H2 ratios. In these F-T reactions, we aimed four points as described in the abstract. To test the catalytic activity for standard biosyngas (gas 4), the hydrogenations of Ar:CO:CO2: H2 (5:26:13:54) were performed. As the author reported in the previous paper [4], in the hydrogen deficient condition, the reaction of CO2 was not occurred at 270 °C (step 1) but it reacts at 300 °C (step 2) in both hydrogen deficient and balanced gas composition with the similar product selectivities. That means that the reaction proceeds through reverse water gas shift (RWGS) reaction, therefore the products come from CO and not from CO2 [5]. Some representative F-T reaction results are summarized in Table 2 at various reaction conditions. From table 2, we could induce the following explanations.

49

(1) The effect of potassium content It is well known that potassium is a good promotor in F-T reaction [6]. We also tested the potassium effect as promotor focused on the aims as describe in abstract. Above 2% of potassium, CO conversion reached more than 90% and olefin selectivity of C2-C4 fraction was very high, more than 80%. While the methane selectivity decreased steadily below to 10% when potassium content in catalysts were over 4%. The WGSR was suppressed with increase of K content. At high K content (8%), the reaction did not occurre under 275 °C and 20 atm (run F8 of Table 2). And a small amount of methane and a large amount of the higher molecular weight fraction were observed comparing to the low potassium containing catalysts (see run Fl). On the other hand, the WGSR proceeded very well at lower pressure, 10 atm (see run F2) [7]. Therefore, the pressure above 20 atm and 4% potassium content are the favorable reaction conditions. Table 2. Reaction results obtained in fixed bed reactor ^ . , , cat. + Composition Temp. P D Rx Catalysts ^ (CQ. ^ . ^ f C ) ( a t m ) XCo Fl ^ " H " " ^

lg + 5g 2.0:1.0:4.0

260

20

X C02

SCH4

SCH8+

Olefm SCM

94.42 -53.30 15.91 29.68 45.90

F2

Fe

"£ U "'4)" K

1

8 + 5 § 2.0:1.0:4.0

275

10

97.12 -57.50 16.12 15.97 69.67

F3

Fe

"^l"^1"K

lg + 5g 2.0:1.0:4.0

275

20

96.21 -49.27 5.11

70.50 78.31

F4

Fe

"^ 1 ~^ 1 " K

lg + 5g 2.0:1.0:4.0

250

20

60.42-74.51 8.52

39.74 79.95

F5

Fe < U

" ^ "^ 1 " K lg + 5g 2.0:1.0:4.0

260

20

96.53 -48.43 4.48

74.35 80.01

lg + 5g 2.0:1.0:4.0

265

20

96.63 -37.75 4.31

75.87 83.25

lg+5g

2.0:1.0:4.0

275

20

93.22 -36.99 4.29

76.15 84.54 76.67 76.76

F6 ^'^'j}'^ F7

Fe
2~ radical anions are formed in the first electronation step and are adsorbed onto the zinc powder. The *CO2~ radical anions are

co 2

e >

'COz" .. c o it e .

> c o + co 3

2

"

I H+ + e" i

> CO+ OFT

I 4 f C ^ > -CH, Hi

- °

:..:«

:

> CH,

- >c2H.

i-CH 2 + 2H+ + 2eScheme 1. A possible mechanism for the reduction of CO2 using metal powders.

59 Table 1 Characteristics of metal powders Metal

Bond radius/A

Ionization potential/eV

Redox poteitial/£o(vs. SHE)

Mg

1.60

7.646

-2.37

Fe

1.24

7.870

-0.440

Zn

1.33

9.394

-0.763

probably adsorbed in C coordination, in which extra electron density is abundant on O atom. H+ in the solution may easily react with the O atom, whereas C atom is bonded to the surface of the zinc powder and the reaction with H+ is prevented. In this manner, CO2H is formed on the metal surface and will be further reduced by electron to CO. On the other hand, bond radius of iron is very similar to bond distance of C=O (1.22 A) as shown in Table 1. Therefore, it is considered that both O atoms in the »CO2~ radical anions can present on the surface of iron powder. Thus, H+ may easily react with the two O atoms followed by the formation of naked C atom. It is expected that the reaction of the naked C atom with H+ will lead the formation of hydrocarbons such as methane. Magnesium is the highest reducing power and ionization tendency, and the lowest ionization potential among three metal powders (Table 1). Therefore, magnesium may be able to reduce readily CO2 to methane due to supply much electron, although bond radius is longer than bond distance of C=O. In fact, amount of formed methane was higher than that of iron powder. Furthermore, the reduction rate of CO2 was fastest among the three metal powders. However, reduced product was only obtained methane as mentioned above. This might be also related to bond radius of magnesium, that is, electronation/protonation of a reactive methylene group may be a favorably reaction than dimerization of it because the each methylene group adsorbed on magnesium surface is relatively remote than proton much present its surroundings. In the case of iron powder, the methylene group lies close together. Therefore, dimerization of the methylene group may be occur and leads the formation of ethylene and ethane, though the major reaction is protonation as can be seen in Fig. 1. Although we expected the tentative mechanism for the reduction of CO2 in which the different reduced products were obtained according to metal powders, more detail reasons could not be completely understood. Furthermore, the reduction of CO2 was not allowed to continue with the reaction time because the surface of metal powders was coated by their oxides. The surface conditions of metal powders after the reduction were observed by XRD. However, XRD patterns only showed zero-valent metals; other metal states were not observed. Therefore, further studies are necessary to reveal the characteristics of metal powders and to obtain a highly reduction efficiency on the reduction of CO2. 4. CONCLUSIONS In the reduction of CO2 using metal powders, amounts and various kinds of reduced products could be revealed in the presence of various metal powders. In the presence of zinc powder, CO was main product and was produced about 200 umol corresponding to 20% of CO2 dissolved in the solution. In the chemical reduction of CO2 using iron powder, the main products from CO2 were methane, ethylene and ethane. However, the reduced product was

60

only methane in the presence of magnesium powder. In the case of iron and magnesium powders, amounts of methane obtained were 2.5 and 4 (amol, respectively. This proposed method is useful, selective and simple for the reduction of CO2. Furthermore, this research can contribute to the large-scale manufacturing of useful products from readily available and inexpensive raw materials. ACKNOWLEDGEMENTS A part of this work was performed at Satellite Venture Business Laboratory (SVBL), Mie University.

REFERENCES [I]

[2] [3] [4] [5] [6] [7] [8] [9] [10] [II] [12] [13] [14] [15] [16] [17] [18] [19] [20] [21]

CD. Keeling and T.P. Whorf, Atmospheric CO2 Records from Sites in the SIO Air Sampling Network. In Trends: A Compendium of Data on Global Change; Carbon Dioxide Information Analysis Center, Oak Ridge National Laboratory: Oak Ridge, TN, 1998. T. Inui, M. Anpo, K. Izui, S. Yanagida and T. Yamaguchi (Eds.), Advances in Chemical Conversions for Mitigating Carbon Dioxide; Elsevier: Amsterdam, 1998. K. Ito, S. Ikeda, T. Iida and A. Nomura, Denki Kagaku 50 (1982) 463. K. Ito, S. Ikeda, N. Yamauti, T. Iida and T. Takagi, Bull. Chem. Soc. Jpn., 58 (1985) 3027. S. Ikeda, T. Takagi and K. Ito, Bull. Chem. Soc. Jpn., 60 (1985) 2517. S. Kaneco, K. Iiba, K. Ohta and T. Mizuno, Int. J. Energy Environ. Econ., 7 (1998) 153. S. Kaneco, K. Iiba, K. Ohta and T. Mizuno, Energy Sources, 21 (1999) 643. S. Kaneco, K. Iiba, K. Ohta and T. Mizuno, J. Solid State Electrochem., 3 (1999) 424. S. Kaneco, K. Iiba, N. Hiei, K. Ohta, T. Mizuno and T. Suzuki, Electrochim. Acta, 44 (1999)4701. S. Kaneco, K. Iiba, S. Suzuki, K. Ohta and T. Mizuno, J. Phys. Chem. B, 103 (1999) 7456. S. Kaneco, K. Iiba, K. Ohta and T. Mizuno, Energy Sources, 22 (2000) 127. S. Kaneco, K. Iiba, M. Yabuuchi, N. Nishio, H. Ohnishi, H. Katsumata, T. Suzuki and K. Ohta, Ind. Eng. Chem. Res., 41 (2002) 5156. C.H. Liao, S.F. Kang and Y.W. Hsu, Wat. Res., 37 (2003) 4109. A. Ghauch, C. Gallet, A. Charef, J. Rima and M. Martin-Bouyer, Chemosphere, 42 (2001)419. D.R. Lide (Ed.), Handbook of Chemistry and Physics, 72nd ed., CRC Press, Boca Raton, FL, 1991. Kagaku Binran-Kiso (Handbook of Chemistry-basic in Japanese), 3rd ed., Chemical Society of Japan, Maruzen, Tokyo, 1984. G. Hochgesand, Ind. Eng. Chem., 62 (1970) 37. S. Kaneco, R. Iwao, H. Katsumata, T. Suzuki and K. Ohta, ITE Lett. Batt. New Technol. Med., 2 (2001) 640. T. Mizuno, A. Naitoh and K. Ohta, J. Electroanal. Chem., 391 (1995) 199. T. Mizuno, K. Ohta, M. Kawamoto and A. Saji, Energy Sources, 19 (1997) 249. A. Naitoh, K. Ohta, T. Mizuno, H. Yoshida, M. Sakai and H. Noda, Electrochim. Acta, 38(1993)2177.


61

Catalytic hydrogenation of CO2 to methanol over Pd/ZnO: Metalsupport interaction Cheol-Hyun Kima, Jae Sung Lee*b, and David L. Trimmc a

S-Oil Corp. R&D center, Ulsan, Korea 689-890,

b

Pohang University of Science and Technology(POSTECH), Pohang, Korea 790330,

c

The University of New South Wales(UNSW), Sydney, Australia

The catalytic hydrogenation of carbon dioxide to methanol over Pd/ZnO catalyst has been studied. The catalyst phase has been examined in depth using several spectroscopic techniques. When reduced above 523 K, ZnOx patches are formed on Pd metal surface by migration of partly reduced ZnOx, and are further reduced to Pd-Zn alloys as the reduction temperature is increased. Keywords

CO2 hydrogenation, Pd catalyst, strong metal support interaction

1. INTRODUCTION Pd/ZnO catalyst has been reported to be active for the methanol synthesis from CO + H2, for vapour phase ester hydrogenation, and for selective hydrogenation of 1,3-butadiene. In these various reactions, metal-support interaction as a result of the pretreatments (mostly reduction) or of the reaction conditions themselves was proposed to be the crucial factor. The purpose of the present study was to elucidate the behavior of Pd and ZnO in the Pd/ZnO methanol synthesis catalyst. This phase behavior has a critical effect on methanol yield in the catalytic hydrogenation of CO2 as reported earlier [l]. 2. EXPERIMENTAL Most of Pd/ZnO catalysts used in the present study were prepared by coprecipitation method. Targeted palladium loadings for prepared catalyst was 5 wt.%. Hydrogenation of CO2 was carried out in a high pressure reactor at various temperatures and 3.0 MPa. The products were analyzed by on-line GC(HP 5890) equipped with a TCD. In order to understand the catalysts system, the Pd/ZnO catalyst was characterized using temperature-programmed methods and spectroscopic techniques including X-ray diffraction(XRD), temperature programmed reduction(TPR), temperature programmed desorption(TPD), X-ray photoelectron spectroscopy(XPS), and X-ray absorption fine structure (XAFS).

62

3. RESULTS AND DISCUSSION The phase composition of the Pd/ZnO catalyst reduced at 523, 573, and 673 K has been investigated by X-ray diffraction(Fig. 1). Except for the sample reduced at 523 K (denoted by Pd/ZnO(523 K) hereafter), peaks of Pd metal could not be seen. At high reduction temperatures, the Pd-Zn alloy was the only Pd-containing phase observed at 28 = 41.5, 44.2° [2-4] and its peak intensity grows with reduction temperature. For Pd/ZnO(523 K), the peaks of Pd"Zn intermetallic phase are very small and a peak of Pd metal (28 = 40.1°) still exists, which means the formation of Pd-Zn alloy has just started. The TPR pattern in Fig. 2 supports above explanation. The formation of water at 353 K is attributed to the complete reduction of PdO and/or the reduction of ZnO to ZnOx(x=0-l). The large water peak which has a maximum at 650 K is ascribed to the formation of Pd'Zn alloy in agreement with the results of Hong et al.[4] and Iwasa et al.[5,6].

Fig. 1 XRD patterns of Pd/ZnO (5 wt.% Pd) catalyst made by coprecipitation and calcined at 773 K

Fig. 2 TPR pattern of Pd/ZnO (5 wt.% Pd) catalyst made by coprecipitation and calcined at 773 K(H2 flow rate= 50 ml/min, ramp rate= 20 K/min)

XPS analysis for Pd/ZnO catalysts reduced at different temperatures showed that binding energies of Pd 3d5/2 were between 335.4 - 335.7 eV(See Table 1). From inspection of binding energy values of Pd 3ds/2 core levels of several supported Pd catalysts reported in the literature, it could be concluded that binding energy values close to 335.0 eV are responsible for Pd metal species. For Pd/ZnO catalysts, several authors reported that the binding energy of Pd in PdZn alloy has a higher value by +0.6 eV[2,7,8]. Therefore the positive binding energy shift of Pd 3d5/2 of present catalysts by +0.4 - +0.7 eV most probably resulted from the formation of Pd-Zn alloys as seen in XRD patterns. On the other hand, Fleisch et al.[9] reported that a transfer of the excess charge of La to

63

Pd particles in a Pd"LaOx surface made Pd more electron-rich state(the maximum negative binding energy shift by decoration of Pd surface with LaOx patches was - 0.7 eV). Hence the binding energy shift of Pd 3ds/2 would increase as the fraction of ZnOx patches decreased when both alloy formation and ZnOx patches are present on Pd surface. For Pd/ZnO(523 K), Pd surfaces would be decorated by many ZnOx patches, which should have caused the negative binding energy shift of Pd as suggested by Fleisch et al.[9]. But the formation of Pd-Zn alloy takes place at the same time, which would increase the binding energy of Pd. Therefore the combined effects of ZnOx decoration and Pd-Zn alloy formation caused only small binding energy shifts of Pd by +0.4 eV. As the reduction temperatures increased, the ZnOx patches would be further reduced and contribute to the Pd-Zn alloy formation, which would increase the binding energy shift of Pd up to +0.7 eV as observed. This binding energy shift of Pd 3ds/2 can be taken as another evidence of ZnOx migration and Pd-Zn alloy formation taking place during catalyst reduction. The XPS/Ar+-sputtering experiments supported further the decoration of ZnOx patches and reduction to Pd-Zn alloy as reduction temperature increased(See Table 2). For the catalysts reduced at 523 K and 573 K, the Pd/Zn atomic ratio increased as the time of sputtering increased. This could be interpreted as indicating that sputtering by Ar+ removes the ZnOx patches which cover the Pd surface. However, for Pd/ZnO(673 K), there was almost no change in Pd/Zn atomic ratio, which could mean that no ZnOx patches remain on the Pd surface as a result of high temperature reduction. The decoration of Pd surface by partly reduced support was reported by Bell and coworkers[lO,ll] for the Pd/La2C>3 catalyst system. They reported that the decoration of Pd surface by LaOx patches caused a binding energy shift of Pd by -0.7 eV[9] and lowering of all CO vibrational frequencies compared to Pd/SiO2[l2]. They also showed the reduction of La2C>3 to LaOx by oxygen adsorption[l3]. Although the particles were relatively large(6 to 18 nm), metal-support interactions were observed. They suggested that reduced LaOx species were present on the surface of Pd particles, where they were believed to donate electrons to Pd[9] and change the relative populations of P d ( l l l ) and Pd(lOO) crystal planes[l2]. Fan and Fujimoto[l4] reported similar effects with Pd/CeO2 in that most of the Pd surface was covered by migrated Ce2Os(partly reduced form of CeO2). They compared Pd/Ce ratios between samples reduced at 673 K (Pd/Ce=0.254) and at 773 K (Pd/Ce=0.023). Likewise, for the present Pd/ZnO catalyst, the decoration of Pd surface by ZnOx patches seems to be the major reason for the decrease in Pd/Zn atomic ratio at the surface. The XANES spectra of Pd/ZnO catalysts reduced at 523, 573, 673, and 773 K and Pd reference materials are presented in Fig. 3 and Fig.4, respectively. The shape of XANES of Pd foil is quite different from those of other reference materials (ionic forms of Pd), i.e. the Pd foil has two strong oscillations at the post-edge region of the absorption spectrum. For Pd/ZnO(523 K), the spectrum has two oscillations although the amplitude is smaller than that of Pd foil. Thus, it could be said that the sample has some Pd metal character, which is consistent with the result of XRD(See Fig. 1). However, other samples reduced above 523 K

64

show quite different shapes of XANES spectra from those of Pd foil and of other ionic forms of Pd reference materials. Combining this with other spectroscopic results, it can be said that the XANES spectra represent those of Pd-Zn alloys and that XANES spectrum for Pd/ZnO(523 K) is a mixed spectrum of Pd metal and a Pd-Zn alloy. Table 1. XPS results of Pd/ZnO catalyst Binding Energy(eV) Pd

0

Binding Energy Shift of Pd5/2(eV)

523

335.4

531.0

+0.4

573

335.6

530.6

+0.6

673

335.7

530.6

+0.7

Reduction Temperature(K)

Table 2. XPS results of Pd/ZnO catalysts with Ar+ sputtering Reduction Temperature(K)

523

573

673

Sputtering Time(min)

Pd/Zn atomic ratio

0

0.0048

2

0.0056

4

0.0066

10

0.0077

0

0.0031

2

0.0038

4

0.0046

10

0.0061

0

0.0058

2

0.0062

4

0.0060

10

0.0065

The small oscillations in absorbance present at 100 — 1000 eV above the absorption edge are isolated from background absorption and Fourier transformed to yield the radial structural function(RSF). Fig. 5 shows the RSFs of palladium reference materials. The EXAFS oscillation function reflects the contribution of neighbor atoms. The peak position in RSF corresponds to an

65

interatomic distance between absorbing and surrounding scatterer atoms displaced from the true distance by a phase shift and the peak intensity is correlated to the average coordination number for the atom at the distance. The Pd-Pd distance in Pd foil appears at 2.55A, which is smaller than that of literature value of 2.75A due to the phase shift. Fig. 6 shows the RSFs of Pd/ZnO catalysts reduced at various temperatures. Two peaks are observed for Pd/ZnO(523 K) with peak maxima located at 2.15Aand 2.52A. Only one peak can be seen for other samples in the range of 2.29A - 2.30A. The interatomic distance of 2.52 A for Pd/ZnO(523 K) is attributed to that of Pd-Pd in Pd metal based on the Pd foil reference value(2.55A). Peaks in the range of 2.15 A - 2.30 A are most probably due to the interatomic distances of Pd-Zn in Pd-Zn alloys. Slight changes in the Pd"Zn interatomic distance for samples reduced at different temperatures are ascribed to the different compositions of Pd and Zn in Pd-Zn alloys. The existence of two peaks in Pd/ZnO(523 K) and the new peak formation which can be assigned to Pd-Zn interatomic distances are more convincing evidence of Pd-Zn alloy formation. The results are consistent with those of XRD, XPS, and XANES.

Fig. 3 XANES spectra of Pd K edge for Pd/ZnO catalyst(5 wt.% Pd) catalyst made by coprecipitation and calcined at 773 K

Fig. 4 XANES spectra of Pd K edge for Pd reference materials

4. CONCLUSION The Pd/ZnO catalyst shows a quite strong metal-support interaction when reduced above 523 K. A schematic model of Pd/ZnO catalyst can be represented by the mixed form of Pd decoration by ZnOx and Pd-Zn alloy formation. Upon reduction, ZnOx patches are formed by migration of partly reduced ZnOx to Pd metal surface and are further reduced to Pd-Zn alloys as the reduction temperature is increased. As mentioned, we have found that there is a good correlation between the amounts of ZnOx patches and the yield of methanol,

66

which has been discussed elsewhere [l].

Fig. 5 Fourier transform of Pd reference materials

Fig. 6 Fourier transform of Pd/ZnO catalysts

Acknowledgement This work has been supported by National R&D Project for Nano Science & Technology, Research Center for Energy Conversion and Storage, and BK-21 project. REFERENCES [1] C.-H. Kim, J.S. Lee, and D.L. Trimm, Topics Catal. 22 (2003) 319. [2] Z. Zsoldos, A. Sarkany, andL. Guczi, J. Catal., 145(1994) 235. [3] A. Sarkany, Z. Zsoldos, B. Furlong, J.W. Hightower, and L. Guczi, J. Catal., 141(1993) 566. [4] C.-T. Hong, C.-T. Yeh, and F.-H. Yu, Appl. Catal., 48(1989), 385. [5] N. Iwasa, S. Masuda, N. Ogawa, and N. Takezawa, Appl. Catal. A, 125(1995) 145. [6] N. Iwasa, T. Mayanagi, N. Ogawa, K. Sakata, and N. Takezawa, Catal. Lett., 54(1998) 119. [7] M.L. Cubeiro and J.L.G. Fierro, J. Catal., 179(1998)150. [8] M.L. Cubeiro and J.L.G. Fierro, Appl. Catal. A, 168(1998) 307. [9] T.H. Fleisch, R.F. Hicks, and A.T. Bell, J. Catal., 87(1984) 398. [10] Y.A. Ryndin, R.F. Hicks, A.T. Bell, and Y.I. Yermakov, J. Catal., 70(1981) 287. [11]R.F. Hicks and A.T. Bell, J. Catal., 90(1984) 205. [12] R.F. Hicks and A.T. Bell, J. Catal., 91(1985) 109. [13] R.F. Hicks, Q-J. Yen, and A.T. Bell, J. Catal., 89(1984) 498. [14] L. Fan and K. Fujimoto, J. Catal., 172(1997) 238.


67

CAMERE Process for methanol synthesis from CO2 hydrogenation Oh-Shim Joo*, Kwang-Deog Jung and Yonsoo, Jung Eco-Nano Research Center, Korea Institute of Science and Technology (KIST), P.O. Box 131, Cheongryang, Seoul 130-650, Korea. Tel: +822-958-5215, FAX: +822-958-5219, E-mail: [email protected] A pilot plant to produce methanol by CO2 hydrogenation has been constructed with the government research fund in participating with POSCO and Korea Electric Power Research Institute (KEPRI). The pilot plant is consisted of a reverse water gas shift reactor and a methanol reactor. Two reactors are serially connected to remove water in the first reactor and then, synthesize methanol in the second reactor. The production capacity of the plant is 100 kg methanol per day. The methanol production yield in CAMERE Process is higher than twice as compared with the yield in the direct hydrogenation of CO2 into methanol without reverse water gas shift reaction. The pilot plant for methanol synthesis from CO2 was combined with the pilot plant for separation of CO2 discharged from a power plant. 1. INTRODUCTION The CAMERE process (Carbon dioxide hydrogenation to form methanol via a reverse-water gas shift reaction) was developed to fix CO2 into methanol. The reverse water gas shift reactor and the methanol synthesis reactor are serially aligned to form methanol by CO2 hydrogenation. Carbon dioxide is firstly converted to CO and H2O via the reverse water gas shift reaction (RWGSR) and then, the water is removed from the reactant gas before injection into the methanol reactors. The higher the conversion of CO2 to CO is in the RWGSR, the higher the methanol productivity can be increased, because methanol yield is dependant on the CO concentration in the CO2/CO/H2 mixture gas. Therefore, the volume of the recycle gas in the methanol synthesis reactors can be minimized by increasing the conversion of CO2 to CO in RWGSR as compared with the direct CO2 hydrogenation into methanol.1 The RWGSR should be carried out at higher temperature than 600 °C to obtain CO2 conversion over 60 % in the thermodynamic point of view.2 Therefore, development of an active and stable catalyst for the RWGSR at higher temperature than 600 °C was a critical requirement for the CAMERE process.

68

The water-gas-shift reaction has been studied intensively for the last several decades in order to adjust H2/CO ratio in the synthesis gas. On the contrary, a reverse-water-gas-shift reaction has attracted little attention because of little demand. The Fe2O3/Cr2O3 catalyst is a well-known commercial catalyst for the water-gas-shift reaction. The commercial catalyst, Fe2O3/Cr2O3, was not a good candidate for the RWGSR of the CAMERE process because of the severe deactivation.3 Deactivation of the catalyst was attributed to the reduction of Fe2O3 to the Fe metal. On the other hand, a new type of catalyst, Z11AI2O4, for the RWGSR was developed, which showed good activity and stability without coke formation.4 In this paper, ZnAl2O4 catalyst was optimized for RWGSR and the pilot plant was operated using the ZnAl2O4 and Cu/ZnO/Al2O3 catalysts. 2. EXPERIMENTAL The ZnAl2O4 catalyst (Zn:Al=l:2 in molar ratio) was prepared by a co-precipitation of the corresponding metal nitrates.5 The catalyst was calcined at 850 °C before the RWGSR. The activity of the ZnAl2O4 catalyst was studied in a test reactor of 3/8 "o.d. for the RWGSR at the temperature range of 400-750 °C with different GHSV (ml/gcat.h). The RWGSR was carried out at atmospheric pressure in the mixed gas with H2/CO2 ratio of 3 in the reaetant gas. The concentration of the exit gas was measured with a gas chromatograph (Shimadzu 8 A instrument with a TCD, Porapak Q column l/8"o.d., 3m long in program temperature mode for analysis of CO, CO2, H2, CH4 and H2O. In the pilot plant, the RWGSR is operated on ZnAl2O4 in the temperature range of 600-700°C and at atmospheric pressure in the pilot plant and methanol is synthesized on CuO/ZnO/Al2O3 under the reaction conditions of 250-300°C and 50-70atm. The feed gas of CO2/H2 (1/3) mixture gas is preheated before RWGSR. The RWGSR is endothermic reaction and three electric heaters are used for the reaction temperature control. The reactor size is 2"(I.D.) x 120cm (length). The methanol synthesis from CO2/CO/H2 mixture gas is exothermic reaction and the reaction temperature is controlled to minimize the hot spot by the exothermic reaction by steam. Four fixed bed reactors for multiple purposes are used for methanol production. A reactor size is 2"(I.D.) x 100cm (length). A two-stage compressor (Diaphragm compressor, Sera MV 4714-IIK) is used for compressing the CO2/CO/H2 mixture gas before injection into the methanol synthesis reactor. The concentrations of gas stream at four points of the pilot plant were monitored with two gas chromatographs and water content in final products was analyzed using karl-fisher (658KF). 3. RESULTS AND DISCUSSION

69

Z11AI2O4 catalysts were prepared by a coprecipitation and the activity of the prepared catalysts for RWGSR was dependent on the pH of the solution (Figure 1). The Z11AI2O4 catalyst prepared at pH=7 shows the highest activity as compared with catalysts prepared at acidic condition. Especially, Z11AI2O4 catalyst prepared at pH =5.4 shows lower activity and broaden diffraction pattern. We obtained the ZnAl2O4 catalyst with the highest activity at pH=7.

Figure 1. CO2 conversion and the X-ray diffractogram of ZnAl2O4 catalyst depending on the preparation pH. (a) pH=5; (b) pH=6.0; (c) pH=6.4; (d) pH=7.0 Figure 2 shows CO2 conversion with respect to GHSV over ZnAl2O4 catalyst prepared at pH=7. The dashed line is the equilibrium conversion for RWGSR. When the reaction temperature is increased, CO2 conversion over Z11AI2O4 catalyst approaches to the equilibrium

conversion.

Most oxide catalysts show high catalytic activity for RWGSR

at

atmosphere

pressure, but are deactivated

rapidly

because

the

RWGSR condition is very reductive above 400°C and the

reactant

ratio

of

23

H2/CO2=3/1. ' Moreover, it should be operated above 600°C to obtain higher CO2 Figure 2. CO2 conversion over ZnAl2O4 with GHSV

conversion

than

60%.

70

Therefore, the stability of the catalyst at the high temperature is very important in the practical point of view. The activity of Z11AI2O4 was rarely decreased at 700°C operation for 9 days with feed rate of 150,000 ml/gcat.h so it turns out to be very stable catalyst as compared with Fe2O3/Cr2O3 and C^CVA^C^ oxide catalyst.3'4 The pilot plant for CAMERE process to obtain methanol from CO2 hydrogenation was constructed based on the detail PFD (Process Flow Diagram) and P&ID (Piping and Instrument Diagram). Figure 3 shows the simple schematic process flow diagram and Figure 4 shows the picture for CAMERE process.

Figure3. A simple schematic process flow diagram for CAMERE Process C0 2 and H2 is mixed in D-105 for the H2/CO2 ratio of 3, which is

injected

into

the

reactor

through HE-101 for RWGSR. After RWGSR, the produced water is removed through HE102 before injection into the diaphragm

compressor.

The

mixed gas of H2/CO2/CO is compressed into the operation pressure for methanol synthesis. Figure 4. Picture for the pilot plant.

71

And the mixed gas from the RWGSR with part of recycled gas is fed into methanol reactor of R-201. Part of recycled gas is fired in the FN-101 to regulate the reaction conditions. The temperature of the four fixed bed reactors was well controlled by steam in the temperature range of 250-300°C for methanol production. The pilot plant has been operated to obtain the optimum reaction conditions and the data for the evaluation of the methanol production cost. To evaluate the effect of the RWGSR on the methanol production yield, the pilot plant was operated in the RWGSR (On) or RWGSR(Off). RWGSR (On) means that the RWGSR was operated at the temperature range of 600-700°C. On the contrary, RWGSR (Off) means that the RWGSR was not operated during methanol production so H3/CO3=3/1 was just injected into the methanol reactor. Table 1 distinctively shows the effect of the reverse water-gas shift reaction on the methanol production yield. The methanol yield in the RWGSR (On) becomes more than twice in comparison with the yield in the RWGSR (Off) at the same reaction conditions. Moreover, the CO2 conversion over the ZnAbO4 catalyst was about 35% in the RWGSR (On). It is worth noting that the reverse water-gas shift reaction shows a significant effect on the increasing of the methanol yield.

RWGSR(On)

RWGSR(Off)

CO2(kgmol/day)

3.34

3.35

3.35

3.34

3.35

3.35

H2(kgmol/day)

10.04

10.09

9.94

10.18

10.15

10.15

CH3OH(kg/day) /100% Yield

107.1

107.3

107.3

107.1

107.3

107.3

CH3OH(kg/day)/ Pilot plant

71.67

72.58

75.46

35.95

34.87

39.82

CH3OH Yield

66.9

67.63

70.46

33.57

32.49

37.11

Pressure(atm)

51.0

51.0

60.7

61.1

61.3

70.9

Table 1. CH3OH Yield for CAMERE process based on RWGSR(On) or (Off) CAMERE Process was simulated using the simulation program of Aspen Plus and the state equation of UNIFAC based on the 2000 tons methanol production in a year. The construction cost for the plant was evaluated based on the Guthrie's Modular Method. The methanol production cost was calculated depending on the hydrogen cost and methanol production capacity as shown in the figure 5. The methanol production cost proportionally increases with hydrogen cost in market and dramatically decreases up to 50,000 tons of methanol production capacity and then, becomes stable with the production capacity. It indicates that methanol can be produced with 300US$/ton from CO2 hydrogenation

72

through CAMERE process if a commercial plant of 100,000ton/year is constructed. It also means we cannot economically produce methanol from CO2 hydrogenation because the methanol is sold at 100-150 US$/ton in the recent market. To become an economical process, the target material of the CAMERE process should be changed into another one having a value added, for example, DME (Dimethyl ether). On

the

other

hand,

the

methanol production cost of 300US$/ton

calculated

here

would be a standard value for carbon dioxide sequestration process. In addition, whenever the carbon tax starts to work, CAMERE sequestrate

process carbon

to dioxide

should be evaluated based on the real situation. Figure 5. Methanol production cost via CAMERE Process depending on hydrogen cost and methanol production capacity. 4. CONCLUSIONS Methanol yield of 70 % was obtained from the pilot plant for CAMERE process. Methanol of 75kg was produced in a day from the pilot plant for which about the 100kg of CO2 was consumed. Based on the results, we estimated the methanol production cost depending on the hydrogen cost and methanol production capacity. Operating cost of about 300USS was requested for lton methanol production through CAMERE process. REFERENCES 1. Joo, O.S. et al, Ind. Eng. Chem. Res., 38(5), 1808, 1999. 2. Park, S.W., Joo, O.S., Jung, K.D., Han, S.H, Appl. Catal. A:general, 211, 81, 2001. 3. Park, S.W., Joo, O.S., Jung, K.D., Kim, H., and Han, S.H, Kor. J. of Chem. Eng, 17(6), 719,2000. 4. Joo, O.S., and Jung, K.D., Bull. Korean Chem. Soc, 24(1), 86, 2003. 5. Joo, O. S., Jung, K. D., Han, S. H., Uhm, S. J., Lee, D. K. and Ihm, S. K., Appl. Catal. A: General 135, 273, 1996.


73

Product Distribution Analysis for Catalytic Reduction of CO2 in a Bench Scale Fixed Bed Reactor Sang-Bong Lee,a Jun-SikKim,a Won-Young Lee,a Kyu-Wan Lee,b and Myoung-Jae Choia* "Advanced Chemical Technology Division, KRICT, PO Box 107, Taejon 305-600, Korea Yanbian University of Science & Technology (YUST), Yanji, Jilin, China *E-mail: [email protected]

b

Hydrogenation of CO2 was carried out over Fe-K/AbOj catalyst in a bench scale fixed bed recycle reactor with an aim to get higher molecular weight hydrocarbons including mainly olefinic fractions. Conversion of CO2 reached 88.2% at the recycle ratio of 5 where reaction temperature, space velocity, and H2/CO2 ratio were 300°C, 2,000ml/gcat.hr and 3, respectively. On increasing recycle ratio, selectivity to C2-C4 olefins tends to decrease and selectivity to higher branched hydrocarbons was increased via C2-C4 olefins oligomerization. Overall hydrogenation mechanism of CO2 was proposed on the basis of the product distribution of oily and aqueous phases. 1. INTODUCTION Direct hydrogenation of CO2 which is called here as modified Fisher-Tropsch (MFT) reaction over Fe based promoted catalysts have been highlighted because of no sulfur and nitrogen containing liquid fuels and chemicals. Similar product distribution of Fisher-Tropsch reaction could be obtained via one-step reaction [1-5]. In our previous work, screening of catalyst, parameter optimization for increasing Cs+ alpha olefins, and de/reactivation of catalysts have been investigated from a process development point of view. Meanwhile, it has been a big problem to separate H2, CO2, CO, and light hydrocarbons, especially CH4, due to its difficulty and high cost. In our process development study for technology of oil production for liquid fuel or chemicals, more simple and economically feasible process is needed. Our suggestion to solve this problem is to skip separation of H2, CO2, CO, and light hydrocarbons but to recycle light gaseous fractions lighter than C5 fraction aiming at liquid fuels or chemicals without sulfur content. In this paper, CO2 hydrogenation in a bench scale fixed bed reactor with recycling of gaseous product and product oriented overall mechanism of MFT reaction are described. 2. EXPERIMENTAL 2.1. Catalyst preparation The catalyst Fe-K/Al2O3 was prepared in a large scale by previously reported method [4] except palletizing. Regular 2x5mm size of the catalyst was prepared by extrusion using 1 wt% of PVA (poly vinyl alcohol) as a binder. 2.2. Hydrogenation of CO2 in a bench scale fixed bed system (Fig 1) 252 g of catalyst was loaded in a fixed bed reactor and was reduced with sufficient H2 for

74

1. Mass flow controller 2. Fixed bed P' reactor 3. Fixed bed 2"J reactor 4. Electric heater 5. Gas-liquid separator 6. Condenser 7. Heat exchanger 8. Back pressure regulator 9. Gas compressor 10. Buffer tank 11. Pressure regulator 12. Wet gas meter

Fig. 1. Flow Diagram of bench scale fixed bed system for CO2 Hydrogenation. 24 hrs at 450°C. CO2 and H2 were passed through the reactor using MFC at elevated temperatures. Pressure was controlled with the aid of BPR. Liquid and oil were trapped in a 3L tank at 5°C and uncondensed gases and hydrocarbons were compressed and kept in a 20L tank at 20 atm and recycled to the reactor according to the flow rate proposed using MFC. Conversion of CO2 (Xco2) and selectivity to CO and CH4 were determined by a gas chromatography on Carbosphere column and TCD by analyzing vent gas. N2 was used as internal standard gas. Light hydrocarbons were analyzed on Poraplot-Q column and FID using the analyzed CH4 as an internal standard gas. Liquid oils and water-soluble organics were analyzed on DB-1 and HP-5 column, respectively. 2.3. Separation of oil product on column chromatograph 5 g of oil product was separated in a silica gel column chromatograph 5 cm-ID x 60 cm-H according to a typical method using hexane and ethylacetate as effluent. Compositions of the separated 4 kinds of fractions with different polarity are analyzed by GC/MS and IR. Carboxylic acids were separated as Na salts followed by acidification with HC1. 3. RESULTS AND DISCUSSION 3.1. Preparation of catalyst and activity test As shown in Table 1, the composition of PVA-bound catalyst was in the proportion of Fe:K:Al2O3=20:7:100 by weight. The BET surface area was measured to be 92 m2/g. The CO2 and H2 uptake were found to be 514 and 20|imol/g of catalyst, respectively. These results were similar to that of reported conventional pressurized catalyst and were expected to show similar activity. 1 wt% of organic binder (PVA) was enough to make uniform extrudate and to give proper strength. Activity test of the new PVA-bound catalyst was conducted under the same conditions as reported elsewhere [6]. Almost the same conversion of CO2 (38%) was obtained, and therefore the catalyst prepared in large scale using PVA binder was confirmed to be usable as like pressure-formed catalyst.

75 Table 1. Composition and properties of Fe-K based catalyst Catalyst

Composition [wt%]

BET surface area [m2/g]

CO 2 uptake [umol/gl

H2 uptake ||nmol/g|

Fe-K/Al 2 O 3

Fe :K:Al2O3 = 20:7: 100

92

514

20

3.2. Mechanism of hydrogenation of CO2 based on the oil product composition Oil layer collected in liquid product separator was separated in 4 spots whose Rf were 0.95, 0.4, 0.2, 0.05, respectively. The amount of the each fraction separated by column chromatographic technology was 85, 3, 5, and 7wt%, respectively. GC and FT-IR data of them were summarized in Figures 2 and 3. The main components of the first fraction analyzed by GC/MS are found to be C6-C25 a-olefins and contains no aromatic hydrocarbons as there is no absorption band around 3,000 cm"1 in FT-IR. The second fraction was found to contain largely C5-C20 2-ketons and aldehydes mixture, judging from GC/MS and the characteristic absorption band of carbonyl compound at 1730 cm"' in FT-IR. The third fraction was found to contain largely C5-C20 alcohols using the same analysis method. In aqueous layer, largely light C1-C5 aldehydes, ketones, alcohols, and acids were detected due to the low solubility of higher ones. At low pressure and contact time, the amount of oil layer tended to decrease and the amount of aqueous layer tended to increase. Their GC and FT-IR data are summarized in Figs 2 and 3, respectively.

Fig. 2. GC of each fraction separated from Fig. 3. FT-IR spectrum of each fraction oily and aqueous phase products. from oily and aqueous phase products.

76

Fig. 4. Plausible overall reaction mechanism of CO2 hydrogenation.

From these results overall plausible mechanism of CO2 hydrogenation is summarized in Figure 4. In the beginning CO 2 is reduced by iron (II) followed by H radical abstraction on the catalyst surface. When the residual H radical attacked carbonyl C and OH, formic acid and CO are formed, respectively. Aldehydes and alcohols can be produced in the next stage. By the same manner Fe-CH2 radical is formed and is regarded as carbon-carbon propagation species. In our reaction system chain propagation is considered as a major route because higher hydrocarbon is major product. Higher a-olefin selectivity to paraffin is attributable to less H2 uptake and no excess H2 in this reaction system. 3.3. CO2 hydrogenation in the fixed bed recycle reactor. As mentioned in the introduction, recycle of gaseous fraction lower than C\ was investigated in order to avoid difficult separation process. As shown in Fig. 5, conversion of CO2 was tended to increase on increasing recycle ratio. Selectivity to CO was tended to decrease probably due to rather high concentration and reactivity of CO than CO2 in the recycling gas as compared to fresh gas. Conversion might be reached above 90% if recycle can be increased above 6, however, continuous feeding was difficult due to the lack of the recycle gases above recycle ratio of 3 where conversion was higher than 70%. Lite-time test of the catalyst was conducted at recycle ratio of 3 and SV of 4,000 due to the reason described above. The catalyst activity maintained for 1,000 hrs (Fig. 6). In order to predict the effect of recycle ratio more precisely, product distribution comparison was summarized in Table 2. Conversion of CO2 in a single reactor was slightly increased to 40.8% at reduced SV 1,000, i.e. low SV is one of a tool to increase conversion. When a series reactor was used having the similar concept of recycle reactor, the CO2 conversion drastically increased to 68.5%. In this case, however, pressure drop was to big to continue the reaction at constant pressure probably due to the additional catalytic layer and

77

this prompted us to investigate recycle reactor. In case of recycle reactor A (R=3) and B (R=6) where SV=4,000, clear increase of CO2 conversion and C$+ selectivity was observed and decrease of CO and light olefin selectivity was observed in comparison with that of single reactor. In case of C (R=5) at low SV of 2,000, CO2 conversion reached 88.2% and also clear decrease of olefin ratio was observed. In comparison of A, B, and C, CH4 and light paraffin and tended to increase while CO, light olefins, and C5"1" tended to decrease as increasing CO2 conversion. These results are explainable on the assumption of accumulation of methane and paraffin in the recycle gas and olefin oligomerization over this catalyst system. The branched hydrocarbons shown in Fig.7 could be one of the evidences for this assumption. A plausible mechanism can be suggested as shown in Fig. 8.

Fig. 5. Effect of recycle ratio on CO2 Fig. 6. Evaluation of long life activity for Conversion. Fe-K /AI2O3 catalyst.

Table 2. Effect of recycle ratio on CO2 Conversion, CO selectivity, hydrocarbon distribution, and olefin ratio Reactor Type Single Series Recycle-A Recycle-B Recycle-C

XcO2

40.8 68.5 62.3 75.6 88.2

Sco 11.1 13.7 4.3 2.2 2.1

c,

7.4 10. 1 4.8 7.3 12. 1

c2

=

5.8 4.9 3.6 3.8 2.7

Hydrocarbon c 2 c3= 1.4 9.6 2.4 6.5 1.1 5.8 2.2 7.4 3.3 4.8

distribution c 3 c4= 1.1 7.0 1.1 3.4 0.8 4.1 1.2 3.8 2.2 3.2

c4 0.9 0.8 1.0 0.8 1.3

>c 5

55.7 57. 1 74.5 71.3 68.3

O/O+P [%] 86.8 77.5 82.3 78.1 61.1

T=300oC, P=1.0MPa, H2/CO2 =3, Single reactor: SV=1000 ml/gcat.hr, Series reactors: SV=1000 ml/gcat.hr, Recycle reactor-A: SVF=1000 ml/gcat.hr R=3, B: S V F = 5 7 0 ml/gcat.hr, R=6, C: SVF=333 ml/gcat.hr, R=5, (A and B: SV=4,000 C: SV=2,000)

78

Fig. 7. Gas chromatogram of oil product from a reaction in recycle C mode (Table 2).

CH2=Cri2 Fe*3-CH2.

"

Fe*3-CH2-CH;,-CH2

pe»2

+

CH 3 -CH 2 =CH 2 alpha oletin

H-CH=CH2 Fe*3-CH2«

R-CH=CH2 Fe*3-CH2-CH2-CH .

R

H2 H Fe*3-CH2-CH2-C-C-CH

R

R

H H2 CH3-CH2-C-C—C=CH2

I

R

R

Branched alpha olefin

Fig. 8. Plausible reaction mechanism of oligomerization and isomerization in CO2 hydrogenation

4. CONCLUSIONS >From our experimental data we can conclude that unconverted feed materials, CO, and formed light hydrocarbons containing C1-C4 can be recycled in order to obtain higher MW hydrocarbons above C5+ fraction. This was possible because a-olefin is oligomerizable over Fe-based catalyst. Redox radical initiation and termination mechanism was suggested for overall reaction and radical propagation mechanism was suggested for light a-olefin oligomerization via CO2 hydrogenation.

REFERENCES 1. Aresta, M., Proceedings of the 4th ICCDU, Kyoto, Japan, Sep. 7-11, Elsevier Science B.V. 114(1997)7. 2. PH. Choi, K.W. Jun, S.J. Lee, M.J. Choi and K.W. Lee, Catal. Letters, 40 (1996) 115. 3. S.S. Nam, H. Kim, G. Kishan, M.J. Choi and K.W. Lee, Appl. Catal., 179 (1999) 155. 4. S.R. Yan, K.W. Jun, J.S. Hong, M.J. Choi, K.W. Lee, Appl. Catal., 194 (2000) 63. 5. M.J. Choi, J.S. Kim, H.K. Kim, S.B. Lee, Y. Kang and K.W. Lee, Korean J. Chem. Eng. 18 (2001) 645. 6. M.J. Choi, J.S. Kim, S.B. Lee, W.Y. Lee and K.W. Lee, GHGT-6, Oct. 1-4, Kyoto, Japan (2002) p.645.


79

Process Evaluation of Biomass to Liquid Fuel Production System with Gasification and Liquid Fuel Synthesis Tomoaki Minowa*, Toshiaki Hanaoka, and Shin-ya Yokoyama Biomass Technology Research Lab, National Institute of Advanced Industrial Science and Technology (AIST), AIST Chugoku, Suehiro, Hiro, Kure, Hiroshima 737-0197 Japan We have proposed biomass to liquid fuel production process with gasification and liquid fuel synthesis (BTL process). The process was experimentally studied using rice straw as a biomass, steam gasification, and FT and Oxo synthesis, supported by New Energy and Industrial Technology Development Organization (NEDO), Japan. Overall mass and energy balance for the process was calculated. Liquid products of 8 wt% on the dry biomass basis was estimated, and its energy yield was around 25% on HHV basis. The process could be operated without energy from outside, when obtained off gas was used for the process. 1. INTRODUCTION Biomass is not only renewable but also sustainable energy or carbon resource, and it's using mitigates CO2 emission by replacing fossil fuel. For the effective biomass utilization, one of the technical options is liquid fuel production from syngas, which comes from biomass gasification. Biomass is gasified for the production of syngas (mixture of H2, CO and CO2), and then liquid fuels can be synthesized from the syngas by Fischer-Tropsch reaction. Obtained liquid fuel can be used instead of transportation fuel, and its demand is worldwide; not only developed countries, but developing countries and rural area. The process had been studied by the international joint research team, Agrofuel Team, supported by New Energy and Industrial Technology Development Organization (NEDO), Japan (1999GP012).1' The proposed process consisted of steam gasification with oxygen, cooling down for heat recovery and cleaning, FT synthesis from the obtained syngas, and further Oxo synthesis from the remained syngas and olefin. hi this paper, we describe the mass and energy balance of the overall process from biomass to liquid fuels by analyzing the data obtained by Agrofuel Team.

2. CALCULATION 2.1. Thermochemical data of materials Rice straw was selected as feedstock, because it is the most abundant agricultural byproduct in Asian countries; about 6 billion tons a year. The properties of the received rice straw is shown in Table 1. * E-mail, [email protected], Fax, +81-823-72-1956

80

Table 1 Properties of rice straw Straw 2 Straw 1 Straw 3 Properties (wt% on wet basis) Moisture 15.0 11.6 9.9 Fixed carbon 13.7 15.0 14.3 Volatile matter 58.4 56.2 59.5 Ash 12.9 13.9 19.6 Eleme:ntal compositiorI (wt% on dry b asis) C 40.9 41.5 37.7 H 5.1 5.2 4.9 0 37.6 36.7 36.8 S 0.09 0.2 0.07 N 0.81 0.43 0.29 Cl 0.36 0.17 0.09 Ash (estimate) 15.18 15.72 21.75 Total 99.90 100.06 101.6 Higher heating value 14.78 13.65 12.9 (MJ/kg on wet basis')

Average 12.3 14.3 58.0 15.5 40.03 5.07 37.03 0.12 0.51 0.207 17.55 100.52 13.78

From the elemental composition and the heating value, its enthalpy was estimated as -4.614 MJ/kg. Straw, 1 kg + O2, 35.9 mol -> CO2, 33.4 mol + H2O, 25.3 mol + 13.78 MJ Other thermochemical data (enthalpy and heat capacity) was collected references.2'3) Table 2 shows thermochemical data. Since no data was found for C7H14O, C10H20O, C, and Ash, they were presumed as shown in Table 2. 2.2. Gasification Gasification was carried out at Nagasaki Institute of Applied Science. To get syngas efficiently, steam gasification with oxygen was selected. The obtained optimum conditions and the results are shown in Table 3. Oxygen of 93 vol% was used for the gasification as an oxidation agent for the partial oxidization to produce reaction heat, and it was assumed to be made from air using PSA in the process. The required energy for PSA was assumed 0.3 kWh/Nm3-O2- The gasification was estimated as exothermic. The carbon conversion to gas was 99%, and carbon of 1% was assumed to form char. Other elements in the straw, S, N and Cl, were assumed to form H2S, NH3 and HC1, respectively. Remained nitrogen gas (N2) in the oxidation agent was assumed no reaction. 2.2. Cooling down for heat recovery and cleaning After the gasification, cooling down should be placed. At the cooling down, heat recovery to make hot steam could be carried out. Efficient of the heat recovery was assumed around 80% to make hot steam of 1,000 °C. Then, condensed wastewater would be obtained. Byproducts, H2S, NH3 and HC1, can be dissolved in water, and therefore, they were assumed to remove in the condensed water. Char and ash were also assumed to remove with the condensed water. Thus, the syngas, which was the gas composition as shown in Table 3, would be introduced to the following FT synthesis.

81

H2 CO

co 2

CH4 C3H6 C3H8 C4H8O C6H12 C6H14

C 7 H, 4 O C 9 H 18 C9H20 C10H20O

H2O H2S NH3 HC1 O2 N2 C, solid Ash

Table 2 Thermochemical data for this study Boiling temp. Cp, J/mol/K; =a+b.T+c.T2 +d.3 Mw Tb, K b c d a 20.4 27.144 9.27E-03 -1.38E-05 7.65E-09 2.016 81.7 30.871 -1.29E-02 2.79E-05 -1.27E-08 28.01 44.01 194.7 19.796 7.34E-02 -5.60E-05 1.72E-08 16.043 111.7 19.252 5.21E-02 1.20E-05 -1.13E-08 3.71 2.35E-01 -1.16E-04 2.21E-08 42.081 225.4 44.097 231.1 -4.225 3.06E-01 -1.59E-04 3.22E-08 348 -3.345 3.63E-01 -2.51E-04 7.45E-08 72.107 84.162 336.6 -1.746 5.31E-01 -2.90E-04 6.05E-08 86.178 341.9 -4.413 5.82E-01 -3.12E-04 6.49E-08 = Cp, C4H8O (wt/wt) 114 420 -3.718 8.12E-01 -4.51E-04 9.71E-08 126.243 424 3.144 6.78E-01 -1.93E-04 -2.98E-07 128.259 = Cp, C4 H8O (wt/wt) 156 18.015 373.2 32.244 1.92E-03 1.06E-05 -3.60E-09 34.08 212.8 31.943 1.44E-03 2.43E-05 -1.18E-08 17.031 239.7 27.316 2.38E-02 1.80E-05 -1.19E-08 36.461 188.1 30.293 -7.20E-04 1.25E-05 -3.90E-09 90.2 28.107 -3.68E-06 1.75E-05 -1.07E-08 31.999 77.4 31.151 -1.36E-03 2.68E-05 -1.17E-08 28.013 Average Cp, 1.256 kJ/kg 12 Average Cp, 1.256 kJ/kg -

Conditions Case 1 Case 2 Gas composition Case 1 Case 2

Evaporated AHf heat kJ/mol kJ/mol 0 -110.6 -393.8 -74.9 20.4 -103.9 -130.2 31.53 -41.7 28.3 -167.3 28.87 -707 36.34 -103.6 36.94 -229.2 -1252 40.69 -242 -20.2 -45.7 -92.4 0 0 0 0

Table 3 Optimum gasification conditions and results Pressure, MPa steam/C Temperature, °C 5 0.1 1,000 5 0.1 1,000 H2 41.2 vol% 46 vol%

CO 22. 1 vol% 23 vol%

CO2 29.4 vol% 25 vol%

CH4 7.3 vol% 5 vol%

O2/C 0.25 0.2 H2/CO 1.86 2.00

2.3. FT synthesis FT synthesis was studied at Korea Research Institute of Chemical Technology. The FT synthesis was carried out at 260 °C and 2 MPa using mixture gas of CO, CO2 and H2 of 1:1:2.5 (vol/vol) ratio. They developed a effective catalyst for the FT synthesis, and one pass reaction was assumed in this study. From the obtained results, the following reaction were assumed based on CO, nCO +2nH2 -> (CH2)n + nH2O, FT reaction, 52.8% CO + H2O -> CO2 + H2, Water-gas shift reaction, 28.9%, and CO -> CO, no reaction, 18.3%. The selectivity of hydrocarbons were Ci (methane), 9.77 C-mol%, C2.4, 30.52 C-mol%, C5-7,

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21.46 C-mol%, and C8+, 38.25 C-mol%. Olefin/paraffin (O/P) ratio were 70.25% for C2-4 hydrocarbons, and the other O/P ratio was assumed as 50%. In the calculation, hydrocarbons was represented as C3, Ce and C9, respectively. Since FT synthesis occurs at high pressure, the compressed energy of syngas is requires, and it was estimated by integral of P.AV. 2.4. Oxo synthesis Oxo synthesis, oxygenate production from plefin, was considered to upgrade the hydrocarbons obtained from FT process for transportation fuel. It was studied at Tokyo University, and carried out at 130 °C and 5 MPa. The following results were obtained, Olefin + H2, CO -> oxygenate, 60% Olefin + H2 -» paraffin, 10% Olefin -> iso-olefin, 20%, and Olefin -> olefin, 10% (no reaction). Overall reaction was as follow, olefin + 0.7 H2 + 0.6 CO -> 0.6 oxygenate + 0.3 olefin + 0.1 paraffin. 3. RESULTS AND DISCUSSION The scale of the process was assumed as 100 dry-ton rice straw a day, and mass and energy balance of overall process were calculated. The results are summarized in Table 4, and Figure 1 shows the flow in case 1 (O2/C = 0.25). Liquid fuel yields were 8 or 9 ton from 100 dry-tons biomass, and around 25% of energy was recovered. Obtained liquid consisted of a half of hydrocarbons and another half of oxygenate. Each reaction process, gasification, FT synthesis and Oxo synthesis, were exothermic. Therefore, the process could be worked without heating energy from outside, if heat could be recovered efficiently by heat exchanger. But the heating energy of 217.5 GJ was required for the drying of rice straw. In case 2 (O2/C = 0.2), surplus heating energy of around 100 GJ was required in the gasification process, because enough heat energy could not be recovered by heat exchanger. In addition, the electric power was required for the PSA operation and compressing, and it was estimated around 100 GJ (electricity). Surplus off gas could be used for the process energy (heat and power), and even if the efficient of power generation from off gas was as low as 20%, the electric power of 175 or 185 GJ could be obtained. Thus surplus off gas had enough energy for the overall process running. Table 4 Results of overall process evaluation (scale, 100 dry-ron) Liquid Energy „ . ~ Required energy for the 1process ... . .j Energy -*— ^—; rbJ in off bgas yield yield Heat Electric power ^(O^/Cr =l 0.25) o o 2

8ton

23 9%

-

J m / ^ GJ, ? G power) ^ (175

217.5 GJ

102.7 GJ

^ . ^ n= -0.2) n (O2/C

9ton

27 3%

G J « nGJ, ^ n(185 power)

313.5 GJ

105.3 GJ

-

83

Figure 1 Mass and energy balance of overall process from biomass to liquid. (Case 1,O2/C = 0.25) 4. CONCLUSION In this study, we estimated overall process of biomass to liquid fuel, which was proposed by Agro-fuel Team supported by NEDO. The mass and energy balance was calculated, and the following conclusion was obtained, 8 or 9 ton of liquid fuel would be produced from 100 dry-tons of rice straw at the condition of O2/C of 0.25 or 0.2, 24% or 27% of energy could be recovered as the liquid fuel,

84

S

the produced liquid would consist of a half of hydrocarbons and another half of oxygenate; it could be used as transportation fuel, S the surplus off gas would have enough energy for the overall process heat and power; thus this process could be operated without energy from outside. To improve the total efficient, other processes, gasification, FT, and so on, should be surveyed in future. ACKNOWLEDGEMENTS This study was done in the international joint research project of "The Production of Clean Fuels from Agricultural Byproducts" supported by New Energy and Industrial Technology Development Organization (NEDO). The authors would like thank Prof. Kyu-Wan Lee of Korea Research Institute of Chemical Technology, Dr. Keiichi Tomishige of University of Tsukuba, Prof. Masayasu Sakai of Nagasaki Institute of Applied Science, Prof. Jing-Lai Zhou of Institute of Coal Chemistry, Chinese Academy of Sciences, Prof. Kaoru Fujimoto of The University of Kitakyushu, and Prof. Noritatsu Tsubaki of Toyama University, who were members of the research team. REFERENCES 1. K. W. Lee, et al., Final Report of International Joint Research Project (2002). 2. JANFILE, Thermochemical data. 3. Handbook of Chemical Engineering, Society of Chemical Engineering, Japan.


85

Novel Catalysts for Gasification of Biomass with High Energy Efficiency Tomohisa Miyazawa, Mohammad Asadullah, Mohammad Nurunnabi, Baitao Li, Kimio Kunimori, Keiichi Tomishige* Institute of Materials Science, University of Tsukuba 1-1-1 Tennodai, Tsukuba, Ibaraki 305-8573, Japan Tel + Fax: 81-298-53-5030, E-mail: [email protected] We have developed novel catalysts for gasification of biomass with much higher energy efficiency than conventional methods (non-catalyst, dolomite, commercial steam reforming Ni catalyst). It is found that novel Rh/CeCVSiCh catalyst exhibited higher performance than conventional other systems, especially at low temperature (873 K). In addition, the catalyst was stable for 4 h in the gasification of cedar wood. 1. INTRODUCTION The use of biomass as an energy resource would lead to decrease the emission of CO2 as well as NOX, SOX, and particulate matters into the atmosphere [1]. This is because the plants, source of biomass, naturally recycle CO2 from the atmosphere during their growth and the biomass derived fuels in various forms are known to be super-clean fuels. The gasification of biomass at around 1073-1223 K to syngas (H2 + CO) can potentially be used either as a gaseous fuel for power generation or as a feedstock for the synthesis of clean transportation fuels such as methanol, dimethyl ether, and Fischer-Tropsch oils, or many other chemicals. The formation of tar and char in the gasification process is the most severe problem. Some nickel based catalysts [2-9], dolomite [10] and olivine [11] catalysts have been found to be active for tar cracking in the primary reactor within the temperature range of 1073-1173 K for dolomite and olivine, and 973-1073 K for nickel based catalysts. However, it has been reported that the Ni-based catalysts were deactivated significantly by carbon deposition on the catalyst surface [2, 3, 6, 12]. Recently we have found that Rh/CeO2 was an effective catalyst in cellulose gasification even at lower temperature than usual [13, 14], but the sintering of CeO2 caused the deactivation, therefore, we have developed stable Rh/CeO2/SiO2 catalyst [15-28]. Here it is demonstrated that Rh/CeO2/SiO2 catalyst is applicable to the gasification of cedar wood. In addition, we evaluate this catalytic gasification in terms of the cold gas efficiency. 2. EXPERIMENTALS The catalysts we have used in this investigation are Rh/CeO2/SiO2 with 1.2 xlO"4 mol Rh/g of catalyst and various contents of CeO2, commercial steam reforming Ni catalyst (G-91, TOYO CCI, catalyst composition: 14 mass% Ni, 65-70 mass% A12O3, 10-14 mass% CaO, and 1.4-1.8 mass% K2O), and dolomite (21.0 mass% MgO, 30.0 mass% CaO, 0.7 mass% SiO2, 0.1 mass% Fe2O3, and 0.5 mass% AI2O3). The dolomite was calcined at 773 K for 3 h before reaction. The

86

CeCVSiC^ was prepared by the incipient wetness method using the aqueous solution of Ce(NH4)2(NO3)6 and SiO2 (Aerosil, 380 m2/g). The loading of CeO2 on SiO2 was in the range of 10-80 mass%. After loading the Ce salt on SiO2, it was dried at 383 K for 12 h following the calcination at 773 K for 3 h under air atmosphere. Then the Rh was loaded on CeO2/SiO2 by impregnation of the support with acetone solution of Rh(C5H7O2)3. The acetone solvent was then evaporated at around 333 K with constant stirring. Then the catalyst thus produced was dried at 383 K for 12 h. The final catalyst was pressed, crushed and sieved to 45-150 um particle size. The mass% of CeO2 in the Rh/CeO2/SiO2 is denoted in the parentheses, such as Rh/CeO2/SiO2(60). In each run, 3 g of catalyst was used and pretreated by a hydrogen flow at 773 K for 0.5 h. The sketch diagram of the single bed reactor is shown in Figure 1. The reactor is made of quartz glass. An inner tube is inserted from the top. Cedar wood was supplied from the top of the reactor together with N2. The biomass feeder consisted of a glass vessel with a small pore at the bottom and it was vibrated by the vibrator and the vibration rate controlled the feeding rate. The gasifier consisted of a fluidized bed section at the middle of the reactor. In the gasification of biomass with oxygen, O2 and N2 were supplied from the bottom of the reactor. The sample of the product gas was collected and analyzed by gas chromatography (GC). The concentration of CO, CO2 and CH4 products was determined by FID-GC equipped with a methanator using a stainless steel column packed with Gasukuropack 54 and the concentration of hydrogen was determined by TCD-GC using a stainless steel column packed with a molecular sieve 13X. The rate of the gas flowing out of the reactor was measured by a soap membrane meter. The amount of solid carbon (coke on the catalyst and char) was determined by the total amount of gas (mainly CO2) formed after stopping the feeding of biomass and cellulose under the air flowing at the reaction temperature. The carbon-based conversion to gas (C-conversion) was calculated by "A/B x 100", where A represents the formation rate of CO + CO2 + CH4 and B represents the total carbon supplying rate of biomass. The C-conversion listed in the tables is the average during the reaction (15~25 min). The yield of solid carbon (coke+char) is calculated by (total CO2 + CO) / (total carbon amount in fed biomass) x 100. The tar yield is defined as (100 C-conv (%-C) - solid carbon(char+coke) yield (%-C)). The feeding rate of biomass or cellulose, N2, and air are described in each result. Equivalence ratio (ER) can be calculated by (supplied oxygen weight) / (dry biomass weight) / (stoichiometric oxygen weight /dry biomass weight for complete combustion). Energy efficiency of biomass gasification can be estimated on the basis of cold gas efficiency. The cold gas efficiency is calculated by the ratio of heat value of biomass combustion to that of produced gas. In this calculation, the external heat energy supplied from the furnace is neglected. The gasification of cedar wood was carried out. The moisture content of the wood was about 10%. This wood was ground with a ball mill to about 0.1-0.3 mm size. The composition of the cedar wood was C 45.99 mass%, H2O 10 mass%,H5.31 mass%, 038.25 mass%,N 0.11 mass%,C10.01 mass%, S 0.02 mass%, and ash 0.3 mass%. 3. RESULTS AND DISCUSSION The C-conv as a function of temperature on different catalysts is compared in Table 1. On Rh/CeO2/SiO2(60) catalyst the C-conv is 88% at 823 K, which jumped to the 98% level at 873 K. At 823 K, this value is 54% over G-91 catalyst which did not attain to 90% level even at 973 K. On the other hand, the C-conv is 43 and 41% on the dolomite catalyst and in the non-catalyst systems, respectively at 823 K. Although the C-conv attain to about 94% at high temperature

87

(1173 K) on dolomite, the value is under 90% in the non-catalyst system at the same temperature. The rest of the carbon in biomass related to the tar and solid carbon which is either deposited on the catalyst surface or exited the reactor with the flow of product gas. The gas formation rate, hydrogen to CO ratio, solid carbon and tar yield at different temperatures are also shown in Table 1. Although the syngas formation rate on all catalysts and non-catalyst systems reasonably increased with increasing temperature, it is much higher on Rh/CeO2/SiO2(60) at all temperatures and this represents the superiority of the Rh/CeO2/SiO2(60) catalyst over all other conventional catalysts. Furthermore, although about 1% solid carbon was formed on Rh/CeO2/SiO2(60) catalyst above 873 K, the tar was not formed at these temperatures at all. However, large amount of the tar and solid carbon was formed in other systems as shown in Table 1. Since the amount of solid carbon on Rh/CeO2/SiO2(60) catalyst was very small even at lower temperature, the catalyst surface can be kept quite clean. The hydrogen to CO ratio slightly decreased with increasing temperature on Rh/CeO2/SiO2(60) and G-91 catalysts. However, it was remarkably increased on dolomite catalyst with increasing temperature and it was same in the non-catalyst system. This explains that the hydrogen formation in these systems is difficult at lower temperatures. In addition, some extent of methane was formed on Rh/CeO2/SiO2(60) and G-91 catalysts, no C2 hydrocarbons were formed, however it was remarkably formed in the latter two systems.

Figure 1. Schematic diagram of continuous-feeding fluidized bed reactor

Figure 2. Dependence of cold gas efficiency on reaction temperature using various methods in the gasification of cedar wood. : Rh/CeO2/SiO2 (60), O : G-91, A : dolomite, X: non-catalyst, Reaction conditions are the same as those in Table 1.

88

Table 1. Comparison of the performance of various catalysts at different temperatures in the gasification of cedar wood. Temp. Formsition rate /îmc)l/min /K

CO

H2

1605 2093 2594 3024 747 1464 1963 2562 566

2290 2616 3209 3456 1094 1960 2522 3433 174

CO2

CH4

c2

H2/CC) C-conv Solid C Tar yield yield /%-C /%-C /%-C

2915 513 1 11 1.4 88 2744 785 0 3 1.3 97 2497 628 1.2 99 0 1 1.1 99 2109 580 1 0 1.5 2255 116 34 12 54 219 1.3 2497 73 15 12 1.3 79 2329 245 10 11 1.3 89 2252 309 9 2 125 0.3 43 1738 20 37 27 731 248 1767 173 65 0.3 873 48 19 33 1423 649 2533 270 99 0.5 923 75 17 8 973 2017 1098 2474 461 180 0.6 2 89 9 1 1073 2061 1335 2631 527 203 0.7 94 5 1173 2403 1566 1939 580 480 0.7 94 1 5 Non549 132 1654 106 38 0.2 41 823 8 51 catalyst 873 849 169 1989 157 119 0.2 54 4 42 1068 253 2037 213 150 0.2 923 60 3 36 1288 338 2084 269 183 0.3 973 67 2 31 1073 1890 603 2102 433 384 0.3 13 84 4 1173 2268 732 2104 414 346 0.3 89 4 7 Conditions: Catalyst 3 g, ER = 0.28, feeding rate 150 mg/min (10% moisture, C 5748 l^mol/min, total H2 4815 umol/min and O2 2208 umol/min), N2 flow 60 ml/min from the top and 68 ml/min from the bottom, O2 flow 42 ml/min from the bottom. 823 Rh/CeO 2 /SiO2 (60) 873 923 973 823 G-91 873 923 973 823 Dolomite

Figure 2 shows the dependence of cold gas efficiency on reaction temperature using various catalysts. It is clear that the gasification of cedar wood over Rh/CeO2/SiO2 was more energy efficient than other catalysts. The efficiency over G-91 is apparently high, however, it is expected that the efficiency decreased with time on stream since the serious deactivation was observed over G-91 although the details of the reaction time dependence is not shown here. In the case of non-catalyst and dolomite, high temperature must be necessary for the effective gasification. The catalyst life in this biomass gasification process with 150 mg/min of feeding rate, 3 g of catalyst, ER around 0.3, and the oxygen supply in part of cocurrent and part of countercurrent systems at 923 K was tested for 4 h as shown in Figure 3. The C-conv was quite constant in the 98-99% level during the reaction period. The formation of CO, CO2, and H2 was almost stable with a subtle fluctuation may be due to the feeding rate fluctuation or experimental error. However, the formation of methane slowly decreased after about 2 h. The catalyst deactivation is mainly related to the solid carbon deposition on the catalyst surface as observed in methane reforming to produce syngas [29-31], and it is known that the catalyst fluidization can decrease

89

the amount of carbon deposition [32-34]. In this case, it was not severe problem over our novel catalyst.

Figure 3. Life test of Rh/CeO2/SiO2(60) in the gasification of cedar wood. Catalyst 3g, biomass 150 mg/min (H2O 10%, C, 5748 umol/min; total H2, 4815 umol/min; total O2, 2208 umol/min), O2 flow 21 ml/min (O2 859 umol/min) from the bottom, O2 flow 21 ml/min (O2 859 umol/min) from the top, N 2 60 ml/min from the top, and 68 ml/min from the bottom, temperature 923 K. * : C-conversion, : H2, : CO, A: CO2, and : CH4. 4. CONCLUSION Rh/CeO2/SiO2 exhibited much higher activity and more excellent stability than conventional systems using steam reforming Ni catalyst, dolomite, and non-catalyst in the gasification of cedar wood with air. Especially, when Rh/CeO2/SiO2 was used, almost complete gasification was available even at 873 K. This means that the energy efficiency can be enhanced by the utilization of the effective catalyst. ACKNOWLEDGEMENT This research was supported by the Future Program of Japan Society for the Promotion of Sciences under the Project "Synthesis of Ecological High Quality of Transportation Fuels" (JSPS-RFTF98P01001) and the 21 st Century COE program under the Ministry of Education, Culture, Sports, Science and Technology. REFERENCES [1] [2] [3]

J.M. Ogden, Ann. Rev. Energy Environ. 24 (1999) 227. L. Garcia, M. L. Salvador, J. Arauzo, R. Bilbao, Energy Fuels 13 (1999) 851. L. Garcia, M. L. Salvador, J. Arauzo, R. Bilbao, Fuel Proc. Tech. 69 (2001) 157.

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[4] [5] [6] [7] [8] [9]

L. Garcia, R. French, S. Czernik, E. Chornet, Appl. Catal. A: Gen., 201 (2000) 225. L. Garcia, M. L. Salvador, J. Arauzo, R. Bilbao, Ind. Eng. Chem. Res. 37 (1998) 3812. J. Arauzo, D. Radlein, J. Piskorz, D. S. Scott, Ind. Eng. Chem. Res. 36 (1997) 67. J. Arauzo, D. Radlein, J. Piskorz, D. S. Scott, Energy Fuels 8 (1994) 1192. C. Courson, E. Makaga, C. Petit, A. Kiennemann, Catal. Today 63 (2000) 427. Y. Tanaka, T. Yamaguchi, K. Yamasaki, A. Ueno, Y. Kotera, Ind. Eng. Chem. Res. 23 (1984)225. [10] J. Gil, M. A. Caballero, J. A. Martin, M. -P. Aznar, J. Corella, Ind. Eng. Chem. Res. 38 (1999) 4226. [11] S. Rapagna, N. Jand, A. Kiennemann, P. U. Foscolo, Biomass Bioenergy 19 (2000) 187. [12] E. G. Baker, L. K. Mudge, M. D. Brown, Ind. Eng. Chem. Res. 26 (1987) 1335. [13] M. Asadullah, K. Tomishige, K. Fujimoto, Catal. Commun. 2 (2001) 63. [14] M. Asadullah, K. Tomishige, K. Fujimoto, K., Ind. Eng. Chem. Res. 25, (2001) 5894. [15] M. Asadullah, S. Ito, K. Kunimori, M. Yamada, K. Tomishige, J. Catal. 208 (2002) 255. [16] M. Asadullah, S. Ito, K. Kunimori, K. Tomishige, Ind. Eng. Chem. Res., 41 (2002) 4567. [17] M. Asadullah, S. Ito, K. Kunimori, M. Yamada, K. Tomishige, Environ. Sci. Tech. 36 (2002) 4476. [18] K. Tomishige, M. Asadullah, S. Ito, K. Kunimori, Kagaku Kogaku Ronbunshu 28 (2002) 666. [19] K. Tomishige, M. Asadullah, S. Koyama, S. Ito, K. Kunimori, J. Jpn. hist. Energy 82 (2003)261. [20] K. Tomishige, T. Miyazawa, M. Asadullah, S. Ito, K. Kunimori, J. Jpn. Petrol. Inst. 46 (2003) 69. [21] M. Asadullah, T. Miyazawa, S. Ito, K. Kunimori, K. Tomishige, Stud. Surf. Sci. Catal. 145 (2003) 307. [22] M. Asadullah, T. Miyazawa, S. Ito, K. Kunimori, M. Yamada, K. Tomishige, Green Chem. 4 (2002) 385. [23] M. Asadullah, T. Miyazawa, S. Ito, K. Kunimori, K. Tomishige, Appl. Catal.A:Gen. 246 (2003) 103. [24] K. Tomishige, M. Asadullah, K. Kunimori, Catal. Surv. Asia, 7 (2003) 219. [25] M. Asadullah, T. Miyazawa, S. Ito, K. Kunimori, K. Tomishige, Energy Fuels 17 (2003) 842. [26] K. Tomishige, T. Miyazawa, M. Asadullah, S. Ito, K. Kunimori, J. Jpn. Petrol, hist. 46 (2003) 322. [27] K. Tomishige, T. Miyazawa, M. Asadullah, S. Ito, K. Kunimori, Green Chem. 5 (2003) 399. [28] M. Asadullah, T. Miyazawa, S. Ito, K. Kunimori, M. Yamada, K. Tomishige, Appl. Catal. A:Gen. 255 (2003) 169. [29] K. Tomishige, Y. Chen, K. Fujimoto, J. Catal. 181 (1999) 91. [30] K. Tomishige, K. Fujimoto, Catal. Surv. Jpn. 2 (1998) 3. [31] K. Tomishige, K. Fujimoto, J. Jpn. Petrol. Inst. 44 (2001) 65. [32] K. Tomishige, Y. Matsuo, Y. Sekine, K. Fujimoto, Catal. Commun. 2 (2001) 11. [33] K. Tomishige, Y. Matsuo, Y Yoshinaga, Y. Sekine, M. Asadullah, K. Fujimoto, Appl. Catal. A: Gen. 223 (2002) 225. [34] Y. Matsuo, Y. Yoshinaga, Y. Sekine, K. Tomishige, K. Fujimoto, Catal. Today 63 (2000) 439.


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Improving carbon utilization in biomass hydrocarbons via Fischer-Tropsch Synthesis

conversion

to

synthetic

D. Unruh*, M. Rohde, G. Schaub Engler-Bunte-Institut, Universitat Karlsruhe (TH), Engler-Bunte-Ring-1, D-76131 Karlsruhe, Germany *corresponding author [email protected] Biomass is gaining interest as a feedstock for hydrocarbon synthesis via Fischer-Tropsch (FT) synthesis because fossil CO2 emissions can be significantly reduced in the transportation sector. The overall carbon utilization that can be achieved by common FT synthesis is rather low (e.g. 135 kg Diesel fuel /1 wood) as synthesis gases from biomass gasification generally have a high CO2 content. The conversion of CO2 in FT synthesis could improve the overall carbon utilization, but this is limited by the availability of H2 and the CO2/CO shift equilibrium. The present study investigates the application of H2O-selective ceramic membranes as a novel method to enhance CO2 conversion in FT synthesis. A ID fixed bed reactor model with integrated membrane was developed to quantify the effect of H2O removal under the assumption of an ideal membrane. Model calculations illustrate that the in-situ removal of H2O from the reaction mixture by a membrane increases CO2 conversion. Initial results of permeation experiments are also presented.

1. INTRODUCTION The use of biomass as a fuel source will help to reduce fossil CO2 emissions and allows the substitution of petroleum. E.g. in Europe, 5.75 % of all transportation fuels should be substituted by biofuels by 2010. One way to obtain high quality transportation fuels is to gasify biomass (Eq. 1) and subsequently to convert the derived synthesis gas to long-chain hydrocarbons by Fischer-Tropsch (FT) synthesis (Eq. 2). The hydrocarbon products have high cetane numbers and are free of sulfur and aromatics and therefore exhibit low particle emission under combustion. CHi.6Oo.7 + v O 2 + CO + 2 H2

WH2O

-

xCO + y H 2 + zCO 2 -(CH2)- + H2O

(1) (2)

The overall process exhibits a significant loss in energy content and product yield because carbon is removed as CO2 to increase the H/C ratio. At the same time the energy requirement of the overall process is significant and is usually met by burning part of the feedstock. As an example based on an optimistic estimate of 50 % energy efficiency [1], one ton of wood (with 60 % organic content) yields 135 kg of Diesel fuel (Eq. 3). 100CH 16 O 07 + 45 O2

-

40 -(CH2)- + 60 CO2 + 40 H2O

(3)

If all energy requirements are be met by other sources, chemical stoichiometry would allow a 75 % increase in hydrocarbon yields (Eq. 4).

92 70-(CH 2 )- + 30CO 2 + 10H 2 O (4) 100CHi.6Oo.7 As the regional availability of biomass is a limiting factor in countries with high population densities (such as Western Europe), the challenge to increase product yields and carbon utilization is obvious. Here, it is necessary to overcome the stoichiometric constraints by adding H2 or CH4 (Eq. 5 and Eq. 6), which would result in significant increases in hydrocarbon yields from biomass. 100CHi.6Oo.7 + 90H 2 100-(CH2)- + 70H 2 O (5) * 190-(CH2)- + 70H 2 O (6) 100CHi.6Oo.7 + 90CH4 However, conversion of CO2 formed in gasification (Eq.l) does not only require sufficient amounts of H2, but in addition, chemical equilibrium and kinetic limitations of the CO2 shift according to Eq. 7 must be overcome to form CO. Hydrocarbons are then formed from CO in the consecutive FT reaction (Eq. 2). CO2 + H2 . CO + H2O (7) The CO2 shift (Eq. 7) and FT reaction (Eq. 2) can either occur in the same reactor, if Fe is used as a shift-active FT catalyst, or in two separate reactors (Fig. 1). In this case, Co could be used as a catalyst. The Co catalyst is not shift-active at FT conditions, but it is regarded as an attractive FT catalyst [2]. This paper discusses strategies for overcoming chemical equilibrium and kinetic limitations for CO2 conversion and presents studies regarding a novel membrane reactor that should enhance both CO2 shift and FT reactions simultaneously in the same reactor (Eqs. 2 and 7) by in-situ removal of H2O. Biomass Gasification

CHI. 6 OQ.7

*

CO2 + CO + H2

r

; CO/CO 2

"f ~*"| H2

shift

- CO2 + i CO + H2

FT - Synthesis ( CH z)

i -'

w/o CO2 removal

Fig. 1. Flow scheme of the conversion of biomass to liquid hydrocarbons, the shift step is optional.

2. IMPROVING CARBON UTILIZATION - CONVERSION OF CO 2 At temperatures typical for the FT synthesis of long-chain hydrocarbons (200 - 250 °C), chemical equilibrium favors the side of CO 2 in CO 2 /CO shift reaction (Eq. 7). There are two possibilities to overcome this equilibrium limitation. One is to increase temperature favoring CO 2 conversion [3], though this has detrimental effects on FT product selectivities. The other is to remove CO (by FT reaction) and H 2 O (e.g. with a H 2 O-selective membrane) thereby enhancing the conversion of CO 2 to CO. In-situ removal of H 2 O during FT synthesis has - besides the equilibrium effect on the CO 2 /CO shift reaction - other positive effects in the presence of Fe catalysts: (i) decrease in kinetic inhibition of FT reaction by H 2 O [4], (ii) higher partial pressures of reactants, (iii) higher space time, and (iv) slower catalyst deactivation due to reoxidation by H 2 O [5].

*"

93

3. APPLICATION OF A H2O-SELECTIVE MEMBRANE For a numerical and an experimental study, a fixed bed reactor with a membrane wall in lab-scale geometry was chosen. Figure 2 shows the arrangement, where the FT catalyst is filled into an annular channel. The inner wall of the annular channel is a glass tube (thermocouple), the outer wall a cylindrical, H2O-selective membrane. The membrane (length 300 mm, outer diameter 10 mm, wall thickness 1.5 mm) is fixed in a steel tube with 13 mm inner diameter. The outer annular channel between steel tube and membrane is used as sweepgas channel. Fig. 2. Lab-scale fixed bed membrane reactor with insitu H2O removal, membrane as reactor wall and separation unit, annular channel filled with catalyst particles, feed gas: H2/CO/CO2. H2O is removed via the membrane into the sweep-gas channel (left channel). The membrane used consists of a ceramic (AI2O3) body produced by Inocermic GmbH (Hermsdorf), with a functional, hydrophilic layer, based on Si(OH)x, prepared at the Fraunhofer Institut fur Angewandte Chemie in Berlin [6].

4. MATHEMATICAL MODELLING STUDIES OF A MEMBRANE REACTOR For the purpose of studying the limiting factors for membrane application, onedimensional (ID) and two-dimensional (2D) mathematical models for a fixed bed reactor configuration have been developed [7]. These models include the kinetics of FT and CO2 shift reactions determined in earlier experimental studies [8]. They should represent the effects of in-situ removal of H2O, determine permeance requirements and identify promising experimental conditions. 4.1 Assumptions The model includes the following assumptions: (i) The lab-scale reactor is nearly isothermal ( 2 K), as it has small radial dimensions and the catalyst is diluted with quartz sand, (ii) For the lab-scale reactor, axial dispersion is neglected, (iii) Inter- and intraparticle transport limitations are considered to be negligible, given the small particles used [8]. (iv) A constant catalyst bed porosity is assumed over the reactor cross section, (v) A pressure drop of less than 0.5 % over the catalyst bed is calculated with the Ergun equation, (vi) Ideal membrane behaviour is assumed for the ID calculations, i.e. only H2O permeates. 4.2 Membrane transport model The transport through the membrane is described by a phenomenological approach, which simplifies the complex membrane transport (Eq. 8). Here, p,,R and p,,p are the partial pressures on the retentate and permeate sides of the membrane. With the assumption that either the sweep-gas flow is high in comparison to the permeation flow or vacuum is applied at the permeate side, the partial pressures on the permeate side are set equal to zero. This means that accumulation of the permeating components in the sweep gas is not considered. So the driving force through the membrane is at its maximum.

94

4.3 Mass balance and initial conditions The following model equation is derived (Eq. 9) by applying the assumptions to a differential mass balance for a single component i. Eq. 9 is an initial value problem with the inlet flows as initial conditions, which can be solved by a Runge-Kutta method. The kinetic reaction model from Riedel et al. [8] is applied. C3H8 is chosen as a model component for the complex hydrocarbon product [8]. ri,=Peri-(plR-plP) (8) mem

i f = ^P--^(P,-P.) A -/ W

W

4.4 Results of numerical simulations Numerical simulations indicate that CO2 conversion at typical FT conditions strongly depends on variations in space velocity, temperature, membrane area / reactor volume (or catalyst mass) ratio, syngas composition and membrane permeances and selectivities. Figure 3 shows the effects of membrane application and variations in temperature and space velocity on CO2 conversion, calculated with the ID model. For typical lab-scale space velocities (GHSVmod ~ 0.5 cmV(g s)), temperature and H2O removal have a strong influence on CO2 conversion. In a conventional lab-scale reactor, CO2 is produced indicated by a negative conversion. At temperatures below 350 °C, CO2 is converted to a low degree. A temperature increase, however changes the product distribution (e.g. to higher methane yields) and leads to an increase in carbon formation. On the other hand in a membrane reactor, 40 % of CO2 is converted at T = 250 °C and GHSVmod = 0.5 cm3/(g s). At 350 °C, CO2 can be converted completely under these idealized conditions. For high gas velocities, the influence of membrane application decreases, as the residence time decreases (leading to lower overall conversions, lower H2O partial pressures and lower transmembrane flows in comparison to the synthesis gas flows). Fig. 3. Calculated effect of in-situ H2Oremoval on CO2 conversion in model syngas from biomass gasification, simulated with kinetic model of FT fixed bed reactor, feed gas pm/pco/pco2: 7.1/1.7/1.2 bar, AmemfVM: 850 m2/m3 (labscale geometry), PerH2o / Perm = 10"7 / 0 mol/(s Pa m2), p = 10 bar, kinetic data for 100Fe/13Al2O3/10Cu/10K [8] Typical lab-scale space velocity: GHSVmod = Vin n /met ~ 0.5 cmV(g s)

95

5. EXPERIMENTAL DATA OF MEMBRANE PERMEANCE In parallel experiments, permeances of key components through a hydrophilic ceramic membrane [6] were measured under non-reacting and reacting conditions. The experimentally determined permeances can be implemented into the mathematical model and help to identify the potentials and the limits of membrane application. Figure 4 shows single gas permeance as a function of temperature. The experimental results show that the membrane is not ideally selective under these conditions, but H2O permeates preferentially. The F^-permeance shows a strong increase with temperature. First experiments with H2O indicate a similar increase for the FÔ-permeance. Therefore, the theoretical separation factor a = Perô/Perm determined by single gas experiments remains relatively low. The other gases show the expected behaviour. The Permeances increase in relation to temperature and depend on molecular size and shape. Future binary and multi-component permeation experiments will show if separation factors increase, e.g., due to pore blocking (by H2O condensation) and competing adsorption. In Figure 5, data points from the single gas experiments and from the literature are compared with simulation results gained with the 2D model [7]. The comparison shows that no advantages regarding hydrocarbon yields and CO2 conversion can be anticipated with the membranes investigated so far. The highest H2-permeance allowed in this respect would be 3*10"9 mol / (Pa s m2). Membrane application would result in a decrease in hydrocarbon yield (due to a loss of H2). H2O permeance should ideally be higher than l*10"7 mol / (Pa s m2) to gain higher hydrocarbons yields in comparison to a reactor without membrane. It is important to note that these calculations were done with maximum driving forces for H2O and H2. If gases rich in H2 are used as sweep gas, the driving force for H2 could be lowered, leading to lower H2 fluxes (or allowing higher H2 permeances).

Fig. 4. Single gas permeance data in lab-scale measurement (Pressure increase measured for given volume flow of each gas), membrane type: see section 3

Fig. 5. Collection of permeance data with various membranes and calculated range required (based on 2D-simulations, shadowed area, AY> 10%) for the use of an iron catalyst in lab-scale fixed bed reactors [7], exp. Data (T = 100 - 300 °C): circles: Espinoza et al. [9], squares: Kolsch et al. [10], diamond: Aoki et al. [11], triangle: own exp. data

96

6. CONCLUSIONS Model calculations suggest that the application of a H2O-selective membrane for in-situ H2O removal can significantly improve CO2 conversion in FT synthesis. However, none of the membranes investigated so far have met the specific requirements necessary to obtain the desired positive effects. Although H2O permeance values are high, the respective H2 permeance values are not low enough to prevent significant loss in H2 Therefore, the future investigations will focus on the (i) identification of ranges of membrane permeances and selectivities for various reactor geometries, (ii) effect of sweep gas composition on achievable fluxes, and (iii effect of temperature on membrane stability. ACKNOWLEDGEMENT Financial support from the German Government (BMBF, ReFuelNet, Forderzeichen 015F0118) and discussions with Dr. Kolsch are gratefully acknowledged. REFERENCES 1. Tijmensen M.J.A., Faaij A.P.C., Hamelinck C.N., van Hardeveld M.R.M., Biomass & Bioenergy, 23 (2002) 129 (2) 2. Schulz H., Beck K., Erich E., Stud. Surf. Sci. Catal. 52 (1988) 457 3. Rohde M., Unruh D., Schaub G., Fischer-Tropsch synthesis with CO2-containing syngas from biomass - Kinetic analysis of fixed bed reactor model experiments, Proceedings 7th International Conference on CO2 Utilization, Seoul, October 2003 4. Zimmerman W. H., Bukur D. B., Can. J. Chem. Eng., 68 (1990) 292 5. O'Brien R. J., Xu L , Spicer R. L , Milburn D. R, Davis B. H., Catalysis Today, 36 (1997)325 6. Kolsch P., Noack M., Caro J., Weh L., Keramische Zeitschrift, 51 (1999) 112 7. Unruh D., Rohde M., Schaub G., In-situ Removal of H2O during Fischer-Tropsch Synthesis - a Modelling Study, Proceedings of the DGMK Conference "Chances for Innovative Processes at the Interface between Refining and Petrochemistry", October, 9 11, 2002, Berlin, Germany 8. Riedel T., Schaub G., Jun K.-W., Lee K.-W., Ind. Eng. Chem. Res., 40 (2001) 1355 9. Espinoza R. L., Santamaria J. M., Menendez M. A., Coronas J., Irusta S., Production of hydrocarbons, International Publication Number: WO 99/64380 (1999) 10. Kolsch P., Noack M., Druska P., Miiller D., Toussaint P., Caro J., Chemie Ingenieur Technik, 70(1998)860 11. Aoki K., Kusakabe K , Morooka S., Ind. Eng. Chem. Res, 39 (2000) 2245 Nomenclature r radius, radial coordinate, mm Amem membrane surface, m2 Pi partial pressure of i, Pa Vin_n volume flow, mVs

z axial coordinate, mm A cross section, m2 T temperature, K xmod = l/GHSVm0d

Vbed volume of catalyst bed, m3 p total pressure, Pa GHSVmod space velocity, m7(kg s) P density, kg/m3

Per Permeance, mol/(Pa s m2) mcat catalyst mass, kg ry reaction rate, mol/(kg s) v,j stoichiometric factor n, molar flow, mol/s Indices: i component, j reaction, P permeate, R retentate, mod modified, in inlet, n norm


97

Fischer-Tropsch synthesis with CO2-containing syngas from biomass Kinetic analysis of fixed bed reactor model experiments M. Rohdea*, D. Unruh", P. Piasa, K.-W. Leeb and G. Schauba a

Engler-Bunte-Institut, Universitat Karlsruhe (TH), Engler-Bunte-Ring 1, D-76131 Karlsruhe, Germany * corresponding author [email protected] Korea Research Institute of Chemical Technology, Daejon, Korea

Biomass conversion into synthetic hydrocarbons via Fischer-Tropsch (FT) synthesis has the potential to produce high-value transportation fuels or petrochemicals from non-fossil sources. This theoretical work studies the effect of CO2 inlet concentration on CO2 conversion and hydrocarbon yield. A simple fixed bed plug flow reactor model is applied and validated by independent experimental data. This study illustrates that the rate equations and parameters derived by Riedel et al. [2] for pure CO2 hydrogenation are applicable for the entire range of CO/CO2 mixtures. Additionally, case studies are carried out concerning the effects of H2 availability and reactor temperature on CO2 conversion and hydrocarbon yield. 1. INTRODUCTION Typical biomass gasification gases can contain large amounts of CO2. A measure for the CO2 content in a syngas is the ratio Zco2 = PC02/ (pco+ Pcm)- Zco2 can vary between 0.05 and 0.7, depending on feedstock, gasification process and gas cleaning process [1]. Regarding the hydrogenation of CO/CO2 mixtures, the amount of H2 in the syngas is essential, expressed by the ratio ZH2 = Pm/ (2pco + 3pco2), which can be derived from the stoichiometric Equations 1 and 3. Syngases with zm< 1 are called H2-deficient, syngases with xm= 1 are called balanced. Earlier investigations regarding CO2 hydrogenation under FT conditions on a K-promoted Fe-catalyst showed that conversion of CO2 to long-chained hydrocarbons is possible. Riedel et al. [2] proposed the reaction network in Figure 1 and developed a kinetic reaction model to describe the experimental data with a K-promoted Fe-catalyst (Table 2). It is essential that the Fe-catalyst exhibits a sufficient high CO2/CO shift reaction activity which can be attained by promotion with K. The main reaction path of the CO2 hydrogenation is a two-step mechanism with CO as intermediate as shown in Figure 1. In the first step, CO2 is converted to CO by the CO2 shift reaction (Eq. 2), and in the subsequent step, the intermediate CO reacts via FT CO 2 s ^2)y CO

N 3

\ -*

L

L

- ^ - (CH2)

CO

+ 2 H2

->

(CH2)

+

H2O

AHR«0

(1)

co 2

+

^

co

+

H2O

AHR>O

(2)

CO2

+ 3 H2

->

(CH2)

AHR«0

(3)

H2

+ 2 H2O

Fig. 1. Reaction network of CO2 hydrogenation proposed by [2] for K-promoted Fe-catalyst. (1) FT reaction, (2) CO2/CO shift equilibrium, (3) direct CO2 hydrogenation

98

reaction to hydrocarbons (Eq. 1). The direct CO2 hydrogenation is not significant (Eq. 3). In addition to the hydrogenation of CO2/H2 syngas, Lee et al. [3, 4] studied the effect of CO2 concentration in the feed syngas regarding the application of FT synthesis to biomass derived syngases i.e. Biomass-to-Liquids (BTL) [1]. The experiments showed that (i) CO2 is converted at high inlet concentrations, however (ii) CO2 is formed at low inlet concentrations. The aim of the present study is to model these experimental results with a mathematical fixed bed reactor model, applying a kinetic approach derived from a similar catalyst [2]. 2. EXPERIMENTAL DATA BASE Lee et al. [3, 4] systematically varied zCo2 in the range from 0...1 for balanced syngases (ZH2 = 1) and performed an additional experiment for an Eh-deficient syngas (ZH2 = 0.43). These experiments were carried out in a lab-scale fixed bed reactor with 0.1-1 g of Kpromoted standard catalyst Fe/Al2O3/Cu/K (100/16/6/4) diluted with quartz sand [2, 3]. The isothermal reactor was operated at 300°C, 1 MPa and xmod = 2 g s/cm3. Table 1 Experimental data [3, 4] for FT synthesis in lab-scale fixed bed reactor with variable CO, CO2 and H2 inlet concentrations. Fe/Al2O3/Cu/K (100/16/6/4), T = 300°C, p = 1 MPa, zCo2 = (0...1), ZH2 = (1, 0-43), Tmod = 2 g s/cm3. Feed gas composition (vol-%)

Conversion (%)

(-)

CO

CO2

H2

ZCO2

Zffi

XCO2

Xco

0 6.68 14.40 22.7 34.13 11.42

25.19 20.05 14.40 7.97 0 34.24

74.81 73.27 71.18 69.33 65.87 54.33

1 0.75 0.50 0.26 0 0.75

0.99 1 0.99 1 0.96 0.43

36 30 11 -43 0 1

0 87 95 97 99 81

Selectivity (%) Xc 32 44 53 61 69 21

SHC.C

91 97 99 99 70 99

Yield (calc.) (%)

Yco.c CO: 9 3.4 3.2 2.5 2 1 CO2: 30 5.8 Si,c

YCO2.C

YHC.C

62 53 45 37 30 74

35 43 52 60 69 21

3. MATHEMATICAL MODEL The chosen lab-scale fixed bed reactor has the following properties: (i) it is nearly isothermal due to high catalyst's dilution by quartz sand, (ii) axial dispersion is negligible because of large Bodenstein numbers (Bo> 100), (iii) internal and external mass transfer limitations and (iv) channeling effects are not significant because the catalyst particles are very small (dp < 100 (im) [2, 3]. Therefore, a simple one-dimensional, isothermal plug flow reactor model for the gas phase is sufficient (Eq. 4).

^ dmcat

=Z d M j

m

ca t =0:

ni=ni0

j= E ? ' ^ ^

H C

(4)

J - r i , LU2-SH

This simple initial boundary value problem of 5 coupled ordinary differential equations is solved with MATLAB applying a Runge-Kutta method for stiff systems. The general form of the rate equations for FT reaction and CO2 shift reaction (Eq. 5-7) and the kinetic parameter values are summarized in Table 2. In order to fit the numerical results to the experimental data, two parameters dj have been added to Equation 4. Propane (C3Hg) has a molecular

99

Table 2 General form of rate equations for Fischer-Tropsch reaction (FT) and CO2 shift reaction (CO2-SH) and kinetic parameter values [2, 5]. Fe/Al2O3/Cu/K (100/13/10/10), T = 300, 330, 360°C, p = 1 MPa, PH2/PC02 = 3:1, xmod = 0.042-21.4 g s/cm3, FT

CXVSH

.

33

^5

J.CO2

"

2.7

7.4

EAJ

(J/mol)

72000

55000

k'0J

^

S

0.748

0.01384

a b

/ _ , '

PCOPHZ

PcO "*" aFT,H2oPH2O "*" ^FT,CO2PcO2 ^

Pc02PH2-PcoPH20KP,CO-SH

_y

C02 SH

"

C02 SH

"

p c o +a S H H 2 O p H 2 O +b S H C O 2 p C O 2

log K PC0 _ SH = ( 2 0 7 3 / T - 2 . 0 2 9 )

( )

(7)

weight similar to the average of the organic products formed and is used as a model hydrocarbon (HC) [2]. 4. RESULTS OF MODEL CALCULATIONS AND DISCUSSION In Figure 2, the carbon monoxide and hydrocarbon yields are plotted versus the modified residence time for the balanced CO2 syngas. The plot of YCo,c exhibits the typical behaviour of an intermediate product in a consecutive reaction which was experimentally confirmed by [2]. The original Riedel kinetics predicts too high carbon monoxide yields as well as too low hydrocarbon yields regarding the experimental data. The catalyst of Lee et al. appears more active in this respect. By adjusting the parameters dj, it is possible to fit the experimental data (see Figure 2 and 3). An important result of the parameter adjustment is that an increase in the CO2-shift rate constant has only a slight influence on the overall conversion and hydrocarbon yield. In comparison, a slight increase in the FT activity leads to relatively large shift of the profiles. This proves that the FT reaction itself is the limiting step in the reaction network. For further discussion, two new characteristic numbers are introduced (see Fig. 3): a critical CO2 inlet concentration Zco2,crit and the maximum possible CO2 conversion Xco2,max under the given conditions. From Figure 3, the following conclusions can be drawn: (i) below zco2,crit CO2 does not react, but additional CO2 is formed by the CO-shift reaction i.e. XCo2 is negative.

Fig. 2. Carbon monoxide yield YCo,c and hydrocarbon yield YHc,c versus modified residence time xmod for synthesis gas PH2/PC02 = 3:1, (zCo2 = 1, ZH2 = 0.99), calculated with original Riedel kinetics ) and with fitted Riedel kinetics (—). Fitting parameter: dFT = 4, dCo2-sH = 6. (o) and (>) experimental data, see Table 1. T = 300°C, p = 1 MPa, xmod = 2 g s/cm3.

100

Fig. 3. Conversion and yield as a function of CO2 inlet concentration of balanced syngases (ZH2 = 1)- (0), (>) and ) experimental data, see Table 1. Trends calculated with fitted Riedel kinetics. Fitting parameter: dFT = 4, dco2-SH = 6. T = 300°C, p = 1 MPa, tmod = 2 g s/cm3. (ii) Above Zco2,crit CO2 is converted into CO which is further hydrogenated to hydrocarbons by the FT reaction. Therefore, syngases with Zco2 < zCo2,cnt should not be converted over an CO2/CO shift active catalyst, (iii) The total carbon conversion X c and the total hydrocarbon yield YHC.C decrease linearly with increasing CO2 content in the feed gas. This is attributed to the limited conversion of CO2, restricted by equilibrium constraints and by reaction conditions. Hence, Xccc.max is limited to a low value, (iv) CO is converted primarily and almost completely.

Fig. 4. Effect of H2 deficiency in syngas on conversion. Balanced (ZH2 = 1, —) and H2deficient (zH2= 0.43, ) syngases, (o), (t>) and ) experimental data, see Table 1. Conditions see Fig. 3.

Fig. 5. General effect of temperature on conversion for balanced syngases (ZH2= 1). T = 350°C , 300°C (—), 250°C (—). Riedel kinetic extrapolated for T = 250°C. Conditions see Fig. 3.

101 The fixed bed reactor model applying a fitted Riedel kinetics describes the experimental data quantitively. The following two case studies examine the influence of H2-availability and of temperature on zCo2,crit and on XCo2,max- Figure 4 shows the effect of H2-defiency, calculated for a H2-deficient syngas with Zm~ 0.43. The calculated curves come close to the experimental data. Zco2,crit shifts to higher CO2 inlet concentrations and XCo2,max reduces significantly. As a consequence, CO2 conversion becomes even more difficult. Though H2 is the limiting reactant, H2 is not completly consumed. XCo2 is suppressed by equilibrium constraints since a lower H2 partial pressure favors the side of CO2 (see Eq. 3). Therefore, a balanced syngas is essential because of (i) stoichiometric reasons and (ii) the effect of H2 partial pressure on the CO2/CO equilibrium. An increase of the reaction temperature by 50 K leads a lower zco2,mt and to a higher Xco2,max- At T = 350°C, CO2 is converted above zco2= 0.27 and Xco2,max >s around 0.53. Therefore, higher temperatures are favourable for the CO2 conversion, but lead to product distributions shifted to short-chain hydrocarbons. The main effect of the temperature is attributed to CO/CO2 shift equilibrium (KPCO-SH) and not to the increase of the rate constants. 5. CONCLUSIONS The following conclusions can be drawn from this kinetic analysis of fixed bed reactor model experiments: (i) experimental trends of conversion and yield versus CO2 inlet concentration are represented very well though the kinetics was determined a) for a catalyst with different K content and b) only for a balanced CO2 syngas. The fixed bed reactor model could be validated by independent experimental data, (ii) The limiting reaction is the FT reaction which means a further increase in the catalyst's shift activity is not required, (iii) The model is also capable to predict the measured effects of ^-deficiency, (iv) An increase of temperature is favourable for the CO2 conversion. The effects can be attributed mainly to the CO/CO2 shift equilibrium (KPCO-SH) where increasing temperature favors the side of CO. ACKNOWLEDGEMENT Financial support from the German Government (BMBF) is gratefully acknowledged (ReFuelNet, contract number 015F0118). NOTATIONS aj,H2O

-

bj,CO2

-

Bo

uL

convective

D^

diffusive

Dax dp

AHR EAJ

K k' o .

k' = k'o . e x p f - ^ - 1

^ RTj

Inhibition coeffient for H2O in reaction j Inhibition coeffient for CO2 in reaction j Bodenstein number

Drt

m2/s m J/mol J/mol mol/(s kg Pa) mol/(s kg Pa)

Fitting parameter for reaction j Axial dispersion coefficient Catalyst particle diameter Heat of reaction Apparent activation energy Reaction rate constant for reaction j Preexponiental factor in Arrhenius law

102

TV^P.CO-SH

_ PCO2PH2

v

^P.CO-SH ~

Equilibrium constant of the CO-shift t-

PHZOPCO

reaction

L meat N, Nc.i lii

m kg mol/s

u

P pi R / j

r

Pa Pa J/(molK)

p, = P(n ; /22ri;) 8.314 '

1

1 | dNj | \~T~\ VyV dt ) j

r =

j

mcat Tmod

x

mod = ,/.

K

Mass catalyst specific reaction rate of v . . . reactionj Isothermal reactor temperature

g s/cm 3

Modified residence time Gas velocity Conversion

mol/(skg) v

T —

Reactor length Mass catalyst Number of moles of i Number of carbon atoms in i Molar flow of i Stoichiometric coefficient of i in reaction j Total pressure Partial pressure of i Ideal gas constant

&J

Mn.NTP

u Xj

Xs =(ri i i n ~n{)lhlM

m/s -

Yi,c

Xc =

-

Yield of i related to total carbon fed

-

Selectivity of i related to k fed

=—; n

Sjk

+

n

CO2,in

N c i rij S; k = -— : —— n

k,in- n k

_( ZCO2=

Z

H2 =

Pco2

1

IPCO2+PCOJM

f ZH2

CO.in

Molar fraction of CO 2 in CO and CO 2

"

- *e feed

-

Stoichiometric weighted molar fraction ofH 2 in CO and CO 2 in the

\

n

^ !v3pCO2+2PcoJm

fted

REFERENCES 1. 2. 3. 4. 5.

Schaub, G., Unruh, D., Rohde, M., Synthetic Hydrocarbons Fuels and CO 2 Utilization, Proceedings 7th International Conference on CO2 Utilization, Seoul, October 2003 Riedel, T., Schaub, G., Jun, K.-W., Lee, K.-W., Kinetics of CO2 Hydrogenation on a K-Promoted Fe Catalyst, Ind. Eng. Chem. Res. 40 (2001) 1355-1363 Ryu, J.-S., Lee, K.-W., Choi, M.-J., Yoo, H.-S., The Synthesis of Clean Fuels from CO2 Rich Biosyngas, submitted to Korean J. Chem. Eng. (2003) Lee, K.-W., Personal communication (2002) Riedel, T., Reaktionen von CO 2 bei der Fischer-Tropsch Synthese - Kinetik und Selektivitat, Dissertation Universitat Karlsruhe (TH), Shaker Verlag 2003, Aachen


103

Hydrogen Production from Woody Biomass by Novel Gasification Using CO 2 Sorbent T. Hanaoka"'*, S. Fujimotoa, T. Yoshidaa, K. Kameib, M. Haradab, Y. Suzuki0, S. Yokoyamaa, T. Minowaa a

Biomass Technology Research Laboratory, National Institute of Advanced Industrial Science and Technology (AIST), Suehiro 2-2-2, Hiro, Kure, Hiroshima, 737-0197, Japan b Center for Coal Utilization, Japan (CCUJ), Daikyouchou 24, Shinjuku-ku, Tokyo, 160-0015, Japan c Clean Fuel Research Group, National Institute of Advanced Industrial Science and Technology (AIST), Onogawa 16-1, Tsukuba, Ibaraki, 305-8569, Japan In H2 production from woody biomass by steam gasification using CaO as a CO2 sorbent, the effect of reaction parameters such as the molar ratio of CaO to carbon in the woody biomass ([Ca]/[C]), reaction pressure, reaction temperature, and molar ratio of oxygen to carbon in the woody biomass ([O2]/[C]) on H2 yield was investigated. In the absence of CaO, the product gas contained CO2. On the other hand, in the presence of CaO ([Ca]/[C]=1, 2, and 4), no CO2 was detected in the product gas. At a [Ca]/[C] of 2, the maximum yield of H2 was obtained. The H2 yield was largely dependent on the reaction pressure, and exhibited the maximum value at 0.6 MPa. It is noteworthy that H2 could be obtained from woody biomass at a much lower pressure compared to other carbonaceous materials such as coal (> 12 MPa) and heavy oil (> 4.2 MPa) in steam gasification using a CO2 sorbent. H2 yield increased with increasing reaction temperature. The addition of oxygen had negative effects on the H2 production. Woody biomass is the one of the most appropriate carbonaceous materials in H2 production by steam gasification using CaO as a CO2 sorbent, taking the reaction pressure into account. 1. INTRODUCTION Hydrogen is a very promising fuel from the standpoint of the control of global warming because it generates only H2O and no CO2 when it is used [i], For example, there is a high efficiency device such as fuel cell. It is important to develop technology which produces hydrogen from renewable energy resources, taking the fossil fuel saving into account. Accordingly, it is earnestly hoped that a sustainable society would be realized in the future by the utilization of hydrogen derived from renewable energy. Biomass is a renewable energy resource, and is the only renewable organic material. Hence, it is extremely important to develop technology which produces hydrogen from biomass [ii]. In recent years, Lin et al. proposed a novel gasification method for H2 production from

*t.hanaoka@,aist. go. i p "Clean Gas Production from Biomass" sponsored by the Ministry of Economy, Trade and Industry

104

Moreover, this method could be carbonaceous materials using a CO2 sorbent [iii-vi]. applied to various carbonaceous materials such as coal, oil, biomass, and plastic [iii-vii]. This method for H2 production is a new technique integrating steam gasification of carbonaceous materials, water gas shift reaction and CO2 absorption using CaO as a CO2 sorbent in a single reactor. The conventional method consists of three processes (reaction (l)-(3)). However, in this method, hydrogen can be obtained in a single reactor (reaction (4))Carbonaceous materials + H2O -> CO + H2 CO + H 2 O ^ C O 2 + H 2 CaO + CO2 * CaCO3 Carbonaceous materials + CaO + 2H2O -> 2H2 + CaCO3

(1) (2) (3) (4)

In the present work, in order to obtain basic information, the effect of the reaction parameters such as the effect of the molar ratio of CaO as a CO2 sorbent to carbon in the woody biomass ([Ca]/[C]), reaction pressure, reaction temperature, and molar ratio of oxygen to carbon in the woody biomas ([O2]/[C]) was investigated on H2 yield using a batch reactor. 2. EXPERIMENTAL Japanese oak (0.106-0.250 mm) was used as a woody biomass sample. The analysis of carbon and hydrogen content in Japanese oak was performed on an elemental analyzer (Carlo Erba, NA-1500) and ash content was evaluated by burning in air in a muffle furnace at 873 K. The elemental composition was: C =48.9 wt%, H = 6.2 wt% and O = 44.7 wt% (by difference). Its ash content was 0.2 wt%. Ca(OH)2 powder (Wako Pure Chemical Industries, Ltd.) was used as a CO2 sorbent . Dehydration from Ca(OH)2 occurs at about 673 K; therefore, during heating up to the reaction temperatures (>873 K) in some runs, Ca(OH)2 will change into CaO. At a high pressure, Ca(OH)2 is stable without dehydration; therefore, in the present study, both CaO and Ca(OH)2 absorb CO2 as a CO2 sorbent. A conventional autoclave was employed as a batch reactor. Its inner volume was 50 cm3. Temperature and pressure in the autoclave were measured by a thermocouple and a pressure transducer, respectively. Japanese oak, Ca(OH)2, and distilled water was charged into the autoclave with a C/H2O molar ratio of 1/6. Before the runs, the gas phase in the autoclave was flushed with nitrogen, if necessary, oxygen was added to the gas phase. The reaction was started by heating the autoclave using an electric furnace. After the desired temperature had been reached, it was held for 10 min, and the autoclave was then removed from the furnace and cooled by a fan. When the temperature in the autoclave dropped to room temperature, the valve was opened to collect the product gas in a gas collection bag. Its volume was measured by a wet gas meter (W-NK-0.5, Shinagawa). The reaction mixture remaining in the autoclave was washed with distilled water to separate the filtrate (water soluble compounds) and solid by filtration. The solid including CaCO3 was added to an aqueous solution of HC1 to release CO2 according to equation (5): CaCO3 + 2HC1 -» CaCl2 + H2O + CO2

(5)

The volume of gas including CO2 obtained by HC1 addition was measured through the gas meter, and the gas was collected into another gas collection bag. The residue after releasing CO2 was then separated by filtration to obtain the filtrate and residue on a filter

105

paper. The residue was dried at 378 K overnight. The product gas and obtained gas by HC1 addition were analyzed by gas chromatography with TCD (GC323, GL Sciences) and FID (GC353B, GL Sciences). The amount of inorganic carbon (CO2) in both filtrates was measured by a TOC meter (TOC-5000A, Shimadzu). The CO2 obtained by HC1 addition and CO2 dissolved in the filtrates were regarded as CO2 fixed in CaO. 3. RESULTS The effect of the molar ratio of CaO as a CO2 sorbent to carbon in the woody biomass ([Ca]/[C]) on H2 yield was investigated in the range between 0 and 4. Table 1 shows the gas yield for each ratio. The reaction temperature was 923 K, and the reaction pressure was 0.6 MPa. The product gas in the absence of CaO ([Ca]/[C]=0) contained not only H2 and hydrocarbon but also CO2. On the other hand, in the presence of CaO ([Ca]/[C]=1, 2, and 4), the product contained no CO2, and mainly contained H2. Small amounts of hydrocarbons such as CH4, C2Hg, and C3H8 were detected, and the gas obtained by HC1 addition contained only CO2. Almost no residue was obtained in each run. The H2 yield in the absence of CaO was lower than those in the presence of CaO, and exhibited maximum valuesata[Ca]/[C]of2. The effect of the reaction pressure on H2 yield was investigated in the range between 0.3 MPa to 8.4 MPa. Table 2 shows the gas yield for each pressure. The reaction temperature was 923 K, and [Ca]/[C] was 2. The H2 yield increased with increasing reaction pressure, and exhibited maximum values at 0.6 MPa, and decreased with increasing reaction pressures higher than 0.6 MPa. Table 1 Effect of [Ca]/[C] on gas yield [Ca]/[C] 0 1 2 4 a)

Pn)duct gas ]/ield / mL,(STP)g-v /ood"1 , CH4 others CO2a) C0 2 H2 484 725 817 556

114 214 187 114

9 32 44 26

232 0 0 0

0 345 477 482

CO2 fixed in CaO; 923 K, 0.6 MPa, holding time; 10 min

Table 2 Effect of reaction pressure on gas yield Prc»duct gas 1'field 1 mL (STP) g-iivood"1 CH4 " others CO 2 a ) CO2

Reaction pressi. MPa

H2

0.3 0.6 1.3 2.0 6.4 8.4

518 817 602 511 352 350

a)

148 187 130 131 157 160

10 44 30 32 44 41

0 0 0 0 0 0

283

477 268 244 179 185

CO2 fixed in CaO; 923 K, [Ca]/[C]=2, holding time; 10 min

106

Table 3 Effect of reaction temperature on gas yield Product gas yield / mL (STP) g-wood"1 Reaction temp. K CH4 others CO2 CO2a) H2 873 923 973 a)

496 817 875

109 187 159

34 44 14

0 0 0

342 477 363

C 0 2 fixed in CaO; 0.6 MPa, [Ca]/[C]=2, holding time; 10 min

Table 4 Effect of [O2]/[C] on gas yield [O2]/[C] H2

a)

Prc>duct gas i>deld / mL (;STP) g-vrood"1 others CH4 CO2 CO2a)

0 0.025

817 602

187 110

44 27

0 0

477 420

0.5 0.6

657 584

91 100

18 18

0 0

338 301

CO2 fixed in CaO; 923 K, 0.6 MPa, [Ca]/[C]=2, holding time; 10 min

The effect of reaction temperature on H2 yield was investigated at 873, 923, and 973 K. Table 3 shows the gas yield for each temperature. The reaction pressure was 0.6-0.7 MPa, and [Ca]/[C] was 2. The H2 yield increased with increasing reaction temperature. In order to restrict the generation of hydrocarbon, the effect of the molar ratio of oxygen to carbon in the woody biomass ([O2]/[C]) was investigated in the range between 0 and 0.6. Table 4 shows the gas yield for each ratio. The reaction temperature was 923 K, reaction pressure was 0.6 MPa, and [Ca]/[C] was 2. The yield of hydrocarbon such as CH4, and others (C2H6, and C3H8) decreased with increasing [O2]/[C], whereas the H2 yield also decreased. 4. DISCUSSION CaO plays the role not only of a CO2 sorbent, but also that of a catalyst for biomass gasification; therefore, both effects of CaO must be discussed. At a [Ca]/[C] higher than unity, the product gas contained no CO2. These results indicate that the amount of CaO as a CO2 sorbent at a [Ca]/[C] higher than unity would be enough to absorb CO2. Some researchers reported that CaO catalyzed the pyrolysis of biomass [viii-x]. Figure 1 shows the reaction mechanism in the H2 production from biomass. In the present work, the H2 yield in the presence of CaO were higher than that in the absence of CaO. Both woody biomass and CaO are solid phase. Thus, CaO would hardly affect the first decomposition of biomass such as reactions (7) and (8), and would play a promotive role in reaction (10). The H2 yield at a [Ca]/[C] of 2 were higher than that at a [Ca]/[C] of unity. This result would be attributed to the increase of the contact frequency between tar (vapor) and CaO in the range between 0 and 2. On the other hand, at a [Ca]/[C] of above 2, since

107

the excess CaO might restrict diffusion of tar, the H2 yield decreased. This is an interesting phenomenon, and more detailed study is necessary. These results suggest that the [Ca]/[C] ratio has an optimum value for H2 production from woody biomass by novel gasification using CaO. The H2 yield was largely dependent on the reaction pressure. At more than 0.6 MPa, the H2 yield decreased with increasing reaction pressure. In the reactions (7), (10), and (11) in Figure 1, the molar number increases. In these reactions, a low pressure is favorable for H2 production. Consequently, because the pyrolysis and steam reforming would be the rate-determining step at more than 0.6 MPa, the H2 yield decreased with increasing reaction pressure. On the other hand, up to 0.6 MPa, the reaction pressure would clearly promote the H2 production; however, in order to clarify the effect of pressure, further study is necessary. The H2 yield increased with increasing reaction temperature. This would be attributed to the promotion of endothermic reactions such as reactions (7), (10), and (11). Higher ratio of [O2]/[C] led to the decrease of not only hydrocarbon yield but also H2 yield; moreover, CO2 yield also decreased with increasing [O2]/[C]. The addition of oxygen had negative effects on the H2 production from biomass using CaO. Lin et al. reported H2 production from coal using CaO at more than 12 MPa [v], and Sato et al. reported H2 production from heavy oil using CaO at more than 4.2 MPa [vii]. At 923 K, the H2 yield in coal was 1.06 H2-mol/C-mol in the feedstock at 60 MPa (with NaOH as catalyst), and that in heavy oil was 0.24 H2-mol/C-mol in the feedstock at 4.2 MPa. On the contrary, in the present work, at 923 K, 0.6 MPa, the H2 yield was 0.89 H2-mol/C-mol in the feedstock It is noteworthy that H2 production was observed even at less than 2.0 MPa and the maximum yield of H2 could be obtained at a pressure as low as 0.6 MPa. The reactivity of woody biomass was quite different from those of coal and heavy oil, and woody biomass is one of the most appropriate materials in the H2 production using CaO. In order to enhance H2 yield, we intend to study the effect of reaction pressure and [Ca]/[C] and the addition of a catalyst for gasification, in further detail.

*

(6)

{

n / Biomass ^-

Char

Char

V) / /

S

\

T*

(VaPOT)

\(8) X^

\o'+CO e + C O 2 ^ O + CO + e'

(2) (3)

In fact, the objectives of the addition of CO2 into the plasma methane conversion are to increase the methane conversion, to change the product distribution and increase the yield of objective products and to reduce the specific energy consumption [18,19,28]. Figure 2 shows the relative energies required for the generation of methyl radicals from the reactions between

112

methane and H, O, OH and O". The energy barriers for these reactions were 49.13 (TS2), 46.44(755), 7.11 (TS4) and 39.67(TS5) kJ/mol, respectively. After the initiation of discharge reactions, the oxygen species will play an important role in the generation of methyl radicals. From the literature [36-41], there are two possibilities for CO2 dissociation. One is via CO2" and the other is via CO, as shown in eqs. (2) and (3). The experiment showed that CO2" could be formed during the electrochemical reduction of CO2 [36-39] CO2" can also be formed within the discharges. The calculated reaction enthalpy to form CO2"is 76.07 kJ/mol, which is slightly higher than the experimental value (48.14-62.76 kJ/mol [40]).

Figure 2. Relative energies for reactions between some species (H, O, OH and 0") and CH4 Figure 3 presents the reaction scheme of CO2 dissociation. Evidently, the dissociation via CO needs much higher energy than the attachment to form CO2".

Figure 3. The reaction scheme of CO2 dissociation 3.2. The formation of oxygen-containing species To produce oxygenates from methane and carbon dioxide using cold plasmas, in addition to methyl radicals, some oxygen-containing species like OH and COOH are very necessary too. These oxygen-containing active species are normally generated from the secondary reactions or other multi-step reactions, including the recombination of radicals and dissociation of

113 products. As discussed below, the formation of OH, OH" and COOH" is one of principal pathways to synthesize the ultimate oxygenated products. The reaction energies for the formation of OH, OH" and COOH" (from the reactions O+H—>OH, O" +H->OH" and H+CO2"->COOH") are -465.68, -459.18 and -45.18 kJ/mol, respectively. These reactions can easily occur since they are exothermic. Moreover, from the recombination of the radicals, like CH3 and CH2, the radical chain reactions are started and higher hydrocarbons and oxygenates are thereby produced. These reactions were also exothermic, for example, the reaction energies for CH3+CH2->CH3CH2 and CH3CH2+CH2->CH3CH2CH2 are -448.11 and -430.95 kJ/mol. 3.3. The pathways for the formation of higher hydrocarbons According to the radical chain mechanism, the initiate reaction step to form the higher hydrocarbons is the methane dissociation by the electrons or the hydrogen abstraction by some active species (O, H and OH). The reactions include CH4+(e, O, H, OH, O")—>CH3+(H+e', OH, H2, H2O,OH"). Then the chain growth reactions proceed as follows: CH 3 +CH 3 ^CH 3 CH 3 CH3CH3+e-»CH3CH2+H+e' CH3CH3+H->CH3CH2+H2 CH3CH2+CH3->CH3CH2CH3 CH3CH2CH3+e^CH3CH2CH2+H+e' (straight) CH3CH2CH3+e->-CH3CHCH3+H+e' (branched) CH3CH2CH2+CH3->CH3CH2CH2CH3 (straight) CH 3 CHCH 3 +CH 3 ^ CH3CH(CH3)CH3 (branched) 2CH3CH2-+CH3CH2CH2CH3 CH3CH2CH2CH3+e->CH3CH2CH2CH2+H+e' (straight) CH3CH2CH2CH3+e^CH3CH2CHCH3+H+e' (branched) CH3CH2CH2CH2+CH3->CH3CH2CH2CH2CH3 (straight) CH3CH2CHCH3+CH3^CH3CH2CH(CH3)CH3 (branched)

(4) (5) (6) (7) (8) (9) (10) (11) (12) (13) (14) (15) (16)

Such reactions are responsible for the synthesis of higher hydrocarbons within discharges. In addition, alkenes and acetylene can be produced from the CH2 and CH radicals. At the same time, the formed hydrocarbon radicals will involve in these reactions to produce oxygenates, like alcohols and acids, which will be explained in detail in the following sections. Table 1 shows the enthalpies and entropies of part of important reactions for the synthesis of higher hydrocarbons. It suggests these reactions easily occur under the condition of cold plasmas. 3.4. The pathways for the formation of alcohol products From the above analysis, it can be seen that there exist OH, OH" and a great amount of hydrocarbon radicals, e.g., CH3, CH3CH2 and so on. Therefore a series of alcohol products can be easily synthesized within cold plasmas. Methanol can also be obtained by the reaction CH4+O—»CH3OH. Another important pathway to synthesize alcohols is via the dissociation of lower alcohols, e.g., CH3OH +e->CH2OH + H +e'. The reaction energy is 414.63 kJ/mol. The formed CH2OH can react with hydrocarbon radicals to generate higher alcohols. The enthalpies and entropies of the reactions to synthesize alcohols are presented in Table 2. Obviously, the

114

main pathways to form alcohols via OH, OH" and CH 2 OH are thermodynamically favored. TABLE 1: Main thermodynamic parameters of the synthesize main hydrocarbon products reactions at 298.15 K

Reaction CH 3 +CH 3 ^CH 3 CH 3 CH 3 CH 2 +CH 3 ^CH 3 CH 2 CH 3 CH 3 CH 3 +e^CH 3 CH 2 +H+e' CH3CH3+H->CH3CH2+H2 CH3CH2CH3+e->CH3CH2CH2+H+e' CH3CH2CH3+e->CH3CHCH3+H+e' CH3CH2+CH3CH2^CH3CH2CH2CH3 CH3CH2CH2+CH3^CH3CH2CH2CH3 CH3CH2CH2CH3+e-^CH3CH2CH2CH2+H+e' CH3CH2CH2CH3+e-»CH3CH2CHCH3+H+e' CH3CH2CH2CH2+CH3^CH3CH2CH2CH2CH3 CH3CH2CHCH3+CH3->CH3CH2CH(CH3)CH3 CH3CHCH3+CH3-^CH3CH(CH3)CH3 CH3CH2CH2+CH3CH2^CH3CH2CH2CH2CH3 CH3CHCH3+CH3CH2^CH3CH(CH3)CH2CH3 2CH 3 CH 2 CH 2 ^C 6 H 14 C5H1I+C5H,,-»C10H22 CH 2 +CH 2 ^CH 2 CH 2 CH 2 +CH 3 ^CH 2 CH 2 +H CH 2 CH 2 ^CH 2 CH +H CH2CH+CH3->CH2CHCH3 CH+CH^C 2 H 2

AH (kJ/mol) .370.28 -359.99 429.75(AE) -10.66 (AE) 443.65 (AE) 425.14 (AE) -343.34 -366.78 442.20 (AE) 425.99 (AE) -360.04 -348.00 -351.55 -349.34 -332.05 .355.35 -450.73 .727.03 -276.73(AE) 534.88 (AE) -472.909 -1111.77

AS J/mol-K -170.28 -164.76

-171.73 -174.67

-172.95 -190.13 -188.93 -177.98 -191.68 -168.21 -187.33 -153.06

-166.14 24.23

TABLE 2: Main thermodynamic parameters of the synthesize main alcohol products reactions at 298.15 K Reaction AH (kJ/mol) AS J/mol-K -419.25(AE) CH 4 +O->CH 3 OH CH 3 +OH->CH 3 OH -392.87 -136.98 CH 3 CH 2 +OH-^CH 3 CH 2 OH -389.42 -141.47 CH 3 CH 2 CH 2 +OH^CH 3 CH 2 CH 2 OH -394.11 -140.33 CH 3 +OH"->CH 3 OH+e -273.02 -65.08 CH 3 CH 2 +OH-->CH 3 CH 2 OH+e -269.76 -69.65 CH 3 CH 2 CH 2 +OH"->CH 3 CH 2 CH 2 OH+e -274.47 -68.65 CH 2 OH+CH 3 ->CH 3 CH 2 OH -171.23 -352.93 CH 2 OH+CH 3 CH 2 ^CH 3 CH 2 CH 2 OH -167.16 -334.18 3.5. The pathways for the formation of acid products The acids are principally synthesized via two pathways with CO 2 " or CO. The synthesis of

115

formic acid via CO2" route can also be carried out by electrochemical reduction of CO2 [36-39] Kaneco et al. thought that CO2~ is produced through electron transfer during electrochemical reduction of CO2 and then COOH is generated from CO2". In this work, the carboxyl was generated by the reaction CO2" +H—>COOH" [43], Then the acids were obtained by the following reactions: H+COOH-^HCOOH+e CH3+COOH"^CH3COOH+e CH 3 CH 2 +COOH^CH 3 CH 2 COOH+e

(17) (18) (19)

Another reaction route via CO is similar to the catalytic route [44]. For example, the formation of acetic acid via CO route can be expressed by the following reactions:

Or

CH 3 +CO^CH 3 CO CH 3 CO+O^CH 3 COO CH3CO+O"^CH3COO" CH3COO+H->€H3COOH CH 3 COO+H^CH 3 COOH+e CH3CO+OH->CH3COOH CH3CO+OH"-»CH3COOH+e

(20) (21) (22) (23) (24) (25) (26)

From the above analysis, it can be seen that the carbon atom of carboxyl comes from CO2 or CO, the dissociated product of CO2. Table 3 presents some important thermodynamic parameters for reactions of syntheses of acids. Between the two routes (via CO2" and CO) to synthesize acids, it can be seen that the route via CO2" is more thermodynamically favored than that via CO. The energy requirement to form CO2" from CO2 is much lower than that to form CO. The reaction steps via CO2" are also less than via CO. Most of CO generated has become the ultimate product before forming acids. TABLE 3: Main thermodynamic parameters of the synthesize main acids products reactions at 298.15 K Reaction AS J/mol-K AH (kJ/mol) H+COOH"->HCOOH+e -484.40 (AE) CH3+COOH"^CH3COOH+e -433.78 -177.07 CH 3 CH 2 +COOH^CH 3 CH 2 COOH+e -416.21 -191.83 -104.35 (AE) CH 3 +CO^CH 3 CO -509.82 (AE) CH3CO+O-»CH3COO CH3COO+H->CH3COOH -439.60(AE) CH3CO+OH->CH3COOH -464.20 -155.41 CH 3 CO+OH-^CH 3 COOH+e -344.83 -83.65 CH3CH2+CO->CH3CH2CO -75.17 -155.13 -467.53 (AE) CH3CH2CO+O-*CH3CH2COO -477.71 (AE) CH3CH2COO+H->CH3CH2COO+H

116

CH3+CH2COOH^CH3CH2COOH CH3CH2+CH2COOH->CH3CH2CH2COOH

-348.326 -331.586

-183.22 -182.52

In addition, higher acids can be also obtained via the dissociation of lower acids. For example, acetic acid can be dissociated: CH3COOH + e—>CH2COOH + H + e', with which the energy needed is 425.93 kJ/mol. Then the C3 and C4 acids can be formed by the following reactions: CH3+CH2COOH->CH3CH2COOH CH 3 CH 2 +CH 2 COOH^CH 3 CH 2 CH 2 COOH

(27) (28)

The reaction thermodynamic parameters were also presented in Table 3. It can be seen that these reactions can easily occur once the formation of CH2COOH. 3.6. The main reactions to produce CO and H2 Carbon monoxide is principally generated from the dissociation of CO 2 , as shown in Figure 3. The experiment also showed the selectivity of CO was almost proportional to CO2 concentration in the feed [28]. In addition, the reaction between the hydrocarbon radicals and oxygen atom can also produce CO [45]. The relative energies for the formation of CO and H2 via the reaction CH3+O is presented in Figure 4. The intermediate CH3O formed firstly, then the intermediate is converted into H2COH or H+H2CO via the TS6 and 757. The former requires more energy. Of course, the H2COH can be converted into either H+H2CO or H2+HOC through TS8 and TS9. The H+H2CO or H2+HOC pass the transition state (TS11 and TS10) and is converted into H, CO and H2. The two processes can also be described as these reactions: CH2+O—>CO+H2 and CH+O—>CO+H. Moreover, we thought the further oxidation of hydrocarbon and oxygenate products or the further oxidation of carbon-containing radicals will induce an extra production of CO.

Figure 4. Relative energies for the formation CO and H2 via the reaction CH3+O Regarding the hydrogen production, the recombination of H atom into H2 is the main

117 source of hydrogen, with which the energy change is -440.58 kJ/mol. H2 can be obtained during the hydrocarbon decomposition process. 4. CONCLUSIONS The present investigation has confirmed that the principal obstacle for synthesis from methane and carbon dioxide using cold plasmas is the dissociation of the feed gases. Thermodynamically, once the feed gases are dissociated, a series of products can be then easily produced. The pathways of the formation of various products, including syngas, higher hydrocarbons and oxygenates, have been also investigated. Upon the present study, CO mainly comes from the dissociation of CO2 or hydrocarbon and from the reaction of CH3+O. Oxygenates are produced via the reactions of hydrocarbon radicals and the oxygen-containing species (e.g., OH, OH' and COOH") or via the dissociation of lower oxygenates. The latter also follows the radical chain growth mechanism. The results obtained from the present DFT study are very helpful for the further improvement in the synthesis from CH4 and CO2 via cold plasmas. ACKNOWLEDGMENT The authors are grateful for supports from the Key Fundamental Research Project of Ministry of Science and Technology of China (973, G1999022402), Major Research Foundation of Ministry of Education of China and ABB Switzerland Ltd. REFERENCES 1. Choudhary, V.R.; Baneriee, S.; Rajput, A.M. Appl. Catal. A 2002, 234(1), 259-270. 2. Verma, S.S. Energ.Conver.and.Manag. 2002, 43(15) 1999-2008. 3. Tofan, C ; Klvana, D.; Kirchnerova, 3. Appl. Catal.B: Envim.2002, 36(4) 311-323. 4. Suib, S. L.; Zerger, R. P. J. Catal. 1993,139, 383-391. 5. Liu, C ; Mallinson, R. G; Lobban, L. L. Appl. Catal. A 1999, 778, 17-27. 6. Liu, C ; Lobban, L. L.; Mallinson, R. G. J. Catal 1998,779, 326- 334. 7. Liu, C ; Mallinson, R. G; Lobban, L. L. Appl. Catal. A 1997,164, 21-31. 8. Zhang, J.-Q.; Zhang, J.-S.; Yang, Y.-J.; and Liu, Q. Energy Fuels 2003, 17(1), 54-59. 9. Yao, S. L.; Ouyang, F.; Nakayama, A.; Suzuki, E.; Okumoto, M.; Mizuno, A. Energy Fuels 2000, 14, 910-914. 10. Kado, S.; Sekine, Y; Fujimoto, K. Chem. Commun. 1999, 2485- 2486. 11. Oumghar, A.; Legrand, J. C ; Diamy, A. M.; Turillon, N. Plasma Chem. Plasma Process. 1995, 75,87-107. 12. Zhu, A.; Gong, W.; Zhang, X.; Zhang, B. Sci. China B (in Chinese) 2000, 30 (2), 167-173. 13. Okumoto, M.; Rajanikanth, B. S.; Katsura, S.; Mizuno, A. IEEE Trans. Ind. Appl 1998, 34 (5), 940-944. 14. Matsumoto, H.; Tanabe, S.; Okitsu, K.; Hayashi, Y; Suib, S. L. J. Phys. Chem. A 2001, 105(21), 5304-5308. 15. Yao, S. L.; Ouyang, F.; Nakayama, A.; Suzuki, E. Trans. Mater. Res. Soc. Jpn. 2000, 25 (1), 373-376. 16. Bugaev, S. P.; Kozyrev, A. V.; Kuvshinov, V. A.; Sochugov, N. S.; Khryapov, P. A. Plasma Chem. Plasma Process. 1998,18 (2), 247- 262. 17. Larkin, D. W.; Caldwell, T. A.; Lobban, L. L.; Mallinson, R. G Energy Fuels 1998, 72,

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740-744. 18. Eliasson, B.; Liu, C ; Kogelschatz, U. Ind. Eng. Chem. Res. 2000, 39 (5), 1221-1227. 19. Liu, C. J.; Xue, B.; Eliasson, B.; He, R; Li, Y.; Xu, G. H. Plasma Chem. Plasma Processing 2001, 21 (3), 301-310. 20. Li, Y.; Xu, G. H.; Liu, C. J.; Eliasson, B.; Xue, B. Z. Energy Fuels 2001, 75 (2), 299-303. 21. Czernichowski, A. Private communications. 22. Mutaf-Yardimci, O.; Saveliev, A. V.; Fridman, A. A.; Kennedy, L. A. Int. J. Hydrogen Energy 1998, 23 (12), 1109-1111. 23. Kado, S.; Urasaki, K.; Sekine, Y; Fujimoto, K. Chem. Commun. 2001, 415-416. 24. Gesser, H. D.; Hunter, N. R.; Probawono, D. Plasma Chem. Plasma Process. 1998,18 (2), 241-245. 25. Bromberg, L.; Cohn, D. R.; Rabinovich, A. Energy Fuels 1998, 72,11-18. 26. Chang, M.; Huang, C. J. Adv. Oxid. TechnoL 1999, 4 (3), 333-338. 27. Liu, C.-J.; Li, Y; Zhang, Y.-R; Wang, Y; Zou, J.; Eliasson, B.; Xue, B. Chem. Lett. 2001, 1304-1305. 28. Li,Y; Liu, C.-J.; Eliasson, B.; Wang,Y Energy Fuels 2002,16,864-870 29. Okumoto, M ; Su, Z.; Katsura, S.; Mizuno, A. IEEE Trans. Ind. App\. 1999, 35 (5), 1205-1210. 30. Delley, B. J. Chem. Phys 1990, 92, 508-517. 31. Delley, B. J. Chem. Phys 2000, 113, 7756-7764. 32. Perdew, J. P.; Wang, Y Phys. Rev B 1992, 45, 13244-13249. 33. Perdew, J. P.; Wang, Y Phys. Rev B 1986, 33, 8800-8805. 34. Hehre, W.J.; Radom, L.; Pople, J.A. Ab initio molecular orbit theory. 1986, pp 275-278. 35. Zhang, Y-R; Li, Y; Liu, C.-J.; Eliasson, B. Influence of electrode configuration on direct methane conversion with CO2 as a co-reactant using dielectric-barrier discharges. In ACS Symposium Series, No. 852: Utilization of Greenhouse Gases, Ed. Liu, C.-J.; Mallinson, R.; Aresta, M. 2003. pp 100-115 36. Kaneco, S.; Iiba, K.; Ohta, K.; Mizuno.T.; Saji, A. J.Electroanal.Chem. 1998, 441, 215-220. 37. Kaneco, S.; Iiba, K.; Ohta, K; Mizuno.T.; Saji, A. Electrochim.Acta. 1998, 44, 573-578. 38. Mizuno, T.; Kawamoto, M.; Kaneco, S.; and Ohta, K. Electrochim.Acta 1998 43, 899-907. 39. Kaneco, S.; Iiba, K.; Hiei, N.-H.; Ohta, K; Mizuno, T.; Suzuki, T. Electrochim.Acta 1999 44,4701-4706. 40. Compton, R.N.; Reinhardt, P.W.; Cooper, C. D. J.Chem.Phys. 1975 63, 3821-3826. 41. Liu, C ; Marafee, A.; Hill, B.; Xu, G; Mallinson, R.; Lobban,L. Ind.Eng.Chem.Res. 1996 35,3295-3301. 42. Schroder, D.; Schalley, C.A.; Harvey, J.N.; Schwarz, H. Inter. J.Mass.Spectro. 1999 185 , 25-35. 43. Wang, J.-G; Liu, C.-J.; Zhang, Y.-R; Eliasson, B. Chem.Phys.Lett. 2003 368, 313-318. 44. Moser, W.R.; Marshik-Guerts, B.J.; Okrasinski, S. J. / Mol. Catal. A:Chem 1999 143, 71-83. 45. Preses, J.M.; Fockenberg, C ; Flynn , G.W. J.Phys. Chem. A. 2000 104, 6758-6763.


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Synthesis Gas Production from CO2 and H2O with Nonthermal Plasma Shigeru Futamura* and Hajime Kabashima National Institute of Advanced Industrial Science and Technology AISTTsukuba West, 16-1 Onogawa, Tsukuba, Ibaraki, 305-8569 Japan E-mail: s-futamuraffi.aist.go.jp ABSTRACT: Synthesis gas was produced from CO2 and H2O with nonthermal plasma. A ferroelectric packed-bed reactor worked much better than a silent discharge plasma reactor. CO: and H2O competitively reacted to give CO and H2, respectively. Arbitrary molar ratios of H; to CO were obtained by controlling that of H2O to CO2. Energy conversion efficiency decreased with water content, and its maxima were observed in its functions of reactor energy density. 1. INTRODUCTION The average global ground temperature has been increasing by C since 1861. and the contribution of CO2 occupies ca 64 %.' Recognition of the CO2 influence on the accelerated global warming' stimulated researches relevant to the development of technologies for the capture, sequestration,2 and fixation of CO2J 4 along with those for energy conservation. Many of CO? fixation technologies are still in the phase of fundamental research. Photochemical reduction of CO 2 to CO is achieved with some of rhenium complexes/ but UV light is necessary in the complicated processes of CO2 reduction and the quantum yield is lower than unity. Direct methanol synthesis from CO2 is possible with Cu/ZnO-based catalysts at 200 ~ 300°C.4 However, this reaction is not cost-effective because a high pressure of H2 [5.0 MPa of H2/CO (3:1 v/v)] derived from electrolysis is assumed to be the H2 source. Preceding processes for plasma reforming of methane are carried out at rather high temperatures5 6. Recently, we have been applying nonthermal plasma to H2 production from small molecules such as water,78 methane,8 and methanol,8 and steam reforming of light paraffins such as methane,910 ethane,10 propane,10 and 2,2-dimethylpropane10 at ambient temperature. In H2 production, the substrate reactivity decreases in the order: methanol > methane > water. In the steam reforming of light paraffins, methane is the most desirable substrate because the highest H2 yield is obtained with almost quantitative recovery of carbon atoms in the presence of sufficient water. The technical merits of nonthermal plasma consist in quick response and convenient operation at ambient temperature. It has been shown that the energy efficiency of the reactor is highly affected by the plasma-generating methods.7""' These results urged us to explore the possibility of H2O-CO2 as an alternative substrate

120

system to give synthesis gas in the reforming with nonthermal plasma. This paper will present the novel results on the reactor effect on the reaction behavior of H2O-CCK product yields as functions of reactor energy density (RED [kJ/L]), water concentration dependence of product yields, [H2]/[CO], carbon balance, and energy conversion efficiency, and the successful results of a continuous operation. 2. EXPERIMENTAL A ferroelectric packed-bed reactor (FPR) and a silent discharge reactor (SDR) used in this research were described in detail elsewhere."' 2 Gas flow rate was set at 100 mL min"1 (residence time 44.5 s) and 50 mL min"1 (residence time 3 s) for FPR and SDR, respectively. The both reactors employed 50 Hz ac and high voltage up to 8 kV was applied to both of them. No breakdowns occurred during operations within their maximum voltages. CO2 balanced with N2 in a standard gas cylinder was introduced to the reactor through a Teflon tube by adjusting the CO2 concentration and flow rate with sets of mass flow controllers after humidification in a water-bubbler in a thermostatic bath. Water concentrations were determined with a dew point hygrometer, and they were controlled within the range of 0.5 ~ 2.5 %. Reforming of H2O and CO2 was carried out at room temperature and an atmospheric pressure after the CO2 was saturated in H2O, and its concentration was kept constant on the inlet of the reactor. H2 and methane were quantified by a TCD-GC with a packed column of Molecular Sieve 13X. CO, CO2, ethane, ethylene, and acetylene were quantified by a GC equipped with a TCD and a FID with a packed column of Porapak Q+N and Molecular Sieve 13X. As a measure of energy density for FPR and SDR, RED will be used later (1), where Power denotes the plug-in power. The yields of H,, CO, and O, were defined in (2), (3). and (4), respectively. RED (kJ / L) = A X Power (kW) / [Gas flow rate (L/min) / 60] (A = 0.1 and 0.2 for FPR and SDR)

(1)

H2 yield (mol%) = 100 X [H2 concentration (ppm) / initial H2O concentration (ppm)]

(2)

CO yield (mol%) = 100 X [CO concentration (ppm) / initial CO2 concentration (ppm)]

(3)

O2 yield (mol%) = 200 X [O2 concentration (ppm) /{[H2O]ini,,ai + [CO]in,tiai (ppm)}]

(4)

3. RESULTS AND DISCUSSION Table 1 shows the reactor effect on the reaction behavior of H2O -CO2 in N2. A 24.6fold higher CO2 conversion was obtained with FPR than with SDR. Since the lifetime of the energetic electron is shorter than 100 ns, these data cannot be rationalized by the residence time difference between FPR and SDR. Similar types of barrier discharge plasmas are

121

attained in both the reactors, but it is considered that electrons with higher energies are more highly populated in FPR than in SDR.7"10 Based on these results, only FPR was used as a reactor in the subsequent reactions. Table 1 Reactor Effect on CO2 Conversion and Product Distribution Reactor CO2 Conv. Product yield (mol%) [HJ/[CO] (mol%) H7 CO O, FPR 12.4 2.1 12.3 11.8 2.8 SDR 0.5 0.7 >0.5 [H,O ] = 2.0 %; [CO,] = 1.0 %; background gas N2; Q = 100 and 50 mL / minfor FPR and SDR, respectively. RED = 1.20 kJ / L. CO2 conversion increases with RED (Fig. 1). Water addition suppresses CO2 conversion, but its decrement levels off at > [H20] = 1.0 %. Similarly, H2 formation from H2O was suppressed by CO2, but its additive effect was much smaller than that of H2O toward CO formation from CO2. Irrespective of the reactor type, comparable yields were obtained for H2 and CO. In the case of FPR, the O2 yield was much lower than those Reactor energy density (kJ/L) of H2 and CO. In the case of SDR, Fig. 1 Additive effect of water on CO: conversion the O2 amount was below the detection limit. The yields of H2, CO, and O2 increase with RED in the case of FPR (Fig. 2). O2 reacts with H or H2 to give back H2O and also quenches CO2 deoxygenation.13 Therefore, some adsorbents of O2 or its transfer materials such as membranes should be housed inside of the reactor to increase the yields ofH 2 and CO. Figure 3 shows the plot of CO yield vs. CO? conversion. The slope of the dotted line is unity, and the data points on this line show that the carbon atoms in the reacted CO2 are quantitatively recovered as CO. In the absence of H2O, CO yield saturates at 85 % with an increase in RED. Meanwhile, a good carbon recovery is obtained on addition of H2(). The stoichiometry of the CO2 deoxygenation process shows that a half amount of O2 is produced compared to that of CO, but the data in Fig. 3 suggest that some unidentified reactions of CO2 and CO occur in the absence of H2O. H?O itself or its oxygen atoms may be involved in the oxidation of carbonaceous materials derived from CO2.

122

Figure 4 shows the molar ratio of H2 to CO as a function of RED. This ratio depends on the H2O concentration, but almost the same ratios are obtained, irrespective of RED. These data suggest that compositions of synthesis gases from H2O-CO2 mixtures can be controlled by selecting pertinent ratios for initial concentrations of H2O and CO2. This trend sharply contrasts with those observed in the steam reforming of light paraffins.7"10 In this reaction, CO2 and H2O are deoxygenated independently from each other. Fig. 5 shows a plausible mechanism for the reforming of H2O and CO2 in nonthermal plasma [(1) and (2)]. They competitively react to give H2 and CO, respectively. The data in Figs. 1 and 4 suggest almost no interaction between CO2 and H2O. O2 is formed from oxygen atoms in (3), but its lower yields suggest some pathways for their dissipation [(4)]. CO deoxygenation [(2)] rapidly proceeds in nonthermal plasma in the absence of O2.13 However, the reaction of CO and O(3P) is much slower [(6)]. High carbon recoveries are obtained even in the absence of water, and the carbon atoms in the reacted CO2 are almost quantitatively

Fig 3 Carbon balance in the reaction of CO2-H:O in N:

recovered as CO in the presence of H2O. Since the bond strength of CO is 11.10 eV. carbon formation from CO is neglected under the reaction conditions. O('D) decays to O(JP) in the presence of excessive CO2 [(5)], but this is one of the active oxygen species to consume H2 [(9)] along with OH radical [(10)]. CO2 is regenerated from the reaction of CO and H2O in nonthermal plasma [(8)], but the efficiency of this reaction is not necessarily high.'J In this research, the energy conversion efficiency (ECE) in this reaction was defined in (5). ECE (%) = 100 X {E(H2) [W]+ E(CO) [W]} / EK{orm [W]

(5)

E(H2), E(CO), and Ereform denote the heats of combustion for H2 and CO formed, and the energy consumed in the reforming of H2O and CO2, respectively.

123

ECE decreases with water content, but almost no change is observed in the efficiencies at [H2O] > 2.0 % (Fig. 6). Lower efficiencies in the presence of water reflect the reversibility of the H2O reaction and formation of active oxygen species such as OH radical and oxygen atoms. Interestingly, maxima are observed for ECE as functions of RED. Only modest ECE is obtained in this reaction, but the increase in voltage frequency can improve ECE as in the case of methanol reforming.14

Fig. 4 Additive effect of water on [Ik]/[CO]

The continuous operation of FPR i

r

in the reforming of H 2 O-CO 2 in N2 for 5 h (RED at 1.19 to 1.26 k J / L ) gives c o 2 - ^ - c o + o ^0 the constant yields of H 2 and CO in o + o

2

l

A

.

x,

Mcm-'moiccuic'V) -O + H2

i

^

* H 2 O=^=OH + H

r

r

about 12 % and 10 %, respectively + (Fig. 7), as observed in the steam ' ~ 7 reforming of methane. O('D> + co 2 —-O( 3 P) + c o 2 CO + O( 3 P)

4. CONCLUSION ,

.,

-CO2

(2) (3)

(5)

7.4\ioIJ

(6)

1.1 X 10 '"

_ ,

...

CO^~C

+ O

In this process, target compositions of synthesis gases could be obtained CO + OH—-Products ÔH + H by selecting pertinent ratios of H 2 O H 2 + O ( i D ) and CO2 depending on purposes such ,,

(1)

,

r-

as H 2 utilization and preparation 01 a feedstock for methanol Synthesis. At

H 2 + OH

-H2O+H

(7)

(8) (9)

(10)

2.4x 10" 1.1 x 10"1 5.5X10 1 2

Fig. 5 Plausible mechanism for the reforming of CO2 and I12O

this moment, initial concentrations of H2O and CO? cannot be increased due to the occurrence of the backward reaction of H2 and O2. Rapid separation of the products is mandatory to increase the potential of this method. Energy conversion efficiency is not satisfactory at this stage, and increase in voltage frequency is required to improve the efficiency. References (1) 2nd Assessment Report of IPCC, 1997. (2) Kajishima, T.; Saito, T.; Nagaosa, R.; Kosugi, S. Energy, 1997, 22, 257. (3) Hawecker, J.; Lehn, J.-M.; Ziessel, R. Helv. Chim. Ada, 1986, 69, 1990.

124

(4)

(5)

(6)

(7)

(8)

(9)

(10)

(11)

(12)

(13)

(14)

Saito, M ; Fujitani, T.; Takeuchi, M.; Watanabe, T. Appl. Catal. A: General, 1996, 138,311. Bromberg, L.; Cohn, D. R.; Rabinovich, A.; Energy Fuels, 1998, 12, 11. Zhou, L. M.; Xue, B.; Kogelschatz, U.; Eliasson, B.; Energy Fuels, 1998, 12, 1191. Kabashima, H.; Einaga, H.; Futamura, S. Chem. Lett., 2001, 1314. Kabashima, H.; Einaga, H.; Futamura, S. IEEE Trans. Ind. Applicat., 2003, 39, 340.

Fig. 6 Effects of RED and water content on energy conversion efficiency

Kabashima, H.; Futamura, S. Chem. Lett., 2002, 1108. Futamura, S.; Kabashima, H.; Einaga, H. IEEE Trans. Ind. Applicat., submitted for publication. Futamura, S.; Zhang, A.; Yamamoto, T. IEEE Trans. Ind. Applicat., 2000,36, 1507. Einaga, E.; Ibusuki, T.; Futamura, S. IEEE Trans. Ind. Applicat., 2001, 37, Fig. 7 Continuous production of synthesis gas with 1476. nonthermal plasma Futamura, S.; Zhang, A.; Yamamoto, T.; IEEE Trans. Ind. Applicat., 1999, 35, 760. Futamura, S.; Kabashima, H. IEEE Trans. Ind. Applicat., submitted for publication.


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CO 2 Reduction by Blast Furnace Top Gas Recycling Combined with Waste Hydrocarbon Gasification Jae-Ou Choi* and Jang-Gyu Kim** Environment & Energy Research Center, Research Institute of Industrial Science and Technology(RIST) RO.Box 135 Pohang Kyungbuk, 790-600, Korea *iochoi(q),rist.re.kr, **[email protected] Carbon has been used for a reducing agent in the conventional blast furnace. Decreasing carbon usage by both using hydrogen source and improving the energy efficiency is the first target and CO2 squestration is another way to cut CO2 emission. Experimental evidences show that hydrogen substitutes the coke as a reducing agent by way of stack gas injection. Correlation shows that the coke consumption can be lowered proportionally to the effective(CO+H2) volume. With this correlation the coke consumption in a blast furnace can be estimated while hydrogen-rich gas is supplied from the gasification of the waste mixture. In the estimation, pure oxygen was used to increase effective(CO+H2) volume. The use of oxygen has other advantages such as increasing blast furnace efficiency and emitting nearly pure CO2 after combustion with oxygen. By recycling the blast furnace top gas through the gasifier, the carbon utilization efficiency is increased and the top gas is also recycled to the tuyere for the purpose of raceway temperature control and the optimum blast furnace operation. 1. INTRODUCTION While steel plays a key role in modern society, the steel industry is one of the major sources of global warming gases. This problem has been attacked conventionally by improving efficiency of energy and resource use and this conventional method is now known to have limited potential for repression of CO2 emission by 5-10%. [1] But in the course of time some radical concepts have been proposed which aim at drastic curtailment of CO2 emission through more efficient use of carbon as a reducing agent in the blast furnace or in the smelting reduction process by in-process recycling of the furnace top gas. Top gas recycling is a technology in which a part of the hot blast injected into the blast

126

furnace through tuyeres is replaced by the recycled furnace top gas. Several modes have been suggested to the top gas recycling technology and they can be classified into three categories; 1) Simple blast replacement: the top gas is recycled without addition of other gases in order to supply sufficient tuyere gas volume; 2) Oxygen enriched blast replacement: part of the blast gas is replaced by the recycled top gas maintaining total tuyere gas volume and oxygen is enriched up to the same partial pressure as in the normal blast gas; 3) Hot reducing gas (HRG) injection: similar to the oxygen enriched blast replacement except CO2 is removed from the recycled top gas. In case of HRG Austin et al. predicted that the production efficiency can be improved 25% and the fuel rate can be lowered 20%. [2] A modified HRG system has been proposed by the Arcelor Group and it is thought to be operated with the 100% recycling of the top gas and complete substitution of air by oxygen, and, therefore, no use of hot blast. [3] This process is characterized by the substantial reduction of the coke and coal, which may count CO2 emission reduction of 30% or more. Another characteristic of this process is the separation and sequestration of the CO2 that will effect the complete prevention of CO2 emission. Here is another suggestion aimed at the efficient use of carbon as reduction agent via top gas recycling and the use of hydrogen-rich waste as additional reducing agent via gasification, the top gas is recycled through the gasifier to the blast furnace. The purpose of this suggestion is to achieve substantial CO2 emission reduction without the use of a big amount of additional oxygen and CO2 separation because the CO2 separation and the oxygen production are deemed to be technically burdensome in most steel plants. 2. Top Gas Recycling Combined with Wastes Gasification (TRCG) waste 5.7ton/d coke 200kg/d

.

2.1. Gasifier desien and operation In order to design a proper gasifier for waste material many kinds of wastes were considered and a mixture of them was chosen as a standard material. Because most waste materials contain more

carbon

than

hydrogen,

the

gasifier

operation should be controlled so that the carbon should remain in the form of solid if the gasifier is to be operated as a hydrogen-rich gas producer. Fig. 1 shows the material balance on the basis of 5.7ton/d charge. The charge composition was Fig. 1 Material balance for the gasifier

assumed to be the mixture of the various wastes

127

such as municipal sludge, PE chips, automobile scrap shredder residue, waste plastic mix, tire chips, wood pieces. And additionally, iron making coke or coal was mixed into the charge as a supplementary fuel. The detailed composition of the waste mixture as weight fraction is as follows: C 0.629, H 0.083, N 0.0096, O 0.0786, Fe 0.695, Fe2O3 0.0214, CaO 0.0244, SiO2 0.0486, A12O3 0.0235, MgO 0.005, TiO2 0.0064, BaO 0.001, Na2O 0.0014, K2O 0.002, ZnO 0.0036, PbO 0.0004, SO3 0.004, Cl 0.0054, P2O5 0.0002, Cu 0.001, and Cr2O3+V2O5 0.0005. The thermodynamic calculation for conversion equilibrium of the material with the temperature shown in the fig. 1 shows that the product gas composition Fig. 2 Schematic diagram of experimental

as volume fraction showed CO 0.2, H2

gasifier

0.783, CH4 0.006, CO2 0.001, H2O 0.005, and HC1 0.004.

A gasifier that is composed of drying part and gasifying part was designed for the test. In the drying part the charge is dried by the part of heated gas from gasifying region and the gas product from this region can be recycled to the bottom of gasifier or treated separately. Dry gas is taken out from the gasification part of the furnace and mostly moisture free and, therefore, can be injected directly into the blast furnace as reducing gas. Fig. 2 shows a schematic diagram of experimental set-up of the gasifier. The largest inside diameter of the gasifier is 30cm and length of gasifying part is 85cm and that of drying part is 45cm. The mixtures of polypropylene, polystyrene, lump coke, coke breeze, and shredder light fraction with various ratios were charged continuously at the top of the gasifier. The temperature of the gasifier was kept around 1000°C at the hottest point of the gasifier and in the range of 200-500°C at the gas outlet. Pure oxygen is used to maintain the proper temperature in the gasifier auto-thermally. In all cases the result showed that almost all CO2 injected was transformed into other chemicals such as CO, methane, and various organic compounds. The metal components of the charge are thought to have catalytic function to promote the reaction.

128

The conversion of CO2 in the gasifier expected

is well

from

the

chemical equilibrium data with existence

the of

abundant elemental

carbon if the temperature is kept higher than 1100°C and if the carbon is added in the form of coal the reaction rate will be faster because coal has more functional groups that can Fig. 3 The dependence of coke decease on the composition of the gas injected to blast furnace stack

be activated more easily. In the same way, the

hydrocarbons in the gasifier react with CO2 much more easily due to their unstable chemical bonds compared with coke or coal. 2.2. Experimental evidence on stack gas injection to reduce the coke rate In the process of stack gas injection some hydrocarbon fuel is reformed to a mixture of H2 and CO and this gaseous reducer is injected into the lower stack of the blast furnace. This process was investigated by several authors and steel plants and is thought to be a proper candidate for lowering CO2 in the ironmaking process. The results of these experimental investigations have been reviewed in the state-of-art paper of Rhee [4] and are summarized in fig. 3. The data are from the experimental and real blast furnace tests of the different institutions like CRM, NKK, Nippon Steel and U. S. Bureau of Mines already done during 1960's and 70's. The effectiveness index of the stack gas injection in lowering coke consumption in the blast furnace can be expressed in the sense of "effective (CO+ H2) volume per ton hot metal" which is calculated from the content of partial pressures of CO, H2, CO2, and H2O in the injected gas stream. Fig. 3 shows that the coke consumption can be lowered proportionally to the effective (CO+ H2) volume. 2.3. Suppression of CO2 by hydrogen-rich gas injection to the blast furnace with the use of waste gasification Gasification of hydrocarbon is a proper method for production of hydrogen-rich reducing gas and it provides some technical advantages; some substitute fuel such as pulverized coal,

129

waste plastics, and some other hydrocarbon

materials

are

transformed into a gaseous form and it can be injected into the blast furnace more easily than when

those

materials

are

directly injected through the tuyeres

in solid form.

The

preparation of coal and plastic material in the form of pellet or Fig. 4 Material balance for conventional blast furnace

powder is not necessary and extreme high temperature for

rapid conversion is not needed because the large reaction volume in the gasifier gives sufficient reaction time. Another technical advantage is that some part of the top gas, which contains mainly CO2, H2O, CO and H2 can be fed into the gasifier and converted to CO and H2 rich gas. This recycling of hot reducing gas brings the use of carbon and hydrogen to the higher grade. Fig. 5 is an example of the use of gasifier and pure oxygen for the blast furnace and the conventional blast furnace operation data is shown in fig.4 for comparison, hi this case a 400ton-per-ton-hot-metal scale gasifier is used to maintain total energy input into the blast furnace. By limiting oxygen supply and adjusting top gas recycle to the gasifier, large portion of carbon in the gasifier charge can be harvested in the solid form so as to minimize carbon input to the blast furnace. In this case only about 25% of carbon went into the blast furnace and the effective (CO+H2) volume of the gas product is about 240Nm3 per ton hot metal and CO2 emission is lowered by about 30%. The emission from the top of the blast furnace is combusted with pure oxygen to produce pure CO2 and this final emission can be further treated to get a fixed form of CO2. In this calculation total thermal energy used in the blast furnace was maintained the same as in the case of fig. 4. But because there is still no confirmation how the use of pure oxygen in blast furnace can increase the efficiency of the energy consumption, the constant energy consumption for both cases was assumed only for calculation and there is possibility for the energy consumption to be greatly reduced so that a smaller gasifier and less oxygen consumption will be caused. The estimation also indicates smaller gas volume to be handled that the gas handling facilities can be made compact and that the blast furnace has much more flexibility for recycle gas adjustment as a result of pure oxygen usage. In case of the use of the conventional hot blast the nitrogen dilutes effective (CO+H2), and the only way available is to lower coke rate enough to bring about substantial decrease of CO2

130

is to increase gas volume and

a

large

gasifier

is

inevitable. 3. DISCUSSION The main target of TRCG process is to lower coke rate and

in

special

case

to

operate the blast furnace

Fig. 5 Material balance for top gas recycling with gasifier and pure oxygen use instead of hot blast

nitrogen-free

so that the

high

of

purity

CO2 is

harvested as an end product and the facilities are made compact.

The advantage of the utilization of the gasifier is thought to be the coke rate decrease without any other efforts such as CO2 separation and the use of pure oxygen but oxygen will be inevitable if an optimal furnace operation is wanted. But how it will serve the CO2 breakthrough end is a matter of confirmation. For the time being there are two research topics for the oxygen blast furnace. One is how oxygen and coke consumption can be minimized. If the blast furnace condition is optimized with proper top gas circulation and oxygen injection the heat energy generated at the bottom of the blast furnace can be optimally used for the reduction of iron ore and the minimum oxygen and coke consumption can be brought about. The other is how the oxygen can be generated without generating CO2. Nuclear energy and the renewable energy like wind and solar powers are possible ways to solve the problem, but there are many obstacles to be overcome for the extensive utilization of these powers. Much effort, therefore, will be needed in order to accomplish the final goal, the realization of the CO2 reduction breakthrough technology in the field of ironmaking. REFERENCES 1. J. Birat, "The challenge of global wanning to the steel industry, a European viewpoint," IRSID Innovation report (2002). 2. P. R. Austin, H. Nogami, and J. Yagi, ISIJ International, 38 (1998), 239. 3. Arcelor presentation, "Breakthrough Technologies for Abatement of CO2 Emissions," IISI Executive Board Meeting-Rome (2002). 4. P. C. Rhee, Bulletin of Korean Metallurgical Society, 16-2 (1978), 197.


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Microwave-assisted Reactions of Oxiranes with Carbon Dioxide in Ionic Liquids M. Ntichter*, B. Ondruschka, J. Hoffmann, A. Tied Institute of Technical Chemistry and Environmental Chemistry, Friedrich-Schiller-University of Jena, Lessingstr. 12, D-07743 Jena, Germany; e-mail: [email protected] The microwave-assisted reaction of carbon dioxide with different oxiranes in ionic liquids is described. Some work-up possibilities are discussed also from the viewpoint of scaling-up. 1. INTRODUCTION With respect to sustainability and structure-activity-relationships it is necessary to check the effect of substances, reactions and processes on the environment [1], The fixation of carbon dioxide in chemical reactions is an industrial and an environmental problem which challenges many research groups around the world in the last years. Some reactions with using different catalysts have been published, but only the reaction of oxiranes with carbon dioxide is realized as an industrial process [2]. Ethylene carbonate (R = R1 = H) and propylene carbonate (R = CH3, R1 = H) are produced according to figure 1 in great amounts (more than 10 kt/y). These substances are used as solvents in many fields and as starting material for polycarbonates and other plastics. Many catalytic systems are described for the reaction of oxiranes with carbon dioxide, and lately this reaction is also carried out in various ionic liquids [3]. Ionic liquids are novel, socalled "green" solvents, which are entirely composed of ions. These compounds are per definition salts with a melting point below 100 °C [4]. In the fixation of carbon dioxide with oxiranes, the ionic liquids act both as catalyst and as solvent for the starting material and products (Fig. 1). Fig. 1: General reaction scheme of synthesis of cyclic carbonates

Various imidazolium bromides were used in the experiments (Fig. 2) as solvent and catalyst. The variation of the length of the second carbon chain alters the physical characteristics [5] of the used substances (Tab. 1). Our aims in the following discussed experiments were: (I) testing of the using of microwave in this reaction, (II) variation of the carbon dioxide source, (III) bettering of work-up of reaction mixtures, and (IV) scale-up.

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Fig. 2: Quarternisation of methylimidazole

Tab. 1: Imidazolium bromides used abbreviation regular name 1 -n-propyl-3-methylimidazolium bromide [C3MIM]Br 1 -n-butyl-3-methylimidazolium bromide [C4MIM]Br 1 -n-hexyl-3-methylimidazolium bromide [C6MIM]Br 1 -n-octyl-3-methylimidazolium bromide [C8MIM]Br

R= C3H7 C4H9 C 6 H 13

2. EXPERIMENTAL The experiments were carried out in a microwave system ETHOS MR (MLS/Milestone) with two different reactor systems (for details see: www.milestone.com). Small amounts of reaction mixtures (up to 50 mmol educt) were irradiated in the rotor HPR 1000/10 in single reaction or parallel reactions, whereas greater quantities (up 1 mol educt) react in the microwave high pressure reactor uCLAVE. The reactor systems and their limiting parameters are shown in picture 1 and 2. Picture 1: Microwave system ETHOS MR (2.45 GHz) with rotor HPR 1000/10 for 10 parallel reactions (50 ml, 240 °C, 50 bar)

Picture 2: Microwave system ETHOS MR with high pressure reactor (components of reactor system, 500 ml, 240 °C, 60 bar)

Solid carbon dioxide was used with the rotor system HPR 1000/10, whereas gaseous carbon dioxide was dosed from pressure bottles when the reactions were carried out in the microwave autoclave. Carbon dioxide was used in excess (1.2 to 5 fold) in all reactions. The

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pressure in the (J.CLAVE reactor was pressurised to 10 bar and adjusted several times depending on consumption. The conversion is calculated from both the amount of oxiranes consumed and the increase in mass of the reaction mixture at the end of the reaction. 3. RESULTS AND DISCUSSION 3.1. First experiments in microwave field A series of oxiranes was converted in a first assay under reaction conditions which were compared to literature values [3]. The reaction mixtures were worked up as proposed in the literature. Thus, the possibilities of employing microwaves as energy source in this type of reaction were tested, and the feasibility of the reaction with the available equipment was demonstrated (Tab.2). Tab. 2: Reactions of carbon dioxide with oxiranes GmbH Leutkirch, Germany; extraction with MTBE) reactor / ionic liquid oxirane time batch size [min] HPR 1000/10, 30 [C8MIM]Br phenyl0.05 mol phenylHPR 1000/10, 60 0.05 mol phenyl30 uCLAVE, 0.35 mol HPR 1000/10, 30 butyl0.05 mol HPR 1000/10, 30 hydroxy methyl0.05 mol HPR 1000/10, 60 [C4MIM]Br phenyl0.05 mol HPR 1000/10, 60 methyl0.05 mol methyl60 uCLAVE, 0.35 mol

in microwave field (ETHOS MR, MLS temperature [°C] 120

conversion [%] 91

yield* [%] 62

120

>99

77

140

94

38

120

82

63

120

>99

**

140

>99

77

140

95

85

140

91

67

' Isolated product; ** Only oligomeric or polymeric products, no work-up possible by extraction with MTBE.

As expected, the microwave energy is easily absorbed by the reaction mixture [6] and the required reaction temperature was quickly reached. About quantitative conversion of the oxiranes to the products was achieved within 30 min at temperatures of around 120 °C. Only a slight improvement in conversion was observed when the reaction time was extended to 60 min. At this point, some work-up problems were encountered: Although high yields were expected (increase in mass of the reaction mixtures and near quantitative conversion of oxiranes with good selectivities (as determined by GC)), isolating the cyclic carbonates and recycling of the ionic liquid proved to be very difficult. 3.2. Work-up of reaction mixtures Work-up by liquid-liquid extraction with MTBE at room temperature was the first option investigated. This method is relatively simple to realise for experiments in mmol-scale, but it

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is already problematic for reactions carried out in 0.05 mol scale for three reasons, i) The necessary volume of MTBE is high in comparison with the extracted material, ii) the extraction time is very long due to lengthy phase separation of the components and iii) the handling is comparatively difficult due to the high viscosity of ionic liquid. The liquid-liquid extraction (second option) with MTBE at higher temperatures (45 - 55 °C) using a perforator allows for the isolation of the products as well. This process demands long times, a relatively great amount of solvent and a perforator whose dimension matches the respective reaction scale. The vacuum distillation of the whole reaction mixture was tested as a third option [7]. The temperature used in the reaction between carbon dioxide and oxirane was also chosen as the maximum distillation temperature in this process. The use of a vacuum between 0.01 und 0.001 mbar was necessary, for a distillation rate of 1 drop per approx. 5 sec. Thus, the distillation of a 0.5 mol reaction mixture requires under lab conditions around 5 hours. The thermal stress on the reaction mixture is high, so that some degradation of the ionic liquids is detectable, and the products partly oligomerise to products of higher molecular weight, thus decreasing the yield. The extraction of the cyclic carbonates after prior dilution with the four to six fold volume of water was tested as fourth work-up option, analogous to [8]. The dilution step is necessary because only with a degradation of the obviously existent complex between ionic liquid and reactions products a clean and fast extraction is possible. The resulting homogenous mixture was extracted with different organic solvents. The best yields were achieved by using dichloromethane. The yields of the analytical extraction experiments are summarized in Tab. 3. Tab. 3: Extraction of a mixture from the product of the reaction of carbon dioxide with methyloxirane in [C4MIM]Br and water (1:4 v/v) with different solvents(shaking machine Vibramax VXR, IKA, Germany, 1200 rpm, 10 min, 5 ml mixture + 5 ml solvent) relative concentration solvent of propylene carbonate [%] 40.1 toluene 65.2 ethylacetate 82.5 trichlormethane 100 dichloromethane 28.0 MTBE 29.5 diethylether During the extraction, the ionic liquid remains nearly quantitatively in the aqueous phase, whereas the reaction product and the by-products are extracted from the water-ionic liquid mixture (which is now formally a salt solution as opposed to an ionic liquid). An extensive complete extraction of the product is feasible at room temperature and with a comparatively small amount of solvent (200 ml). A positive side-effect of this process is the purification of the ionic liquid from higher molecular by-products of the reaction, which is not possible with the other processes described above. The ionic liquids are quantitatively recyclable by means of concentrating the remaining aqueous phase on rotary evaporator, followed by drying in vacuum. The cleaned and dried ionic liquid is reusable for the described reaction. The cyclic carbonates were further purified by means of vacuum distillation, whereby small amounts of extracted ionic liquids are retained in the distillation residue.

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This work-up process with dichloromethane as extracting solvent was also used in the following experiments. All the above-discussed work-up processes are compared in Tab. 4 for the reaction of methyloxirane with carbon dioxide in [C4MIM]Br. Tab.4: Comparison of work-up processes from the mixture of reaction of carbon dioxide with methyloxirane in presence of [C4MIM]Br (ETHOS MR, uCLAVE, 0.35 mol methyloxirane, 140 °C, 60 min) process solvent time process temperatures yield amount [ml] [°C] [hi [%] extraction with MTBE, 1000 14 room temperature, 30-35 subsequently vacuum 140 distillation 500 hot extraction with 18 50-60, 40 MTBE, subsequently 140 vacuum distillation 8 vacuum distillation 140 45-50 extraction with 200 5 room temperature, 90 dichloromethane after 140 dilution with water, subsequently vacuum distillation 3.3. Experiments with modified work-up and under various reaction conditions The below described experiments were worked up by means of dilution of the reaction mixtures with water, followed by extraction with dichloromethane. The parallel reaction of different oxiranes in a rotor system HPR 1000/10 under otherwise identical conditions [7] shows the different reactivities of the used educts and products (table 5). Table 5: Comparison of reactions of different oxiranes with carbon dioxide in presence of [C6MIM]Br (ETHOS MR, parallel reactions in HPR1000/10, 0.05 mol oxirane, 140 °C, 60 min, extraction with dichloromethane after dilution with water) conversion [%] purity l) [%] yield 1} [%] oxirane 86 methyl>95 97 1 88 2 90 98 chloromethyli) 47 >99 hydroxymethyl3 83 >95 >99 4 butyl79 >95 96 phenyl5 67 6 1 -carbethoxy-2-phenyl- >95 68 >99 cyclohexene oxide 7 41 v ' Isolated yield; ' GC of isolated product; ' Only small amounts of products detectable by GC; Two unknown isomers (m/z=294) are detectable. 5) Unknown substance (59 %) is detectable.

4)

The oxiranes investigated are nearly quantitatively converted under these reactions conditions. Products are isolated in good to excellent yields in the absence of additional functional groups in the starting material molecule (entry 1, 4, 5 in table 5). hi the presence of such a functional group, consecutive reactions are observed, which complicate further work-up after the extraction from water with dichloromethane.

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The comparison of the reaction of methyloxirane with carbon dioxide in different ionic liquids showed that [C3MIM]Br gave the lowest yield under the selected conditions (table 6). Tab. 6: Variation of ionic liquid in reaction of carbon dioxide with methyloxirane (ETHOS MR, uCLAVE, 0.517 mol methyloxirane, 140 °C, 60 min, extraction with dichloromethane after dilution with water) conversion [%] yield * [%] ionic liquid [C3MIM]Br 57 47 [C4MIM]Br 79 66 [C6MIM]Br 93 75 [C8MIM]Br 82 66 "isolated yield after vacuum distillation

A conversion of around 80 % is found for all ionic liquids with a chain length greater than four carbon atoms. Longer chains increase the conversion only marginally. The reason for this effect may be the different solvent capacities of ionic liquids for carbon dioxide and methyloxirane. For example, in [CsMIMJBr, only a part of the components is dissolved, whereas in [C4MIM]Br (and higher homologues), the ionic liquid, starting material and products form a homogeneous phase. Further experiments will show if the maximum conversion shown in table 6 (which is reached at an alkyl chain length of six carbon atoms) is real, or if further improvement of the work-up process will lead to a constant increase in conversion and yield. 4. CONCLUSION Microwave energy can be used advantageously as energy source in the addition of carbon dioxide to oxiranes. This method is usable for non-functionalised and functionalised oxiranes. The main advantages in comparison to the conventional reactions are short heating times to reaction temperatures, high energy efficiency and very well regulation of reaction conditions. For the latter, more work for a detailed description of reaction processing is necessary. Cyclic carbonates react very easily with amines and alcohols. Thus, simple derivatives (ethylene carbonate, propylene carbonate) are good substitutes for phosgene in the synthesis of urea derivates and carbonic acid esters. Higher or functionalized derivatives are valuable intermediates in heterocyclic chemistry and for speciality polymers. REFERENCES [1] [2] [3] [4] [5] [6] [7] [8]

P. T. Anastas, M. M. Kirchhoff, Ace. Chem. Res. 2002, 35, 686-694 J. H. Clement, Ind. Eng. Chem. Res. 2003, 42, 663-674 V. Calo, A. Nacci, A. Monopoli, A. Fanizzi, Org. Letters 2002, 4, 2561-2563 review see: P. Wasserscheid, W. Keim, Angew. Chem. Int. Ed., 2000, 39, 3772-3789 P. Bonhote, A.-P. Dias, N. Papgeorgiou, K. Kalyanasundaram, M. Gratzel, Inorg. Chem. 1996,55,1168-1178 J. Hoffmann, M. Niichter, B. Ondruschka, P. Wasserscheid, Green Chemistry, 2003, 5, 296-299 H. Kawanami, A. Sasaki, K. Matsui, Y. Ikushima, Chem. Comm. 2003, 896-897 F. Shi, Y. Deng, T. SiMa, J. Peng, Y. Gu, B. Qiao, Angew. Chem. Int. Ed., 2003, 42, 3257-3260


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CCh Reforming of n-Heptane on a Ni/AkCb Catalyst K. Johanna Puolakka and A. Outi I. Krause Department of Chemical Technology, Helsinki University of Technology, P.O. Box 6100, FIN-02015 HUT, Finland; e-mail [email protected] CO2 reforming of »-heptane was studied on a 15 wt-% Ni/AkCb catalyst. In addition, thermal cracking of «-heptane was examined. Experimental results were compared with the thermodynamics of «-heptane CO2 reforming and of possible side reactions. The composition of the product at 900 °C under atmospheric pressure was close to that calculated from the thermodynamics.

1. INTRODUCTION Carbon dioxide is widely considered as a greenhouse gas causing global warming. Thus, carbon dioxide storage and catalytic activation for chemical reactions are of great interest. Carbon sequestration in saline aquifers, coalmines, oil and gas wells, and the ocean could be done during the changeover from fossil energy to renewable energy [1]. A potential reaction of carbon dioxide is CO2 reforming, in which the important raw material for chemical industry, synthesis gas (i.e. hydrogen and carbon monoxide) is formed. The CO2 reforming of methane has been already extensively studied, because also methane is regarded as a greenhouse gas [2]. Other potential feedstocks for the CO2 reforming are light, sulphur-free GTL (FischerTropsch) fractions, which are not suitable for gasoline due to low octane numbers. There are only few published studies on CO2 reforming of higher hydrocarbons thus requiring basic experiments with commercial catalyst to study if the reaction is feasible. In this work the reaction was studied with «-heptane (Equation 1) as the model compound for gasoline. C7H16 + 7 CO2 - 8 H2 + 14 CO

AH°298K = +1395kJ/mol

(1)

The typical side reactions of the CO2 reforming are the Boudouard reaction (Equation 2) and the water-gas shift reaction (Equation 3): 2 CO - C + CO2 H2O + CO - H2 + CO2

AH°298K = -172 kJ/mol AH°298K = -41 kJ/mol

(2) (3)

Major challenges in CO2 reforming are the highly endothermic reaction requiring high temperature and the catalyst deactivation by carbon deposition [3].

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2. EXPERIMENTAL Experiments were carried out in a fixed bed flow reactor with an inner diameter of 6 mm. The reactor was heated by a three-zone tube furnace. Reaction products were analyzed with two gas chromatographs. Columns were DB-1 (J&W Scientific) for hydrocarbons and a packed column with activated carbon with 2% squalane for hydrogen, argon, carbon monoxide, carbon dioxide and small hydrocarbons. First, thermal cracking of w-heptane was studied with silicon carbide in a quartz glass reactor between 550 and 800 °C. Catalytic experiments were carried out at furnace temperatures of 700 and 900 °C under atmospheric pressure with a commercial 15 wt-% Ni/AkCb catalyst. The amount of the catalyst was 0.1 g and the particle size was 0.2-0.3 mm. The total feed rate in the experiments was 100 cm« min , and the composition 63 mol-% CO2 and 3 mol-% «-heptane balanced with argon. Threefold stoichiometric excess of carbon dioxide was used to reduce coke formation. The catalyst was reduced at 900 °C with a mixture of H2 and Ar (50% H2) for one hour. In addition, the performance of the catalyst was studied at a higher pressure of 4 bar using a stainless steel reactor (AISI 316), which limited the furnace temperature to 700 °C. The total feed rate was higher, 200 cm« min , and the catalyst reduction was done at 700 °C. Thermodynamic calculations of the CO2 reforming of w-heptane were performed with the HSC Chemistry 3.02 program.

3. RESULTS AND DISCUSSION 3.1. Thermodynamic calculations The composition of the product gas at the thermodynamic equilibrium was calculated as a function of temperature including the side reaction in addition to the »-heptane reforming. E.g. at 700°C the following net stoichiometric equation resulted: C7H16 + 5.98 CO2 ^ 11.72 CO + 5.16 H2 + 0.27 H2O + 1.29 CH4(4) The equilibrium conversions of rc-heptane and CO2 were 100% and 85% respectively. Compared to the basic CO2 reforming (Equation 1) the molar ratio of H2 to CO decreased from 0.57 to 0.44. However, the calculated enthalpy for reaction 4, +1084 kJ/mol, is somewhat smaller than the value for the mere CO2 reforming due to the lower enthalpies of the side reactions. At high temperatures the thermodynamic equilibria of side reactions are on the side of carbon dioxide consumption and carbon monoxide formation. For the reverse water-gas shift and the reverse Boudouard reactions AG < 0 kJ/mol, when temperatures are higher than 830 and 703 °C respectively. Thus, these temperatures define a suitable range for the reaction. Reaction temperatures higher than 703 °C are suitable in coke removal due to the reverse Boudouard reaction. In addition, the higher the temperature the larger the amount of hydrogen which is converted to water in the reverse water-gas shift reaction. This is not desirable. However, as the temperature dependency of AG of the water-gas shift reaction is relatively small, the upper limit of 830 °C is only suggestive.

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3.2. Thermal experiments The conversion of «-heptane without catalyst was studied at the temperatures relevant for the reforming reactions. In these experiments «-heptane was cracked to smaller hydrocarbons, ethene being the predominant product. This was expected because rc-heptane is a good feedstock for steam cracking. No oxygen containing products were detected indicating reforming reaction not taking place. The conversion of «-heptane increased with temperature being 0% at 550°C and close to 100% at 800°C. The molar ratio of methane to ethene was 0.43 at 800°C. This ratio increased also with temperature, being 0.32 at 700°C. 3.2. Catalytic experiments In the catalytic experiments at 700 °C the main products were carbon monoxide (32 cm« min ) and hydrogen. In addition, smaller hydrocarbons (less than 1 cm« min each, methane to ethene ratio 1.09) were formed indicating that only some thermal cracking occurred besides the reforming reactions. The coking of the catalyst was also significant increasing the pressure in the reactor. In addition, the catalyst bed temperature increased indicating decrease of the endothermic reforming reaction. Initially the total conversion of «-heptane was 97%, of which 8% was cracked to lighter hydrocarbons. The initial conversion of carbon dioxide was 28% compared to the stoichiometric 33% (due to the threefold excess of carbon dioxide). At 900 °C the conversion of «-heptane was 100% and that of carbon dioxide 42%. The carbon dioxide conversion being higher than the stoichiometric one indicates that carbon dioxide was consumed in other reactions. Furthermore, the H2 to CO molar ratio of 0.32 was less than the stoichiometric ratio of 0.57 according to Equation 1. These facts point out to the occurrence of the reverse water-gas shift reaction and the reverse Boudouard reaction. The product composition agreed very well with thermodynamics, the calculated ratio of H2 to CO being 0.28 at 900 °C. The catalyst bed was coking only slightly and the pressure increase was much slower compared to the increase at 700 °C. In addition, no smaller hydrocarbons were formed. This all means that higher temperatures are more favourable for the reforming reactions. However, if high H2 to CO ratio is required, higher temperatures are not suitable due to the water-gas shift reaction, when larger than stoichiometric amount of CO2 is used.

Fig. 1. CO2 conversion and the outgoing CO flow at 4 bar and 700 °C.

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According to the thermodynamics coke formation is reduced at higher pressures. In addition, the pressure remains constant, when controlled. In larger scale industrial application the pressure will also be higher than atmospheric. In the experiments performed at 4 bar and 700°C the initial conversions for «-heptane and carbon dioxide were 83% and 13%, respectively. 31% of «-heptane was cracked and the initial molar ratio of methane to ethene was 0.61. The activity of the catalyst decreased clearly during the experiment of 6.5 hours. The conversion of carbon dioxide and the outgoing carbon monoxide flow are shown as a function of time in Figure 1. It seems that increasing the pressure worsened the situation, as was to be expected on the basis of the thermodynamic calculations considering conversion. This means that at higher pressures higher temperatures are needed. 4. CONCLUSIONS The CO2 reforming of «-heptane is a promising way of utilising carbon dioxide. High temperatures are required for the optimal production of synthesis gas. However, the overall reaction remains very endothermic even though the side reactions lower the reaction enthalpy slightly. Combining the CO2 reforming with partial oxidation could provide a solution as it decreases the need for outside heating. This could have a positive effect on the catalyst deactivation, too. ACKNOWLEDGMENTS The financial support for this work from the Technology Development Centre of Finland (TEKES) and Fortum Oil and Gas Oy is gratefully acknowledged.

REFERENCES 1. Hileman, B., How to Reduce Greenhouse Gases, Chem. Eng. News 80 (21) (2002) 37-41. 2. Bradford, M.C.J., Vannice, M.A., CO2 Reforming of CH4, Catal. Rev.-Sci. Eng. 41(1) (1999) 1-42. 3. Rostrup-Nielsen, J.R., Bak Hansen, J.-H., CO2-Reforming of Methane over Transition Metals, J. Catal. 144 (1993) 38-49.


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Hexaaluminate Catalysts of the Novel Process of Syngas Production through Catalytic Oxidation and Steam-CO2 Reforming of Methane 0 . Yokota, T. Tanaka, Z. Hou and T. Yashima Chemical Research Group, Research Institute of Innovative Technology for the Earth, Kizugawadai 9-2, Kizu-cho, Soraku-gun, Kyoto, 619-0292, JAPAN. FAX: +81-774-75-2306. E-mail: [email protected] Novel process of synthesis gas (syngas) production through complete oxidation and steamCO2 reforming of methane using both oxidation and reforming catalysts is proposed. Hexaaluminate catalysts were applied for these catalysts. La-Mn hexaaluminate was found to have high catalytic activity in low temperature region as the combustion catalyst. On the other hand, Ba-Ni hexaaluminate was found to have good resistance against carbon deposition on its surface as the reforming catalyst. The set of these oxidation and reforming catalysts is excellent candidate for our novel process. 1. INTRODUCTION Synthesis gas is an essential industrial feedstock in several catalytic processes, such as the methanol synthesis, the Fischer-Tropsch synthesis, and the ammonia synthesis. The most common favored process for conversion of natural gas to syngas is the steam reforming process. The conventional steam reforming process involves highly endothermic reaction, therefore it requires a large amount of external heating by use of natural gas resources and additionally it leads exhaust of CO2 to the environment. Autothermal reforming process can solve these problems. However, this process usually uses burner for internal combustion which exposes materials to higher temperature than 1700 °C and the reforming catalyst under the burner gets soot coating. We propose the new autothermal reforming process using oxidation and reforming catalysts as shown in Fig. 1. Some cation-substituted hexaaluminate compounds have been reported as catalysts for high temperature combustion under flow of excess O2 in comparison with CH4 [1-2]. However, to the authors knowledge, only Utaka et al have dealt with hexaaluminate (BaNiAln0i9) for combustion catalyst under flow of excess CH4 in comparison with O2 [3] and few works have been carried out for reforming catalysts [4-5]. In this work, Mnsubstituted hexaaluminate was used as a methane oxidation catalyst and Ni-substituted hexaaluminate was chosen as a steam-CO2 reforming catalyst. The catalytic activity of both of combustion and reforming catalysts in proposed new autothermal reforming process was investigated.

Fig. 1. New autothermal reforming process using oxidation and reforming catalysts

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2. EXPERIMENTAL 2.1. Oxidation catalyst Mn-substituted hexaaluminates (XMnAl u 0i9 : X = Ba, La, Sr and Sro.gLao.2 ) were prepared by co-precipitation route using carbonates [6]. The samples were calcined at 500 °C for 5 h after drying at 110°C overnight. The obtained samples were compressed to pellet form and calcined at 1200 °C for 10 h. The specific surface areas of calcined samples were measured by BET method using N2 adsorption. The catalytic combustion activity was tested in a fixed-bed apparatus using a quartz reactor at atmospheric pressure. A gaseous mixture of CH4/O2/H2O = 200/100/400 ml/min was supplied at a space velocity (S.V.) of 38,000 ml/g-cat/h. The reactor (i.d. = 10 mm) was loaded with 1.1 g (10-30 mesh) of catalyst. The initial temperature of combustion was decided by measuring the temperature of sample surface which is suddenly raised up compared with the setting temperature. The product gases were analyzed by on-line gas chromatograph with thermo conductivity detector. 2.2. Reforming catalyst Ba-Ni hexaaluminates (BaNiAln0i9;Ni content amounts to 7 wt%.) were prepared by solid state reaction route. The powder mixture of A12O3, BaCO3, and NiO were ground and compressed to pellet form, then calcined at 1500 °C for 10 h. Each samples were crushed to particles with the size of 300-850 um. Methane reforming reaction with H2O and CO2 was carried out at 700 °C under atmospheric pressure (CH4/H2O/CO2 = 150/100/50 ml/min S.V. = 18,000 ml/g-cat/h, process time = 70 h). Before the reaction occurred, catalyst was reduced by gas mixture of H2 and N2 (H2:N2 = 100:100 ml/min) at 900 °C for 1 h. Conventional Ni/Al2O3 catalysts (A: Ni 18 wt% / A12O3, B: Ni 32 wt%/Al2O3, C: Ni 6wt%/Al2O3-SiO2) were used for comparison. These catalysts were reduced at 700 °C for 1 h by gas mixture of H2 and N2 (H2:N2 = 100:100 ml/min) before the reaction test. 3. RESULTS AND DISCUSSION 3.1. Oxidation catalyst To our knowledge, it is a first time to perform catalytic activity test for combustion of gas mixture with excess of CH4 (CH4/O2= 2/1) above Mn-substituted hexaaluminates. Table 1 shows the initial temperature of combustions for XMnAln0i9 (X = Ba, La, Sr, and Sro.8Lao.2). LaMnAlnOi9 is the most active catalyst for oxidation which starts combustion under 450°C. Sro.8Lao.2AliiOi9 prepared by Machida et al shows the highest catalytic activity at the CH4 oxidation (CH4/air = 1 vol%/99 vol%) among Sr^Lai-jMnAlnOigp]. However in this study, LaMnAluOi9 has the highest activity similar to the results reported by Jang et al [8] and Eguchi et al [9] at the CH4 oxidation (CH4/air = 1 vol%/99 vol%), though the surface area of Table 1. Initial temperature of combustions for XMnAluOig Surface area Initial temperature Produced gases*' 2 Catalysts (m /g) (°C) CO2/CO (vol%)

rice*' (%)

13 510 24/2 BaMnAlnO,9 98 9 430 24/1 LaMnAlnO^ 100 8 500 24/1 100 SrMnAluO19 7 450 Sr0 8La02MnAlnO19 25/0 100 *) The data 20-40 min after combustion and the mean of 2 times measurements

T

1CH4*) (%)

26 25 25 25

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sample shows different values. Every catalyst shows 100 % of O2 conversion except for BaMnAlliiOig and 25-26 % of CH4 conversion 20-40 min after combustion. Since the ratio of reacted CH4 and O2 is 1/2 with the main product of CO2 and the amount of the forming CO is negligible, it is obvious that complete oxidation reaction (CH4 + 2O2 = CO2+ 2H2O) occurs on these hexaaluminate catalysts under flow of CH4/O2/H2O = 2/1/4. 3.2. Reforming catalyst Methane reforming reaction with H2O and CO2 was carried out above BaNiAlnOi9 catalyst. To enhance the catalytic activity of Ni substituted-hexaaluminates some researchers use pre-reduction before reforming reaction [4-5]. However there are no detailed data to study the effect of reduction temperature on the CH4 conversion. Fig. 3 shows the reduction temperature dependence of CH4 conversion. In order to obtain reforming activity, it was confirmed that the BaNiAlnOi9 requires pre-reduction at higher temperature than 700 °C. After pre- Fig. 2. Reduction temperature dependence of CH4 reduction a number of Ni conversion. particles were formed on the surface of the catalyst. Particles sizes were uniform between 20 and 30 nm according to SEM observation. Figure 3 shows the comparison of BaNiAluOig and conventional N1/AI2O3 catalysts (A, B, C). The BaNiAlnOi9 was reduced in H2 flow at 900 °C for 1 h before reforming. As shown in Fig. 4, the CH4 conversion at the beginning of reaction in the presence of BaNiAlnOi9 is not higher than those in presence of other catalysts. However BaNiAlnOig keeps this conversion rate for 70 h and no carbon deposition was detected by TG-DTA measurements. On the other hand, catalyst A and B obtained a large amount of carbon deposition that greatly hindered gas flow within 10 h of reaction. The efficiency of methane conversion over catalyst C decreases with the time and becomes lower than that of BaNiAlnOig after 70 hours. After reaction 10 wt% carbon deposition was detected on the surface that is supposed to be a reason for the decreased activity of catalyst C. 4. CONCLUSION We proposed the process of syngas production through complete oxidation and steam-CO2 reforming of methane using oxidation and reforming catalysts. La-Mn hexaaluminate is excellent candidate for the oxidation catalyst because it shows a high catalytic activity from low temperature region and makes the complete combustion reaction possible. Ba-Ni hexaaluminate is good for the reforming catalyst since it keeps the catalytic activity for longer time than other catalysts owing to no any carbon deposition on its surface.

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Fig. 3. CH4 conversion during 70h reforming test ACKNOWLEDGEMENTS The research was supported by New Energy and Industrial Technology Development Organization (NEDO, Japan). REFERENCES 1. H. Arai, K. Eguchi, and M. Machida, MRS Int'l. Mtg. on Adv. Mats., Mater. Res. Soc, 2 (1989)243 2. M. Machida, K. Eguchi, and H. Arai, J. Catal., 120 (1989) 377. 3. T. Utaka, S. A. Al-Drees, J. Ueda, Y. Iwasa, T. Takeguchi, R. Kikuchi, and K. Eguchi, Appl. Catal. A: General., 247 (2003) 125. 4. M. Machida, T. Teshima, K. Eguchi, and H. Arai, Chem. Lett., 2 (1991) 231. 5. Z. Xu, M. Zhen, Y. Bi, and K. Zhen, Catal. Lett., 64 (2000) 157 6. G. Groppi, C. Cristiani, P. Forzatti, P. L. Villa, Appl. Catal. A: General, 104 (1993) 101.K. 7. M. Machida, K. Eguchi, and H. Arai, J. Catal. 123, (1990), 477. 8. B. W. L. Jang, R. M. Nelson, J. J. Spivey, M. Ocal, R. Oukaci, G. Marcelin, 47 (1999) 103 9. Eguchi, H. Inoue, K. Sekikawa, and H. Arai, Stud. Surf. Sci. Catal. 101 (1996) 417


145

Redox Behavior of Cu-ferrite for CO2 Decomposition Hyun Chang Shin1, Seung Chul Choi2, Chul Sung Kim3, Oh-Shim Joo1 and Kwang-Deog Jung1* 1

Eco-nano Research Center, Korea Institute Science and Technology, Cheongryang P. 0 . Box 131, Seoul, Korea 130-650, FAX: +82-2-958-5219. E-mail: [email protected]

2

Department of Materials Science and Engneering, Ajou University, Seoul, Korea

3

Department of Physics, Kookmin University, Seoul, Korea Cu ferrite was prepared by a coprecipitation method in order to investigate the redox

behavior under CH4 and CO2. The weight change by CH4 reduction and CO2 oxidation was monitored by a microbalance. The redox properties of Cu ferrite were observed by TGA, XRD and Mossbauer spectrometer. 1. INTRODUCTION hi order to solve the energy and the environmental problem, the two-step redox reaction with the metal oxide has been proposed to decompose stable gas such as CO2 and H2O at lower the temperature [1]. The two-step redox reaction is based on the redox process of the metal oxide as follows. MO + H2 MOi.6 + CO2

= =

MOi.s + 6H2O MO

+ CO 2

(1) (2)

where MO is the metal oxide and MO 1.5 is the reduced metal oxide. Tamaura et al. reported that Ni ferrite of metal oxide was highly efficient for the CO2 decomposition using the twostep redox reaction at 300 °C [2]. The aim of this paper is to show a behavior of the structure using XRD and Mossbauer spectrometer for the redox reaction of Cu-ferrite with CH4 and CO2. 2. EXPERIMENTAL CuFe2O4 was synthesized by a coprecipitation method. Requisite quantities of FeCl3 and Q1CI2 were dissolved in distilled water. The precipitate was stirred at 80 °C for 6 h. The product was filtered and washed with distilled water several time and dried at 100 °C for 12 h.

146

The dried Cu ferrite powder was calcined at 900 °C for 2 h. In order to observe the redox behavior of CuFe2O4, Cu ferrite was reduced in CH4 gas and oxidized in CO2 gas with a microbalance (TGA, Cahn Instruments Inc., Cahn 2000). The structural change by the redox treatment was analyzed by XRD and Mossbauer spectrometer. XRD was carried out with CuKal radiation (using a McScience, model M18SHF-SRA) to monitor structural changes of Cu ferrite. Mossbauer spectrum was recorded using a conventional Mossbauer spectrometer of the electromechanical type with a 30 mCi source in a Rh matrix at 4.2 K. 3. RESULTS AND DISCUSSION The weight loss of CuFe2O4 was monitored with TGA during the reduction with CH4 gas up to 800 °C (Fig. 1-a). In the reaction, the weight of CuFe2O4 decreased showing a two-step reduction. One step reduction started from 350 °C and the weight of CuFe2O4 decreased about 8 wt% up to 540 °C. The other step reduction started from 640 °C and the weight of CuFe2O4 decreased about 19.5 wt% at 800 °C. There are 4 oxygen atoms in CuFe2O4, and the total weight percentages of oxygen atoms in CuFe2C>4 are 26.7 %. Therefore, the weight loss of 19.5 wt% meant that 3 oxygen atoms in CuFe2O4 were Fig. 1 Weight Change during the CuFe 2 O 4 with CH 4 (a) and oxidation of the reduced CuFe 2 O 4 with CO 2 (b) TGA analysis

eliminated. The reduced Cu ferrite was oxidized by CO2 decomposition reaction. The weight change of the reduced Cu-ferrite during the CO2 decomposition was monitored by TGA (Fig. 1-b).

With increasing the temperature, the reduced Cu-ferrite was oxidized and the weight was increased by CO2 decomposition (Fig. 1-b). The CO2 decomposition with the reduced Cu ferrite occurred from 300 °C. The weight of the reduced Cu ferrite was increased about 12 wt% up to 800 °C. It was recovered to 92% of the fresh CuFe2O4 weight by the CO2 decomposition. The 12% weight recovery meant that 1.8 oxygen atoms supply to the reduced Cu ferrite. The oxygen shortage of 8% from the original CuFe2O4 meant a 1.2 oxygen atom shortage in CuFe2O4. The XRD pattern of Cu ferrite calcined at 900 °C for 2 h showed in Fig. 2-a. The XRD pattern of Cu ferrite was indexed to a single phase of the tetragonal CuFe2O4 with the spinal structure. After the reduction with CH4 up to 800 °C, the spinel structure of CuFe2O4 disappeared, and a mixed phases of Cu, a-Fe, and FeO was observed (Fig. 2-b). The

147

weight loss at 800 CC was 19.5 wt% in TGA result (Fig. 1-a), which was equivalent to the loss of 3 oxygen atoms in CuFe2O4- The elimination of 3 oxygen atoms in CuFe2O4 should give the mixture of Cu, a-Fe, and FeO phases. The oxidation of the reduced Cu ferrite generated the mixture of metallic Cu and Fe3C>4 (Fig. 2-c). XRD pattern indicates that CO2 gas was decomposed by oxidation of Fe and FeO into Fe3O4, and the metallic Cu was not oxidized and remained in a metallic state even at 800 °C in the CO2 atmosphere. The oxygen shortage of 8% from the fresh CuFe2O4 in CO2 oxidation (Fig. 1-b) resulted from the inability of Cu oxidation with CO2. Mossbauer spectrum of the tetragonal CuFe2O4 Fig.2 XRD patterns of (a) fresh CuFe2O4, (b) with the spinal structure showed in Fig. 3-a. CH4 reduced CuFe2O4 and (c) CO2 oxidation C u F e 2 o 4 w i t h t h e s p i n d structure has the cubic of the reduced CuFe2O4; ) CuFe2O4, ) o /Axr , « T- ^ ,,™r n close-packed arrangement 01 the oxygen ions, with r & J Cu, (A) Fe, ) FeO, and ) Fe3O4. ° ' + 3+ the Cu and Fe ions in two different kind sites [3].These sites have the tetrahedral (A site) and octahedral (B site) oxygen coordination. Cu2+ in CuFe2O4 has a relatively large negative the octahedral site preference energy (OSPE), -15.2 kcal/mole, and it indicates that Cu2+ occupy strongly the octahedral (B) site [4]. So, CuFe2O4 has the inverse spinel structure ([Fe3+]A [Cu2+ Fe + ]B O4). Therefore, Mossbauer spectrum of the prepared CuFe2O4 was composed of two sextets arising from Fe3+ in the tetrahedral (A) and octahedral (B) site. Mossbauer parameters obtained by fitting the spectrum at 4.2 K are shown in Table 1. The isomer shift value of Fe3+ ion in the octahedral site is more positive than that of Fe3+ ion in the tetrahedral site because of the larger bond of Fe3+ and O2". Also, the pattern with a larger hyperfine field exhibits a larger isomer shift, which is consistent with the

nearly

between

general

fields

and

correlation shifts

in

ferrites[5]. It is shown clearly in Table

1. In the

tetragonal shift

was

case of the

CuFe2O4, observed

quadrupole for

the

octahedral site while it was nearly zero

for

the

tetrahedral

sites.

Mossbauer spectrum of the reduced Cu ferrite at 800 °C was given in

Table

L Magnetic nyperIme

field (HF), quadrupole

shift (QS), isomer shift (IS), and area % of each Fe site at 13 Kfor ferrites.HF Ferrites

(kOe)

QS (mm/s)

IS (mm/s)

G (mm/s)

A (%)

511.66

-0.01

0.25

0.48

54.6

Fresh

A site

CuFe 2 O 4

B site

543.43

-0.15

0.38

0.46

45.4

CH4 Reduced Ferrite

A site

340.29

0.001

0.11

0.315

60.6

Bsite

466.59

-0.25

0.80

0.419

39.4

CO2 oxidized Ferrite

A site

506.46

-0.25

0.87

0.99

40.6

-0.03

0.30

0.57

59.4

Bsite

513.27

148

Fig. 3-b and Table 1. Mo'ssbauer spectrum recorded at 4.2 K could be decomposed into two sextets of the strong six peaks corresponding to a-Fe and the weak six peaks corresponding to FeO. After the CO2 decomposition, Mossbauer spectrum of the oxidized Cu ferrite was shown in Fig. 3-c. Mossbauer spectrum at 4.2 K showed clearly two hyperfme magnetic splitting, which was a clear evidence for Fe3O4 existing two sextets of A and B site. The tetrahedral (A) site is due to Fe3+ and the octahedral (B) site is due to Fe3+ and Fe2+. The area ratio of the two patterns for A and B sites was 1.46 (Table 1). 4. CONCLUSION The reduced Cu ferrite by CH4 was applied to Fig.3 4.2 K Mossbauer spectra of (a) fresh CuFe 2 O 4 , (b) CH 4 reduction of CuFe 2 O 4 and (c) CO 2 oxidation of the reduced CuFe 2 O 4 ;

CO2 decomposition in order to avoid the greenhouse effects. Cu ferrite calcined at 900 °C for 2 h after preparation was the tetragonal phase (CuFe2C>4) with the inverse spinel structure. After

the reduction reaction up to 800 °C, the spinel structure changed to the mixture of Cu, Fe, and FeO. In the CO2 decomposition, CO2 gas was decomposed by the oxidation of Fe and FeO. Cu was not oxidized and remained in a metallic state up to 800 °C. During the redox reaction of CuFe2O4, the structural change of the tetragonal CuFe2O4 with inverse spinel structure can be summarized as followings. CuFe2O4

-»

Cu

+

Fe3O4 Cu

+

a-Fe

REFERENCES 1. T. Sano, N. Hasegawa, M. Tsuji and Y. Tamaura, J. Mater. Chem. 6 (1996) 605. 2. T. Kodama, Y. Wada, T, Yamamoto, M. Tsuji, and Y. Tamaura, Mater. Res. Bull. 30 (1995) 1039. 3. S. Krupica, and P. Novak, "Oxide Spinels" in Ferromagnetic Materials, ed. E. P. Wolfarth, North-Holland Publishing Co., 1982, Vol. 3. 4. T. G. Reynolds III, "Ferrite (Magnetic) Ceramics" in Ceramic Materials for Electronics, ed. R. C. Buchanan., 1996. 5. B. J. Evans and S. S. Hafner, J. Phys. Chem. Solids 29 (1968) 1573.


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CO 2 Reforming by CH4 over Ni-YSZ Modified Catalysts Dong Ju Moon , Jong Woo Ryu, Dong Min Kang, Byung Gwon Lee and Byoung Sung Ahn Reaction Media Research Center, Korea Institute of Science & Technology (KIST), P.O. Box 131, Cheongryang, Seoul, 130-650, Korea TEL : +82-02-958-5867, FAX : +82-02-958-5809, E-mail: [email protected] The catalytic reforming of CO2 by CH4 over Ni-YSZ modified catalysts was investigated to check the possibility of development of a high performance catalyst electrode for application in an electrocatalytic reforming in the SOFC system. The catalysts before and after the reaction were characterized by X-ray diffraction (XRD), N2 physisorption and CO chemisorption. It was found that Ni-YSZ-CeO2 catalyst is a promising candidate for development as the catalyst anode for an electrocatalytic reforming in the SOFC system. 1. INTRODUCTION It was known that carbon dioxide (CO2) is a greenhouse effect gas and contributes much to global wanning [1, 3-6]. Therefore the reduction and sequestration of CO2 has been attracting interest from an environmental perspective. The CO2 reforming by CH4 is one of the CO2 conversion methods [2-3]: CO2 + CH4 -» 2CO + 2H2, AH0298 = 247 kJ/mol. This reaction has an advantage of the production of synthesis gas as well as the reduction of greenhouse gas. It is of special interest from an industrial perspective since it produces synthesis gas with a low H2 / CO ratio, which can be preferentially used for Fischer-Tropsch synthesis. Furthermore, both CH4 and CO2 are the cheapest reactants. However, the CO2 reforming by CH4 is a highly energy consuming and coke forming process [4-6]. To solve these problems, author suggested an electrocatalytic internal reforming of CO2 by CH4 to coproduce a syngas and electricity in a solid oxide fuel cell (SOFC) system [2, 8-9]. In this work, we have studied the CO2 reforming by CH4 over Ni-YSZ-CeO2 and NiYSZ-MgO catalysts to check the possibility of development of a high performance catalyst electrode material for application in an electrocatalytic reforming in the SOFC system. 2. EXPERIMENTAL The catalyst material was a mixture of NiO and YSZ (NiO : YSZ = 35 : 65 vol%). NiO powder (99.99%, Sigma-Aldrich Co.) has a diameter ca. 6.2 [m, and YSZ powder (TZ-8Y, TOSOH Co.) consisted of ZrO2 doped with 8 mol% Y2O3. The CeO2 (99.9%, Sigma-Aldrich

150

Co.) or MgO (98%, Nakarai Chemical Co.) powder was added to the mixture. The sample was pulverized into powder by a ball mill. The carbon dioxide reforming by CH4 was carried out in a conventional fixed bed reactor system. Flow rates of reactants were controlled by mass flow controllers [Bronkhorst HI-TEC Co.]. The quartz reactor with an inner diameter of 7 mm was heated in an electric furnace. The reaction temperature was controlled by a PID temperature controller and was monitored by a separated thermocouple placed in the catalyst bed. The product gases were analyzed by an on-line GC [Hewlett Packard Co., HP5890 series II] equipped with a carbosphere column (3.18xl0"3 mO.D. and 2.5 m length) and a thermal conductivity detector. BET surface area and total pore volume of catalyst before and after the reaction were determined from N2 physisorption using a sorption analyzer [Quantachrome Co., Autosorb1C]. Structure of catalyst before and after the reaction was analyzed by an XRD analyzer [Shimazdu Co., XRD-6000] equipped with a Cu Ka source at 40 kV, current of 30 mA and scanning speed of 5 deg/min. 3. RESULTS AND DISCUSSION Figure 1 shows the reaction rates of CO2 and CH4 over Ni-YSZ-CeO2 and Ni-YSZ-MgO catalysts with a time on stream. The catalyst of 0.2 g was charged in the reactor. The catalyst was reduced at 750 °C in H2 for 1 h. The catalytic reforming of a mixture of CO2 (12 vol%) and CH4 (12 vol%) was carried out at the reaction conditions of 800 °C, atmosphere and total flow rate of 20 cc/min for 50 h. It was found that Ni-YSZ-CeO2 catalyst displayed higher reaction rate than Ni-YSZ-MgO catalyst under the tested conditions. The reaction rates of CO2 and CH4 over Ni-YSZ-CeC>2 catalyst were obtained 9.1 and 8.2 mmol/cm2-sec, respectively.

Figure 1. The reaction rates of CO2 and CH4 over Ni-YSZ modified catalysts with a time on stream in the catalytic reforming ofCO 2 withCH 4 at800°C.

Figure 2. The effect of reaction temperature on the reaction rates over Ni-YSZ modified catalysts in the catalytic reforming of CO2 with CH4.

151

Figure 2 shows the effect of reaction temperature on the reaction rates of CO2 and CH4 over Ni-YSZ modified catalysts. It was found that the Ni-YSZ-CeC>2 catalyst showed higher activity than the Ni-YSZ-MgO catalyst at temperature range of 650 —800 °C and the maximum activity was observed at above 800 °C. Table 1. The characteristics of the prepared catalysts before and after the reaction _ . _ BET surface Total pore Active metal Catalyst Conditions , 2, s , , ,s -c , ii \ area (m /g) volume (cc/g) surface area (m /g) before reaction 10.2 0.004 0.098 C 2 *~ after reaction 15.7 0.007 0.076 Ni-YSZM gs O

beforereaction

after reaction

8 9 13.4

°003 0.005

°- 0 8 1 0.059

The BET surface area, total pore volume and active metal surface area of catalysts before and after the catalytic reforming of CO2 by CH4 are summarized in Table 1. It was found that the surface area and total pore volume of Ni- YSZ-CeCh and Ni-YSZ-MgO catalysts after the reaction increased, respectively, and on the other hand, the active metal surface area of catalysts after the reaction decreased. These results were interpreted that the BET surface and total pore volume of catalysts after the reaction were increased by the deposition of carbon on the surface of catalyst, and that the active metal surface area of catalysts after the reaction were decreased by the sintering of Ni, which was caused by the reaction at high temperature. Figure 3 represents the X-ray diffraction patterns of Ni-YSZ-CeO2 and Ni-YSZ-MgO catalysts before and after the catalytic reforming of CO2 by CH4, respectively. It was found that NiC phase on the surface of catalyst was produced after the catalytic CO2 reforming over Ni-YSZ modified catalysts. However, when both catalysts were used as catalyst electrode

Figure 3. XRD patterns of Ni-YSZ modified catalysts before and after the catalytic reforming of CO2 by CH4. : NiO, a: NiC T: YSZ, : CeO2 ( a) before reaction (b) after reaction

152

material of the electrochemical cells (Ni-YSZ-CeO2 or Ni-YSZ-MgO I YSZ | (La,Sr)MnO3) in the SOFC system [9], it was found that the carbon deposited on the surface of catalyst under the catalytic reforming was mainly desorbed to carbon dioxide by the reaction of oxygen ion transferred from the cathode in the electrochemical cell. It was identified that the H2 and CO produced in the electrocatalytic internal reforming of CO2 with CH4 can be applied as reactants in the SOFC system, which operated between 800 °C and 1000°C to produce a syngas and electricity. It was suggested that the syngas released from the anode side can be applied as a raw material for the production of highvalued chemicals. 4. CONCLUSIONS Catalytic reforming of CO2 by CH4 was investigated over Ni-YSZ modified catalysts to check the possibility of development of a high performance catalyst electrode for application in an electrocatalytic reforming in the SOFC system. The Ni-YSZ-CeO2 catalyst displayed higher activity than Ni-YSZ-MgO catalyst at temperature range of 650~800°C and the maximum activity was observed at above 800 °C. It was found that the Ni-YSZ-CeO2 catalyst is a promising candidate for development as a catalyst anode material of an electrocatalytic reforming in the SOFC system. ACKNOWLEDGMENT This paper was performed for the Carbon Dioxide Reduction & Sequestration Center, one of 21st Century Frontier R&D Programs funded by the Ministry of Science & Technology of Korea. REFERENCE 1. T. Shishido, M. Sukenobu, H. Morioka R. Furukawa, J. Shirahase and K. Takenhira, Catal. letters ,73(1) (2001) 21. 2. D. J. Moon, T. Y. Kim, J. H. Kim, S. D. Lee and S. I. Hong, Eng. Conf. Inter., Quebec, Canada, July 2 (2003) 33. 3. S. Wang and G Q. Lu, Energy & Fuels, 12 (1998) 248. 4. T. Takeguchi, Y. Kani, K. Ehuchi and M. Aizwa, J. of Power Sources 112 (2002) 588. 5. C. J. Michael, and M. A. Vannice, J. of Catalysis, 183 (1999) 69. 6. S. B. Tang, F. L. Qiu and S. J. Lu, Catalysis Today, 24 (1995) 253. 7. Z. W. Liu , H. S. Roh and K. W. Jun, J. Ind. Eng. Chem., 9(3) (2003) 267. 8. D. J. Moon and J. W. Ryu, in press to Catalysis Today (2003). 9. D. J. Moon, J. W. Ryu, T. Y. Kim, S. D. Lee and B. G. Lee, in application to Korea Patent & U.S. Patent (2003)


The interation between CO2 and CH4 on Ru-Co-catalysts L.B. Shapovalova, G.D. Zakumbaeva, LA. Shlygina, A.A. Shyrtbaeva

The Institute of Organic Catalysis & Electrochemistry of the Ministry of Education and Science of the Republic of Kazakstan; 142, Kunaev str., Almaty, 480100, Kazakstan; Fax: (007) 3272 915722; E- mail: [email protected]

The interaction between CO2 and CH4 on R11-CO/AI2O3 and Ru-Co/Al2C>3 modified by ZSM-zeolite catalysts of clusters type has been studied with the application of the experimental and quantum-chemical methods. The basic products of CO2+ CH4 interaction are Ci-C4-alcohols and Ci-C4-acids on Ru-Co/Al2O3. The CO and hydrogen are dominant in the CO2+ CH4 interaction products formed over RU-CO/AI2O3+ ZSMzeolite catalyst in these conditions. Quantum-chemical calculation has been shown that CO2 and CH4 molecule can introduce into the mono- and bimetallic Co-, Ru- and Co-Ru -clusters. Quantum-chemical accounts show that the C-H bonds lengthening for "CH4 + CO2 + M-clusters" systems are less then ones for "CH4+ M-clusters". However, binding energy is stronger in "M-clusters+CO2+CH4"-systems than ones in "M-clusters+CO2" or "M-clusters+CH4 1. INTRODUCTION The interaction between CO2 and CH4 has been proposed as one of the most promising technologies for utilization of these two gases. The molecules of the greenhouse gases CO2 and CH4 have very strong energies of C-0 and C-H-bonds. Most of reactions involving CO2 activation contain metal atoms as the active center. The application of multicomponent metallic supported catalytic systems open the wide opportunities for activation of molecules and the chemical syntheses on the CO2 and CH4 base Recently the results of study of interaction between hexene-1 or propylene and carbon dioxide on R11-CO/AI2O3 and RI1-CO/AI2O3 have been published [1, 2, 3]. In this paper the process of interaction between CO2 and CH4 over RU-CO/AI2O3 and RuCo/Al2O3 +ZSM cluster type catalysts has been studied. 2. EXPERIMENTAL The interaction between CO2 and CH4 on Ru-Co/Al2O3 and Ru-Co/Al2O3 modified by ZSM-zeolite catalysts of clusters type has been studied with the application of the experimental and quantum-chemical methods. Quantum-chemical calculation have been made on the basis of ZINDO-1 (quantum-chemical programs Hyperchem-6). The CO2+

153

154 CH4 reaction was carried out in flow type reactor in the range of 473-723 K and pressure variation from 0.1 to 1.0 MPa. Catalysts were prepared by impregnation of support with mixture of RuC^ and Co(NO3)2-6H2O solution. Then they were reduced by hydrogen at 773K during 3 hours, washed from Cl and NO3 ions and dried up in the air at 303-323K. Catalyst was additionally reduced directly in the reactor at temperatures from 473 to 673K during 1 hour before the reaction between CO2 and CH4.The reactant gas mixture consists in CO2, CH4 diluted with Ar. The ratio of CO2:CH4:Ar is 1:0.5:6. The reaction rate was controlled on CO2 and CH4 decrease by using a chromatographic analysis. IR-spectra of reactants adsorbed on catalyst surface were recorded in a Specord IR75 spectrometer in the 1200-3500 cm"1 range. 3. RESULTS AND DISCUSSION The basic products of CO2+ CH4 interaction are Ci-C4-alcohols and Ci-C4-acids and aldehydes on Ru-Co/Al2O3 . For example, the CH4 conversion is 25.9% at T=573K and P=0.35 MPa. The reaction products are methanol (3.8%), formaldehyde (1.4%), ethyl alcohol (2.7%), ethyl (1,4%) and propionic (8,6%) aldehydes, propionic (35,5%), butyric (19.5%), formic(10.8%) and acetic (20.2%) acids. At 0.6 MPa and T=573K the CH4 conversion is 46.8%. The propionic (20.1%) and acetic (31.5%) acids, propionic aldehyde (17.8%), butanol-2 (3.8%), methanol (7.6%) and formaldehyde (1.9%) are mainly products under these conditions. Besides that there are traces of CO, O2, H2 and C2-C4-hydrocarbons. It should be noted that the CO and hydrogen are dominant in the CO2+ CH4 interaction products formed over RU-CO/AI2O3+ ZSM-zeolite catalyst in these conditions. It was calculated the adsorption models of 3 types: "M-cluster + CO2-", "M-cluster + CH4" and "M- cluster + CH4 + CO2" by quantum-chemical ZINDO methods (Table 1). There was used ZINDO method giving more high value of binding energy of molecules in comparison with real ones. But this method allows to estimate the changes in binding energy in complex: "M-cluster + molecule" depending on the cluster nature. The accounts were made in comparable conditions. It has been shown that CO2 molecule can introduce into the mono- and bimetallic Co-, Ru- and CO-Ru-clusters. The .Êbind change shows that binding energy between the atoms of the CO2 molecule and the metal atoms of Ru-, Co- and Ru-Co-clusters increases when the quantity of Ru rises. The insertion of the CO2 molecule into the Co-, Ru- and Co-Ru-clusters changes their configurations and M-M-lengths. Analogical phenomena are observed when the interaction between CH4-molecules and Co-, Ru or Ru-Co-clusters takes place (Table 1). In case CH4-adsorbtion on 4Ruclusters the M-M-lengths are increased from 2.66-3.24 A to 3.29-3.64 A. For 2Co2Ru- clusters the M-M-lengths are 2.67-4.86 A (without CH4) and 3,26-3,72 A (with CH4). 4Co- clusters are exposed to the radical changes by reason of the CH4 adsorption: it was became plane. There are the loosing and cleavage of C-H-bonds of CH4 molecules adsorbed on Co-, Ru- and Co-Ru-clusters. The binding energy of CH4 molecules with mono- and bimetallic clusters are —418.88 kcal/mol (4Ru); -330.1 kcal/mol (2Co-2Ru) and -184.03 kcal/mol (4Co). The strongest bond is observed for system "CH4+4Ru".

155

Table 1 Quantum-chemical accounts of CO2, CH4 and CO2+CH4 over Ru-, Co- and Co-Ruclusters Bond lengths, A2Ebind Ebind Al Ebind Complexes Ml-Ml kcal/mol C-H C-0 -713.74 OCO 1.23 -1001.76 1.09 CH4 2.82-3.84 126.70 Co4 63.29 Co4 2.84-4.11 305.42 2.82-4.35 ColO 2.67-4.86 66.51 Co2 Ru2 2.54-3.20 70.23 Co2 Ru2 2.66-3.24 38.18 Ru4 2.66-3.12 -115.57 RulO 3.39-3.91 -1022.34 -434.60 1.38 Co4 CO2 2.82-3.90 -991.67 -583.35 1.35 C0IOCO2 2.90-3.80 -1165.83 -518.59 1.39 Co2Ru2CO2l 3.10-3.82 -1188.57 -541.34 Co2Ru2CO22 1.38 3.10-3.53 1177.58 530.35 Co2Ru2CO23 1.37 3.09-3.84 1185.79 538.56 Co2Ru2CO24 1.37 2.97-3.80 -1174.28 530.77 Co2 Ru2 CO2 1.41 2.71-3.70 -1274.88 -599.32 1.42 Ru4 CO2 2.80-3.60 -1595.21 -765.90 1.46 RulO CO2 -1122.50 1.14-1.15 2.81-3.58 184.03 Co4 CH4 1.18-1.22 3.28-3.72 1265.32 330.07 Co2Ru2CH4 1.23-1.28 3.29-3.64 -1382.46 418.88 Ru4 CH4 1.10-1.12 3.28-3.66 Co4CO 2 CH 4 -2097.40 445.19 1.37 1.10-1.16 2.93-4.70 Co2Ru2CO2CH4 -2371.67 722.68 1.35 1.10-1.15 2.70-3.67 Ru4CO 2 CH 4 -2484.45 807.13 1.43 A, E b i n d = E bind ( MLnCOz) - EwndCMLn) - E bind (CO 2 ); A E b i n d = E bind ( MLnCHU) - Ebind (ML n ) - E bind (CH 4 ); A E b i n d = E bind ( MUCO2CH4) - E b i n d (ML n ) - E bind (CO 2 ) - Ebind (CH 4 ) *** ****E bind ( ML n CO 2 ) - binding energy of: "M-cluster + CO 2 .", E bind (ML n ) - binding energy of M-cluster, E bind ( ML n CO 2 CH 4 ) - binding energy of "M- cluster + CH 4 + CO 2 ", Ebind (CO2) - binding energy of the CO2 molecule, Ebind (CH4) - binding energy of CH 4 -molecules, Ebind (ML n CH 4 ) - binding energy of "M-cluster + CH 4 " The figure 1 shows the optimized structures of joint adsorption of CH 4 + CO2 on mono- and bimetallic Co-, Ru- and Co-Ru-clusters. Quantum-chemical accounts show that the C-H bonds lengthening for "CH 4 + CO 2 + M-clusters" systems are less then ones for "CH 4 + M-clusters". However, binding energy is stronger in "Mclusters+CO2+CH 4 "-systems than ones in "M-clusters+CCV or "M-clusters+CH 4 ". IR-data of CO2 adsorption on Co-Ru-catalysts are conformed with quantumchemical calculations. Adsorption bands of CO 2 gas (2350 sm"1), CO2ads(1580 and 1440 sm"1). CO ads (1950 sm" 1 ,2020 sm"1) are presented in IR-spectra of adsorbed CO2.

156

Fig. 1 The optimized structures of adsorbed CO2 and CH4 on Co4 (A), Ru4 (B) and Co2Ru2 (C) clusters

The experimental and quantum-chemical accounts show that the mechanism of the interaction between CO2 and CH4 on RU-CO/AI2O3 is very complicated. Probably the first stage of the CO2 + CH4 interaction is the cleavage of molecules bonds: CH4 -> CH3rac + HMC CH3 a f l c

—> CH2aac

+

Haac

CO 2 -> 0 0 ^ + O

w

-» C O T + O W

COajjc ~~* C a a c + Oaac

Syntheses-gas (CO + H2) may be the one of the possible product. This supposition is confirmed with the results of the CO2 + methane interaction on RU-CO/AI2O3+ ZSMzeolite catalyst. IR-Spectroscopy data of NH3 adsorption show that the ZSM-zeolite incorporation into RU-C0/AI2O3 is accompanied by the increase Lewis acidic centers (adsorption bands at 3550, 3400, 1600 cm"') and catalyst's de-structive ability. The CO and hydrogen are main products formed over RU-CO/AI2O3+ ZSM-zeolite catalyst. REFERENCES 1. G.D.Zakumbaeva, L.B.Shapovalova, Japan-FSU Catalysis. Seminar (1994). "Catalytic Science and Technology for 21 Centure Life", Japan (1994) 28. 2. G.D.Zakumbaeva, L.B.Shapovalova, Advances in Chemical Conversions for Mitigating Carbon Dioxide. Studies in Surface Science and Catalysis. Elsevier Science B.V. V. 114 (1998) 171. 3. L.B.Shapovalova, G.D.Zakumbaeva , A.V. Gabdrakipov, LA. Shlygina, A.A. Shyrtbaeva, Applied. Organomet. Chem. V.14 (2000) 853


157

Selective Formation of Light Olefins by the Cracking of Heavy Naphtha over Acid Catalysts Sang Yun Hana, Chul Wee Leea, Jeong Ri Kima'b, Nam Sun Hana'b, Won Choon Choia, ChaeHo Shinb and Yong-Ki Park3* a

Advanced Chemical Technology Division, Korea Research Institute of Chemical Technology (KRICT), Daejeon 305-600, Korea. bDepartment of Chemical Engineering, Chungbuk National University, Chungbuk 361-763, Korea FAX: +82-42-860-7388 E-mail: [email protected] The catalytic cracking of heavy naphtha was carried out over various types of proton form zeolites such as of H-ZSM-5, H-Mor, H-Beta, and SAPO-11 for the selective production of ethylene (C2=) and propylene (C3=). Among them, the H-ZSM-5 showed the highest olefin selectivity and propylene yield. The yield of light olefins and product distribution were also strongly influenced by the concentration and strength of acid sites. Maximum yield of ethylene and propylene was obtained over H-ZSM-5 with the lowest Si/Al ratio. In the optimized reaction conditions, the C^ + C3~ yield of 48.5 wt % with C^ /Cf ratio of 0.90 was obtained. 1. INTRODUCTION Light olefins, such as ethylene and propylene, are important raw materials for various petrochemicals. These olefins are mainly produced by the thermal cracking with steam in many countries. This thermal cracking process consumes about 40 % of the energy required in the entire petrochemical industry. Therefore, global warming issues have stimulated the development of cracking process which maximize energy-saving and minimize CO2 emission. In addition, there is an increasing demand of cracking technologies capable of controlling the composition of olefins through naphtha cracking. While further improvements in the thermal cracking process are obviously possible, one route which appears to offer considerable potential is the application of catalytic cracking. The catalytic cracking has been proposed as an alternative technology of naphtha cracking to obtain the desired olefin yield with much less energy cost and CO2 emission as well as biomass cracking [1,2,3]. Recently, it has been reported that propylene demand is continuously increasing than that of ethylene [4]. In this work, the catalytic cracking of heavy naphtha over various zeolite catalysts has been investigated. Parameters such as the concentration and strength of acid site, pore structure of catalyst and reaction temperature were considered as main factors of catalytic cracking activity. Among the various types of catalyst, H-ZSM-5 was found to be the most effective. Thus, the effects of concentration and strength of acidic site over H-ZSM-5 were investigated. And the reaction conditions such as temperature, contact time, and naphtha/steam ratio were optimized for the maximum yield of olefins. 2. EXPERIMENTAL

158

H-ZSM-5 (Si/Al=20, 25, 40, and 75), H-Beta (Si/Al=150), and H-Mor (Si/Al=12.5) were purchased from Zeolyst Co., and SAPO-11 was synthesized according to the literature [5]. The cracking activity of heavy naphtha over the prepared catalysts was evaluated using a fixed-bed flow reactor made of Inconel under atmospheric pressure. A known amount of catalyst was put into the reactor and then pretreated in helium flow at 500 °C for 1 hr before each reaction. The cracking condition was as follows; Naphtha/steam(wt/wt) = 2, weight hourly space velocity(WHSV) = 5 h'1, temperature = 650°C. As a feed, a heavy naphtha containing 22.0 wt% n-paraffins, 33.2 wt% i-paraffins, 19.8 wt% naphthenes, 11.5 wt% olefins, and 13.6 wt% aromatics and average molecular weight of about 106 g/mol was used. The produced gas and liquid hydrocarbons were separated by cooling at -10°C. The gas product was analyzed by an on-line GC(HP 6890) equipped with NP-l(Alltech) and GSCARBONPLOT(Agilent) columns and the condensed liquid product was analyzed by the other GC equipped with AT™-PETRO(Alltech) column. The selectivity of product was calculated based on the weight ratio of individual component yield to the overall hydrocarbon yield. 3. RESULTS AND DISCUSSION The catalytic cracking of heavy naphtha was carried out at 650 °C over various types of proton form zeolites to determine the effect of concentration, strength of acidic site and pore dimension on the conversion and selectivity to light olefins. As shown in figure 1, the olefin yield of various types of zeolite catalysts was compared with that of steam cracking. Improvement in olefin yield and CflCf ratio was not observed over the acidic catalysts except H-ZSM-5. The yields of Cf + Cf over H-Mor, H-Beta, SAPO-11, and steam cracking were 11.2%, 15.2%, 13.3%, and 12.9%, respectively while that over H-ZSM-5(20) was 37.5%. For the CflCf ratio, a guide to decide the degree of catalytic cracking, the H-ZSM-5 also showed a high improvement and revealed the value of 0.99. This means that the propylene could be produced selectively through the catalytic cracking of heavy naphtha over the H-ZSM-5 without the loss of total olefin yield. For detail investigation of the dependence of concentration and strength of acid sites, cracking reaction was carried out while changing the Si/Al ratio of H-ZSM-5 (Figure 2). As the Si/Al ratio decreased from 75 to 20, the yield of Cf + Cf increased from 29.3 % to 37.5 % together with the increase of propylene yield. And at the low Si/Al ratio

Fig. 1. Gas yield in catalytic cracking of heavy naphtha over various types of H-form zeolites. The number in parenthese indicates Si/Al ratio (temperature=650°C, naphtha/steam(wt/wt)= 2, WHSV=5rf')

Fig. 2. Gas yield in catalytic cracking of heavy naphtha over H-ZSM-5 with different Si/Al ratio (temperature=650°C, naphtha/steam(wt/wt)= 2, WHSV=5h-').

159

Fig. 3. Gas yield as a function of temperature on the catalytic cracking of heavy naphtha over H-ZSM-5(20) at naphtha/steam(wt/wt)= 2andWHSV=2h"'.

Fig. 4. Gas yield as a function of contact time on the catalytic cracking of heavy naphtha over HZSM-5(20) at 650 °C and naphtha/steam (wt/wt)=2.

propylene as well as ethylene was produced preferentially without the loss of total olefm yield. It is thought that it might be caused by the enhanced free radical initiation of hydrocarbons due to the increased number of acid sites and Lewis/Bronsted acid site ratio. According to our previous result, the number of acid sites and the ratio of Lewis/Bronsted acid sites of H-ZSM5 were increased proportional to the aluminum content [6]. Based on the mechanism of catalytic cracking, the high yield of C3~ and Cf olefins at low Si/Al ratio could be explained by the increased stability of secondary or tertiary carbonium ions in the presence of acid sites before {3-scission of hydrocarbons [7]. For the H-ZSM-5(20), the reaction temperature, contact time, and naphtha/steam ratio were optimized to maximize total yield of olefins. As the temperature increased, the ethylene yield was increased gradually up to 675°C but that of propylene showed maximum at 650°C and then began to decrease as shown in figure 3. The decrease of propylene yield above 650°C is thought to be caused by further cracking of the produced propylene due to the enhanced cracking rate at high temperature. The yield of ethylene and propylene was also influenced by the contact time. The yield of total olefin and propylene showed a maximum at a certain WHSV (Fig. 4). As the contact time increased, aromatics in liquid product and methane in gas product increased exponentially. This means that the olefins such as C3" and Cf are an intermediate product and precise contact time control with narrow distribution is required to prevent secondary reactions of the produced olefins such as aromatization or further cracking. Even not shown in this result, it was found that the yield of olefins was also influenced by the naphtha/steam ratio of feed. The steam in the feed played an important role as a diluting agent or an inhibitor for the secondary reaction of produced olefins. When the naphtha/steam ratio increased from 1 to 4, the yield of total olefins was decreased from 37.5 wt% to 36.1 wt% and CO and CO2 were produced in the presence of steam through the reactions of steam reforming and/or coke gasification reaction [8]. The product distribution of gas and liquid products in optimized reaction condition was summarized in Table 1. The naphtha conversion, total gas yield, and total olefin yield were 77.9 wt%, 71.7 wt %, and 48.5 wt %, respectively. The aromatics in the feed was 13.6 wt% but after cracking it increased to 16.5wt%. About 3.0 wt% of aromatics mainly consisted of BTX was formed through the secondary reaction during the cracking process. It is thought that it was caused by the hydride transfer from the produced olefins to a carbonium

160

Table 1. Product distribution in heavy naphtha cracking over H-ZSM-5(20) Feed ( wt % ) Product ( wt % ) (Liquid composition) n-paraffms 22.0 2.94 i-paraffins 33.2 5.32 Naphthenes 19.8 2.43 Olefms 11.5 1.13 Aromatics 13.6 16.5 (Gas compsotion) Ethylene 22.9 Propylene 25.6 Ethylene + Propylene 48.5 Total gas yield 71.7 Ethylene/Propylene ratio 0.90 Reaction conditions: 675"C, WHSV = 2 h'1, naphtha/steam (wt/wt) = 2

intermediate to form paraffin and hydrogen-deficient species and then further transformed into aromatic hydrocarbons and coke on acid sites of catalyst. 4. CONCLUSION Three dimensional medium-sized nanoporous zeolite such as H-ZSM-5 with lower Si/Al ratio was found to be effective for the production of light olefins with high propylene selectivity in the cracking of heavy naphtha. The olefin yield and the product distribution were strongly influenced by the reaction temperature, contact time and naphtha/steam ratio. In the optimum reaction condition, the C2= + C3= yield of 48.5 wt % and C£ ICi° ratio of 0.90 were obtained. ACKNOWLEDGMENTS This research was supported by a grant from Carbon Dioxide Reduction & Sequestration Research Center funded by the Ministry of Science and Technology of Korea. REFERENCES 1. Y. Yoshimura, N. Kijima, T. Hayakawa, K. Murata, K. Suzuki, F. Mizukami, K. Matano, T. Konishi, T. Oikawa, M. Saito, T. Shiojima, K. Shiozawa, K. Wakui, G. Sawada, K. Sato, S. Matsuo and N. Yamaoka, Catalysis Surveys from Japan, 4 (2000) 157. 2. Komatsu, T., Ishihara, H., Fukui, Y, and Yashima, T., Appl. Catal. A., 214 (2001) 103. 3. Y.-Y. Lee and B.-H. Lee, J. Ind. Eng. Chem., 4(4) (1998) 334. 4. C&EN March 17 (2003) 23. 5. K.V.V.S.B.S.R. Murthy, SJ. Kulkarni and S. Kahaja Masthan, Micropor. Mesopor. Mater. 43(2001)201. 6. J.W. Yoo, C.W. Lee, S.-E. Park and J. Ko, Appl. Catal. A. 187(1999) 225. 7. Richard H. McCue, "Catalytic Olefins Production" AIChE spring National Meeting 2003. 8. Basu. B., Kunzru. D., Ind.Eng.Chem.Res., 31 (1992) 146.


161

Effect of additives and a preparation method on catalytic activity of Cu/ZnO/ZrO 2 system in the carbon dioxide hydrogenation to methanol J. Sloczynski, R. Grabowski,1 A. Kozlowska, M. Lachowska,2 J. Skrzypek2 1

Institute of Catalysis and Surface Chemistry, Polish Academy of Sciences, Krakow-Poland Institute of Chemical Engineering, Polish Academy of Sciences, Gliwice-Poland *[email protected]

2

Effects of a preparation method, Cu-loading, and the addition of MgO and MnO on the catalytic activity and the adsorptive properties of Cu/ZnO/ZrC>2 catalysts were studied. The catalysts obtained by a co-precipitation method showed a higher catalytic activity in the methanol synthesis from CO2 in comparison with the catalysts obtained by a citric method. The MnO additive introduced to the catalyst obtained by the citric method led to an increase of the catalytic activity. A simple correlation between the adsorptive properties towards reactants and the catalytic activity was established. 1. INTRODUCTION Synthesis of methanol from CC^has gained an increasing interest in the past decade [1, 2] since it contributes to the reduction of the greenhouse effect. On the other hand, CO2 may be regenerated in the hydrogenation-steam conversion cycle, and hydrogen produced can be used in fuel cells [3-5]. High activity of the catalysts containing copper and zirconia in the hydrogenation of CO2 is well documented [6-10]. hi the present work, possibilities of modifying the catalysts were studied. The aim of these modifications was to change the population of the active centers and to augment the activity of the catalysts. 2. EXPERIMENTAL The oxide precursors containing the same carrier (ZnO/ZrO2 = 3) were obtained by (A): co-precipitation from aqueous solutions containing nitrates of the appropriate metals and sodium carbonate at constant pH=7+8, and subsequent calcination of the obtained precipitate at 623K, (B): complexing the solution components by using citric acid, followed by drying and calcination at 623K. Oxide additives (MnO and MgO) were introduced using the same methods. The catalysts were obtained by the reduction of CuO in the oxide precursors in the stream of 5% H2 in N2 at 523K. The precursors and the catalysts were characterized by the measurements of the specific surface area (BET method, N2 ads. at 77.5K) and the active surface of Cu (reactive adsorption of N2O at 363K). Phase composition and size of the crystallites were determined by XRD and TEM methods, and XPS method was used for the determination of the surface composition. Catalytic activity in the methanol synthesis reaction

162

was determined in a high-pressure flow reactor (493K, 8MPa, SV=3300h~', H2/CO2=3). Adsorption of the methanol synthesis reactants on the reduced catalysts was investigated under static conditions using vacuum microbalance. 3. RESULTS AND DISCUSSION A basic physicochemical characterization of the oxide precursors and the catalysts is presented in table 1. The data indicate that samples A obtained by the co-precipitation have surface areas five times higher than samples B obtained by the citric method. The average size of the CuO crystallites in the oxide precursors and the Cu crystallites in the reduced catalysts is higher for B samples, and the Cu dispersion decreases with the increase of the Cu content in the system. The addition of MgO and in particular of MnO brings about the increase in the dispersion of Cu. Furthermore, addition of MgO prevents sintering of the catalyst.The results of the measurements of the copper surface with the use of the reactive adsorption of N2O confirm the above tendencies. Table 1. Physicochemical characterization of the catalysts. catalyst catalyst surface area [m g" ] oxide catalyst after at % Cu symbol precursor reaction 2 100 76 A, 101 64 10 A2 62.5 89 39 A3 65.7 19.5 16.6 Bi 13.4 64.3 + 7.5Mg 11.8 B2 21.2 67.7 + 2.2Mn 14.8 B3 a - after reduction at 523K

copper surface areaa [m^gCu) 1 ] 50.5 27.7 14.2 6.9 15.1 22.2

copper crystal sizea [nm] 7.9 13.4 39.3 22.3 20.5

The dependence of the space time yield of methanol on temperature is shown in Fig. 1. Catalyst A3 containing 62.5% of Cu, obtained by the co-precipitation method, proved to be the most effective in the practical applications but catalyst B3 with the MnO addition had a similar activity.

163

As demonstrated in Table 2, the addition of MgO and MnO leads to an increase of the specific activity of the catalyst (calculated per m2 of the catalyst). The same table also contains a comparison of the catalytic activities, expressed in several ways for the temperature of 493 K, with the population of the active adsorption centers of the reactants, represented by reactant adsorption factor (RAF), is given in Table 2. Table 2. Catalytic activity and adsorptive properties of the catalysts.

catalyst

Catalytic activitya [mmolMeOH/h] [ ]/gcat [ ]/gCu [ ]/m2cat

Al A2 A3 Bl B2 B3 a-T=493K

1.3 2.2 5.0 2.8 2.7 4.4 p = 8Mpa

92.0 31.1 9.3 4.9 4.5 7.2

0.017 0.034 0.128 0.169 0.230 0.296

TOFxl0J f molec.MeOH _^|

[_

surf.at.Cu " J 5.05 3.25 2.40 0.65 0.31 0.26

RAF dimension less 1.42 1.56 1.95 3.80

GHSV= 3300h-1 H2/CO2=3

The total adsorption of the reactants was a sum of the irreversible adsorption (about 10% of the total adsorption) and the reversible adsorption, which was described by the Langmuir equation. The limits of the reversible adsorption-nj ( for p->oc) obtained from the Langmuir equation, were taken as a measure of the number of active centers for each individual reactant T . On this base an empirical factor RAF reflecting the adsorptive properties of the catalysts could be defined: RAF=(nCo2.ncH3OH)/(nco»nH2o) (1) Our results showed that CO and CO2 adsorbed concurrently on the same centers (probably Cu), and that methanol and water also adsorbed concurrently but on the centers with the participation of the support. Similar results have been obtained by Mortera et al. [11] for the adsorption of CO and CO2 on ZrO2. Also Liu et al. [12] have shown that water adsorbs concurrently in relation to CO2 and methanol, and inhibits the reaction of the methanol synthesis from CO2. It can be expected that every modification of the adsorptive centers by an appropriate preparation method, or by the application of the promoters, which facilitate the adsorption of CO2 at the cost of CO, increase the adsorption of methanol, and restrict the adsorption of water, should lead to the increase of participation of the methanol synthesis reaction: CO2 + 3H2 -> CH3OH + H2O (2) and limit the participation of the RWGS reaction: CO2 + H2 -> CO + H2O (3) The results presented in the table 2 show that the catalytic activity in the reaction of the methanol formation from CO2, related to the unit surface area of the catalyst, increases indeed with the increase of the factor RAF. The catalytic activity of the copper (per g Cu) generally increases with the decrease of the size of the copper crystallites, but obviously other effects also play an important role (eg. the effect of the MnO additive). In fact, with the increase of the specific activity of the catalysts (per m2 of catalyst), TOF (related to one Cu surface atom) decreases, which

164

demonstrates that other active centers than copper account for the reaction of the methanol synthesis from CO2. They probably involve the participation of atoms of the support, which is in accordance with the results of the of the adsorption studies described above. ACKNOWLEDGMENT The work was financed by the State Committee for Scientific Research , KBN under project No7T09C010 21

REFERENCES 1. J. Haggin, Chem. Eng. News. 72 (1994) No. 13, 29 2. Proceedings of the advances in Chemical Conversions for Mitigating Carbon Dioxide" Conference , Kioto 1997, in Studies in Surface Science and Catalysis vol.114, T. Inui, M. Anpo, K. Izui, S. Yanagida, T. Yamaguchi 3. J. P. Breen, J.R.H. Ross, Catal Today. 51 (1999) 521 4. S. Velu, K. Suzuki, M. Okazaki, M.P. Kapoor, T. Osaki and F. Ohashi, J. Catal. 194 (2000) 373 5. P. J. de Wild, M. J. F. M. Verhaak, catal Today. 60 (2000) 3 6. R.A. Koppel, A. Baiker and A. Wokaun, Appl. Catal. A: 84(1992) 77 7. C. Frohlich, R.A. Koppel and A. Baiker, Appl. Catal. A: 106(1993) 275 8. Y. Ma, Q. Sun, D.Wu, Wen-Hao Fan, Yu-Long Zhang, Jing-Fa Deng, Appl. Catal. A: 171 (1998)45 9. J. Wambach, A. Baiker, A. Wokaun, Phys. Chem. Chem. Phys., 1 (1999) 5071 10. E. E. Ortelli, J. Wambach, A. Wokaun, Appl. Catal. A: 216 (2000) 227 11. C. Mortera, L.Orio and C.Emanuel, J.Chem.Soc.Faraday Trans. ,86 (1990)3003 12. G.Liu, D.Willcox, M.Garland, H.H.Kung,J.Catal 96(1985) 251


165

Selective Formation of Ethylene Carbonate from Ethylene Glycol and Carbon Dioxide over CeO2-ZrO2 Solid Solution Catalysts

Keiichi Tomishige*, Hiroaki Yasuda, Mohammad Nurunnabi, Baitao Li, Kimio Kunimori Institute of Materials Science, University of Tsukuba 1-1-1 Tennodai, Tsukuba, Ibaraki 305-8573, Japan Tel + Fax: +81-29-853-5030, E-mail: [email protected] Abstract Ethylene carbonate can be formed from ethylene glycol and CO2 over CeO2-ZrO2 catalysts. The activity of ethylene carbonate formation was much dependent on the catalyst composition and calcination temperature, and it is found that CeC>2-ZrO2 (Ce/(Ce+Zr)=0.33) calcined at high temperature (1273 K) is very effective catalyst. The by-product of diethylene glycol was formed at higher reaction temperature than 423 K, however, at lower temperature, diethylene glycol was not detected. 1. INTRODUCTION Dimethyl carbonate (DMC) has attracted much attention in terms of a non-toxic substitute for dimethyl sulfate and phosgene, which are toxic and corrosive methylating agents [1-3]. In addition, DMC is considered to be an option for meeting the oxygenate specifications for the transportation fuel [4]. The traditional synthesis of DMC used to require phosgene as a reagent. There are two processes based on the oxy-carbonylation of methanol have been already successfully developed up to larger scale: (1) the oxidative carbonylation of CH3OH with carbon monoxide and oxygen catalyzed by cuprous chloride [5], and (2) an oxidative carbonylation process using a palladium catalyst and methyl nitrite promoter [6]. On the other hand, conversion of carbon dioxide to useful industrial compounds has recently raised much interest in view of the so-called "'Sustainable Society" and "'Green Chemistry" [7]. H

H2C^

IÔ H

2C

^ - \

+ CO2

C=O

H2C-0

H3C-O C=O + 2CH3OH

*-

H 2 C—o' H2C—OH I H 2 C-OH

(1)

H 2 C—O H ,C-OH C=O + "| (2) H3C-V H2C-OH

H2C—O + CO,

-

\=O H2C-O/

+

H2O

(3)

The utilization of carbon dioxide as the raw material of the DMC synthesis process has been

166 attempted. One possible reaction scheme is ethylene carbonate formation from ethylene oxide and CO2 (Eq. (1)) [8, 9], and the transesterification of ethylene carbonate (EC) with methanol (Eq. (2)) [10-12].In this system, ethylene glycol (EG) is always co-produced with DMC. If ethylene carbonate can be synthesized from EG and CO2, the reaction scheme becomes more valuable since EG can be reused. However, the synthesis of EC from EG and CO2 (Eq. (3)) has not been reported. Some catalysts have been reported to be effective for the synthesis of DMC from CH3OH and CO2. Recently we have found that ZrC>2 is effective to selective DMC formation from CH3OH and CO2 [13-17]. In addition, it is found as a result of catalyst development that CeO2-ZrO2 solid solution catalysts exhibited much higher activity than ZrO2, especially at lower reaction temperature [18, 19]. Here we demonstrate that CeCb-ZrCb can be applied to the reaction between EG and CO2 to form EC. 2. EXPERIMENTALS CeO2-ZrO2, CeO2 and ZrO2 were prepared by calcining the hydroxides (available from Daiichi Kigenso. Japan) for 3 h under air atmosphere at various temperatures (673-1273 K). The preparation method of the hydroxides was on the basis of the patent [20]. The molar ratios Ce/(Ce + Zr) of the catalysts were 0, 0.33, 0.5 and 1.0. The reaction was carried out in a stainless-steel autoclave reactor with an inner volume of 70 ml. The standard procedure is as follows: 6.20 g EG (100 mmol, dehydrated), 4.92 g CH3CN (solvent, 120 mmol. 99%, dehydrated) and 0.05-0.5 g catalyst were put into an autoclave, and then the reactor was purged with CO2. After that, the autoclave was pressurized with CO2 (200 mmol, 99.99%). The reactor was heated and magnetically stirred constantly during the reaction. Products in liquid phase were analyzed by a gas chromatograph (GC) equipped with FID. The capillary column TC-WAX was used for the separation column. In the liquid phase, EC and diethylene glycol (DEG) were observed as the products. On the other hand, all the products in the gas phase was below the detection limit of FID-GC. The surface area of the catalyst was measured with BET method (N2 adsorption) using Gemini (Micromeritics). 3. RESULTS AND DISCUSSION Figure 1 shows the calcination temperature dependence of EC formation starting from EG and CO2 catalyzed by CeO2, CeO2-ZrO2 (Ce/(Ce + Zr) = 0.33 and 0.5). Although the result over ZrO2 was not shown here, no EC formation was detected over ZrO2 with various calcination temperatures (673-1273 K). BET surface area of the catalysts is also shown in Figure 1. As shown later, the EC amount at the equilibrium level is 1.18 mmol under these reaction conditions. In the case of CeO2, the EC amount reached a maximum on the catalyst calcined at 873 K, and it decreased on CeO2 calcined at higher temperature than 873 K. This tendency can be explained by the low surface area of CeO2 at 1073 K (7 rrT/g). The maximum appeared at 1073 K over CeO2-ZrO2 (Ce/(Ce+Zr)=0.5). The catalyst had rather high surface area 34 m2/g even after the calcinations at 1073 K. This is because ZrO2 addition to CeO2 inhibited the sintering of CeO2. However, the sintering proceeded drastically by the calcination at higher temperature (1273 K). This effect is observed more significantly on CeO2-ZrO2 (Ce/(Ce+Zr)=0.33). It is very interesting that the CeO2-ZrO2 (Ce/(Ce+Zr)=0.33) catalyst calcined at higher temperature exhibited higher activity. It seems that the catalyst with lower surface area exhibited higher activity. In contrast to other catalysts, the BET surface area of CeO2-ZrO2 (Ce/(Ce+Zr)=0.33) was not so low (20 m2/g) even at 1273 K. calcination. It is possible that the surface area of the catalyst can decrease with higher calcination temperature (>1273 K) and it can make the activity lower. At present, the cause of this behavior is not elucidated. The tendency has been also observed in the DMC synthesis from CH3OH and CO2 over CeO2-ZrO2 solid solution [18]. In addition, diethylene glycol (DEG) can not be detected under these reaction conditions. It is concluded that CeO2-ZrC>2 (Ce/(Ce+Zr)=0.33) calcined at 1273 K showed the highest activity. Therefore, we investigated the reaction condition

167

dependence of EC formation using this catalyst. On the basis of the data shown above, we can compare the EC amount with the total amount of Ce and Zr in the catalyst. In 0.05-g CeO2-ZrO2 (Ce/(Ce+Zr)=0.33), the total amount of Ce and Zr ions is estimated to be about 0.36 nimol. which is smaller than EC formation (0.68 mmol). The turnover number (TON) is at least 1.9. and in fact it must be much higher because the number of surface active site is much smaller than the total amount of Ce and Zr ions. Figure 1. Dependence of the F:C formation over CeO2 ( A ), CeO2-ZrO2 (Ce/(Ce+Zr)=0.33) ) and CeO2-ZrO2 (Ce/(Ce+Zr)=0.5) ( I ) catalysts on calcination temperature. Reaction conditions: temperature 423 K, EG : CO2 : CH3CN = 100 : 200 : 120 mmol, time 2 h, 0.05 g-cat. The number near each mark describes BET surface area of the catalyst (m7g). The observed equilibrium level of EC amount =1.18 mmol.

Table 1 shows the dependence of the amount of EC and DEG formation on various reaction conditions over CeO2-ZrO2 (Ce/(Ce+Zr)=0.33) calcined at 1273 K. DEG is an expectable by-product in EG+CO2 reaction (Eq. (4)).

2

H2C — O H I H 2 C—OH

*"

OH

/

H2 \

L

/ C H2

\

H2 / ~ \ C OH H2

+

H,0

(4)

It is clear that the formation of DEG was not observed at lower reaction temperature (383 and 403 K) even after 8 h using 0.5-g cat. This indicates that the formation of EC is very selective in this temperature range. In contrast, at higher reaction temperature, the formation of DEG became more significant. It is found that EC formation is limited by the equilibrium because the EC amount on 0.5 g-cat was not ten times as large as that on 0.05 g-cat in 2-h reaction. This is also supported by the result that the EC amount in 8-h reaction was four times as large as that in 2-h reaction, which is clearly observed in the result at 403 K. Therefore, it is interpreted that the amount of EC over 0.05 g-cat corresponds to the formation rate and that over 0.5 g-cat is close to the equilibrium level. At 403 K, the observed equilibrium level of EC formation is determined to be 1.09 mmol because of no DEG formation. When DEG formation is observed, the EC amount decreased by H2O formed in Eq (4). This behaviour is observed in the results at higher reaction temperature (423-443 K). In the case of the result of 0.5 g-cat, DEG amount increased with the reaction time, on the other hand, EC amount decreased. This relation indicates that EC formation is very sensitive to this kind of by-reactions. In this reaction system, it is possible to form H2O and the oligomers of ether and carbonate. In the case that these reactions proceed, EC formation is strongly inhibited. In fact, we have continued GC analysis

168

for the long time (10 times of EC retention), but nothing was observed. Under the reaction conditions, the formation of oligomers can be neglected. From the results listed in Table 1. we can estimate the EG-based selectivity of EC formation (EC/(EC+2*DEG)). For example, at 423 K, the formation rate of EC can be estimated to be 6.8 mmol/h/g-cat (=0.68 mmol / (2 h*0.05 g)). And the formation rate of DEG can be estimate to be 0.005 mmol/h/g-cat (=2*0.01 mmol / (8h*0.5 g)). The selectivity of EG can be calculated to be higher than 99.9%. Furthermore, the selectivity is higher than at lower reaction temperature like at 383 and 403 K. This estimation demonstrated that the selectivity of EC formation from EG+CO2 reaction over CeO2-ZrC>2 (Ce/(Ce+Zr) = 0.33) calcined at 1273 K was very high. Table 1. Results of EG+CO2 reaction over CeO2-ZrO2 (Ce/(Ce+Zr)=0.33) calcined at 1273 K under various reaction conditions. Reaction temperature Catalyst weight Reaction time Formation amount/ mmol /h /K EC DEG /g 2 383 n. d. 0.05 0.16 383 2 n. d. 0.5 0.86 403 2 0.42 n. d. 0.05 403 0.96 0.5 2 n. d. 403 1.09 0.5 8 n. d. 423 n. d. 0.05 2 0.68 423 n. d. 0.5 1.18 2 423 0.01 0.5 1.12 8 443 n. d. 0.05 0.78 2 0.01 0.5 443 1.07 2 0.04 443 0.5 0.90 8 463 0.05 0.83 2 0.02 463 0.5 0.99 2 0.04 463 0.5 0.68 8 0.11 Reaction conditions: EC]:CO2:CH3CN=100 . in. d.; not detected by FID-GC. 4. REFERENCES [I] [2] [3] [4] [5] [6] [7] [8] [9] [10] [II] [12] [13] [14] [15] [16] [17] [18] [19] [20]

Y. Ono, Appl. Catal. A. 155(1997)133. M. Aresta and E. Quaranta. CHEMTECH, (1997) 32. A.-A. Shaikh and S. Sivaram. Chem. Rev., 96 (1996) 951. M.A. Pacheco and C.L. Marshall. Energy Fuels, 11 (1997) 2. U. Romano, R. Tesei, MM. Mauri and P. Rebora, Ind. Eng. Chem. Prod. Res. Dev., 19 (1980) 396. T. Matsuzaki and A. Nakamura, Catal. Surv. Jpn.. 1 (1997) 77. D. Delledonne, F. Rivetti and U. Romano. Appl. Catal. A, 221 (2001) 241. X.B.Lu. X.J. Feng and R. He. Appl. Catal. A. 234 (2002) 25. J. Gao and S.H. Zhong, PROGRESS IN CHEMISTRY, 14 (2002) 107. B.M. Bhanage, S. Fujita. Y. Ikushima and M. Arai, Appl. Catal. A. 219(2001) 259. M.S. Han, B.G. Lee, B.S. Ahn, K.Y. Park and S.I. Hong, React. Kinet. Catal. Lett., 73 (2001) 33. T. Tatsumi, Y. Watanabe and K.A. Koyano, Chem. Commun., (1996) 2281. K. Tomishige, T. Sakaihori, Y. Ikeda and K. Fujimoto, Catal. Lett.. 58 (1999) 225. K. Tomishige, Y. Ikeda, T. Sakaihori and K. Fujimoto. J. Catal., 192 (2000) 355. Y. Ikeda, T. Sakahori, K. Tomishige and K. Fujimoto, Catal. Lett.. 66 (2000) 59. Y. Ikeda, K. Fujimoto, K. Tomishige. J. Phys. Chem. B105 (2001) 10653. Y. Ikeda, Y. Furusawa, K. Tomishige, K. Fujimoto, ACS sym. Ser.. 809(2002)71. K. Tomishige, Y. Furusawa, Y. Ikeda, M. Asadullah, K. Fujimoto, Catal. Lett., 76 (2002) 71. K. Tomishige and K. Kunimori. Appl. Catal. A, 237(2002) 103. Jpn. Kokai Tokkyo Koho, 1999, 11-292538.


169

Effects of Palladium Particle Size in Hydrogenation of Carbon Dioxide to Methanol over Pd/ZnO Catalysts Deogseong Lee, Jun-Yearb Lee, and Jae Sung Lee Department of Chemical Engineering, Pohang University of Science and Technology, san 31 Hyoja-dong, Pohang 790-330, Korea. E-mail: [email protected]. FAX: +82-54-279-5528 Pd/ZnO catalysts with various Pd particle sizes were prepared by water-in-oil microemulsion method and investigated as catalysts for CO2 to methanol. The Pd particle size of the catalyst could be controlled by varying pH of microemulsions, regardless of Pd content. As pH varied from 1 to 7, the average Pd particle size increased from 2.1 to 10.1 nm. The distribution of metal particle sizes was considerably narrower compared with that of the catalysts obtained by the conventional precipitation method. The effects of Pd particle size were complicated by the formation of Pd-Zn alloy and coverage of Pd particles by ZnO, both of which were pronounced for small particles. Thus methanol selectivity increased with decreasing Pd particle size, showing a maximum at 3.4 nm, and then decreased for smaller particles. This selectivity pattern showed an excellent correspondence with the pattern of surface Pd fractions measured by EPMA. 1. INTRODUCTION Reducing the metal particle size in supported catalysts enlarges the metal surface area and enhances the catalytic activity. Variation of size between 1 nm to 10 nm also changes activity per surface atom(turnover rate) for structure-sensitive reactions[l]. However, the catalysts prepared by conventional impregnation and precipitation method have a broad size distribution, which often makes the size dependant catalytic phenomena equivocal. It has been demonstrated in recent years that the chemical techniques relating colloidal metal particles provided better control over the metal particle size[2,3]. Nearly monodispersed nanoparticles can be synthesized in microemulsions and their sizes are variable according to preparation conditions of the microemulsions. In this study, Pd/ZnO catalysts were prepared with various Pd particle sizes using water-inoil microemulsion and effects of the particle size were studied on the catalytic activity for the hydrogenation of CO2 to methanol.

170

2. EXPERIMENTAL The water-in-oil microemulsion systems were composed of polyoxyethyleneglycol-pnonylphenyl

ether(NP-5)/cyclohexane/aqueous

palladium

nitrate

solution.

The

microemulsions were prepared by injecting the aqueous palladium nitrate solution into the NP-5 solution in cyclohexane. Zinc oxide powder was added as a catalyst support and hydrazine was used for precipitation. Palladium particle size was controlled by varying pH of the preparation solution. The precipitate was separated from the solution, washed with ethanol and water, dried at 383 K, and then calcined in air at 773 K for 3 h. The catalysts prepared by this method will be denoted by (MEx) where x is pH of the solution. In addition, the conventional catalysts were prepared by using co-precipitation (COP) and depositionprecipitation (DEP). The targeted palladium loading for all prepared catalysts was 5 wt.%. The amount of palladium in the prepared catalysts was characterized by inductively coupled plasma mass spectrometry(ICP-MS). Surface area was measured by the N 2 BET method. The metal particle of the catalysts was characterized by X-ray diffraction(XRD) and transmission electron microscopy(TEM). To find distribution of metals on catalytic surface, EPMA analysis was performed. The catalytic hydrogenation of carbon dioxide was carried out using a fixed bed flow reactor(SUS-316, i.d.=ll mm). The standard reaction condition was: T=533 K, P=3.0 MPa with H2/CO2(3/1 vol.%), and space velocity=12000/h. Before each run, the catalyst was reduced in situ at 573 K for 2 h under H2 flow. The gaseous reactants and products were analyzed by a gas chromatograph with TCD. 3. RESULTS AND DISCUSSION The metal particle size in microemulsion depends upon many factors; micelle size, metal ion concentration in micelle, concentration of hydrazine, pH, etc[2,3]. In this study, pH was chosen as a size-controlling factor of Pd particles. As shown in the Table 1, Pd particle size was directly proportional to pH. As pH varied from 1 to 7, the average Pd particle size measured by XRD peak broadening varied from 2.1 to 10.1 nm. It was in a fairly good agreement with the results obtained by TEM photographs of the catalysts. Pd particle sizes of the MEx catalysts were much smaller than those of COP and DEP catalysts. More importantly, the distribution of metal particle sizes of MEx catalysts observed by TEM was considerably narrower compared with that obtained for COP and DEP catalysts. MEx catalysts also showed good performance in the hydrogenation of carbon dioxide. Of particular interest was variation of methanol selectivity with Pd particle sizes. The selectivity of methanol increased as Pd particle size decreased. This trend continued until Pd

171

Table 1 Characteristics of catalysts Catalyst

Pd

Surface

Particle size (nm)

CO2 hydrogenation

content

area

XRD

Conv. (%)

(wt%)

TEM

Sel. (%)

2

ME1

4.9

(m /g) 15.4

ME2

4.9

15.5

3.4

2.5

8.4

62.3

ME4

4.8

15.1

4.2

6.4

8.2

47.5

2.1

1.8

4.1

10.7

ME6

4.9

15.2

6.8

7.2

7.8

35.5

ME7

5.0

15.5

10.4

10.2

10.5

27.1

DEP

4.4

12.8

-

35.0

8.3

50.1

COP

3.2

24.1

-

18.5

5.1

34.7

particle diameter reached 3.4 nm, and then the selectivity started to decreased. Overall CO2 conversion did not show such a systematic trend. Fig.l shows the XRD patterns of MEx catalysts. The XRD patterns indicated that the intensity of Pd(lll) peak decreased while Pd-Zn alloy peak increased with decrease in solution pH. In hydrogenation of CO2 to methanol, Pd-Zn alloy is not the active phase[4]. Thus, when the Pd particle size becomes small, the formation of the Pd-Zn alloy is easier and wastes Pd that could be otherwise used to form the active phase. The synthesis of methanol over a Pd/ZnO catalysts is best provided by a geometry involving small ZnOx islands decorating Pd surfaces[4]. Partly-reduced ZnOx migrates onto the surface of Pd particle during reduction procedure. There exists an optimum level of

Fig. 1. XRD patterns of the catalysts (

Pd(l 11),

Pd-Zn alloy)

Fig.2. A correlation between the pH in microemulsion and the surface Pd fractions measured by EPMA or selectivity of methanol

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ZnOx decoration to give maximum methanol yield[4]. The formation of Pd-Zn alloy leads to reduction of the decoration and a loss of methanol formation. Fig.2 shows the atomic fraction of Pd on catalyst surface measured by EPMA. The amount of Pd on the catalyst surface shows a remarkable correspondence with methanol selectivity of the catalysts. The trend of Pd-Zn alloy fraction shows a smooth decrease with Pd particle sizes. This trend together with the reduction in surface Pd fraction with Pd sizes accounts for the decrease in methanol selectivity for ME2 - ME7, but not the sudden decrease for ME1. There must be another important factor determining methanol selectivity that is closely related to surface Pd fractions determined by EPMA. This factor could be the surface coverage of Pd by ZnOx which might increase for smaller Pd particles. Thus, in addition to Pd-Zn alloy formation, some small Pd particles are heavily covered by ZnOx, and both effects lead to decreased activity and methanol selectivity in CO2 hydrogenation. 4. CONCLUSION We were able to prepare Pd/ZnO catalysts with various palladium particle sizes by the microemulsion method by controlling pH of the preparation solution. The methanol selectivity of the catalyst was directly influenced by the Pd particle size in the catalysts. Selectivity increased in general when the particle size became smaller. However, when it became too small, the formation of Pd-Zn alloy was pronounced and some Pd particles were covered by ZnO, which led to reduction of the active Pd on the catalyst surface and loss of selectivity. Therefore, there was an optimum palladium particle size around 3.4 nm that exhibited the highest activity for the hydrogenation of CO2 to methanol. ACKNOWLEDGMENT This work has been supported by National R&D Project for Nano Science & Technology, Research Center for Energy Conversion and Storage and Brain Korea-21 Project. REFERENCES 1. M. Boudart and G.Djega-Maridassou, Kinetics of Heterogeneous Catalytic Reactions, Priceton University Press, Princeton, New Jersey, 1984. 2. M. Kishida, K. Umakoshi, J.-i. Ishiyama, H. Nagata, and K. Wakabayashi, Catalysis Today 29, 355-359(1996). 3. H. H. Ingelsten, J.-C. Beziat, K. Bergkvist, A. Palmqvist, M. Skoglundh, Q. Hu, L. K. L. Falk, and K. Holmberg, Langmuir 18, 1811-1818 (2002). 4. C.-H. Kim, J. S. Lee, and D. L. Trimm, Topics in Catalysis 22, 319-324 (2003).


173

Ga, Mn and Mg promoted copper/zinc/zirconia - catalysts for hydrogenation of carbon dioxide to methanol M. Lachowska, J. Skrzypek Institute of Chemical Engineering, Polish Academy of Sciences, PL 44-100 Gliwice, ul. Baltycka 5, Poland e-mail: m.lach(a>iich. gliwice.pl Ga, Mn and Mg - promoted copper/zinc/zirconia catalysts have been investigated. The catalysts were prepared using the method of thermal decomposition of citrates. Four catalysts of the same copper/zinc contents and different contents of zirconium and gallium, manganese and magnesia were studied. The comparison Ga, Mn, Mg doped catalysts with Cu/Zn/Zr catalyst has been given. Doping of Ga to Cu/Zn/Zr catalyst increases the rate of methanol production. Ga promoted copper/zinc/zirconia catalyst exhibits remarkable high stability and high selectivity. The crude methanol does not contain practically any by-products other than water. 1. INTRODUCTION The growing demand for methanol and associated requirement for improved reaction efficiency has given impetus to the search for ways to make further improvements to these catalysts. The effective utilisation of CO2 is of a great significance from the point of view of global environmental protection. However, the Cu/ZnO/AkOs catalyst having a considerably high activity for methanol synthesis from syngas, exhibits rather poor activity and stability for hydrogenation of CO2. Therefore, many efforts have recently been made to develop new catalysts for methanol synthesis from CO2+H2 [1]. Literature reports concerning the advantageous influence of zirconium have been appearing since several years [2, 3], there also appeared attempts of modification copper containing catalysts with metals doping [4, 5]. 2. EXPERIMETAL Ga, Mn, Mg doped copper/zinc/zirconia - catalysts have been investigated. The catalysts have been prepared in our laboratory using the method of decomposition organic complexes of citric acid containing metallic components of catalyst. The require amounts of nitrates of Cu, Zn, Zr, Ga, Mn, Mg were added in small portion under stirring to corresponding 2 M solution of citric acid. The solution obtained was evaporated in a revolving flask in vacuum overnight at 373 K, dried at 403 K for 8 h and calcinated on air for 1 h subsequently 423, 473, 523, 573 and 623 K. The resultant powder were palletising, crushing and sizing to obtain 0.8-1 mm graining. Experiments were carried out in a fixed bed continuous flow pressure reactor. The apparatus scheme is shown in the Fig. 1. The catalyst (2 g) was placed in stainless tube reactor.

174

The feed gas was deoxidized with BTS deoxidizer and dehydrate with the molecular sieves 5 A. The catalyst was reduced in a stream of diluted hydrogen (10% H2 in N2) at 473 K and atmospheric pressure and stabilized in the mixture of reactants by raising the temperature by steps of 30° from 473 to 623 K every 2 h and pressure 4 MPa. The parameters of activity testing were following: Pressure : 4 and 8 MPa Temperature range : 453 - 493 K GHSV : 1 600 - 10 000 1/h The synthesis gas feed composition : H2-66, CO2-22, N2-12 vol.%. The composition of the catalysts we tested is presented in the Table 1. Table 1 Composition of the studied catalysts Catalyst CuO ZnO ZrO2 MeO (wt%) (wt%) (wt%) (wt%) No 1 Cu/Zn/Zr 62.4 25.0 12.6 3.9 No 2 - Ga-Cu/Zn/Zr 4.5 65.3 26.3 3.9 4.5 No 3 - Mn-Cu/Zn/Zr 65.3 26.3 4.5 3.9 No 4 - Mg-Cu/Zn/Zr 65.3 26.3

Figure 1. Apparatus scheme 3. RESULTS AND DISCUSSION The main reactions that occur in carbon dioxide hydrogenation are following: CO2 + 3H2 CH3OH + H2O - a direct synthesis of methanol from CO2 and H2. CO2 + H2 CO + H2O - the reverse water gas shift reaction.

(1) (2)

175 The results of experiments are shown at the Fig. 2-5 and Table 2. All studied catalysts show good activity and selectivity. The only reaction products were methanol, water and carbon monoxide, no other products, such as higher alcohols, ethers, even methane were detected. Table 2 Methanol yields and selectivity obtained for the examined catalysts Catalyst WMeoH(for493 K) S (for 493 K) S max (for453K) Cu/Zn/Zr No 1 No 2 - Ga-Cu/Zn/Zr No 3 _ Mn-Cu/Zn/Zr No 4 - Mg-Cu/Zn/Zr

83 184 92 87

60 69 63 63

83 87 100 99

Selectivity was calculated as the ratio of CO2 that reacted to MeOH, to whole reacted CO2. All investigated catalysts show very high selectivity. The most interesting from the application point of view is Ga-Cu/Zn/Zr catalyst (No 2) that shows the highest methanol yields and over 85% maximum value of selectivity, Mn-Cu/Zn/Zr catalyst (No 3) exhibits the best maximum value of selectivity (100%), but about 50% lower methanol yields.

Figure 2. The temperature dependence of methanol yield for GHSV = 3400 1/h, P = 8 MPa for the catalysts: No 1 ; No 2 (O); No 3 (O); No 4 (A)

Figure 3. The temperature dependence of methanol selectivity for GHSV = 3400 1/h, P = 8 MPa for the catalysts: No 1 ; No 2 (O); No 3 (O); No 4 (A)

176

Figure 4. The GHSV dependence of methanol yield for T = 473 K, P = 8 MPa for the catalysts: No 1 ; No 2 (O); No 4 (A)

Figure 5. The GHSV dependence of methanol selectivity for T = 473 K, P = 8 MPa for the catalysts: No 1 ; No 2 (O);No4(A)

4. CONCLUSIONS Ga, Mn, Mg promoted copper/zinc/zirconia catalysts, prepared using the method of thermal decomposition of citrates, show excellent selectivity. The only liquid products of the reaction are methanol and water, thus it is easy to distil this raw methanol into an extremely pure finished product. Maximum methanol yields of 184 g-kgcat'^h"1 and very good (about 70%) selectivity were obtained at temperature 493 K for Ga-Cu/Zn/Zr catalyst. It is reasonable to continue more detail investigations of selected Ga-Cu/Zn/Zr catalyst. ACKNOWLEDGMENT "The work was financed by the State Committee for Scientific Research, KBN under project No7T09C010 21p01" REFERENCES 1. L. Jinyao, S. Jiangliu, H. Dehua, Z. Qijian, W. Xiaohui, L. Yu and Z. Qiming, Appl. Catal. A:General218(2001)113. 2. S.L. Jae, I.M. Kwang, H.L. Si, Y.L. Sang and G.K. Young, Cat. Lett. 34 (1995) 93. 3. S.-K. Ihm, Y.-K. Park, J.-K. Jeon, K.-C. Park and D.-K. Lee, Proceedings of the Fourtht Intern. Conf. on Carbon Dioxide Utilization, Kyoto, Japan P-052 (1997). 4. K. Ushikoshi, K. Mori, T. Watanabe, M. Takeuchi, Proceedings of the Fourtht Intern. Conf. on Carbon Dioxide Utilization, Kyoto, Japan O-40 (1997). 5. J.S. Lee, K.I. Moon, S. H. Lee, S.Y. Lee, Y.G. Kim, Cat. Lett. 34 (1995) 93.


177

Catalytic Hydrogenation of Carbon Dioxide to Light Olefins in a Fluidized Bed Reactor Jun-Sik Kim, a * Sang-Bong Lee,a Myoung-Jae Choi," Yong Kang,b and Kyu-Wan Leec "Environment and Resources Group, KRICT, P.O. Box 107, Daejeon 305-600, Korea b Dept. of Chemical Engineering, Chungnam National University, Daejeon 305-764, Korea c Yanbian University of Science & Technology (YUST), Yanji, Jilin, China Catalytic hydrogenation of CO2 to light olefins was studied in a fluidized bed reactor and compared with that in a fixed bed at identical experimental conditions. It has been found that performance of fluidized bed reactor is superior to the fixed bed reactor with respect to CO2 conversion and light olefin selectivity due to relatively homogeneous and high temperature reaction. 1. INRODUCTION CO2 utilization is one of the advanced research areas on account of global warming cause by annual rise in the atmospheric CO2. It is also low price and abundantly available energy source in nature. Hydrogenation of CO2 to value added chemicals has been remained to be attraction of chemist and engineers. Many catalysts such as K, V, Cr, Mn and Zn promoted iron oxide have been studied, however Fe-K catalyst showed prominent results for CO2 conversion and olefins selectivity [1, 2]. Fluidized beds can be utilized efficiently for the multiphase reaction, because it can realize higher heat and mass transfer rate due to the effective contacting and mixing in comparison with any other contacting mode [3], It is reported that Fischer-Tropsch synthesis can be effectively carried out in fluidized bed reactors in view of syn-gas conversion and selectivity to valuable chemicals such as light olefins and oxygenates. Economic feasibility is another added advantage [4]. In the present paper, we are reporting our results of CO2 hydrogenation over Fc-Cu-K-AI catalyst. The catalytic performance of fluidized bed reactor is compared with that of fixed bed reactor. The effects of process parameters (temperature, pressure, space velocity, H2/CO2 ratio and gas velocity) on CO 2 hydrogenation activity and olefin selectivity in the product are presented using a fluidized bed reactor. 2. EXPERIMENTAL The catalyst Fe-Cu-K-Al was synthesized from their respective nitrate salts by precipitation and incipient wetness impregnation [2]. The physical and chemical properties of the catalysts are given in Table 1. Experiments were carried out in a stainless steel column (0.052 m ID x 1.5 m high) as shown in Fig. 1. The feed gases (CO2, H2) were used from cylinders. Catalyst particles (dp=75-90/zm) were supported on a perforated plate containing 19 evenly spaced holes of 1.5mm diameter, which served as a mixed gas distributor. Three zone heaters were installed each at calming section, main column and freeboard region. The temperature was controlled by PID controller. Reaction pressure was maintained by back pressure regulator and measured by pressure sensors. The products passing through a heat exchanger were condensed in the cold product separator. A digital bubble flow meter was used to measure the *To whom correspondence should be addressed (e-mail: [email protected]).

178

exit gas flow rate. The analysis of the reaction products (gas and condensed liquid) was carried out employing gas chromatography. Table 1. Physical and chemical properties of Fe-Cu-K-Al catalyst BET surface area CO2 chemisorption Density, p Fe-Cu-K-Al [pnol/g] [wt%] [kg/m3] [m2/g] 100:6.6:6:15.7

160.2

265.1

3,250

Umf [m/s] 0.012

3. RESULTS & DISCUSSION The effect of gas velocity on CO2 hydrogenation have been considered as important operating variable in a fluidized bed reactor. The effects of gas velocity on the pressure drop in the fluidized bed, which is used to find a minimum fluidization velocity (Umf). From the intersection of pressure drop, Umf of the Fe-Cu-K-Al catalyst is determined at 0.012 m/s (Table 1.). It is found that the space-time-yield attained maximum values with increasing gas velocity in the fluidized bed reactor because the increase of gas velocity leads to the increase of intensity of contacting and mixing between the gas reactant and the catalyst (Fig. 2). It is interesting to note that a similar trend can be also observed in bubble column slurry reactor using Fe-K catalyst for F-T synthesis [6]. This can be explained by the flow regime transition from homogeneous fluidization to heterogeneous fluidization states. It has been pointed out that the unified approach relies on analogy in the hydrodynamic behavior between the gassolid and gas-slurry systems using fine particles [7]. The reaction temperature, pressure, space velocity and H2/CO2 ratio have profound effect on CO2 hydrogenation. It has been found that CO2conversion increased but the CO selectivity

Fig. 1. Experimental apparatus Fig. 2. Effects of gas velocity on the 1. gas mixer 5. cyclone STY(T=300°C, P=lMPa 2. wind box 6. Cold separator 3. distributor 7. BPR decreased, with increasing reaction temperature, pressure and H2/CO2 ratio in the reactor,

179

Fig. 4. Axial temperature distribution in a Fig. 3. Effects of reaction temperature on fixed and fluidized bed. CO2 Conv., CO and CH4 selectivity (P=lMPa, SV=2000ml/gca,.hr). whereas CO2 conversion decreased with increasing space velocity [2, 3J. Among these operating variables, reaction temperature is considered to be most important parameter to promote CO2 hydrogenation because of the CO2 equilibrium conversion. Effects of temperature on the CO2 conversion, CO and CH4 selectivity are shown in Fig. 3. It is clear that CO2 conversion increases and CO selectivity decreases with increasing temperature. This figure also shows that the selectivity of CO is highest at 250°C. but it decreases with increasing temperature upto 325°C and became stable with further increase temperature. Fluidized beds are very attractive for CO2 hydrogenation to hydrocarbons (MFT, modified Fischer-Tropsch) due to their excellent heat transfer rate. MFT is a highly exothermic reaction as well as FT synthesis. The axial temperature distributions in the fixed and fluidized beds are explained in Fig. 4. There is a temperature difference (AT) of 2°C in the expanded bed region of the fluidized bed whereas, in the case of fixed bed reactors there is a 'hot spot' occurs at the inlet of gaseous reactants. The CO 2 conversion and products selectivity are shown in Table 2. The CO2 conversion and olefin selectivity have exhibited higher values in the fluidized bed reactor than those in the fixed bed reactor. It is interesting to note that the fluidized bed reactor exhibited higher oxygenates selectivity than the fixed bed reactor. It can be attributed to relatively short residence time and higher temperature causing no further hydrogenation of oxygenates. Among oxygenates, ethanol was predominantly formed (45wt%) in the fluidized bed reactor. When fluidized bed is used for CO2 hydrogenation, the product selectivities are found comparable to those of F-T synthesis [4]. It seems that in the fluidized bed, light hydrocarbons are favorable whereas relatively heavy hydrocarbons are formed in a fixed bed reactor. Molecular weight distribution of hydrocarbons as a function of the carbon number is shown in Fig. 5. From the weight of liquid hydrocarbons, chain growth probability (a) was calculated 0.85 and 0.71 in the fixed and fluidized bed, respectively. The olefin-(l) content in linear hydrocarbons is found to be much higher in the fluidized bed than that in the fixed bed is shown in Fig. 6. The influences of reactor type on product selectivity could be due to the more homogeneous contacting and mixing between the feed gas phase and the solid catalyst particles in fluidized bed reactors at relatively high temperature. Table 2. Carbon dioxide conversion and product selectivity (P=lMPa, SV=2000ml/gcat.hr) Reactor

I

CO^

CO

Oxy!

HCs

Hydrocarbon distribution [C-mol%]

180

Type

Conv. [C-mol%]

Fixed Bed (300°C) Fluidized Bed (325°C)

35.6 48.9

Sel. Sel. Sel. c, c2= c 2 c3= c 3 C4= C4 >C 5 [C-mol%] [C-mol% ] [C-mol% ] Sel. Sel. Sel. Sel. Sel. Sel. Sel. Sel. 9.6 10.9 5.8 3.2 11.6 2.9 9.3 2.4 53.9 3.5 86.9 8.8

28.8

Fig. 5. Molecular weight distribution of hydrocarbons as a function of the carbon number (P=lMPa, SV=2000ml/gcat.hr).

62.4

Fig.

12.2 7.8

6.

1.6

14.1 0.2

8.2

0.8

55.1

Olefin-(l) content in linear hydrocarbons as a function of the carbon number (P=lMPa, SV=2000 ml/gcat.hr).

4. CONCLUSIONS The hydrogenation of CO2 has been investigated in a fluidized bed reactor by using Fe-CuK-Al catalyst as a fluidized solid phase. The space-time-yield showed maximum value at Uc=0.02m/s with the increase of gas velocity. The CO2 conversion increased with rise in reaction temperature whereas the reverse trend was observed for CO selectivity. Fluidized bed reactor showed better catalytic performance than that of fixed bed reactor in view of CO2 conversion and selectivity to light olefins and oxygenates. REFERENCES 1. X. Xiaoding and J. A. Moulijin, Energy & Fuels, 10 (1996) 305. 2. S. R. Yan, K. W. Jun, J. S. Hong, M. J. Choi and K. W. Lee, Appl. Catal. A, 16 (1999) 357. 3. J. S. Kim, H. K. Kim, S. B. Lee, M. J. Choi, K. W. Lee and Y. Kang, Korean J. Chem. Eng., 18(2001)463. 4. M. E. Dry, Appl. Catal. A, 138 (1996) 319. 5. P. H. Choi, K. W. Jun, S. J. Lee, M. J. Choi and K. W. Lee, Catal. Letters, 40 (1996) 115. 6. E. Sanders, S. Ledakowicz, and W. D. Deckwer, Can. J. Chem. Eng., 64 (1986) 133. 7. R. Krishna, J. Ellenberger and S. T. Sie, Chem. Eng. Sci., 51 (1996) 2041.


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Synthesis of Dimethyl Carbonate from CH3OH and CO2 with Ce l x Zr x O 2 Catalysts G B. Hana, N. K. Park3, J. H. Juna, W. C. Changb, B. G Leec, B. S. Ahnc, S. O. Ryua and T. J. Leea* National Research Laboratory, School of Chemical Engineering & Technology, Yeungnam University, Gyeongsan, 712-749, Korea b

Pung Nam Semicon Tech. Co., Chilgok, 718-922, Korea

'Environmental & Process Technology Division, KIST, Seoul, 136-791, Korea

The direct synthesis of dimethyl carbonate (DMC) from methanol and CO2 is considered as the most attractive process among various alternative methods to produce DMC because its simplicity, lower toxicity and utilization of low cost CO2. The selective synthesis of DMC from methanol and CO2 was performed effectively with Cei-xZrxO2 catalysts prepared by the citric complexation technique. Formation of solid-solution of Cei_xZrxO2 was confirmed in XRD analysis. Rate formation of DMC was strongly dependent on the content of Zr in CeixZrxO2 and Ceo.sZro.2O2 catalyst showed the best reactivity in this study. It was concluded that the reactivity of the catalyst was related to the acid-base sites on the Cei_xZrxO2 surface from the results of the TPD of TEA and CO2. The mechanism of DMC formation was proposed from the experimental results. 1. INTRODUCTION Dimethyl carbonate (DMC) has been used as an intermediate to synthesize polycarbonate, polyurethane, medicine and agricultural chemicals. Commercial processes of DMC synthesis utilize the reaction between methanol and phosgene and lead high yields in DMC synthesis. However, those processes could cause serious problems because of their complex multi-step and of the use of the toxic phosgene and carbon monoxide in the processes. In these days the direct synthesis of DMC from CH3OH and low-cost CO2, which is known as a global warming gas, has been introduced to solve those problems. It was reported that MgO, zeolite, Y-AI2O3, hydrotalcite and basic metal oxides were used by many research groups for the DMC synthesis [1]. In this study, however, Cei.xZrxO2 solid solution catalysts were prepared by using CeO2 and ZrO2 for the possible improvement of activity and selectivity. Effects of the mixed oxide catalysts on the reaction of DMC synthesis were also investigated. 2. EXPERIMENTAL The Cei_xZrxO2 catalysts were formulated by a citric complexation technique with cerium nitrate hexahydrate [Ce(NO3)3-6H2O, Aldrich Chem.] and zirconium nitrate oxide dehydrate

182

[ZrO(NO3)2-2H2O, Kanto Chem.] as precursors of catalysts. These catalysts were prepared in the several different molar ratios of Ce to Zr(x=0, 0.1, 0.2, 0.5, 0.8, 1.0). The crystal structures of catalysts were analyzed by XRD spectra (D/MAX-2500, Rigaku). The surface area of the catalyst was measured by BET (ASAP 2010, Micrometrics) method. For the DMC synthesis 247 mmol of methanol, 114 mmol of CO2 and 0.05~0.35g of catalyst were feed into the stainless steel autoclave. Then, the reactor was pressurized with CO2 to 65 atm and it was heated up to 150°C. The reaction time was 2 hours. Methanol (Kanto Chemical, 99.8%) and CO2 (domestics, 99.9%) were used without further purification. The products were analyzed by gas chromatogragh (Shimadzu 14-B, FID and TCD) at room temperature.

3. RESULTS AND DISCUSSION Cei_xZrxO2 catalysts were calcined at 600°C and then those were analyzed by XRD to confirm the appropriate formulation of Cei.xZrxO2 catalysts. It is known that CeO2 has the cubic structure and the excellent oxygen storage capacity and that ZrO2 has the tetragonal structure and the excellent oxygen mobility [2]. The peaks corresponding to the crystal face in XRD were gradually shifted and broadened with increase of Zr content (Fig. 1). It is indicated that the cubic structure of pure CeO2 was converted to the tetragonal structure of ZrO2 as ZrO2 was added into it. There might be the optimimum value for x in Cei_xZrxO2 catalysts to maximize both oxygen storage capacity and oxygen mobility.

Fig. 1 XRD patterns of Ce].xZrxO2 . , , j . ,nnor> catalysts calcined at 600 G

Fig. 2 The dependence or Zr content on DMC formation and surface area of Cei.

xZrxO2 catalyst calcined at 600°C To investigate the effect of the molar ratio of Ce to Zr on the DMC synthesis, the values of x were increased from 0 to 1.0. As the amount of ZrO2 increased, BET surface area of the catalyst also increased except the pure ZrO2. However, it was not observed the direct relationship between the DMC formation and the BET surface area of catalyst. Trovarelli et al. reported that Ce,.xZrxO2 with x < 0.5 had the higher oxygen storage capacity and the better oxygen mobility than those of pure CeO2 [3]. The activity of Cei.xZrxO2 catalyst might be higher than the others because the mobility of the lattice oxygen was improved by the addition of Zr to CeO2 in the range of x < 0.5 (Fig. 2). This result is consistent with the argument of Trovarelli et al.. It is supposed that the oxygen mobility was proportional to the amounts of basic sites on the surface of catalysts. The more amounts of basic sites, the higher reactivity of catalyst. It is indicated that the reactivity was improved by the increase of oxygen mobility. Tomishige et al. reported that Ceo.2Zro.sO2 catalyst having the highest surface area showed the higher activity than others for the DMC synthesis [4, 5]. However, Ce0 gZr02O2 catalyst showed the highest activity in this work. It is believed that the different catalytic properties of

183

two studies were caused by the reason of that the calcination temperatures and the preparation methods of catalyst were different from each other. The influential factors of the Ceo.sZro.2O2 formula for the DMC formulation were also investigated in this experiment. The optimized amount of catalyst and reaction time were in the range of 0.1~0.15g and about 2 hours, respectively. Dependence of DMC formation on CO2 pressure was investigated for the Ceo.sZro.2O2 catalyst calcined at 400°C. Since methoxy carbonate as the reaction intermediate was easily formed upon CO2 insertion, more amount of DMC was formulated with the CO2 pressure increase as shown in Fig.3. The dependence of the concentration of methanol on the DMC formation was also investigated. Fig. 4 shows that the amount of methanol was proportional to the yield of DMC. Considering the dependence of the reaction rate on the concentrations of the reactants, the partial reaction orders seemed to be nearly a 1st order for CO2 and a half order for methanol. The stronger adsorption strength of methanol than that of CO2 was observed in the result of the experiment. While CO2 was adsorbed without any dissociation, methanol was adsorbed after dissociation due to the different adsorption strength of each other.

Fig. 3 Dependence o f reaction rate on CO2 pressure. Reaction conditions: 150D for 2 hrs. Amount: CH3OH=247mmol, Ceo.sZro.2O2 calcined at 400°C=0.1g

Fig. 4 Dependence of reaction rate on the concentration. Reaction conditions: 150°C, CO2 initial pressure=30atm. Amount: Ceo.sZro.2O2 calcined at 400 °C=0.1g

M e 0 H

Fig. 5 Profile of CO2-TPD on Ce,.xZrxO2 Fig. 6 Profile of TEA-TPD on Cei.xZrxO2 catalvsts catalvsts Table 1 Adsorption amoi nt of CO-i on Cei.vZrÔi TPD(temperature programmed CO desorption) experiment was conducted in Catalysts (nmol/O.SgCat.) order to find the relationship between the CeO 2 3.13 reactivity of Cei_xZrxO2 catalysts and the Ce 0 S Zr 0 5 O 2 12.25 characteristics of acidic-basic sites. Fig. 5 Ce08Zr02O2 20.61

184

shows the profile of TPD of CO2 adsorbed on catalysts. The adsorption amount of CO2 on Cei_xZrxO2 was listed in Table 1. The Ceo.8Zro.2O2 formula showed the highest adsorption and strength among Cei.xZrxO2 catalysts as the result of TPD experiment. Tomishige et al. reported that the yield of DMC increased as both the amount and the strength of the basic site on the catalyst increased [6]. This result is well agreed with their result. TPD experiment was conducted by a cahn balance after dipping each catalyst in triethylamine(TEA) for overnight in order to investigate amount and distribution of acid site on catalyst. Fig. 6 shows the result of TEA-TPD on catalyst. Even though the acid strength for each catalyst was similar, the amount of acid site was different from each other. Since acid site of Ceo.8Zro.2O2 was more than any other Cei_xZrxO2 catalyst just as the characteristics of base, it was indicated that the formation of DMC would be intimately related to the amount and strength of the neighboring acid and base sites. Active sites consist of M=O as a basic site and of Mn+-OH (below Mn+) as an acidic site in the mixed oxide catalyst. Pure CeC>2, which has a cubic structure, was composed of Ce=O as a basic site and of Ce4+ as an acidic site. As adding ZrC"2 to CeC>2, Ce4+ on the surface of mixed oxide catalyst could be converted into Ce=O, subsequently, oxygen mobility increased. Thus, the synthesis of DMC could be easily progressed due to the synergetic effect among Zr4+, Ce4+, and the increased basic site. Reaction mechanism of DMC synthesis was proposed in this study as follows: Initially, methanol was CH3OH + 2 ce=o -> o=ce-ocH3 ce-OH dissociated into methoxy radical and tnen c co2 + o=ce-ocHs - o=ce-oco-ocH3 ° 2 was activated at the base site(Ce=O). On the CH3OH + 2Ce"(ZO-Ce--CH 3 (Zr-CH 3 )

o=ce-oco-ocH3 + ce4*-CH3(Zr4*-CH3)

+

Ce".OH(Zr4*-OH)

other

4

* (CH^co + ce=o ce^zr *)

side;

methanol

Was

activated to methyl radical at the acid sites(Ce and Zr ).

4. CONCLUSION The catalysts for the direct synthesis of DMC are required to have both the base and the acid sites. In this study Cei-xZrxO2 composite oxides was proposed as the best catalyst for the direct synthesis of DMC. The optimized formula of the catalyst was obtained by controlling the ratio of Ce to Zr in order to increase the acid-base sites and the intensity of catalyst. Since the partial reaction orders were a 1st order with respect to the CO2 pressure and a half order with respect to the CH3OH concentration, the adsorption intensity of CH3OH was stronger than that of CO2 over CeCh catalyst. While CO2 was not dissociated, CH30H was dissociated into methoxy and methyl radicals. From the experimental results, a reaction mechanism, in which the neighboring acid-base sites exist together, was also suggested in this study for the description of those reactions.

REFERENCES 1. B.M.Bhanage,S.I.Fujita,Y.Ikushima,and M. Arai, Appl. Catal. A: General, 219 (2001) 259. 2. J. H. Lee, T. J. Lee, and K. S. Kim, 2nd Asia-Pacific Congress on Catalysis, (2000) PB28. 3. A. Trovarelli, F. Zamer, J. Llorca, C. D. Leitenburg, G. Dolcetti, and J. T. Kiss, J. Catal., 169(1997)490. 4. K. Tomishige, and K. Kunimori, Appl. Catal. A: General, 237 (2002) 103. 5. K. Tomishige, Y. Furusawa, Y. Ikeda, M. Asadullah, and K. Fujimoto, Catal. Lett., 76 (2001)71. 6. K. Tomishige, Y. Ikeda, T. Sakaihori, and K. Fujimoto, J. Catal., 192 (2000) 355.


185

The effect of catalyst pore structure change into the selectivity and conversion of CO 2 hydrogenation over Fe-K/Y-Al2O3 Sung-Chul Lee*, Jae-Hun Jang^ Byung-Yong Lee*, Taewhan Jeong* and Suk-Jin Choungt * College of Environment and Applied Chemistry, Kyung Hee University, Gyeonggi 449-701, Korea *Jinro Distillers Co. LTD., Gyeonggi 425-866, Korea E-mail: [email protected] The chemical fixation of CO2 has been recently paid an attention as one of the effective CO2 utilizing technologies. In recent years, Fe-K/y-AkCh catalysts for CO2 hydrogenation showed remarkably high activity and selectivity into long-chain hydrocarbons and light olefins. The study is focused on investigating into the effects of pore structure change of catalyst in CO2 hydrogenation. The pore structure is controlled by polyethylene glycol (PEG) that is added in forming of catalysts. As these results, it was found out that Fe-KJy-AhO} with PEG had excellent catalytic activity of CO2 hydrogenation. Fe-K/y-Al2O3 with PEG exhibited low methane (8.3 mol%), C2-C4 selectivity (13.4mol%) than Fe-K/y-Al2O3 without PEG. It is expected that the pore, which is generated by PEG, affect on the selectivity of high molecular weight hydrocarbons. In the case of Fe-K/y-AÔi with PEG, the interstitial space between pellets is filled with high molecular weight hydrocarbons. Therefore, the intraparticle olefin diffusivity decreases with increasing molecular size. 1. INTRODUCTION In recent years, global warming caused by CO2 emission had been recognized as an urgent problem. As a result, the chemical fixation of CO2 gained attention as one of the effective carbon dioxide utilizing technologies, and fixation by chemical method was still a problem to be solved. The main problem CO2 hydrogenation was that its products were common and were low valuable materials, such as methanol and methane [1, 2, 3]. However, CO2 was valuable as an infinite carbon source. Trovarelli et al. suggested that the hydrogenation of CO2 to hydrocarbons proceeded through the formation of CO as the intermediate [4]. According to other authors [5, 6], CO2 hydrogenation on metal catalysts occurred through a consecutive mechanism in which CO2 was first converted to CO by the reverse water gas shift (RWGS) reaction, and then CO was hydrogenated to hydrocarbons. Therefore, the chain growth mechanism of hydrocarbon synthesis from carbon dioxide was similar to that of the FT reaction. Iglesia et al. proposed [7, 8] that increased intraparticle residence time and olefin concentration gradients reflect diffusional restrictions that lower the rate of removal of reactive olefins from catalyst pores. Secondary reactions altered FT synthesis selectivity by chemical transformations of these primary products. In many cases, high CO and water concentrations during FT synthesis inhibited secondary reactions of hydrocarbons. The reaction schemes in CO2 hydrogenation are generally consistent with the phenomena observed in CO hydrogenation [5].

186

This study is focused on developing a better understanding of the effect of organic binders on pore structure of iron-based catalysts in CO2 hydrogenation. Especially, the purpose of this study is to see the role of pore structure, which is produced by organic binders, on the activity and selectivity in CO2 hydrogenation. 2. EXPERIMENTAL Fe-K/y-Al2O3 catalysts are prepared by the impregnation of y-Al2O3 (A, B and C type) with aqueous solutions of Fe(NO3)3-9H2O and K2CO3 with or without PEG. The nominal catalyst composition was 1.00Fe/0.35K/5.00Al2O3. The impregnated catalysts were homogeneously mixed, and a small amount of water was added. Pellets of 3 mm in length were extruded through a 1.5 mm-diameter die. These pellets were dried at 393K for 12 h and calcined at 773K for 24 h in air. CO2 hydrogenation was carried out in a bench scale fixed bed reactor (1.6 cm-IDx60 cm-High). The reaction and internal standard gases (CO2, H2, N2, He) were taken from cylinders and their flow rates were controlled by MFC (mass flow controller, Brooks Co.). Reaction temperature was controlled at 573K and reaction pressure was maintained at lOatm by BPR (back pressure regulator, Tescom Co.). 21.0 g of catalysts were filled up and the flow rate of the mixed gas was 2,000 ml/g-cat.h at STP. The composition of the gas was H2/CO2 = 3/1. The liquid products were separated from gas products in the gas-liquid separator and condenser. The exit gas flow rate was measured by a digital bubble flow meter to evaluate the reaction conversion. The gaseous products were analyzed by two kinds of on-line GC-TCD using internal standard gases, N2 and He, to check the consumption of CO2 and H2, respectively. With the data obtained from the GC-TCD analysis, the conversion of CO2 and the yield of CO and CH4 were calculated. The peak areas of organic products were referred to the peak areas of CH4 (GC-TCD); the organic product selectivity and yields were determined from the GC-FID analysis. 3. RESULTS AND DISCUSSION Reactants and products must diffuse through high-molecular-weight liquid hydrocarbons during CO2 hydrogenation. The liquid phase may be confined to the mesoporous structure within catalyst pellets or extend to the outer surface and the interstitial spaces between pellets.

Figure 1. Pore size distribution of iron based catalysts.

Figure 2. Conversion and hydrocarbon selectivity of CO2 hydrogenation in iron based catalysts.

187

The pore size distribution profiles of catalyst appear in Fig. 1. The distribution of A type catalysts are fairly narrow with a maximum at approximately 4.5nm. The pore size distribution of B type catalyst is similar to that of A type catalyst. However, that of C type catalyst is shifted toward large pores. The catalytic activities and selectivity for Fe-K/y-Al2O3 catalysts are compared in Fig. 2. In the case of A and B type catalysts, the conversion of CO2 is approximately 39%. However, when C type catalyst is used, it shows a little high activity compared with those obtained over other catalysts. Comparison of the product selectivity for catalysts was performed under the same reaction conditions. For the A type catalyst, the main products are higher hydrocarbons such as C5+ hydrocarbon. The B type catalyst tends to be similar to the A type catalyst, but it appear that C type catalyst has higher C2-C4 selectivity (40.5%) than A type catalyst. From these results, the CO2 conversion remains almost unchanged as the pore diameter. However, the C5+ selectivity is dramatically changed with pore diameter. Table 1. Physical properties of Fe-K/y-Al2O3 (A type) with and without PEG Fe-K/y-Al2O3 without PEG

Fe-K/y-Al2O3 with PEG

Total pore area (m2/g)

76.4

114.3

Average pore diameter (nm)

16

18.9

Bulk density (g/ml)

1.58

1.20

Porosity (%)

49.6

64.8

The catalytic activities and selectivity of Fe-K/y-Al2C>3 with and without PEG are compared in Fig. 3. When Fe-K/y-Al2O3 without PEG was used, the conversion of CO2 was 39.9%, while in the case of Fe-K/y-Al2O3 with PEG the conversion of CO2 was 45.4%. CO2 conversion was higher in the presence of PEG, as compared with just Fe-K/y-Al2O3. A comparison of product selectivity for catalysts with and without PEG was done at same reaction conditions. Fe-K/y-Al2O3 without PEG showed higher CO (10.0 C-mol%) and C2-C4 hydrocarbons (17.5 C-mol%) compared with Fe-K/y-Al2O3 with PEG. For Fe-K/y-Al2O3 with PEG, the amount of lighter hydrocarbons decreased, but that of higher hydrocarbons increased. The main products over Fe-K/y-Al2O3 without PEG were lighter hydrocarbons such as C1-C4 hydrocarbons, and those over Fe-K/y-Al2O3 with PEG were higher hydrocarbons such as C5+ hydrocarbons. The PEG, used as an additive and sublimed off prior to the high temperature treatment, controlled the bulk density and pore diameter Figure 3. Conversion and hydrocarbon of catalysts. The physical properties of Feselectivity of CO2 hydrogenation in iron K/y-Al2O with and without PEG are shown 3 based catalysts in Table 1. The pore size distribution of the

188

catalysts in the range of micro pore was similar but that of catalysts in the range of large pore was greater for Fe-K/y-Al2O3 with PEG. This result showed that the pore of Fe-K/y-Al2O3 was increased by addition of the PEG. From these data, it was confirmed that the high CO2 conversion and hydrocarbon selectivity on iron-based catalyst relates to the pore size distributions of Fe-K/y-AhCh with and without PEG. It has been postulated that the diffusion of molecules through liquid, which is filling product in the intraparticle, is much slower than gas-phase transport. As a result, slow removal of reactive products can also modify selectivity by enhancing the rate of their secondary reactions.

4. CONCLUSION In the case of Fe-K/y-Al2O3 with PEG, CO2 conversion and higher hydrocarbon selectivity is increased. The increase in catalytic porosity is caused by the addition of PEG. These results show that Fe-K/y-A^Ch with the large porosity exhibits high CO2 conversion and higher hydrocarbon selectivity. It has been postulated that the diffusion of molecules through liquid, which is filling product in the intraparticle, is much slower than gas-phase transport. As a result, slow removal of reactive products can also modify selectivity by enhancing the rate of their secondary reactions. It was suggested that the increased catalyst's porosity enhances the probability of CO2 uptake and chain propagation, therefore, CO2 conversion and higher hydrocarbon selectivity are increased. ACNKOWDGMENT This work was supported by Korea Energy Management Corporation (2001-C-CD02-P-03). The authors thank for financial support. REFERENCES [1] J. Nerlov, S. Sckerl, J. Wambach, I. Chorkendorff, Appl. Catal. A 191 (2000) 97. [2] F. Tihay, A. C. Roger, G Pourroy, A. Kiennemann, Energ. Fuel. 16 (2002) 1. [3] T. Kim, S. Vijayalakshmi, S. Son, S. Ryu, J. Kim, J. Ind. Eng. Chem. 9 (2003) 481. [4] A. Trovarelli, C. Mustazza, G. Dolcetti, Appl. Catal. 65 (1990) 129. [5] C. Kuei, M. Lee, Can. J. Chem. Eng. 69 (1991) 347. [6] M. J. L. Gines, A. J. Marchi, C. R. Apesteguia, Appl. Catal. A 154 (1997) 155. [7] E. Iglesia, S. C. Reyes, R. J. Madon, J. Catal. 129 (1991) 238. [8] E. Iglesia, S. C. Reyes, R. J. Madon, S. L. Soled, Adv. Catal. 39 (1993) 221.


189

Tri-reforming of CH4 Using CO2 for Production of Synthesis Gas to Dimethyl Ether Seung-Ho Leea, Wonihl Choa, Woo-Sung Jua, Young-Soon Baeka, Jong-San Changb and SangEon Parkb a

LNG Technology Research Center, R&D Division, Korea Gas Corporation, 973 Dongchun-

Dong, Yeonsu-gu, Incheon 406-130, Korea. b

Catalysis Center for Molecular Engineering, Korea Research Institute of Chemical

Technology (KRICT), P.O. Box 107, Yuseong-gu, Daejeon 305-600, Korea. In this study, we report the results obtained on tri-reforming of methane over the Ni/ZrCh based catalyst in order to restrain the carbon deposition and to evaluate the catalytic performance. Results of tri-reforming of CH4 by Ni/Ce-ZrO2 in the three catalysts (Ni/CeZrO2, Ni/ZrO2 and Haldor Topsoe R67-7H) are showed that the carbon on the reactor wall and the surface of catalyst were reduced dramatically. It was found that the weak acidic site, basic site and redox ability of Ce-ZrO2 play an important role in tri-reforming of methane conversion. Carbon deposition depends not only on the nature of support, but also on the oxidant as like steam or oxygen. 1. INTRODUCTION The methane or natural gas reforming is one of the processes to produce synthesis gas with desired H2/CO ratios for clean fuels and chemicals. The catalytic reforming of methane with carbon dioxide to produce synthesis gas with lower H2/CO ratio is a challenging approach for the chemical utilization of natural gas and carbon dioxide; consequently, considerable effort has been devoted to the development of high performance catalysts [1-5]. Thus we have developed modified zirconia supported Ni catalysts for carbon dioxide reforming (CDR), steam reforming (SRM) and partial oxidation (POM). We have shown that Ni/Ce-ZrO2 precoated AI2O3 is highly active and stable in POM and SRM [6]. In present work, we refer to tri-reforming of methane with desirable H2/CO ratios by simultaneous oxyCO2-steam reforming of methane for the bench scale process [7,8]. 2. EXPERIMENTAL Support materials employed in this study were Y-A1 2 O 3 (SBET= 234 m2/g) and 6-AI2O3 (SBET= 167 m2/g), which was prepared by calcining y-Al2O3 at 900 °C for 6 h. Ce-ZrO2 modified 0-AI2O3 support were prepared by the incipient wetness method (CeO2 : ZrO2 :

190

AI2O3 = 1 : 4 : 95 w/w). Supported Ni catalysts (3 wt% Ni) were prepared by the impregnation method using appropriate amounts of Ni(NC>3)2 '6H2O onto supports followed by drying at 100 °C and calcining at 550°C for 6 h in air [6]. All the reforming reactions (SRM, POM, CDR, and TRM) were carried out in a fixed-bed reactor, which made of an Inconel 800H alloy tube with dimension of 1.7" I.D. and 24" in length. The catalysts were reduced in the reactor with 5% (v/v) H2/N2 at 700 °C for 2 h before the reaction. 3. RESULTS AND DISCUSSION New catalyst design based on Ce-ZrO2 support for the CO2 reforming was sequentially prepared by the following steps: the modification of Ce-ZrO2, the selection of a promoter on active Ni species, and simultaneous alteration of the support and Ni surface. According to previous report [9], Ni/Ce-ZrO2 exhibited the highest activity in POM and SRM among the tested catalysts including Ni/MgO and Ni/MgAl2O4 which have been considered as the best catalyst in POM and SRM, respectively. Fig. 1 shows CH4 and CO2 conversion with time on stream in CDR at 800 °C. Ni/Ce-ZrO2 represented the highest activity as well as stability. CH4 and CO2 conversion were 92 and 93% after 20 h, respectively, with a H2/CO ratio of 0.94. Ni/y-AbOs and commercial catalyst of reforming, which main component is Ni/MgO, however, deactivated with time on stream rapidly. Because of phase transformation of alumina, Ni/y-Al2O3 has deactivated at high temperature. Otherwise, it seemed that the activity of commercial catalyst fade away due to carbon formation.

Fig. 1. CH4 and CO2 conversion with time on stream in CDR. (A) Ni/Ce-ZrO2, (B) Ni/MgO, (C) Haldor-Topsoe(R-67-7H). (Reaction conditions: CH4/CO2/N2 = 1/1/1, GHSV = 72,000 ml/h.g, T = 800 °C, P = 1 atm, Filled symbol: CH4, Non-Filled CO2 conv.%)

Fig. 2. TPR patterns of Ni/CeZrO2/ e-Al2O3 catalysts with different nickel loading.

Fig. 2 shows the TPR patterns of Ni/Ce-ZrO2/0-Al2O3 catalysts with different nickel loading. Catalysts are loaded up to 2% with Ni, the two reduction peaks of NiOx and MAI2O4 are observed. When the loading amount of Ni is over 3%, the reduction peaks of NiOx are found sharply. The three reduction peaks are observed in the Ni amount range of 6% to 15%. The presence of ceria in Ni/Ce-ZrO2 catalyst has beneficial effects on the catalyst performance such as improving the catalyst stability and enhancing concentration of the

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highly mobile oxygen species. According to previous results, the role of ceria in this catalyst is assigned to form a thermally stable solid solution with zirconia and to give high capacity of oxygen storage, It is considered that the partially reduced ceria site produces active oxygen species by CO2 dissociation, which reacts with the deposited carbon with the help of oxygen spillover from the support onto the Ni sites. In consequence of reactions of synthesis of syngas, Ni/Ce-ZrC>2 catalyst appeared high activity as well as high stability in the CO2 reforming. Fig. 3 shows the experimental results for the effect of reaction temperature in the CDR and TRM respectively. The conversion of CH4 and CO2 was slightly enhanced with increasing temperature. In the case of CDR, H2/CO ratio maintained at 1. As the temperature was raised, the H2/CO ratio was slightly increased in the TRM.

Fig. 3. The effect of reaction temperature in the (a) CDR (CH4/CO2=1/1) and (b) TRM (CH4/CO2/H2O/O2=1.5/1/1/1) on the conversion, synthesis gas yield and H2/CO product over the Ni/Ce-ZrO2/e-Al2O3 catalyst. (Reaction condition: SV=3000 hr', P=2.5 atrn) Table 1 summarizes the reaction activities among the four kinds of combined reforming of methane over Ni/Ce-ZrC>2 catalyst at 800 °C. Tri-reforming reaction showed an efficient conversion of methane compared to other reactions; however, CO2 conversion exhibited low reactivity for conversion around 82.3%. And then it was found a suitable ratio of H2/CO (=1.44) for gaseous DME synthesis. In case of CDR, it showed a high conversion comparatively. However, it was found that a coking onto catalysts as well as difficult problems for temperature control due to generation of heat on real processes during reactions. Furthermore, the catalysts can be oxidized on steam easily on long time run. Thus, it can be seen that tri-reforming was one of the most promising approaches to the efficient process without coking deposition and generation of heat. Specially, it can be considered that the efficient process for chemical products from syngas with the suitable ratio of H2/CO at a reasonable reactants ratio on reaction process.

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Table 1. Comparison of the activities among the four kinds of combined reforming reaction of methane. Reaction 1)CDR 2) CDR + POM 3) CDR + SRM 4) CDR+SRM+POM

H2O+O2/ CO2 0.0 1.0 1.0 2.0

XCH4 (%)

93.4 94.8 96.5 95.6

Xco2 (%) 90.5 87.6 85.3 82.3

H2/CO ratio 0.98 1.19 1.35 1.44

4. CONCLUSIONS The tri-reforming of CH4 by Ni/Ce-ZrO2 catalyst displays an excellent performance for the coke on the reactor wall and the surface of catalyst. It was found that the weak acidic site, basic site and redox ability of Ce-ZrC>2 play an important role in tri-reforming of methane conversion. Also, we optimized the ratio of methane, steam, carbon dioxide and oxygen for the syngas used DME process. The coke formation decreased rapidly by increasing steam molar ratio. Carbon deposition depends not only on the nature of support, but also on the oxidant as like steam or oxygen. We obtained high efficient catalytic performances on Ni/Ce-ZrO2 catalyst compared to commercial catalyst along with stability for production of syngas. From these founding, this catalyst can be attributed to the efficient catalyst for DME production process (50Kg/day) that is now under testing at our company. ACKNOWLEDGEMENT This Research was performed for the Greenhouse Gas Research Center, one of the Critical Technology-21 Programs, funded by the Ministry of Science and Technology of Korea.

REFERENCES l.M.C.J. Bradford and M.A. Vannice, Catal. Rev.-Sci. Eng. 41 (1999) 1. 2. J.R. Rostrup-Nielsen and J-H. Bak Hansen, J. Catal. 144 (1993) 38. 3.Z.-W. Liu, H.-S. Roh and K.-W. Jun, J. Ind. Eng. Chem. 9 (2003) 753. 4.H.-S. Roh, Z.-W. Liu, H.S. Potdar, J.W. Kim and K.W. Jun, J. Ind. Eng. Chem. 9 (2003) 762. 5.Z.-W. Liu, H.-S. Roh, K.W. Jun, H.S. Potdar and M. Ji, J. Ind. Eng. Chem. 9 (2003) 576. 6.H.-S. Roh, K.-W. Jun, S.-C. Baek and S.-E. Park, Bull. Korean Chem. Soc. 23 (2002) 793. 7. W. Pan, J. Zheng and C. Song, Am. Chem. Soc. Div. Fuel Chem. Prep. 47 (2002) 262. 8.T. Inui, Catalysis 16 (2002) 133. 9.H.-S. Roh, K.-W. Jun, W.-S. Dong, S.-E. Park and Y.-S. Baek, Catal. Lett. 74 (2001) 31.


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Carbon Dioxide Reduction Technology with SOFC System Dong Ju Moon*, Jong Woo Ryu and Sang Deuk Lee Reaction Media Research Center, Korea Institute of Science & Technology (KIST), P.O. Box 131, Cheongryang, Seoul, 130-650, Korea. TEL: +82-02-958-5867, FAX : +82-02-958-5809, E-mail: dimoon@,kist.re.kr The electrocatalytic reforming of CO2 with CH4 was carried out in a solid oxide fuel cell (SOFC) system using an electrochemical cell (Ni-YSZ-(MgO or CeO2) I YSZ I (La,Sr) MnO3). It was found that reaction rates of CO2 and CH4 under the closed-circuit condition were more stable than those of the open-circuit. The results suggest that the electrocatalytic internal reforming of CO2 with CH4 in the SOFC system is an attractive process to coproduce electricity and a syngas with the reduction of CO2. 1. INTRODUCTION The CO2 catalytic reforming by CH4, so-called dry reforming, is an attractive conversion technology because of the possibility of enhancing natural gas utilization with the sequestration of CO2 [1]. However this reaction has two serious problems. The reaction is an intensively endothermic reaction, which consumes much energy and the other is the deactivation of catalyst by carbon formation [1-3]. To overcome these problems, much effort has been focused on the development of catalyst which show high activity and resistance against coke deposition during long-term operation. One of these efforts is an electrocatalytic reforming in SOFC. In our previous works [4-7], it was reported that the electrocatalytic reforming of CO2 by CH4 in a SOFC system has some advantages over the catalytic reforming. Because the syngas generated by an internal reforming can be used as fuels for power generation in SOFC, it was considered that the electrocatalytic internal reforming of CO2 by CH4 is an attractive option for improving the energy utilization of fuels. In this work, the cogeneration of the syngas and electricity by the electrocatalytic internal reforming of CO2 by CH4 was investigated over the electrochemical cells (Ni-YSZ-CeO2 and Ni-YSZ-MgO I YSZ I (La,Sr)MnO3) in the SOFC system. 2. EXPERIMENTAL The electrochemical cell was prepared by a tape casting method. The thickness and area of catalyst electrode layer were ca. 20 IM and 2.25 cm2 (1.5 cmxl.5 cm), respectively. The electrocatalytic reforming system consists of feed supply, electrical furnace, electrochemical

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cell, electrocatalytic reactor and GC analysis sections. A mixture of 12.5 vol% CO2 and 12.5 vol% CH4 was passed through the anode chamber with a flow rate of 20 ml/min, while air (20 ml/min) was passed through the cathode side. The outlet gas from the anode side was analyzed by an on-line GC [H.P. Co., HP5890 series II] equipped with a carbosphere column (3.18xl0"3 m O.D. and 2.5 m length) and a thermal conductivity detector. Pt wire (O.D. = 0.5 x 10"3 m) was used to connect both electrodes to an electrical circuit for controlling the oxygen flux across the YSZ electrolyte. The electrochemical cell was sealed onto the alumina tube (O.D. = 0.025 m, ID. = 0.019 m) using a pyrex glass (O.D. = 0.025 m, I.D. = 0.021 m, h = 0.003 m). The electrocatalytic reactor was operated at 800 °C for 6 h under the open- and close-circuit conditions. 3. RESULTS AND DISCUSSION Figure 1 shows the electrocatalytic reforming of CO2 with CH4 in a SOFC system. The electric power was generated by electrons released in the reactions of CO with oxygen ion, H2 with oxygen ion and surface carbon with oxygen ion. The stability of catalyst electrode was maintained by the reaction of oxygen ion transferred from cathode with the surface carbon formed in the internal CO2 reforming by CH4 in the SOFC system. Figure 2 shows the reaction rates of CH4 and CO2 over electrochemical cells, Ni-based anodes I YSZ I (La,Sr)MnO3 with time-onstream under the open- and the closed-circuit conditions at 800 °C. It was found that the NiYSZ-CeO2 catalyst anode displayed higher reaction rates than the Ni-YSZ-MgO in the SOFC system. Under the open-circuit, the reaction rates of CH4 and CO2 drastically Figure 1. Electrocatalytic reforming of CO2 decreased, respectively. It was considered with CH4 in a SOFC system. that the decreases in catalytic activity were caused by the carbon deposition on catalyst electrode. On the other hand, the reaction rates of CH4 and CO2 were stable after undergoing the electrocatalytic reaction for 5 h under the closed-circuit. The current density was slightly decreased with time-on-stream, and then it became stable. Figure 3 shows the performance of current voltage and power density with current density in the electrocatalytic cells (different anodes I YSZ I (La,Sr)MnO3) at 800 °C when CH4 and CO2 were used as reactants. The open-circuit voltage (OCV) over the Ni-YSZ-MgO catalyst electrode was 0.96 V. It had a maximum power density at 30 mA. Considering electrode area, the power of 37 mA was produced in the single cell. The open-circuit voltage over the NiYSZ-CeO2 catalyst electrode was 1.02 V.

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(a) Ni-YSZ-CeO2 I YSZ I (La,Sr)MnO3

(b) Ni-YSZ-MgO I YSZ I (La,Sr)MnO3

Figure 2. Reaction rates of CH4 and CO2, and current density in the electrochemical cells (Ni-based anodes | YSZ I (La,Sr)MnO3) with a time on stream under the open- and the closed-circuit at 800 °C. Table 1 shows the product distributions for the electrocatalytic reaction of CO2 and CH4, and the amount of coke formed at 800 °C for 5 h in the electrocatalytic cell (Ni-YSZ-CeO2 I YSZ I (La,Sr)MnO3) under the open- and the closed-circuit conditions. The amount of coke formed under the open-circuit was ca. 104 mg C/gcat for 5 h. But the amount of coke under the closed-circuit drastically decreased compared to that under the open-circuit. The results interpreted that the carbon deposited on the surface of anode catalyst under the open circuit was mainly desorbed to carbon dioxide by the reaction of oxygen ion transferred from the cathode with surface carbon. Table 1. The product distributions for electrocatalytic reforming and the amount of coke formed at 800 °C for 5 h in the Ni-YSZ-CeO2 I YSZ | (La,Sr)MnO3 system Product distributions (mol%) Coke formation Reaction mode CO CH4 CO2 H2 m g C / gcatalyst 16.46 50.79 17.43 15.32 104 Open-circuit 28.76 29.81 12.24 29.19 Closed-circuit Flow rate Fuei = 20 ml/min : CH4 (12.5 vol%), CO2 (12.5 vol%), Flow rate Air = 20 ml/min. Figure 4 shows SEM images for the surface of Ni-YSZ-CeO2 anode after the reaction under the closed circuit in the electrocatalytic cell system. It was reported that surface carbon under the open-circuit was accumulated on the surface of anode and it filled up pore [4]. It was found from SEM images that the amount of carbon accumulated on the surface of electrochemical cell (Ni-YSZ-CeO2 I YSZ | (La,Sr)MnO3) was desorbed by the reaction of oxygen ions pumped from the counter electrode.

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Figure 3. The performance of current voltage Figure 4. SEM image for the surface of and power density with current density in the catalyst anode after the reaction in the electrochemical cells (Ni based anodes I electrochemical cell (Ni-YSZ-CeO2 I YSZ YSZ | (La,Sr)MnO3) at 800 °C. I (La,Sr)MnO3) under the closed-circuit. 4. CONCLUSIONS The Ni-YSZ-CeO2 catalyst anode displayed higher reaction rates than Ni-YSZ-MgO in the Ni based anodes I YSZ I (La,Sr)MnO3 system. The reaction rates of CH4 and CO2, and the current density under the closed-circuit were stable after undergoing the electrocatalytic reaction for 5 h, whereas those of CH4 and CO2 under the open circuit slowly decreased. The results suggest that the electrocatalytic internal reforming of CO2 with CH4 in SOFC system is an attractive process to co-produce electricity and a syngas with the reduction of CO2. ACKNOWLEDGEMENT This paper was performed for the Carbon Dioxide Reduction & Sequestration Center, one of 21st Century Frontier R&D Programs funded by the MOST of Korea. REFERENCE 1. S. L. Douvartzides and F. A. Coutelieris, J. of Power Sources, 114 (2003) 203. 2. S. W. Hong, S. M. Oh and D. W. Park, J. Ind Eng. Chem., 7(6) (2001) 410. 3. J. Liu anc S. A. Barnett, Solid State Ionics 158 (2003) 11. 4. D. J. Moon and J. W. Ryu, in press to Catalysis Today (2003). 5. D. J. Moon, T. Y. Kim, J. H. Kim, S. D. Lee and S. I. Hong, Eng. Conf. Inter., Qubec, Canada, July 2(2003) 33. 6. D. J. Moon, J. Y. Yoo, J. M. Park, S. W. Nam and S. D. Lee, The 9th Korea-Japan Symp. on Catalysis, Pohang, Korea, P-77 (2003) 271. 7. D. J. Moon, J. W. Ryu, T. Y. Kim B. G. Lee and S. D. Lee, in application to Korea Patent and U.S. Patent (2003).


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Synthesis of dimethyl carbonate from urea and methanol over metal oxides Mouhua Wang, Ning Zhao, Wei Wei, Yuhan Sun* State Key Laboratory of Coal Conversion, Institute of Coal Chemistry Chinese Academy of Sciences, Taiyuan, 030001 PR China A series of metal oxide were used as solid catalysts for synthesis of dimethyl carbonate from urea and methanol. Among them, ZnO was shown to be the highest active catalyst. It was also found that the activity of other oxides such as CaO, MgO, ZrC-2 was closely related to their basic strength. 1. INTRODUCTION In recent years, much attention has been paid to dimethyl carbonate (DMC). As an environmentally benign building block : ' 2 , DMC shows versatile chemical property and is mainly used for methylation and methoxycarbonylation as a safe substitute for dimethyl sulfate or methyl halides, and it is also used as a solvent to replace halogenated solvents such as chlorobenzene '4. In addition, DMC is believed to be a good additive of gasoline in the future due to its higher oxygen content (53%) and good blending octane ((R+M)/2=105). Dimethyl carbonate was usually produced from methanol and phosgene in a concentrated sodium hydroxide solution5. Because of the use of phosgene, DMC synthesis has been limited in industry. Thus, a non-phosgene process for preparing DMC by oxidative carbonylation of methanol in liquid phase or gas phase was put on stream in the EniChem-Ravenna factory using CuCl as the catalysts6"8. However, the shortcomings of this process were low production rate and high cost of the separation. Another non-phosgene process for the synthesis of DMC was the transesterification from methanol and propylene(ethylene) carbonate which was prepared from CO2 and propylene(ethylene) oxide via cyclic addition. The reaction of cyclic carbonate with methanol was limited by equilibrium. Moreover, the reactant involved in the process (such as epoxides) was explosive and expensive. Recently, urea methanolysis method has been used to prepare of dimethyl carbonate. Firstly, urea and methanol reacted to produce carbamate and ammonia, and then carbamate and methanol were further converted into DMC and ammonia. Ammonia could be recycled to produce urea with CO2 (see Scheme 1). Obviously, this was an economically benign process. *To whom all corresponding should be addressed, E-mail: yhsun(a),sxicc.ac.cn or weiwei@,sxicc.ac.cn: Fax: +86-351-4041153; Tel: +86-351-4053801

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However, the related study except several U.S. patents lo ' n ' 12 for this process has not been reported till now. Homogenous catalysts such as some complexes of organotin were used with high boiling electron donor compound as the solvent (such as bidentate ligands). However, the preparation of the catalyst and the solvent was very difficult and costly. In addition, the separation of the products and the catalyst was also complicated. Thus, homogeneous catalysts were not satisfactory for industrial use. Considering the reaction thermodynamic and the reaction mechanism over homogenous catalysts, solid catalysts (such as CaO and MgO and ZrCh) were developed for synthesis of dimethyl carbonate from urea and methanol here. CO2

+

2NH3-«

I -H 2 O I +CH3OH +CH3OH CH3O-C(O)-NH2 H 2 N-C(O)-NH 2 -NH3 Fig. 1. Synthesis of DMC by urea methanolysis method

;

I I CH3O-C(O)-OCH3 + NH 3

2. EXPERIMENTAL The catalysts, such as ZnO, CaO, MgO, ZrO 2 , were prepared by thermal decomposition of zinc carbonate at 500°C for 4 hours, calcium carbonate at 800°C for 2 hours, magnesium hydroxide at 600°C for 3 hours and zirconium hydroxide at 350°Cin nitrogen atmosphere, respectively. All experiments was carried out in a stainless steel autoclave with an inner volume of 250 ml provided with a magnetic stirrer and an electric heater, a flux column connected with the autoclave which refluxed the methanol and removed ammonia from the reaction system. The products were analyzed by GC. CO2-TPD measurement was performed at a heat rate of 10K min"1 under nitrogen flow (50 mL min"1), and CO 2 desorbed was detected by a BALZA Q-Mass spectrometer. 3. RESULTS AND DISCUSSION 3.1. Synthesis of MC from urea and methanol Figure 2 shows the effect of temperature and reaction time on the yield of methyl carbamate (MC). It was found that methyl carbamate was easily prepared from methanol and urea even if no catalyst was added. Under the reaction condition of 160°C and 10 atm, the methyl carbamate yield was over 98% in 2 hours. Obviously, the reaction rate increased with temperature. The MC yields hardly changed at higher temperature than 160°C even though the reaction time was prolonged from 2 hours to 10 hours. Trace amount of DMC in products (lower than 1%, see table 1, entry 11) was also found when the reaction temperature was higher than 160°C. It was concluded that the first step in figure 1 performed quickly and completely even without catalyst while the second step was slow which might be due to the limitation of thermodynamic equilibrium. As a result, the main purpose in the present work

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was the effect of metal oxide on the synthesis of DMC from MC and methanol. 3.2. Synthesis of DMC from MC and methanol Table 1 lists the performance of solid catalysts for the synthesis of DMC from MC (or urea) and methanol. DMC could be effectively synthesized over the metal oxides. The yield of DMC was 40% or so for ZnO, which was much better than for the homogenous catalysts reported elsewhere 10 '"' 12 (entry 10). The activity of CaO for DMC synthesis was lower than that of ZnO but higher than MgO and ZrO2. The DMC yield for CaO was almost the same whether urea or MC as the feed (entry 3 and 4), suggesting that the rate determine step of synthesis of DMC from urea was the second step. However, urea as the feed led to the different DMC yield form MC as the feed with ZnO as the catalyst (entry 8 and 9), implying that the conversion of urea and MC into DMC followed different reaction mechanisms over ZnO. It is worth to note that N-methyl methyl carbamate (NMMC) was detected by GC-MS at high temperature, which could be produced by the further reaction of DMC with MC because DMC was a good methylating reagent2. So, the side reaction of producing NMMC as well as decomposition of DMC leaded the phenomenon of high MC conversion but low DMC yield at high temperature, especially above 180°C.

Fig. 2. Results of reactions without catalyst (0.2 mol urea and 4 mol methanol, MC yield based on urea)

Fig. 3. CO2-TPD profiles on oxide catalysts

Figure 3 depicts the TPD profiles of CO2 adsorbed on ZnO, ZrO2, MgO and CaO. Temperature programmed desorption (TPD) of CO2 was usually used to measure the number and strength of basic sites, and high desorption temperature indicated high basic strength B . No CO2 adsorption peak was observed on ZnO. For the others, the order of CO2 uptake temperature was as follows: CaO> MgO > ZrO2. This strongly suggested that the performance of catalysts was not closely related to their basicity, and the reaction mechanisms over those metal oxides might be different. ZnO had the highest activity although the activity of CaO, MgO and ZrO2 was in accordance with their basic strength, indicating that the reaction over ZnO took place via a different route.

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4. CONCLUSION The synthesis of dimethyl carbonate form urea and methanol via two consecutive steps was elucidated through blank reaction. The second step, through which DMC was produced by reacting intermediate MC with methanol, could be carried out with metal oxides as the catalysts. Among those oxides, ZnO showed the highest activity for DMC synthesis. CO2-TPD revealed that the activity of catalysts was not closely related to their basicity, and the reaction mechanisms over different metal oxides might be different. Table No. 1 2 3 4 5 6 7 8 9 10 11

1. Results over metal oxides Catalyst Weight, g Temp. °C Reaction time, h 1 CaO 160 10 1 180 CaO 10 1 200 CaO 10 200 11 1 CaO* 1 182 10 MgO 1 200 10 MgO 180 1.5 10 ZrO 2 1 180 10 ZnO 10 180 1 ZnO* 8 180 3 SnOBu2* 180 10 Void* -

MC conv.% 22 39 70 34 14 44 16 9 42 60 92

DMC yield% 8 15 10 11 8 17 5 6 37 16 0.8

Reaction conditions: 0.2 mol MC, 4 mol methanol. *Reactants here are urea and methanol. Values in column MC conversion here mean MC yields; DMC yield is based on urea or MC. REFERENCES 1. P. Tundo, and M. Selva, Chemtech, 1995,25(5), 31-35. 2. Yoshio Ono. Applied Catalysis A: General, 1997, 155, 133-166. 3. A. G. Shaikh, Chem. Rev. 1996, 96, 951-976. 4. E. Suzuki, M. Akiyama and Y. Ono., J. Chem. Soc, Chem. Commun.,1992, 136-137. 5. H. Murdock,/. Phys. Chem., 1919,23,508-516. 6. U. Romono, R. Tesei, et al, Ind. Eng. Chem. Prod. Res. &Dev.,\9S0, 19(3), 396-403. 7. J. Haggin, C&EN, 1987,65(44), 27-30. 8. S. K. King,/. Catal, 1996,161(2),530-538. 9. J. Haggin, C&EN, 1992, 70(18), 25-26; 10. R.Y. Saleh, et al, US patent No.5565603 (1996). 11. J. Y. Ryu, et al, US patent No. 5902894 (1999). 12. J. Y. Ryu, et al, US patent 6392078 (2002). 13. H. Hattori, Chem. Rev., 1995, 95, 537-558.


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Development of heterogeneous catalyst system for esterification of free fatty acid contained in used vegetable oil Hak-Ju Kim, Bo-Seung Kang, Min-Ju Kim, Deog-Keun Kim*, Jin-Suck Lee^ and Kwan-Young Lee* Department of Chemical and Biological Engineering, Korea University, Seoul 136-701, Korea Biomass Research Center, Korea Institute of Energy Research, Yuseong 305-600, Taejeon, Korea g author, FAX: +82-2-926-6102, E-mail: [email protected]

t

Biodiesel produced by the transesterification of vegetable oils (VOs) is a promising alternative fuel to diesel which can overcome limited fossil fuel resources and environmental concern. In this work, an environmentally benign process for the production of biodiesel from used VOs by heterogeneous catalyst has been developed. Free fatty acids contained in used VOs, which cause several severe problems for the transesterification catalyzed by base catalyst, were converted to fatty acid methyl esters (FAME). The activity of several heterogeneous acid catalysts was investigated. Sulfated zirconia showed the highest activity among them. The optimization of the reaction conditions was also carried out. 1. INTRODUCTION Over recent few decades, many efforts have been made to develop clean alternative fuel in many countries. Among the many possible sources, biodiesel fuel derived from vegetable oil (VOs) has attracted attention as a promising substitute for conventional diesel fuels. The increasing use of petroleum continues to increase air pollution levels and accelerate global warming problems caused by CO2. Sharmer et al. estimated that, if 1 kg of pure biodiesel was used as fuel instead of fossil fuel, CO2 production would be reduced by 3.2 kg [1], Furthermore, biodiesel contains no sulfuric compound, which leads to zero-production of SOx. The production of soot can also be reduced due to the property of oxygenated fuel, as well. Biodiesel is generally produced by the transesterification of triglyceride (VOs) to methyl esters with methanol using sodium or potassium hydroxide dissolved in methanol as catalyst (eq. 1). Recently, we have developed a Na/NaOH/y-Al2O3 heterogeneous base catalyst for the production of biodiesel, which showed almost the same activity under the optimized reaction conditions compared to the conventional homogeneous NaOH catalyst [2]. H ! H— C— 00R I H — C — OOR' H-C—00R' H

+

Catalyst 1 r.Hjfffl — »*

H I H - C — OH I H-C—OH H-C — O H

ROOCH, 4-

ROOCH 3

n\

R"OOCH3

H

triglyceride (VOs) methanol glycerol methyl ester (biodiesel) Despite of the environmental advantages of the biodiesel, production cost of biodiesel is pretty high compared to conventional diesel fuel. For the reduction of the production cost,

202

used vegetable oils should be used as feedstock. However, to utilize used VOs as feedstock was thought to be able to cause serious troubles in manufacturing process. First of all, the water content in the used VOs might be considered. Water can shift the equilibrium conversion to the reverse side giving to a decrease of production yield. From our previous research, we could conclude that the water contained in the used VOs was small enough and gave no influence on the biodiesel production yield. However, the free fatty acids (FFA) contained in used VOs are known to cause severe problems for the transesterification catalyzed by base catalyst. FFA can react with the base catalyst (neutralization reaction), which brings to a loss of catalyst and a production of soap as by-product which accelerates the deactivation of the base catalyst. To overcome these problems, FFA contained in used VOs should be removed or converted to inert material. In this work, we tried to convert FFA to fatty acid methyl esters (FAME) (eq. 2), which are the components of biodiesel themselves. This kind of esterification reaction is known to be catalyzed by acid catalyst. Activity of heterogeneous acid catalysts such as sulfated zirconia, sulfated alumina, zeolites, Cs substituted heteropoly acid, tungsten oxide zirconia, was investigated. The optimization of the reaction conditions was also carried out. Free fatty acids (FFA) + methanol —> Fatty acid methyl esters (FAME) + water

(2)

2. EXPERIMENTAL For the preparation of sulfated alumina, Y-AI2O3 (Sumitomo) was pretreated at 550 °C for 12h and then added to 0.8M, 3.6M of H2SO4 solution which was stirred for lh. The powder was filtered and dried at 110°C for 12h and calcined at 550°C for 3h. H type zeolites were purchased from ZEOLYST. The sulfated zirconia was prepared by the dehydration of H2SO4 and Zr(OH)4. The tungsten oxide zirconia was prepared by adding Zr(OH)4 to ammonium metatungstate (Aldrich) solution and stirring for 16h [3]. Zr(OH)4 was purchased or prepared by the conventional method reported elsewhere [4,5] for the preparation of both sulfated zirconia and tungsten oxide zirconia. The Cs substituted heteropoly acid catalyst was prepared by the method reported in previous work [6]. H3PW12O40 (Wako) was dried at 100 °C for 12h to obtain a hexahydrate form. CSCO3 solution was added dropwise at lml/min and stirred for 24h. The powder was dried and calcined at 300 °C for 3h in the atmosphere of a flow of He at 100 ml/min. The BET surface area and pore volume of the prepared catalysts were measured using ASAP 2100 BET surface analyzer. The esterification was carried out in a 100 ml autoclave batch reactor. The used vegetable oil was replicated by adding 4wt% of oleic acid to 50g soybean oil. Generally used oil contains FFA of 1 to 10 wt% depending on the sources. Methanol was loaded in excess (oleic acid : methanol = 1: 19.4 mole ratio). The conversion of oleic acid to methyl ester was calculated according to the amount of the residual oleic acid estimated by means of titration with 0.1N KOH standard solution according to the equation below. _ ,„., Initial Acid Value - Final Acid Value i n A 0 / Converswn(%) = x 100% Initial Acid Value ...... M Acid Value =

xAxFxN

M: Molecular weight of KOH A: Amount of KOH standard solution used for titration F: Concentration coefficient of the KOH standard solution

(3)

203

N: Normal concentration of the KOH standard solution S: Sample weight 3. RESULTS AND DISCUSSION The activities, BET surface areas, average pore sizes and the Hammett acidity of the catalysts tested are listed in Table 1. Sulfated zirconia, reported to have the strongest Hammett acidity among the catalysts surveyed, showed the highest activity. Tungsten oxide zirconia also showed pretty high catalytic activity. H type zeolite had an average pore size of 5 A, which was thought to be too narrow for the diffusion of the triglyceride molecule into the pore. As a matter of fact, inspite of the large surface area and strong acidity, H type zeolite showed poor activity. Sulfated alumina was prepared by physically depositing sulfate functional group on the catalyst support. Therefore, the acid amount was difficult to be controlled and thought to be lower than that of sulfated zirconia, which explains its poor activity. Table 1. Catalytic activity and physical properties of the catalysts Conversion Surface Area Average Pore Hammett Catalysts Acidity Size (A) (m2/g) (%)at60°C c) 141.8 Sulfated Alumina (0.8M) 55.3 43.4 -14.5 132.7 Sulfated Alumina (3.6M) 60.2 33.4 596.0 H type Zeolite 5.0 12.8 a) 81.5 Sulfated Zirconia 88.8 81.8 -16.1 123.9 Sulfated Zirconiab) 42.4 86.9 Tungsten Oxide Zirconiaa 58.9 129.3 83.9 -14.5 Tungsten Oxide Zirconiab 56.7 89.3 86.4 104.1 -12.4 Cs-Heteropoly acid 55.4 73.1 c) "'prepared Zr(OH)4 "'purchased Zr(OH)4 10gofcat alyst was used. The effects of the reaction temperature and the catalyst loading were investigated for the sulfated zirconia catalyst which showed the highest activity among the catalysts tested. The catalytic activity according to the reaction temperature is shown in Figure 1. At temperature above 100°C, the activity tended to decrease. Methanol and triglyceride are miscible forming a single phase below the temperature of methanol boiling point (60 °C). But above that temperature, methanol started to vaporize, forming 2 phases, and the miscibility between the reactants, methanol and triglyceride, inevitably decreased. This decrease of miscibility between reactants brought the diminution of mass transfer as well as reaction rate. Considering the limitation of mass transfer at high temperature, the effect of stirring speed was investigated. The catalytic activity was measured at 120 °C, when the limitation of mass transfer is critical, varying the stirring speed from 300 rpm to 600 rpm. However, the variation of stirring speed had little influence on the catalytic activity. Therefore, the optimum reaction temperature was finally concluded to be in the range of the boiling point of methanol (60 °C). The effect of the catalyst loading on the catalytic activity is presented in Figure 2. The activity tended to increase in accordance with the catalyst loading. This phenomenon could be explained by an increase of acid amount present in the reaction system. However, the experiment of catalyst loading above lOg could not be accomplished since a problem of

204

stirring occurred by an increase of viscosity of the reactant and catalyst slurry.

Fig.l Effect of reaction temperature (sulfated zirconiab lOg, 600 rpm, 2h)

Fig.2 Effect of catalyst loading (sulfated zirconiab,75 °C, 600 rpm, 2h)

4. CONCLUSION For the production of biodiesel from used vegetable oil, free fatty acid contained in the feedstock must be removed or converted to inert material for long catalyst life and process maintenance. Considering the efficiency of the usage of low material, converting free fatty acid to fatty acid methyl ester is one of the most useful methods. Sulfated zirconia and tungsten oxide zirconia were found highly active for converting the free fatty acid to its methyl ester. The optimum reaction temperature was found to be in the range of the methanol boiling point, while at higher temperature mass transfer was limited by the separation of the phase due to the vaporization of methanol. Stirring speed had little effect on the activity. The activity of the catalyst increased according to the catalyst loading. ACKNOWLEDGEMENTS Financial support by the Grant-in-aid from the Industrial Waste Recycling R&D Center and the Ministry of Science and Technology is highly appreciated. The authors also appreciate the support by research grants from the Korea Science and Engineering Foundation (KOSEF) through the Applied Rheology Center (ARC) at Korea University. REFERENCES 1. K. Scharmer, Umweltaspekte bei Herstellung und Verwendung von RME, In RME Hearing, Ministry for Agriculture, Vienna, Austria (1993). 2. H.-J. Kim, B.-S. Kang, M.-J. Kim, D.-K. Kim, J.-S. Lee, and K.-Y. Lee, EuropaCat-IV, Innsbruck, Austria (2003) Bl .089. 3. J.R. Sohn and M.Y. Park, J. Kor. Ind. Eng. Chem., 10(2) (1999) 247-251. 4. D. Corma, V. Forties, M.I. Juan-Rajadell and J.M. Lopez Nieto, Appl. Catal. A: General, 116(1994)151-163. 5. G.D. Yadav and J.J. Nair, Microporous andMesoporous Materials, 33 (1991) 1-48. 6. J.I. Yang, D.W. Lee, J.H. Lee and K.Y. Lee, Appl. Catal. A: General, 194-195 (2000) 123127.


205

Investigation of Synthesis Gas Production from Natural Gas and CO2 by Microwave Plasma Technology Wonihl Choa'*, Woo-Sung Jua, Seung-Ho Leea, Young-Soon Baeka, and Young Chai Kimb a

LNG Technology Research Center, R&D Division, Korea Gas Corporation, 973 DongchunDong, Yeonsu-Gu, Incheon 406-130, Korea. b

Department of Chemical Engineering, Seongdong-gu, Seoul, 133-791, Korea

Hanyang

University,

17

Haengdang-dong,

In a microwave plasma catalytic reaction, synthsis gas production from natural gas and carbon dioxide was carried out with 2.45 GHz microwave plasma. The aim of this study was to investigate the effects of the feed ratio(CH4:CO2), the syngas ratio(H2:CO) and with or without the present of Ni catalyst in a microwave plasma system. Results of carbon dioxide reforming of methane indicated that a microwave plasma catalytic reaction provided the more effective reaction temperature than a conventional reforming process. In the present of Ni catalyst, the ratio of synthesis gas represented the lower value than the plasma reaction without catalyst. An increase in the CH4:CO2 ratio causes an increase in ethane and hydrogen selectivity indicating that an oxidative coupling pathway may play a role under these conditions. 1. INTRODUCTION The conversion and utilization of methane and CO2 can be one of a promising way of reducing the greenhouse gas emission. At present, the reforming of methane with CO2 to produce synthesis gas is a challenging approach for chemical utilization of natural gas and carbon dioxide. Synthesis gas, consisting of hydrogen and carbon monoxide, is an important raw material in natural gas industry as well as petroluem plant. This process would be to recycle the carbon in a chemical process and synthesis gas could be make useful products such as dimethyl ether(DME) or oxygenates [1-4], Non thermal plasma technology have been extensively investigated for cleanup of flue gas to remove volatile organic compounds (VOCs), NOx, and SO2. Recently the reduction of CO2 from flue gas using a corona discharge has been studied by several research group. A plasma catalytic reaction has been found to be of great interest for CO2 reduction[5-7]. Because of the problem of energy carriers, special attention has been paid to dissociation of CO2 at moderate pressure (6.7-26.7 kPa). The minimum energy consumption for the formation of a CO molecule is 3.8 eV/mol and the maximum energy efficiecy is 80%. It has been noted that a microwave discharge is more efficiency than radio frequency discharge in relation to the dissociation of CO2[8]. In this study, we have used the microwave plasma torch and catalytic reactor installed on Ni-based catalyst. We report that our recent results of plamsa catalytic reaction for CO2 reforming of methane over Ni catalyst in a microwave plasma system.

206

2. EXPERIMENTAL In order to investigate the plasma catalytic reaction, the reactor was designed to be able to insert the catalyst pellet on moving plate. The supported Ni catalyst was prepared by impregnation of 5 wt%-Ni on gamma AI2O3 using nickel nitrate (Ni(NO3)6H2O) solution. The dried catalyst was calcined at 500°C for 4 h in air and reduced for 2 h under H2 stream. Prior to characterization and plasma-catalytic reactivity measurements, samples were annealed at 473K in the presence of N2. To produce synthesis gas and others (mainly C2 products) from methane, plasma catalytic reactions were carried out with microwave plasma (2.45GHz, iplas Co.) from 2.5 to 13.2 kPa of operation pressure. The reactor of 1" O.D. quartz tube was connected to a microwave waveguide and resonator. The feed gas system used mass flow controllers, to regulate the flow rates of CH4 and CO2 gases. The analysis of product was analyzed by gas chromatography (Young-in Scientific Co. Model 680D). 3. RESULTS AND DISCUSSION Figure 1 shows that the decomposition of CH4 progresses more easily than that of CO2 and the CH4 conversion increases with increasing CH4/CO2 ratio, whereas the CO2 decreases. The carbon dioxide (CO2) may have decomposed to CO and oxygen in the plasma. The conversion of CH4 and CO2 was almost steady above l.OkW of applied power in this plasma system. The synthesis gas ratio was 1.67 at l.OkW of power, 4.0kPa of pressure and lOOml/min of flow rate.

Figure 1. The effect of applied power of microwave plasma on the conversion of CH 4 and CO 2 (a), mole ratio of CH4/CO2 = 1.0, and the effect of CH 4 /CO 2 mole ratio on the conversion of CH4/CO2 (b), microwave power =1.0 kW, operation pressure = 4.0 kPa, total flow rate = 100 ml/min.

Figure 2 shows the effect of catalysis on the conversion of CH4/CO2 in the plasma. The results showed that the conversion of CH4/CO2 decreased with stream of time at the present of catalyst, whereas the one of CH4/CO2 remained uniform conversion value at plasma reaction system. In case of the catalysis, it may happen to coke over Ni catalyst and cause the deactivation of catalyst. However, the ratio of synthesis gas reduced to 1.0 at the present of Ni catalyst, which was CO2 more easily transfer to CO on the surface of catalyst. It is considered that a dissociation

207

of CO2 breaks out as a result of single collisions of molecules with electrons and provides the formation of oxygen atoms and CO molecules. The contribution of dissociative attachment is small and is additionally compansated by the reverse process of detachment in collisions with CO molecules; that is the only reaction capable of competing eith dissociation by the electron impact in dissociation of CO2 over Ni catalyst is the following reaction[9]: CO2 + CO (a3 n,v) -> 2CO + O,

k = 3.4xlO"I7mV

(1)

In this case, the CO molecules must be excited by an electron impact or formed directly as a result of dissociation of CO2 over Ni catalyst.

Figure 2. Conversions of CH4/CO2 and H2/CO ratio with and without the present of Ni catalyst, microwave power = l.OkW, operation pressure = 4.0kPa, total flow rate =

The results for the comparison on selectivity with the change of CH4:CO2 ratio is shown in Table 1. As the CH4:CO2 ratio increase, the CH4 conversion slowly increases whereas the CO2 conversion decreases. In these test, it is noted that one of the product is ethane as a result of methane coupling reaction. In a microwave and rf plasma of methane or mixtures of methane with inert gases, the main stable products of dissociation are ethane, ethylene, acethylene and hydrogen. It is evident that the synthesis of above products take place during recommbination of radicals generated in the primary acts of dissociation. In case of this experiments, it seem that the formation of ethane is caused by the recombination of the radical (carbene, CH2) from methane over Ni metal[9,10]: CH2 + CH4 -> (C2H6)* -

C2H6

(2)

The reactions of the carbenes implantation take place almost without activation energy. Therefore, a more energetic process may be the recombination of the radical reaction than the abstract an oxygen atom from CO2 (C=O bonding energy is 3.8 eV/mol). It is evidence in this study that the selectivity of C2H6 and hydrogen is larger than the one of CO. The highest fractional conversion of CO2 occurs at a 1:2 of CH4:CO2 ratio.

208 Table 1. Comparison on the conversion and the selectivity with the change of CH 4 :CO 2 ratio at microwave power is 1.0 kW, operation pressure is 10 kPa and total flow rate is 100 ml/min. CH 4 :CO 2 ratio

CH 4 Conv., %

CO 2 Conv., %

C 2 H 6 Select.,%

CO Select., %

H 2 Select., %

67.3 68.2 70.5

61.2 60.4 58.5

24 31 42

42 32 17

32 37 41

1 :2 1 :1 2: 1

4. CONCLUSIONS In this study, a microwave plasma-catalytic reaction was investigated for CO2 reforming of methane over Ni catalyst. The reduction of CO2 takes place during the recombination of CO2 and the formation of CO. The synthesis gas ratio was affected by catalyst, however, the catalysis have the deactivation due to coke deposition. Under the feed of methane-rich, the recombination of the radical reaction is prior to the abstract of oxygen atom from CO2 at present of catalyst. In this experiment, it observed the like-behavior of the methane oxidation coupling in CO2 reforming of methane. A general carbon conversion of 67-70% and synthesis gas raio in the product gas of 0.82.4 are achieved over Ni catalyst. REFERENCES 1. P. R. Stuart, R. J. Munz and W. H. Gauvin, Plasma Chemistry and Plasma Processing, 4(4) (1984) 297. 2. K. Supat, A. Kruapong, S. Chavadej, L. L. Lobban and G. Mallinson, Energy & Fuels, 17 (2003) 474. 3. S. Wang and G. Q. Lu, Applied Catalysis A: General, 169 (1998) 271. 4. S.-H. Lee, W. Cho, W.-S. Ju, B.-H. Cho, Y.-C. Lee and Y.-S. Baek, Catalysis Today, in press (2003) 5. C.-J. Lui, L. L. Lobban and G. Mallison, Greenhouse Gas Control Technologies, Elsevier Science Ltd, 1999. 6. D.W. Larkin, M. Leethochawalit, S. Chavadej, T.A. Caldwell, L.L. Lobban and Richard G. Mallinson, Greenhouse Gas Control Technologies, Elsevier Science Ltd, 1999. 7. S.-W. Hong, S.-M. Oh, D.-W. Park, and G.-J. Kim, J. Ind. Eng. Chem., 7(6) (2001) 6410. 8. L. S. Polak and Y A. Lebedev, Plasma Chemistry, UK, Cambridge International Science Publishing, 1998. 9. V. D. Rusanov and A. A. Fridman, Physics of chemically active plasma, Moscow, Nauka, 1984. 10. H. Suhr, Plasma Chemistry and Plasma Processing, 9(1) (1989) 7S.


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CO2 Hydrate Kinetics in Electrolyte Solutions Containing Clay Minerals Jong-won Lee1, Jeasung Park1,, Moon-Kyoon Chun1, Kyung-min1 Lee, Jae-Hyoung Lee2, JooMyung Kang2 and Huen Lee1' Department of Chemical and Biomolecular Engineering, Korea Advanced Institute of Science and Technology (KAIST), 373-1, Guseong-dong, Yuseong-gu, Daejeon 305-701, Korea 2 The School of Civil, Urban & Geo-Systems Engineering Seoul National University, Seoul 151-742, Korea FAX: +82-42-869-3910. E-mail: [email protected] Understanding the phase behavior and formation kinetics of CO2 hydrate is essential for developing the sequestration process of CO2 into the deep ocean and its feasibility. Threephase equilibria of solid hydrate, liquid water, and vapor were determined for aqueous mixtures containing CO2 and NaCl/clay to examine the effect of both ocean electrolytes and sediments on hydrate stability. Due to the capillary effect by clay pores and inhibition effect by NaCl the corresponding hydrate formation pressure appeared to be a little higher than that required for simple and pure hydrate at specified temperature. In addition, the hydrate formation kinetics of CO2 in pure water and aqueous NaCl solutions with or without clay mineral were also measured at various conditions. 1. INTRODUCTION The thermodynamic properties of clathrate hydrates and phase equilibria of various hydrate-forming systems have been investigated extensively and compiled historically in the various sources. However, much effort has been focused on the formation and decomposition kinetics of gas hydrates by many researchers in very recent years. Englezos et al. measured the formation kinetics of gas hydrates made from methane, ethane, and mixtures of methane and ethane in the aqueous solution[l]. They also proposed a mechanistic model of the formation kinetics with only one adjustable parameter representing the rate constant for the hydrate particle growth. Dholabhai et al. obtained the formation kinetics of methane hydrate in various aqueous electrolyte solutions [2]. The mechanism for gas hydrate nucleation from ice and the kinetics of hydrate formation were proposed by Sloan and Fleyfel [3]. Bishnoi and Natarajan reviewed the formation and decomposition of gas hydrates [4]. Chun and Lee investigated experimentally the kinetics of carbon dioxide hydrate formation[5]. Shindo and coworkers studied the kinetics and mechanism of CO2 hydrate formation at the interface between adjacent semi-infinite reservoirs of liquid CO2 and water[6]. Teng et al. derived a kinetic model of hydrate formation on the surface of a CO2 droplet in high-pressure and lowtemperature water[7]. Recently, the ocean disposal of CO2 recovered from power plants in the form of CO2 hydrate has been focused on by many investigators as a means to mitigate the global warming[8]. The gas hydrate deposits in the deep ocean and in the permafrost regions have been significant as vast sources of natural gas, and it will be planned to produce the natural

210

gas from those deposits commercially. Therefore, comprehensive information on the formation and decomposition kinetics of gas hydrates as well as the thermodynamics and structural properties becomes essential in solving the related problems generated from the technological point of view. In these connections the purposes of this work were mainly to investigate the phase equilibria and the formation kinetics of the CO2 hydrate in the pure water and aqueous NaCl solutions. 2. EXPERIMENTAL The hydrate-forming reactor with an internal volume of about 140 cm3 was made of a type 316-stainless-steel. It had two sight glasses equipped at the bottom and one side of the reactor that allowed a visual observation of hydrate formation inside the reactor. A PCLabCard PCL-711B data acquisition system and a PC486 computer were used to get the flow rate signals from the mass flowmeter. The formation kinetic experiments of the CO2 hydrate were conducted under isothermal and isobaric conditions. Considering the unique phase behavior of a hydrate-forming system, the kinetic experiments were performed at pressures of 2.5, 3.0 and 3.4 MPa and temperatures of 273.15, 274.15 and 275.15 K. 3. RESULTS AND DISCUSSION Three-phase equilibria of hydrate, liquid water and vapor (H-Lw-V) for CO2-H2O-NaCl (or clay) system are measured. The salt effect on the formation of carbon dioxide hydrate was confirmed and H—Lw—V phase boundaries are given in Figure 1. As expected, the addition of NaCl inhibited hydrate formation.

Figure 1. Hydrate phase equilibria The experimental results obtained from the formation kinetics of the CO2 hydrates in the pure water and several aqueous NaCl solutions are plotted in Figure 2 in order to examine the effect of electrolyte concentration on hydrate formation rate at the conditions of 273.15 K and 3.4 MPa. The similar behavior was also observed at other temperature and pressure conditions. Even if only a small amount of NaCl exists in the water, the hydrate formation rate largely decreases when compared with that in the pure water.

211

Figure 2. Formation kinetics curves of CO2 hydrate in the various NaCl solution

Figure 3. Formation kinetics of CO2, water, NaCl and clay mineral mixture

For comparison, four different solutions of pure water were attempted to examine the influencing degree of electrolyte and porous medium and are shown in Figure 3. All solutions showed the similar pattern that the formation rate of CO2 hydrate decreased with time and kept nearly constant after 1 hr. Interestingly, the initial formation rate of the 3 wt % clay solution was found to be a little higher than that of pure water. This rate difference between two solutions can be understood from the fact that clay can act as a nucleation initiator and thus enhance the formation rate at the initial stage of hydration. For the 3 wt % NaCl solution, the final consumed amount of carbon dioxide were 0.0896 moles per mole of water and further its initial formation rate was found to be quite slower than that for the solution without salt. For the 3 wt % clay/3 wt % NaCl solution, the final carbon dioxide consumption appeared to be a little smaller than that of the 3 wt % NaCl solution without clay, but its initial formation rate increased because of accelerated nucleation by the presence of clay. 4. CONCLUSION Two kinds of experiments were carried to understand the physicochemical characteristics of CO2 hydrate in the deep ocean floor. First, three phase equilibria were determined for solution containing CO2, water and NaCl (or clay). The addition of NaCl inhibited hydrate formation as expected. The overall results of hydrate phase equilibria indicate that hydrate formation by injecting the liquefied CO2 into ocean sediments requires a higher pressure than that needed in pure water phase. Secondly, to secure one of the most important process information, the formation kinetic experiments of CO2 hydrate were attempted. The simple and unique experimental system was developed in this work to measure the formation kinetics of CO2 hydrate in the pure water and aqueous NaCl solutions with/without clay. Through prescribed experiments, clay was found to play a considerable role in increasing the initial formation rate by inducing nucleation as an acting inhibitor. On the contrary, the addition of electrolytes inhibited hydrate formation and furthermore reduced its forming rate. This work gives new information regarding the basic phase behavior and the kinetic formation of CO2 hydrate. These results can be used to determine the optimal condition for the sequestration of CO2 into the deep ocean.

212

ACKNOWLEDGEMENT This research was performed for the Carbon Dioxide Reduction & Sequestration Center, one of the 21st Century Frontier R&D Programs, funded by the Ministry of Science and Technology of Korea, and also partially supported by the Brain Korea 21 Project REFERENCES 1. Englezos, P., Kalogerakis, N., Dholabhai, P.D. and Bishnoi, P.R., Chem. Eng. Sci., 42(11), 2647 (1987). 2. Dholabhai, P.D., Kalogerakis, N. and Bishnoi, PR., Can. J. Chem. Eng., 71, 68 (1993). 3. Sloan, E.D. and Fleyfel, F., AIChE J., 37, 1281 (1991). 4. Bishnoi, P.R. and Natarajan, V., Fluid Phase Equilib., 117, 168 (1996). 5. Chun, M.-K. and Lee, H., Korean J. Chem. Eng., 13, 620 (1996). 6. Shindo, Y., Lund, P.C., Fujioka, Y. and Komiyama, H., Energy Convers. Mgmt., 34, 1073 (1993). 7. Teng, H., Kinoshita, CM. and Masutani, S.M., Chem. Eng. Sci., 50, 559 (1995). 8. Seo, Y.-T., Kang, S.-P., Lee H., Lee, C.-S. and Sung, W.-M., Korean J. Chem. Eng., 17(6), 659 (2000).


213

New catalysts for the conversion of urea into carbamates and carbonates with Cl and C2 alcohols Michele Aresta*a, Angela Dibenedetto", Cira Devitaa, Olga A. Bourovab, Oleg N. Chupakhinb a

Department of Chemistry and METEA Research Center, Universita degli Studi di Bari, via Celso Ulpiani 27, 70126 Bari, Italy; b

Ural State Technical University, 19, Mira St., 620002, Ekaterinburg, Russia

The use of urea as an active form of carbon dioxide is a way to substitute phosgene in the chemical industry. This approach is particularly useful when the synthesis of urethanes or organic carbonates is considered. So far, only long chain alcohols have been successfully used in such reactions. In this paper, new catalysts are described that are able to convert urea into carbamates and carbonates, using either methanol or ethanol to afford dimethyl- and diethylcarbonate, respectively. In the best conditions, 92% conversion of urea into carbamate has been observed, using either methanol or ethanol. The resulting urethanes can be further converted into the relevant carbonates in a second step. The two step reaction reveals to be particularly useful as it makes easier the separation process. 1. INTRODUCTION Nowadays, many research efforts are directed towards developing synthetic technologies that may replace toxic chemicals such as phosgene and dimethyl sulfate (dms) with environmentally friendly reagents in the chemical industry. Alkylcarbonates are good candidates for a number of applications and, among them, dimethyl and diethyl carbonates are receiving the greatest attention. Dimethylcarbonate (DMC) is a versatile reagent that can replace dms in methylation [1] and phosgene in carboxymethylation reactions [2], for the synthesis of Pharmaceuticals, agrochemicals, dyes and polymers. It can also function as a polar solvent or additive for fuels [3]. Diethyl carbonate (DEC) is an excellent solvent in the mid-boiling range and an intermediate for the synthesis of various pharmaceuticals. Also, both carbonates are used in transesterification reactions for the production of aromatic-carbamates and -carbonic acid esters such as diphenyl carbonate, or cyclic carbonates [4]. Carbamates are important final products and intermediates in the synthesis of fertilisers, pesticides and isocyanates, that are usually prepared by reacting alcohols with phosgene [5]. Recently, new methods for the production of DMC have been developed that do not employ phosgene and are based on the oxidative carbonylation of methanol with either copper catalysts [6], or also using palladium(II) as catalyst and NO as a promoter [7]. These methodologies suffer of some drawbacks essentially linked to the use of either chlorides, that may attack the metal parts of the reactors, or NO, that may form explosive mixtures with air. Therefore, alternative technologies are sought that may use safer operative conditions, such as those based on the use of CO2 for the direct carboxylation of alcohols [8-10]. This approach

214

appears quite interesting, but suffers thermodynamic limitations [10]. Among the innovative methodologies that have a low environmental impact, the use of urea appears quite interesting. As a matter of fact, urea can be considered an active form of carbon dioxide and has been shown to react with alcohols via a urethane intermediate [Eq. (1)] [11] ROH O=C(NH2)2

ROH H2NCOOR

-NH3

O=C(OR)2

(1)

-NH3

Essentially, primary or secondary long-chain alcohols have been used so far to form in a first step a urethane using metal salts like zinc acetate and lead acetate [12] as catalysts. The prolongation of the reaction or the addition of triphenylphosphine as cocatalyst leads to the formation of carbonates. [Eq. (2)] The second step may have both a low yield and selectivity because of a concurrent reaction in which the urethane is implied. The formation of isocyanuric acid is a major drawback.

g f ' O=C(OR)2 + NH3 H2NCOOR—

(2) R0H

> HNCO

->1/3(HNCO)3

As Eq. 2 shows, the urethane may decompose before reacting with the alcohol to form isocyanic and isocyanuric acid [12], which is observed using catalytic systems such as zinc acetate/triphenylphosphine. According to the literature data such reaction occurs at a much less extent when aluminium or cobalt species are used in combination with the suitable cocatalysts [11]. As reported above, the majority of papers describes the use of long chain alcohols such as 2ethyl-hexanol, 1-pentanol, cyclohexanol, 3-methylhexanol, isononyl alcohol, 6-methyl-2heptyl alcohol and recently the use of glycols has been demonstrated to afford cyclic carbonates [13]. Dibutyltin oxide, dibutyltin dimethoxide, diisobutylaluminum hydride/tributyl phosphite and triphenyltin chloride have been used as catalysts with an appreciable yield of carbonates [4]. So far, only very little information has been made available in the literature about the alcoholysis of urea with methanol or ethanol. Such reactions are of great importance in view of the large industrial use that DMC and DEC may have or considering their use as fuel additives. The objective of this work was to develop and to explore the role of new catalysts for the conversion of urea into DMC and DEC by reaction with methanol or ethanol, respectively. For comparison, we have used either Zn(acetate)2 or other catalysts analogous of those described in the literature to be active with long-chain alcohols. The new metal complexes used as catalysts in this work bear the ligand 3-(pyridyl-2)-5-cyano-6-phenyl-l,2,4 triazine or its derivatives. 1,2,4-triazines are the analogous of bipyridines that find an extended use in catalysis [14]. Triazines easily form complexes with transition metals. The synthesis and characterization of new complexes of 3-(pyridyl-2)-6-(4-tolyl)-1,2,4 triazine and 6-cyano-5phenyl-2(2'-pyridyl)-3,4-cyclo-pentenopyridine with zinc (II) chloride, of 3-(pyridyl-2)-5cyano-6-phenyl-1,2,4 triazine with copper (I), iron (II), zinc (II) chlorides, and of 3-(pyridyl-

215 2)-5-(2,2,2-tri-fluoroethoxy)-6-phenyl-l,2,4,-triazine and 3-(pyridyl-2)-5-(2,2,3,3-tetrafluoropropoxy)-6-phenyl-l,2,4-triazine with nickel(II) chloride are also reported. 2. EXPERIMENTAL PART 2.1. Materials All air sensitive compounds were manipulated under an atmosphere of dinitrogen. The catalytic reactions under pressure of CO2 were run in a stainless-steel autoclave of 100 mL. The gases N 2 (99.99 %) and CO2 (99.99 %) were purchased from Rivoira. Vacuum-inert-gas lines and syringe-techniques were utilised for handling air-sensitive compounds. The heterocyclic ligands of 1,2,4-triazine were synthesised according to the literature [15]. All other chemicals were commercial products. Solvents were purified following the standard literature procedure [16] and stored under nitrogen. IR spectra in the range 4000-200 cm"1 in Nujol with Csl or KBr plates were recorded on a Perkin-Elmer Model 883 spectrometer. GC analyses of carbamates and carbonates were performed using a Hewlett Packard HP 6850 (capillary column: 30 m, MDN-5S, layer film 0.25|j.m, FID) instrument and GC-MS analyses were carried out using a Shimadzu GCMS-QP 5050 (capillary column: 60 m, MDN-5S, layer film 0.25um) apparatus. The C, H, N elemental analyses were obtained using an elemental analyser mod. 1106 Carlo Erba. Metals were determined using the Perkin Elmer Atomic Absorption Spectrometer 3110. The chlorine analysis was carried out by means of combustion of a weighted sample followed by titration of the chloride with AgNC-3 10"2 M using a E436 Metrohm Herisau Potentiograph. *H NMR spectra of ligands were recorded on a Bruker 250 spectrometer and those of the complexes on a Varian XL 200 MHz instrument. 2.2. General method for the synthesis of complexes with the ligands: 6-cyano-5-phenyl-2(pyridyl-2')-3,4-cyclopentenopyridine (cppc), 3-(pyridyl-2')-6-(4-tolyl)-l,2,4-triazine (ptt) and 5-cyano-6-phenyl-3-(pyridyl-2')-l,2,4 triazine (cppt) To a solution of the ligand (1 mmol) in boiling dry acetonitrile under N2, the solution of ZnCl2 or CuCl (1 mmol) in dry acetonitrile was added dropwise. The mixture was refluxed for one hour under stirring. After cooling, a white solid was obtained, that was separated and washed with cold acetonitrile. The isolated complex was characterized by IR, 'H NMR, and C, H, Cl, N, Zn or Cu elemental analysis as reported below. 2.2.1. Characterization of ZnCl2(cppc) White crystals, (yield: 0.13g, 32.59%). IR (Nujol, KBr) 2240 (u, O N ) , 1535 (u, C=N) cm"'. 'H NMR (DMSO-d6,): 5 2.02-2.2 (m, 2H, Hcydopenten), 2.82-2.9 (m, 2H, Hcyciopenten), 3.53-3.6 (m, 2H, Hcyciopen,en), 7.46-7.53 (m, 6H), 7.95-8.03 (m, 1H, Hheteroaromatic), 8.22-8.26 (d, 1H, Hheteroaromatic), 8.71-8.73 (m, 1H, Hheteroaromatic). Anal. Calc. for CzoH^ClzNjZn: C=55.39 %; H=3.48%, Cl=16.35%, N=9.69%, Zn=15.07. Found: C=54.98%, H=3.58%, Cl=16.83%, N=9.47,Zn=15.0%. 2.2.2. Characterization ofZnCl2(ptt) Yellow crystals, (yield: 0.3087g, 68.26%). IR (Nujol, Csl) 1580 (u, C=N), 350, 332 (u, ZnCl) cm'1. 'H NMR (DMSO-d6): 8 2.41-2.46 (s, 3H, Hmethyi), 7.42-7.46 (d, 2H, Haromatlc), 7.637.66 (m, 1H), 8.03-8.21 (m, 3H), 8.48-8.52 (m, 1H, Hheteroaromatic), 8.82-8.84 (m, 1H, Hheteroaromatic), 9.52 (s, 1H, Hheteroaromatic). Anal. Calc. for C15H12Cl2N4Zn: C=46.85; H=3.14; Cl=18.44, N=14.57, Zn=17.0. Found: C=47.45, H=3.31, Cl=18.15; N=14.11, Zn=17.13 %.

216

2.2.3. Characterization ofCuCl(cppt) Violet crystals, (yield: 0.3525g; 51.02%). IR (Nujol, KBr) 2228 (u, O N ) , 1591 (u, C=C), 1526 (o, C=N) cm"1. Anal. Calc. for C15H9ClCuN5: C=50.28, H=2.53, Cl=9.9, N=19.54, Cu=17.73. Found: C=49.92, H=2.81, Cl=9.7, N=18.9, Cu=17.2 %. 2.3. Synthesis and characterization of FeCl2(cppt) 5-Cyano-3-(pyridyl-2)-1,2,4 triazine (0.5 g, 2 mmol) was dissolved in boiling dry acetonitrile under N2, the solution of FeCl2*1.5 THF (0.453 g, 2 mmol) in dry acetonitrile was added dropwise. The solution was refluxed under stirring for twenty minutes. The solvent was evaporated using a vacuum pump and the residual solid was washed with toluene and filtered to afford of 0.3958 g (yield=53.14 %) of the FeCl2(cppt) complex. Dark blue crystals. IR (Nujol, KBr) 2234 (u, O N ) , 1606 (u, C=C), 1578 (u, C=N), 1532 (o, C=N) , 1498 (u, C=N) cm"1. Anal. Calc. for C15H9C12FeN5: C=46.67, H=2.35, Cl=18.37, Fe=14.47, N=18.14. Found: C=46.08, H=2.63, Cl=17.9, Fe=14.5, N=18.05 %. 2.4. Synthesis and characterization of ZnCl2(cppt)2 To the boiling solution of 5-cyano-6-phenyl-3-(pyridyl-2')-l,2,4 triazine (0.5 g, 2 mmol) in dry acetonitrile under an atmosphere of nitrogen, the solution of ZnCl2 (0.263 g, 2 mmol) in dry acetonitrile was added dropwise. The reaction mixture was refluxed under stirring for 1.5 hours. The obtained solution was concentrated and kept at a low temperature (243 K) for one night. The formed precipitate was washed with cold acetonitrile (253 K) and isolated by filtration. After drying, 0.3187 g (yield=25.24 %) of the ZnCl2(cppt)2 were obtained. Brown crystals: IR (Nujol, KBr) 2249(u, O N ) , 1599(u, C=C), 1577(u, C=N), 1509(u, C=N), 1451(u, C=N) cm"1. Anal. Calc. for C3oHi8Cl2NioZn: C=55.03, H=2.77, Cl=10.8, N=21.39, Zn=21.39. Found: C=55.39 , H=2.55, Cl=11.4, N=15.52, Zn=20.8 %. 2.5. General method for the synthesis of complexes of 6-phenyl-3-(pyridyl-2)- 5-(2,2,2trifluoroethoxy)-l,2,4-triazine (pptt) and 6-phenyl-3-(pyridyl-2)-5-(2,2,3,3tetrafluoropropoxy)-l,2,4,-triazine (pptot) with nickel (II) chloride NiCl2 (0.26 g, 2 mmol) was dissolved in boiling absolute methanol (50 mL) under an atmosphere of nitrogen. The solution of the ligand (pptt or pptot) (2 mmol) in absolute methanol (20 mL) was added dropwise. The reaction mixture was refluxed under stirring for 45 minutes. The obtained solution was concentrated and kept at 243 K for one night. The unreacted nickel (II) chloride was removed by filtration. The solvent was evaporated under vacuum, and the residual solid was separated and washed with toluene to afford the relevant complex. 2.5.1. Characterization of NiCl2(pptt) Dark green crystals. Yield: 0.6892 g (74.23%) IR (Nujol, KBr) 1604(u, O C ) , 1577(u, O N ) , 1531(u, O N ) , 1495(u, O N ) , 1278(u, Cheteroaromatic-O), 1164(u, Caliphatic-O) cm"1. Anal. Calc. for Ci6H11Cl2F3N4Ni0: O41.61, H=2.4, N=12.13, C1=15.35, Ni=12.70. Found: 0 4 1 . 0 2 , H=2.39, N=12.21, Cl=15.68, Ni=12.87 %. 2.5.2 Characterization ofMCl2(pptot) NiCl2 (pptot) was obtained in a similar way from NiCl2 (2 mmol) and pptot (2 mmol). Green crystals. Yield: 0.48135 g (59.42%). IR (Nujol, KBr) 1630(u, O N ) , 1600(u, O N ) , 1531(u, O N ) , 1263(u, Cheteroaromatic-O), 1104(u, Caiiphatio-0) cm"1. Anal. Calc. for C17Hi2Cl2F4N4Ni0:

217

C=41.34, H=2.45, N=11.34, Cl=14.35, Ni=l 1.88. Found: C=41.2, H=2.38, N=11.42, Cl=13.9, Ni=11.5%. 2.6. General procedure for the catalytic reaction of conversion of urea into carbamates and carbonates in a single step. Urea (0.5g, 8.33 mmol) in EtOH (10 mL) and the catalyst (see Table 1) were heated in a stainless-steel autoclave (V = 100 mL) under a CO2 pressure as indicated in Table 1. After cooling, the amount of ethyl carbamate and diethyl carbonate were determined by GC-MS. Table 1. Catalyst, mg N Alcohol Pressure, Time, T, Yield, % Yield, % 10 mL h K urethanes DEC,DMC Cocatalyst, mg CO2, atm 393 1 MeOH 20 8 trace Trace Zn (OAc)2; 25 mg 20 22 393 trace 2 EtOH 60.8 Zn (OAc)2; 50 mg 20 17 433 83.12 trace 3 EtOH Zn (OAc)2; 56 mg 17 483 8.37 33.32 4 EtOH 20 Zn (OAc)2; 56.05 mg 2.9 20 65 483 10.39 5 EtOH Zn (OAc)2; 55.55 mg 21 483 91.39 1.4 6 EtOH 20 ZnCl2(cppt)2; 50.45 mg 20 17 483 ZnCl2(cppc), 49.1 mg 0.85 77.53 7 EtOH 20 17 483 trace 85.2 8 EtOH ZnCl2(ptt), 50 mg 62 393 FeCl 2 *1.5THF;31.8mg trace 9 MeOH 20 Trace 22 393 20 FeCl2(cppt); 48.35 mg trace 29.52 10 EtOH 11 EtOH 20 21 483 FeCl2(cppt); 55.3 mg 42.13 2.68 18 413 51.04 CuCl(cppt);53.1mg trace 12 EtOH 20 21 483 CuCl(cppt); 54 mg trace 13 EtOH 20 76.61 14 EtOH 20 17 483 NiCl 2 ,52mg 69.23 0.74 1.3 20 17 483 69.8 15 EtOH NiCl 2 ,55.5mg;NaH,23mg 20 84.43 5.1 16 EtOH 21 483 (PEt2Ph)2NiCl2; 58.3 mg 17 483 45.14 4.9 20 17 EtOH NiCl 2 , 49.05 mg; LiAlH4, 9.95 mg;PPh 3 , 103.4 mg (1:1:1) 17 483 20 20 5.1 18 EtOH NiCl 2 , 51.25 mg;LiAlH4, 37 mg;PPh 3 , 200.45 mg (1:2:2) 17 483 1 80.37 6.49 19 EtOH L1AIH4, 17.4 mg 17 483 9.86 20 50 NiCl2(pptt) 52.55 mg 20 EtOH 1 0.96 17 483 89.23 51.55 mg 1 85.1 0.95 17 483 21 EtOH NiCl2(pptot), 50.55 mg 65 413 74.4 trace 20 EtOH 22 Sc(CF3SO3)3 67.5 mg 18 483 75.4 trace 20 Sc(CF3SO3)3 33.75 mg; 23 EtOH PyridineO.OlmL 2.7. Isolation of the intermediate urethane from the reaction of urea with methanol or ethanol Urea (0.5g, 8.33 mmol) in EtOH (30 mL) or MeOH was heated in a stainless-steel autoclave (V = 100 mL) at 423 K under 1 atm of CO2 for 17 h. After cooling, the reaction mixture was dried under vacuum and the residue was treated with anhydrous diethyl ether to extract the urethane that was isolated and further reacted using the appropriate catalyst to afford the carbonate.

218

2.8. Conversion of the urethane into the relevant carbonate by reaction with the parent alcohol 2.5 g of the urethane isolated as reported in § 2.7 was reacted with the parent alcohol in a stainless steel autoclave in the conditions reported in Table 2 to afford the relevant carbonate that was isolated from the reaction mixture (the isolated yield is given). Table 2. N Alcohol,

mL 1 2 3 4

10 10 10 10

5

10

6

10

7

10

8 9 10 11 12

10 10 10 10 10

Catalyst, mg Cocatalyst, mg Zn(OAc)2, 52.05 mg LiAlH4, 21.25 mg L1AIH4, 14.5mg LiAlH4, 17.35 mg; PPh3, 123 mg LiAlH4, 14.8 mg; PPh3, 103.6 mg (C2H5CH(CH3)O)3A1, 432.2 mg (C2H5CH(CH3)O)3A1, 374.25 mg A1(OC2H5)3, 50 mg NiCl2(pptt), 51.5 mg ZnCl2(ptt), 16.25 mg ZnCl2(cppc), 49.87 mg ZnCl2(cppt)2, 65.3 mg

Pressure Time, T, h K CO2, atm 20 17 483 1 17 483 20 17 483 1 17 483

Yield, % Yield, % DEC DMC 2.2 2.1 18 6.4 19.6

20

17

483

6.3

1

65

423

Traces

1

17

483

18.5

1 1 1 1 1

17 17 17 17 17

483 483 483 483 483

11.5 29.3 9.7 14.0 25.3

15.5 30.4

27.2

3. RESULTS AND DISCUSSION The triazine derivative ligands easily react with the anhydrous metal halides FeCl2, CuCl, NiCl2, ZnCl2 to afford 1:1 adducts [or 1:2 in the case of ZnCl2(cppt)2] that have been fully characterized in the solid state and solution. The ! H NMR spectra of the free ligand and the relevant complex show that the bonding of the ligand to the metal occurs through the interaction of two nitrogen atoms, namely N-2 of the triazine system and N of the pyridyl moiety, that give rise to a penta-atomic ring upon reaction with the metal. (Fig. 1)

Figure 1. This is true also when the ligand bears a potential bonding site as the CN-group. In fact the IR spectrum of the Zn(cppc)Cl2 and Cu(cppt)Cl complexes do not show any sensible shift of the CN-vibration, showing that there is not co-ordination to the metal centre (See the

219

Experimental Part). All synthesized complexes have been checked for their catalytic activity in the alcoholysis of urea. We have used Zn(0Ac)2 as a test catalyst, as it has already been described to be an active promoter of the reactions discussed in this paper. Entries 1-4 (Table 1) show the effect of time and temperature on the reaction. It is quite evident that a temperature of at least 433 K is necessary in order the first step of the reaction may take place, and the conversion of the urethane into carbonate requires at least 483 K, confirming the findings of other authors [2]. These reaction conditions have been used in the rest of the experiments. Interestingly, when ZnCl2-complexes with the ligands cppt, cppc, and ptt (Entries 6-8) are used, the yield of urethane is greatly increased with respect to Zn(0Ac)2. That the ligands have a positive effect on the metal catalytic activity if much more evident with Fe (Entries 9-11). Cu(I) complex with the cppt ligand confirms the data observed with Zn and Fe. We have used anhydrous NiCk and found that it is active in the alcoholysis of urea. The addition of NaH does not improve the yield. Hydrides have been used by other authors to assist the proton-elimination from urea and form a M-NHCONH2 moiety that has been suggested to be the active species in the alcoholysis. We have evidence of the fact that the hydride in the reaction conditions is completely converted by the solvent alcohol into the alkoxo compound NaOR with formation of H2. As a matter of fact, a positive influence of the hydride has not been observed in our studies (Entries 14-15, and 16-18). Conversely, the Nicatalytic properties are improved when the diethylphenylphosphine complex is used, that seems to be more active than the triphenylphosphine, also in presence of L1AIH4 (Entries 1618). The latter per se shows a catalytic activity (Entry 19). It should be emphasized that in alcohol it is converted into Al(0R)3 and LiOR, the former being the real catalytic species, as shown by the use of Al(0R)3 alone. The Ni-complexes with the pptt and pptot fluorurated ligands (Entries 20-21) are also good catalysts for the first step of the alcoholysis of urea to afford the urethane, a reaction that is not affected by the presence of carbon dioxide. We have tried to carry out the reactions under carbon dioxide pressure in order to verify if it were possible to recycle ammonia that is released in the alcoholysis. So far we have not observed any beneficial effect of the presence of carbon dioxide. Naked ions like Sc3+ or its complexes with pyridine (Entries 22-23) do not represent a better catalytic system. Therefore, the complexes Ni(pptt)Cl2 (Entry 20b) and Zn(cppf)2Cl2 (Entry 6) represent the best options as they can afford yields of the order of 90% in urethane. Considered all the results, we have adopted a two step procedure for the synthesis of the carbonate, with a first step carried out at 433 K with isolation of the urethane and its further conversion into the carbonate at 483 K. The results reported in Table 2 show some interesting features. Al-trialkoxo species are able to catalyse the alcoholysis of urethanes (Entries 2-8). It is worth to emphasize that whatever alkoxo complex is used, it reacts with the excess methanol or ethanol to afford Al(0Me) 3 and Al(OEt)3, respectively. This has been proved with Al(iso-butoxo)3 that is converted into the above mentioned species in the relevant alcohol with release of iso-butanol measured by GCMS. Similarly, LiAlH4 converts into LiOR and A1(OR)3 as reported above. The effective catalyst is Al(OEt)3 in EtOH and Al(0Me) 3 in MeOH. CO2 has a negative effect on the reaction yield (Entries 2, 3 and 4, 5). Positive results are obtained using Zn and Ni complexes with the ligands pptt, ptt, cppc and cppt. In particular, Ni(pptt)Cl2 and Zn(cppt)2Cl2 show the most interesting results (Entries 9 and 12, Table 2). The reaction mechanism is not yet clear, but one can assume that the alcohol reacts in a concerted way with the co-ordinated urea to afford NH3 and the urethane (first step) or the carbonate (second step). (Fig. 2)

220

Figure 2. Then role of the ligands on the metal may be relevant to the charge delocalisation that may make the urea carbonyl carbon more positive and prone to the nucleophilic attack by the alcohol. In a concerted way, the proton may be transferred to the -NH2 group causing the release of ammonia. Studies are still in progress for the elucidation of the mechanism. REFERENCES 1. U. Romano, R. Tesei, M.M. Mauri, P. Rebora, Ind. Eng. Chem. Prod. Res. Dev. 9 (1980) 396. 2. M. Aresta, A. Dibenedetto, E. Quaranta, Green Chemistry 1 (1999), 237. 3. M. Pacheco, C.L. Marshall, Energy & Fuels, 11 (1997) 2. 4. E.N. Suciu, B.Kuhlmann, G.A. Knudsen, R.C. Michaelson, J. Organomet Chem., 556 (1998)41. 5. M. Aresta, E. Quaranta, CHEMTECH 27 (1997) 32. 6. U. Romano, F. Rivetti, D. Delle Donne, Appl. Catal. A: Gen. 221 (2001) 241 and references therein. 7. K. Nishihira, S. Tanaka, Y. Nishida, N. Manada, T. Karafuji, M. Marukami, Ube Industries, US Patent 5 869 729 (1999). 8. a) J.C. Choi, T. Sakakura, T. Sako, J. Am. Chem. Soc, 121 (1999) 3793; b) D. BallivetTkatchenko, H. Chermette, T. Jerphagnon in Environmental Challenges and Greenhouse Gas Control for Fossil Fuel Utilization in the 21 st Century, M.M. Maroto-Valer et al. Eds, Kluwer Academic/Plenum Publishers, New York (2002) 371; c) D. Ballivet-Tkatchenko, O. Douteau, S. Stutzmann, Organometallics 19 (2000) 4563. 9. K. Tomishige, T. Sakaihori, Y. Ikeda, K. Fujimoto, Catal. Lett. 58 (1999) 225. 10. M. Aresta, A. Dibenedetto and C. Pastore, Inorg. Chem., 42 (2003) 3256. 11. P. Ball, H. Fuellmann and W. Heitz, Angew. Chem. In. Ed. Eng 19 (1980) 718. 12. M. Paquin, Z. Naturforschg, 1 (1946) 518. 13. W.Y. Sao, G.P. Speranza US Patent 5349077 (1994). 14. G. Chelucci, R.P. Thummel, Chem. Rev. 102 (2002) 3129. 15. V.N. Kozhevnikov, D.N. Kozhevnikov, T.V. Nikitina, V.L. Rusinov, O.N. Chupakhin, M. Zabel, and B. Koenig., J. Org. Chem. 68 (2003) 2882. 16. D.D. Perrin, W.L.F. Armarego, D.R. Perrin, Purification of Laboratory Chemicals, 2nd Edition Pergamon Press: Oxford England, 1986.


221

Group 5 (V, Nb and Ta) element-alkoxides as catalysts in the transesterification of ethylene-carbonate with methanol, ethanol and allyl alcohol M. Aresta, A. Dibenedetto*, and C. Pastore Department of Chemistry and METEA Research Center, Universita degli Studi di Bari, via Celso Ulpiani 27, 70126 Bari, Italy Group 5 (V, Nb, Ta) element-compounds have been used as catalysts for the transesterification of ethylene carbonate with methanol, ethanol or allyl alcohol to afford the relevant dialkyl carbonates. Vanadium species are the most active among the catalysts tested, with niobium showing a close reactivity. Both show a long life-time and can be recycled several times without loss of activity. Tantalum derivatives are the less active in all tested reactions. V-catalysts compare well with the catalysts industrially used so far. The present and previous studies show that V and Nb compounds may be used as catalysts in a wide range of reactions for the synthesis of carbonates from carbon dioxide. 1. INTRODUCTION The trans-esterification reaction is a synthetic route that responds to the principles of atom economy, waste reduction at source, and safety. It is advantageously used to synthesize some particular carbonates (Scheme 1, routes a-b) or carbamates (Scheme 1, route g) that otherwise

RCH=CH2 + 1/2 O2 + CO2 \d

R'NHCOOR H 2 O |g n r a , k R'NH 2 p*- R 0 H e ^ 2 R O H + CO2

R.

R^

\ >

0

\T7

+ CO2 - 7 "

Y O

;

^ ^ (RO)2co R 0H

"

b

^

k . 2 ROH V ^ 2 R O H + (H2N)2CO i 2NH3

(R"O)2CO Scheme 1. Synthesis of cyclic and linear carbonates from carbon dioxide and their interconversion by transesterification. would require the use of conventional methods based on phosgene [1]. In particular, industrial plants produce diphenylcarbonate (DPC) by transesterification of dimethylcarbonate (DMC) with phenol [2]. Noteworthy, DMC is produced by oxidative carbonylation of methanol [3] [Cu] 2 CH3OH + CO + Vi O2

(CH3O)2CO + H2O

(1)

222

[Eq. 1] avoiding, thus, the use of phosgene. Ti(OPh)4 is used as catalyst in the transesterification that proceeds in two steps (Eq. 2a,b) with the formation of an intermediate, that is the mixed methyl phenyl carbonate (MPC) (Eq. 2). Ti(OPh) 4

Ti(OPh) 4 (a)

(b)

PhOH + DMC — ^ - * - MeOC(O)OPh + PhOH —v. r

MPC

MeOH

> DPC

(2)

* MeOH

Coupling the trans-esterification with the carboxylation of epoxide (Scheme 1, route c) [4] or with the one-pot synthesis of cyclic carbonates from olefins, oxygen and CO2 (Scheme 1, route d) [5], results in new phosgene-free ways to the synthesis of linear carbonates. The exploitation of the trans-esterification reaction for the synthesis of carbonates may receive a further push-forward if new routes to linear carbonates will be developed, like the direct carboxylation of alcohol (Scheme 1, route e) [6] and the reaction of urea with alcohols (Scheme 1, route f) [7]. Such reactions have had so far a strict alcohol-requirement: the carboxylation of alcohols requires methanol or ethanol, while the alcoholysis of urea performs much better with high boiling alcohols than with a Cl- or C2- alcohol. The trans-esterification reaction is also used for the synthesis of carbamates by carbomethoxylation of aliphatic and aromatic amines with DMC, DPC, and MPC (Scheme 1, route g) [le-h]. As a matter of fact, the trans-esterification reaction has a great potential, as Scheme 1 shows. Although it has been known for long, it has found seldom industrial application, because of the availability of the very economic phosgene-based processes. Because of the existing barrier to the further expansion of phosgene production, the trans-esterification process may play a new key role in the chemical industry. So far, various catalysts have been used [8], essentially applied to the conversion of alkylene carbonate. Recently, we have found that niobium(V) alkoxides are catalysts for the direct carboxylation of alcohols [9]. Now, we describe and compare the reactivity of V-, Nb- and Ta-(V) alkoxides in selected transesterification reactions, implying ethylene carbonate (EC) and aliphatic C1-C3 alcohols. 2. EXPERIMENTAL SECTION All products used in this work were RP Aldrich reagents. All solvents, and in particular alcohols, were purified, dried and stored using the literature methods [10]. The alkoxides ([Nb(OMe)5]2, [Nb(OEt)5]2, [Nb(OAllyl)5]2 and [NbO(OEt)3]2) were synthesized as previously reported [9,11]. 'H NMR and 13C NMR spectra were recorded on a Varian instrument (200 MHz) and IR spectra were obtained with a 883 Perkin-Elmer spectrometer. The reaction of trans-esterification was carried out in an autoclave with an inner volume of 70 mL, equipped with an electrical heating jacket and a valve for the continuous withdrawal of liquid samples. Qualitative analyses of the reaction mixture were done with a Shimadzu QP5050A GC-MS. For quantitative determinations, a GC-FID instrument equipped with a capillary column HP-5 was used, with naphthalene as internal standard. 2.1. Synthesis of [Ta(OMe)5]2 To a yellow suspension of TaCls in toluene (7.75 g, 21.64 mmol in 150 mL), 4.5 mL (111.23 mmol) of methanol was added dropwise obtaining a colorless solution. After 1 h of stirring, NEt3 (15 mL, 108.2 mmol) was added. [Ta(OMe)s]2 was isolated as a white solid by distillation under vacuum of the solution obtained after the filtration of the reaction mixture

223

(5.64 g, 77.6% yield). Anal calcd for C 5 Hi 5 Ta0 5 : C 17.87, H 4.49, Ta 53.83 %. Found C 17.7, H 4.3, Ta 53.0 %. Infrared in hexachlorobutadiene: bands at 2922 (s) (C-H stretching of terminal methoxo-group), 2820 (s) (C-H stretching of bridging methoxo-group), 1456 (m), 1159 (vs), 1109 (vs) (C-0 stretching of terminal methoxo-group), 1016 (s) (C-0 stretching of bridging methoxo-group), 812 (m), 722 (m), 499 (vs) cm"1. *H NMR: 4.66 ppm (singlet, terminal OCH3), 4.42 ppm (singlet, bridging OC// 3 ); integral ratio 3.9/1. 2.2. Synthesis of VO(OiPr)3 To a dark reddish solution of VOCI3 in benzene (6.53 g, 37,69 mmol in 150 mL) 8.5 mL of isopropanol (111.02 mmol) was added followed by NEt3 (15 mL, 108.2 mmol) obtaining a precipitate of NEt3*HCl that was filtered. VO(OiPr)3, a transparent viscous liquid, was isolated by distillation under vacuum of the filtered reaction solution (6.37 g, 69.3% yield). Anal calcd for C9H21VO4: C 44.27, H 8.67, V 20.86 %. Found C 44.0, H 8.5, V 20.7 %. Infrared (neat compound): bands at 2974 (s), 2930 (m), 2889 (m), 1464 (m), 1451 (m), 1380 (s), 1366 (s), 1326 (s), 1165 (m), 1103 (vs), 1004 (vs), 974 (vs), 859 (s), 795 (s), 748 (vs) and 666 (vs) cm"1. 'H NMR: 5.46 ppm (multiplet, CH), 1.66 ppm (multiplet, CH3). 2.3. Trans-esterification reaction: general procedure The molar composition of the reaction mixture used for the trans-esterification of ethylene carbonate with an alcohol was: alcohol/ethylene carbonate/catalyst =300/7/1. The temperature of the reaction was 420 K and the reaction time changed according to the alcohol. If the reaction was carried in presence of CO2, the same molar ratio as above was used and the autoclave was charged with CO2 up to 5 MPa at room temperature. In all reactions in which vanadium was used as catalyst, VO(OiPr)3 was charged, but the active species resulted to be VO(OR)3 with R depending on the alcohol ROH used. In fact, when an excess of the alcohol ROH was added to VO(OiPr)3 the exchange of alkoxo groups took place instantaneously with the quantitative release of isopropanol as demonstrated by the GC analyses of the mixture. 3. RESULTS AND DISCUSSION In this work, we have compared the activity of three classes of catalysts, namely pentalkoxo- [M(OR)5, M = Nb or Ta], oxo-trialkoxo-metal complexes [MO(OR)3, M = Nb, V], or metal-oxides [Nb 2 O 5 ,, Nb2O4 Nb2O3, NbO, V2O5, V2O4, V2O3, and TiO2] in the transesterification of ethylene carbonate with MeOH, EtOH or Allyl-OH. The choice of ethylene carbonate was due to the fact that it can effectively be synthetized from the relevant epoxide and CO2 using several catalysts [4], including Nb2Os or Nb(IV)- and Nb(V)-complexes. Alternatively, cyclic carbonates can be obtained by oxidative carboxylation of olefins with Nb2Os [5]. Therefore, it appeared of interest to determine whether or not Nb (and other Group 5 metals) were able to catalyze not only the synthesis of cyclic carbonates, but also their conversion into linear carbonates, affording, thus, a wide range of linear and cyclic carbonates by using a sequence of reactions not based on the use of phosgene. In the following paragraphs, we discuss the reactivity of Nb, V and Ta complexes, and compare their efficiency to the literature data relevant to other catalytic systems. It is worth to note that complexes V(OR)s are not stable and could not be isolated so far. Therefore, we have used [Nb(OR)5]2 and [NbO(OR)3]2 as test complexes, and compared the activity of stable V or Ta complexes with their Nb analogues. In the trans-esterification reaction, [Nb(OR)5]2 is more active than [NbO(OR)3]2 similarly to what we have found in the oxidative carboxylation of

224

olefins [9] (Scheme 1, Id). As it will be described below, the V-complexes show an interesting reactivity and reveal to be the best among Group 5 elements. 3.1. DMC synthesis from methanol and ethylene carbonate This reaction was used as a test for comparing the reactivity of the three metal derivatives. In fact the methanol/ethylene carbonate is the most reactive combination than any other and results to be a good test system. The results are reported in Figure 1, that shows that V0(0R)3 and Nb-alkoxides are more active than [NbO(OR)3]2 and Ta-alkoxides. In particular the turnover frequency per hour (TOF h"1) during the first 4 h is respectively 1 mol of DMC per mol of catalyst per h for the V complex, 0.9 for [Nb(OMe)5]2 and only 0.1 for [NbO(OMe)3]2 and Ta-systems. Figure 1 shows that an equilibrium is reached that corresponds to a conversion of 80% of the initial carbonate in the case of V- and Nb-catalysts. The reaction can be continued for hours without any significant change in the conversion. With Ta the reaction is much slower. If EC is reloaded after 7 hours, the conversion slightly increases, showing that the elimination of the glycol is a key step. The best performance observed is a turnover number (TON) of 5.22 with VO(OiPr)3. The catalysts maintain their catalytic activity for several cycles, can be isolated and reloaded without loosing their activity. On the basis of these data, and considering the low activity of Ta-species, we have tested only V and Nb alkoxides as catalysts in other trans-esterification reactions.

Figure 1. Profile of the transesterification reaction of ethylene carbonate with methanoi

Figure 2. Profile of the trans-esterification reaction of ethylene carbonate with EtOH to afford DEC or Allyl-OH to afford DAC.

3.2. Synthesis of Di-Ethyl-Carbonate (DEC) and Di-AUyl-Carbonate (DAC) from ethylene carbonate and the parent alcohol The reaction of ethanol with ethylene carbonate proceeds with a slower kinetics than the reaction of methanol. Figure 2 shows the reaction profile recorded for Nb and V. In this case the TOFh"1 is 0.57 for niobium and 0.75 for vanadium. According to literature data, DAC is prepared from DMC and allyl alcohol [13]. We have shown that ethylene carbonate can be also used in trans-esterification reaction. [Nb(OAllyl)5]2 shows a TOF number of 0.62. With respect to the use of DMC, our system shows a better performance in terms of lower amount of catalyst used and recyclability.

225

3.3. Utilization of metal oxides as trans-esterification catalysts As we have shown that NlÔs and V2O5 are catalysts in the oxidative carboxylation of olefins [14] or in the carboxylation of epoxides [15], we have tested such compounds in the transesterification of EC. We have also used oxides in various oxidation states. In particular, Nb(II, III, IV, and V), V(III, IV and V) and Ti(IV)-oxides were used for the transesterification of EC with methanol. All of the tested oxides have shown some catalytic activity, sometimes comparable with that of the alkoxides. Figure 3 shows the reaction profile using the various oxides as catalysts. The best performance was shown by V2O5 that was able to convert EC at a level of 60% after 3 hours, with a TOFh"1 equal to 0.95. V2O5 is as effective as T1O2 that is used in several industrial applications. Both oxides tend to react with the alcohol causing its oxidation to the relevant aldehyde with consequent catalyst reduction and deactivation. In fact, as Fig. 3 shows V2O3 and V2O4 are worse catalysts than V2O5. Among Nb-oxides, NbC>2 is the most active.

Figure 3. Profile of trans-esterification of EC with methanol (or ethanol) using metal-oxides. 3.4. Effect of the presence of CO2 We have previously shown that [Nb(OR)5]2 reacts with CO2 to afford the hemi-carbonate Nb[OC(O)OR](OR)4 [9]. In order to verify if such carboxylated species has catalytic properties, we have investigated the reaction of methanol, ethanol and allyl alcohol with ethylene carbonate in presence of CO2 and [Nb(OR)s]. We have found that CO2 does not have any influence on the TON, demonstrating that the emicarbonate that is formed under these conditions has not better catalytic properties than the Nb-pentalkoxo complexes. 4. CONCLUSIONS The results reported above show that all Group 5 metal-pentalkoxide, oxo-trialkoxide or oxides are catalysts in the trans-esterification of ethylene carbonate with various aliphatic alcohols. Among them, the V-compounds are the most active. VO(OR)3 is quite stable and can be recycled up to ten times. V2O5 is also active, but slowly reacts with the alcohol with consequent oxidation of the latter to aldheyde and reduction of the metal to V(IV)or(III) that are worse catalysts than V(V). Both VO(OR)3 and V2O5 show an activity that results to be almost equal to that of conventional catalysts as titanium, tin, lead, magnesium and thallium. Such finding is of interest for the versatility of V-catalysts that can promote different reactions such as the synthesis of cyclic carbonates and their conversion. In particular,

226

V0(0iPr)3 has been shown to be active in several carboxylation reactions, like the carboxylation of alcohols [16] and their trans-esterification, while V(V)-oxide and Nb(V)oxide are a good catalysts for the oxidative carboxylation of olefins [14] and the transesterification of the resulting cyclic carbonates. Such versability opens new safe routes to organic carbonates. REFERENCES 1.

2.

3. 4.

5. 6.

7. 8. 9. 10. 11.

12.

13. 14. 15. 16.

a) A. A. G. Shaikh, S. Sivaram, Chem. Rev. 96 (1996) 951 and references therein b) T. Tatsumi, Y. Watanabe, K. A. Koyano, Chem. Comm. (1996) 2281; c) B. M. Bhanage, S. Fujita, Y. Hcushima, M. Arai, App. Catal. A: General 219 (2001) 259; d) B. M. Bhanage, S. Fujita, Y. He, Y. Ikushima, M. Shirai, K. Torii, M. Arai, Cat. Lett. 83 3-4 (2002) 137; e) M. Aresta and A. Dibenedetto, Chemistry-A Eur. J. 8 (2002) 685; f) A. Dibenedetto, M. Narracci, M. Aresta, C. Fragale, Green Chemistry 4 (2002) 439; g) M. Aresta, A. Dibenedetto, E. Quaranta, Tetrahedron 54 (1998) 14145; h) M. Aresta, A. Dibenedetto and E. Quaranta, Green Chemistry (1999)237. a) G. Illuminati, U. Romano, R. Tesei (Snam Progetti SPA) DE Patent No 75-2528412 (1985); b) G. Illuminati, U. Romano, R. Tesei (Snam Progetti SPA) RO Patent No 75-82648 (1980); c) S. Fukuoka, M. Kawamura, K. Komiya, M. Tojo, H. Hachiya, K. Hasegawa, M. Aminaka, H. Okamoto, I. Fukawa, S. Konno, Green Chemistry 5 (2003) 497. a) U. Romano, R. Tesei, M. M. Massi, P. Rebora, Ind. Eng. Chem. Prod. Res. Dev. 19 (1980) 396; b) U. Romano, S. Chim. Ind. (Milan) (1993) 75. a) M. Aresta, A. Dibenedetto, L. Gianfrate, C. Pastore, J. Mol. Catal. A: Chemical 204-205 (2003) 245 and references therein; b) M. Aresta, A. Dibenedetto, L. Gianfrate, C. Pastore, Appl. Catal. A: General 255 (2003) 5. M. Aresta and A. Dibenedetto, J. Mol. Catal. 182-183 (2002) 399. a) M. Aresta, A. Dibenedetto, C. Pastore, XXI Congresso Nazionale della Societa Chimica Italiana, Torino, 22-27 Giugno 2003, IN CO 031; b) J.C. Choi, T. Sakakura, T. Sako, J. Am. Chem. Soc. 121 (1999) 3793; c) D. Ballivet-Tkatchenko, O. Douteau, S. Stutzmann, Organometallics 19 (2000) 4563. a) M. Aresta, A. Dibenedetto, C. Devita, O. Bourova, O. N. Chapukin, Studies on Surface Science and Catalysis this volume (2004) p xxx and references therein. S. Sakai, S. Furusawa, H. Matsunaga, T. Fujinami, J.Chem.Soc. Chem. Commun. (1975) 265. M. Aresta, A. Dibenedetto, C. Pastore, Inorg. Chem. 42 10 (2003) 3256. D.D. Perrin, W.L.F. Armarego, D.R. Perrin, Purification of Laboratory Chemicals, Pergamon Press, England, Oxford 1986. a) C.G. Barraclough, D.C. Bradley, J. Lewis, I. M. Thomas, J. Chem. Soc. (1961) 2601; b) D.C. Bradley, B.N. Chakravarti, W. Wardlaw, J. Chem. Soc. (1956) 2381; c) L. G. Hubert-Pfalzgraf, J. G. Riess, Inorg. Chim. Acta41 (1980) 111. a) L.I. Bobyleva, S.I. Kryukov, B.N. Bobyleva, A.G. Liakumovich, A.A. Surovtsev, O.P. Karpov, R.A. Akhmedyanova, S.A. Koveva, US Patent 1 781 218, (1992) (Yaroslavskij Polit. Institut); b) N. Limura, M. Takagi, H. Iwane, J. Ookago, JP Patent 07 267 944 (1995) (Mitsubishi); c) D.J. Darensbourg, M.W. Holtcamp, Coordination Chem. Rev. 153 (1996) 155; d) R. L. Paddock, T. Nguven, J. Am. Chem. Soc. 123 (2001) 11498. U. Romano (Anic S.p.A, Italy). DE Patent No 82-3219237. M. Aresta, A. Dibenedetto, I. Tommasi, Appl. Organomet. Chem. 14 (2000) 799. M. Aresta, A. Dibenedetto, C. Pastore, in preparation. M. Aresta, ICCDU VII, 12-16 October 2003, Korea, Seoul, 45.


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Hydroformylation with Carbon Dioxide using Ionic Liquid Media Ken-ichi Tominaga and Yoshiyuki Sasaki National Institute of Advanced Industrial Science and Technology 16-1 Onogawa, Tsukuba, Ibaraki 305-8569, Japan Ruthenium complexes dissolved in 1,3-dialkylimidazolium chlorides effectively catalyzed the hydroformylation of various kinds of alkenes with carbon dioxide to give the corresponding alcohols. Compared with the conventional reaction, the reaction proceeded in the biphasic system, where the chemoselectivity in the hydroformylation was improved due to the hydrogenation of aldehydes being occurred in preference to undesired hydrogenation of 1-hexene. The ionic phase was recyclable as a catalyst after extracting the products. 1. INTRODUCTION In recent years, much attention has been focused on Green Chemistry, which is a guide to design chemical products and processes that reduce or eliminate the use and generation of hazardous substances [1]. In the 12 principles of Green Chemistry, it is proposed that "synthetic methodologies should be designed to use substances that possess little or no toxicity to human health and the environment", and that "a raw material of feedstock should be renewable". Carbon dioxide (CO2) is one of such "green" raw materials. In the course of our study on utilization of CO2 as an alternative of highly toxic CO [2-8], one of attractive targets was hydroformylation, which was a widely used chemical process in industry and produced more than six million tons of chemicals every year. Although CO2 have been used as a reaction medium for hydroformylation [9], we have recently found the first example of its use as a reactant (eq. 1) [7]. (1)

This novel hydroformylation is based on the hydrogenation of CO catalyzed

228

Scheme 1

by tetranuclear ruthenium complex [4]. Its reaction mechanism is proposed as shown in Scheme 1. A key step is deprotonation of a hydride complex with halide anions to give an active complex to coordinate CO2. Subsequent electrophilic attack of protons converts the coordinated CO2 to CO. The reaction rate depends on proton affinity of the halide anions of additive salts and increases in the order of I" < Br" < Cl". Although it had been concerned to be difficult to apply this reaction to hydroformylation due to the simultaneous hydrogenation of substrates, we have found that this catalyst system was effective for the hydroformylation of cyclohexene with CO2, in which the corresponding alcohol was formed in the yield of 88 % [7]. However, the application of this reaction was almost limited to internal alkenes, since terminal alkenes were more susceptible to hydrogenation. Besides, the most effective solvent, NMP (N-methyl- 2-pyrrolidone), has so high boiling point that there were difficulties associated with separating products by distillation. In this context, we have preliminarily investigated the biphasic catalysis using ruthenium complex in ionic liquids and have found it to be promising to overcome these problems [8]. Here, we report a comparison of this improved reaction with a conventional reaction in NMP and the factors which cause the efficiency. 2. EXPERIMENTAL Catalytic experiments were carried out in a 50 mL stainless steal autoclave. Typically, a toluene suspension (5.0 mL) of Ru3(CO)i2 (1.0 mmol), [bmim]Cl (1.0 g), and alkene (5.0 mmol) was placed in the autoclave, into which CO2 (4.0 MPa) and H2 (4.0 MPa) were introduced. The reactor was then heated at 140 °C for 30 h with stirring. After the reaction, the reactor was cooled to room temperature and depressurized. The products were identified with GC-MS (Shimadzu QP-5050) and were quantitatively analyzed with GC (GL-Science GC-353). For the recycle experiments, the resulting mixture was evaporated to remove volatiles, followed by extracting products with diethylether (5.0mL x 3), and then the residue was used as a catalyst for the next run. The experiments for the conventional reactions were carried in NMP (8.0 mL). The conditions were the same as stated above except for using [PPN]C1 (0.4 mmol) in place of [bmimjCl.

229

3. RESULTS AND DISCUSSION In the preliminarily study, the hydroformylation of 1-hexene with CO2 was carried out using Ru3(CO)i2 as a catalyst precursor and two-phase system consisting of ionic liquids and organic co-solvents (eq. 2) [8]. The conversion of 1-hexene and reaction selectivity were considerably depends on the co-miscibility of these two phases, which is affected by many factors such as polarity, aromaticity, alkyl chain length, and so on. Among the two-phase systems investigated, the most effective one was toluene/[bmim]Cl (bmim = l-buthyl-3-methylimidazolium) system. The less co-miscible systems reduced the conversion and the more co-miscible ones reduced the chemoselectivity to the hydroformylation. Table 1 Hydroformylation of alkenes with CO2 in biphasic system and in NMPa Entry 1

1-hexene

2 3

Cyclohexene

a-methylstyrene

C

Alkane

Toluene/[bmim] C1

97

84

0

11

NMP

98

64

5

24

Toluene/[bmim] Cl

80

76

0

3

100

88

2

6

78

50

0

22

NMP

4 5

Solvents

Substrates

Yieldb / % Conv. /% Alcohol Aldehyde

Toluene/[bmim]Cl

6

Benzene/[bmim]Cl

88

62

2

22

7

NMP

100

56

1

39

Toluene/[bmim]Cl

100

47

0

49

Benzene/[bmim]Cl

100

60

0

38

NMP

100

10

0

90

8 C

9

10

Styrene

a

Reaction conditions are described in the experimental section. Based on the substrates introduced. cAt 120 °C.

(2) The representative results of the hydroformylation of various kinds of alkenes are summarized in Table 1. The biphasic catalysis in toluene/[bmim]Cl system could afford better yields and chemoselectivities in the hydroformylation of 1-hexene to heptanols (entry 1) than the conventional reaction in NMP (entry 2) due to undesired hydrogenation of the substrates being inhibited. Unfortunately, both reactions lacked the regioselectivity, so that the produced

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heptanols consisted of almost equimolar n-heptanol and 2-methyl-hexanol. Although the conversion was slightly decreased, the chemoselectivity in the hydroformylation of cyclehexene reached 95 % (entry 3). Styrene compounds were more susceptible to hydrogenation (entries 5 and 8), however, benzene/[bmim]Cl system enables the reaction to proceed at lower temperature to give the products in better yields (entries 6 and 9).

Fig 1. Time courses of the hydroformylation of 1-hexene with CO2 catalyzed by ruthenium complex in biphasic system (a) and in NMP (b). 9 = the yield of alcohol, = the yield of aldehyde, = the yield of hexane, = the amount of 1-hexene recovered. The reaction condidions are described in the experimental section. Figure 1 (a) shows the time course of hydroformylation of 1-hexene with CO2 in biphasic system. It is noteworthy that the yield of alcohol increased as the reaction proceeded, whereas the yield of hexane was not almost changed throughout the reaction, and that little amount of aldehydes were detected in the reaction. These results suggest that the hydrogenation of 1-hexene was occurred almost only in the initial stage, while the hydrogenation of aldehydes was readily occurred in preference to the hydrogenation of 1-hexene when they were formed. In contrast, in the conventional reaction in NMP (Figure 1 (b)), the aldehydes remained unhydrogenated, but the hydrogenation of 1-hexene proceeded throughout the reaction. Thus, the selectivity to hydroformylation appears to be correlated with the selectivity in the hydrogenation in each catalysis; in the ionic phase, the aldehydes were more susceptible to the hydrogenation than 1-hexene, so that the remaining 1-hexene could be hydroformylated selectively. In the homogeneous catalytic processes, one of major disadvantages is the separation of the products from the reaction mixture including catalysts. Recent development in the biphasic reaction using ionic liquids [10,11], however, enables product separation by extraction and recycling the catalyst phase. Also in this biphasic reaction, the produced heptanols could be extracted with diethylether after removing volatile compounds. The residue consisting of ruthenium complex and ionic liquids could be used as a catalyst phase for the next run. Figure 2 shows the results of catalyst recycling. Although the catalyst phase was recyclable,

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Fig 2. Recycle experiments for the hydroformylation of 1-hexene with with CO2a. a Reaction conditions are the same as described in the experimental section.

its catalytic activity was gradually decreased probably due to the accumulation of water formed by the conversion of CO2 and the catalyst leaching to the extraction phase. Our study is now in progress to improve the catalyst recycling. In summery, the present biphasic catalysis of ruthenium complex dissolved in 1,3-dialkylimidazolium chlorides has following advantages in the hydroformylation of alkenes with CO2; its chemoselectivity is improved due to the hydrogenation of aldehydes inhibiting the undesired hydrogenation of 1-hexene, and it enables product separation by simple extraction, after which the catalyst phase is recyclable. 4. ACKNOWLEDGMENT This study was supported by Industrial Technology Research Grant Program in '02 from NEDO of Japan. REFERENCES 1. P.T. Anastas and M.M Kirchhoff, Ace. Chem. Res., 35 (2002) 686. 2. K. Tominaga, Y. Sasaki, M. Kawai, T. Watanabe, and M. Saito, J. Chem. Soc. Chem. Commun., (1993) 629. 3. K. Tominaga, Y. Sasaki, M. Saito, K. Hagihara, and T. Watanabe, J. Mol. Catal., 89 (1994) 51. 4. K. Tominaga, Y. Sasaki, K. Hagihara, T. Watanabe, and M. Saito, Chem. Lett., (1994) 1391. 5. K. Tominaga, Y. Sasaki, T. Watanabe, and M.Saito, Bull. Chem. Soc. Jpn., 68 (1995) 2837. 6. K. Tominaga, Y. Sasaki, T. Watanabe, and M. Saito, Std. Surf. Sci. Catal., 114 (1998) 495. 7. K. Tominaga and Y. Sasaki, Catal. Commun., 1 (2000) 1. 8. K. Tominaga and Y. Sasaki, Chem. Lett., 33 (2004) 14. 9. W. Leitner, Ace. Chem. Res., 35 (2002) 746.

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10. T. Welton, Chem. Rev., 99 (1999) 2071. 11. J. Dupont, R. F. de Souza, and P. A. Z. Suarez, Chem. Rev., 102 (2002) 2002.


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Ionic liquid-derived imidazolium metal halides for the coupling reaction of ethylene oxide and CO 2 O-Sung Kwon,a Jelliarko Palgunadi,a Hyunjoo Lee,a Jin Yong Bae,a Byoung Sung Ahn,a Hoon Sik Kima'* and Sang-Hoon Leeb a

Reaction Media Research Center, Korea Institute of Science and Technology, 39-1, Hawolgokdong, Seongbukgu, Seoul, 136-791, Korea

Fax: +82-2-958-5859. E-mail: khsfg.kist.re.kr

b

Korea Institute of Enviromental Science and Technology, 613-2, Bulgwangdong, Eunpyeonggu, Seoul, Korea A series of ionic liquid-derived imidazolium metal halide complexes,

(l-R-3-

methylimidazolium)bMXaYb (M = Zn, Fe, Mn, In; R = CH3> C2H5, «-C4H9, CH2C6H5; X, Y = Cl, Br; a, b = 2, 3) were prepared by reacting metal halide, MXa with an equivalent of (1-R3-methylimidazolium)Y and their activities were tested for the coupling reactions of ethylene oxide and CO2 to produce ethylene carbonate. Among (l-R-3-methylimidazolium)bMXaYb tested, imidazolium zinc tetrahalide exhibited the highest activities. The catalytic activity of imidazolium metal halide was greatly influenced by the nature of halide ligands. 1. INTRODUCTION From the viewpoint of the preservation of environment and the utilization of cheap carbon resource, the chemistry of carbon dioxide has received much attention (1-5). The synthesis of cyclic carbonates by the coupling reactions of epoxides with carbon dioxide has been attracting increasing interest due to the enlarged application of cyclic carbonates in electrolytes, aprotic polar solvents, and raw materials in a wide range of chemical reactions (6). Accordingly, substantial literatures on the catalyst development for these coupling reactions have been reported (7-10). While the advances have been significant, there still remain many subjects to be improved in terms of catalytic activity, stability, and recovery of catalyst. Recently, there have been a considerable number of articles regarding the use of ionic liquids in separation, extractions, electrochemistry, and immobilization of homogeneous catalysts (11).

234

With a hope to develop highly active catalysts using ionic liquids, we have reacted imidazolium halides with metal halides to prepare metal halide complexes containing imidazolium cations. Herein, we report the synthesis and reactivity of a series of ionic liquidbased imidazolium metal halide complexes consisting of l-alkyl-3-methylimidazolium cations and metal halide anion. 2. EXPERIMENTAL 2.1 Chemicals All manipulations for the preparation of imidazolium metal halides were carried out under argon atmosphere unless otherwise stated. Solvents were freshly distilled before use according to the literature procedure. Ethylene oxide was purchased from Honam Petrochemical Co. and used as received. Propylene oxide, 1-methylimidazole derivatives, and metal halides were purchased from Aldrich Chemical company and used without further purification. CO2 was purchased from Sin Yang gas. The l-R-3-methylimidazolium halides were prepared according to the literature procedure (12). 2.2 Synthesis of (1,3-dimethylimidazolium)2ZnBr2Cl2A solution of (l,3-dimethylimidazolium)Cl (2.10 g, 15.83 mmol) in methylene chloride (25 mL) was mixed with a solution of ZnBr2 (1.78 g, 7.90 mmol) in tetrahydrofurane (25 mL) and refluxed for 3 h. After cooling to room temperature, the white precipitate was collected by filtration. Yield: 93.9 %; elemental analysis cald.(%) for CioH^B^CySUZn: C 24.49, H 3.70, Br 32.59, Cl 14.46, N 11.43, Zn 13.33; found: C 24.23, H 3.75, Br 31.83, Cl 14.18, N 11.18, Zn 13.43; All other imidazolium metal halides were similarly prepared. 2.3 Coupling reactions of ethylene oxide and CO2 The coupling reactions were conducted in a 200 mL stainless-steel bomb reactor equipped with a magnet bar and an electrical heater. The reactor was charged with an appropriate catalyst and an epoxide and pressurized with CO2 (1.4 MPa). The bomb was then heated to a specified reaction temperature with the addition of CO2 from a reservoir tank to maintain a constant pressure. After the reaction, the bomb was cooled to room temperature, and the remaining epoxide was removed using an aspirator. The product mixture was analyzed by Hewlett Packard 6890 gas chromatography equipped with a flame ionized detector and a DBwax column, and Hewlett Packard 6890-5973 MSD GC-mass spectrometry.

235

3. RESULTS AND DISCUSSION As shown

in equation

(1), a series of imidazolium

metal halides,

(l-R-3-

methylimidazolium)bMXaYb (M = Zn, Fe, Mn, In; R = CH3> C2H5, «-C4H9, CH2C6H5; X, Y = Cl, Br; a, b = 2, 3) were prepared by reacting MXa (X = Cl, Br) with (l-R-3methylimidazolium)Y (R = CH3, C2H5, ra-C4H9, CH2C6H5; Y = Cl, Br) at an ambient temperature or at reflux depending on the nature of halide groups.

MXa

+

b[ ^ © ^ I Y

[^N^N-ÎMX^

(1)

R = CH 3 , C 2 H 5 , «-C 4 H 9 , CH 2 C 6 H 5 ; X, Y = Cl, Br; a, b = 2, 3 M = Zn, Fe, Mn, In Interestingly, (l-R-3-methylimidazolium)bMXaYb was also obtained when MXa and (1-R3-methylimidazolium)X

were

reacted.

For

example,

(l-butyl-3-

methylimidazolium)2ZnBr2Cl2, can be obtained either from the reaction of ZnCl2 and (1butyl-3-methylimidazolium)Br

or

from

the

reaction

of

ZnBr2

and

(l-butyl-3-

methylimidazolium)Cl. The catalytic activities of various imidazolium metal halides were evaluated for the coupling reactions of CO2 and epoxides at 100 °C for 1 h. As shown in Table 1, imidazolium halides and metal halides were hardly able to catalyze the coupling reactions when they were used alone. However, the combination of metal halide with imidazolium halide exhibited much higher activity for the coupling reactions (entry 1-3). Among (l-R-3methylimidazolium)bMXaYb tested, imidazolium zinc tetrahalide exhibited the highest activities. The catalytic activities of imidazolium metal halides were greatly influenced by the nature of halide ions bonded to metal atom, suggesting the importance of the nucleophilicity of halide ions. The nucleophilicity of halide ions is likely to be enhanced by the presence of imidazolium cations. The dissociation of a halide ion and the following attack of the dissociated halide ion on the carbon atom of the coordinated ethylene oxide would take place more easily for more nucleophilic bromide ion. In contrast, the substitution on imidazolium cations was found to have a negligible effect on the catalytic activity for this reaction.

236

Table 1 Catalytic activities of imidazolium metal halides for the coupling reactions of CO2 and ethylene oxidea

a

Entry

Catalyst

m. p. 1;°c)

TOF(/!";)b

1

(l-butyl-3-methylimidazolium)Cl

73

10

2C

MXa

3

(l-butyl-3-methylimidazolium)2ZnBr2C]2

1 3 5 - 139

2112

4

(l,3-dimethylimidazolium)2ZnBr2Cl2

186 ~ 189

2697

5

(l-ethyl-3-methylimidazolium)2ZnBr2Cl2

167 ~ 170

2137

6

(l-ethyl-3-methylimidazolium)2ZnBr4

183 ~ 185

3588

7

(l-ethyl-3-methylimidazolium)2MnBr2Cl2

8 0 - 84

210

8

(l-benzyl-3-methylimidazolium)2MnBr4

Viscous liquid

2527

9e

(l-butyl-3-methylimidazolium)3FeBr3Cl3

7 4 - 77

697

10e

(l-ethyl-3-methylimidazolium)2FeBr4

Viscous liquid

923

11 e

(l-butyl-3-methylimidazolium)3lnBr3Cl3

4 8 - 50

790

n.r.d

Reactions were carried out at 100 °C and 3.5 MPa of CO2 for 1 h. Molar ratios of ethylene

oxide to zinc were set at EO/Zn = 5000. b TOF(h~'): moles of alkylene carbonate/moles of catalyst/hour.c MXa = ZnBr2, FeBr3, InBr3. d No reaction. e Molar ratios of ethylene oxide to metal were set at EO/M = 1000. Previous reports by Kisch on the synthesis of cyclic carbonates from CO2 and alkylene oxide suggested the parallel requirement of both Lewis acid-activation of alkylene oxide and Lewis base-activation of CO2 (13-15). In the case of imidazolium metal halide-catalyzed reaction, metal center is expected to act as a Lewis acid for the coordination of an alkylene oxide and halide ion is expected to act as a Lewis base to ring-open the coordinated alkylene oxide. The plausible mechanism for the coupling reaction catalyzed by an imidazolium metal halide [MX2Y2] is depicted in scheme 1.

237

Scheme 1. Proposed reaction mechanism for the coupling reaction of ethyl ene oxide with CO2 catalyzed by (l-R-3-methylimidazolium)2MX2Y2. 4. CONCLUSION A

series

of

ionic

liquid-based

imidazolium

metal

halide,

(l-R-3-

methylimidazolium)bMXaYb (M = Zn, Fe, Mn, In; R = CH3> C2H5, «-C4H9, CH2C6H5; X, Y = Cl, Br; a, b = 2, 3), prepared from the reaction of MXa with (l-R-3-methylimidazolium)Y were effective for the coupling reaction of epoxide and C 0 2 . The catalytic activity of (l-R-3methylimidazolium)bMXaYb was found to increase with increasing nucleophilicity of halide ion. But, the substitution on the imidazolium cation showed a negligible effect on the catalytic activity. ACKNOWLEDGMENT This work was financially supported by the Ministry of Environment of Korea and Honam Petrochemical Co.

REFERENCES 1. A. Behr, Carbon Dioxide Activation by MetalCompl exes, VCH, New York, 1988. 2. D.J. Darensbourg, M.W. Holtcamp, Coord. Chem. Rev., 153 (1996) 155. 3. H. S. Roh, A. W. Liu, H. S., Potdar, J. W. Kim, K. W. Jun, J. Ind. Eng. Chem., 9 (2003) 762. 4. Z. W. Liu, H. S. Roh, K. W. Jun, H. S. Potdar, M. Ji, J. Ind. Eng. Chem., 9 (2003) 576. 5. Z. W. Liu, H. S. Roh, K. W. Jun, J. Ind. Eng. Chem., 9 (2003) 267.

238

5. Shaikh, Chem. Rev., 96 (1996) 951. 6. H. S. Kim, J. J. Kim, H. N. Kwon, M. J. Chung, B. G. Lee, H. G. Jang, J. Catal., 205 (2002) 226. 7. R. Paddock, S. Nguyen, J. Am. Chem. Soc, 123 (2001) 11498. 9. H. S. Kim, J. J. Kim, B. G. Lee, O. S. Jung, H. G. Jang, S. O. Kang, Angew. Chem. Int. Ed., 39 (2002) 4096. 10. K.Yamaguchi, K. Ebitani, T. Yoshida, H. Yoshida, K. Kaneda K, J. Am. Chem. Soc, 121 (1999)4526. 11. R. D. Rogers, K. R. Seddon, "Ionic liquid: Industrial Applications to Green Chemistry", ACS symposium series 818, American Chemical Society, Washington DC, 2002. 12. M. Hasan, I. V. Kozhevnikov, M. R. H. Siddiqui, A. Steiner, N. Winterton, Inorg. Chem., 38 (1999) 5637. 13. W. Dumler, H. Kisch, Chem. Ber. 123 (1990) 277. 14. H. Kisch, R. Millini, I. J. Wang, Chem. Ber., 119 (1986) 1090. 15. M. Ratzenhofer, H. Kisch, Angew. Chem. Int. Ed. Engl., 19 (1980) 317.


239

Aliphatic polycarbonate synthesis by alternating copolymerization of carbon dioxide with cyclohexene oxide using heterogeneous zinc complex II Kim*, Min Ju Yi, Seung Hoon Pyun, Dae Won Park, Chang-Sik Ha Division of Chemical Engineering, Pusan National University, Jangjeon-dong, Geumjeong-gu, Busan, 609-735, Korea; e-mail:[email protected] Heterogeneous double metal cyanide (DMC) catalyst was prepared by using K3[Co(CN)6] and ZnC^, together with tertiary butyl alcohol and polytetramethylene ether glycol as complexing reagents. Multi metal cyanide (MMC) catalyst was also prepared by incorporating additional K4Fe(CN)6 in DMC catalyst. All catalysts characterized by various spectroscopic methods were utilized for the copolymerization of cyclohexene oxide and CO2. Both catalysts showed high activity to yield copolymers of narrow polydispersity and moderate molecular weight. MMC catalyst was much more reactive to CO2 than DMC catalyst, resulting in higher carbonate linkage. 1. INTRODUCTION Degradable polymers are materials that undergo chain-cleavage reaction in specific environments. In particular, biodegradable polymers, such as polycarbonates and polyesters, have draw considerable recent interest as materials in biomedical, pharmaceutical, and agricultural applications, as well as environmentally benign packaging.1' The most important commercial polycarbonate is poly(carbonyldioxy-l,4-phenylene) based on bisphenol A. Because of the toxicity of phosgene, the development of environmentally benign processes has attracted significant research. There has been considerable recent interest in the development of catalysts for the alternating copolymerization of carbon dioxide with epoxides to produce aliphatic polycarbonates.2' Because of the low cost and accessibility of the monomers and the attractive properties of polycarbonates, the development of new, efficient catalysts for this polymerization process is a significant scientific goal. In 1969, Inoue first discovered that ZnEt2/H2O catalyzed the copolymerization of propylene oxide and CO2 to give poly(propylene carbonate).3' Inoue's discovery has been the impetus of international efforts in obtaining more efficient catalysts for polycarbonate production. The bulk of this work occurred throughout the 1970s and early 1980s.3'4' Presently there is a commercial process based on this insoluble Zn catalyst.5' Double metal cyanide (DMC) catalysts based on Zn3[Co(CN)6] are well-known catalyst for the polymerization of epoxides and the synthesis of propylene oxide based polyether polyols which are used in a wide range of polyurethane applications.6' The catalyst was originally discovered by General Tire Inc. in the 1960s.7' DMC catalyst was revisited starting in the middle of 1980's, with improvements made by mainly ARCO.8' Recent improvements have made DMC catalysts much more attractive for commercial manufacture of polyether polyols. Compared with conventional KOH catalysts, DMC catalysts give high-quality PPG

240

products that have low level of unsaturation, narrow molecular weight distribution, and low viscosity.6' While DMC catalysts offer significant advantages for the ring-opening polymerization of epoxides, no reports have been published on the copolymerization of epoxide with CO2 thereby. In this report, we have investigated a possibility of DMC catalysts to utilize them as catalysts for the copolymerization. Preliminary results on the effect of the type of DMC catalysts and polymerization parameters on the activity and the structure of resulting copolymers are to be reported. Copolymerization behaviors of multi-metal cyanide (MMC) catalyst prepared by introducing additional Fe complex together with Zn and Co are also to be compared. 2. EXPERIMENTAL 2.1. Material All materials such as potassium hexacyanocobatate(III) (K3[Co(CN)6]), potassium ferrocyanide(II) trihydrate (K4Fe(CN)6 3H2O), Zn chloride (ZnCy, and tertiary butyl alcohol ('BuOH) were purchased from Aldrich and used without further purification. Cyclohexene oxide (CHO, Aldrich) was distilled and stored over Linde type 4A molecular sieves. Polymerization grade of polytetramethylene ether glycol (molecular weight = 1800; PTMEG) was donated by BASF Korea Ltd. 2.2. Preparation of catalysts The DMC catalyst using PTMEG as a co-complexing agent has been prepared according to the following procedures. K3[Co(CN)6] (3.32 g, 0.01 mol) is dissolved in distilled water (40 mL) in a beaker (solution 1). ZnC^ (13.63 g, 0.1 mol) dissolved in distilled water (100 mL) and 'BuOH (20 mL) in a second beaker (solution 2). A third beaker contains solution 3: a mixture of distilled water (1 mL), 'BuOH (20 mL) and PTMEG (3.5 g). Solution 2 is added to solution 1 over 60 min. at 50 °C with mixing using a mechanical stirrer. Solution 3 is then added and the mixture is stirred for 3 min. The mixture is centrifuged. The resulting catalyst cake is dried at 60 °C under vacuum (30 inHg) to a constant weight. The MMC catalyst using PTMEG as a co-complexing agent has been prepared by the same procedure except using K3[Co(CN)6] (1.7 g, 5mmol) and K4Fe(CN)6 3H2O (2.1 g, 5 mmol) mixture for solution 1. 2.3. Copolymerization of cyclohexene oxide with CO2 Copolymerization of CHO and CO2 was carried out by using 30 mL glass reactor at various temperatures. The reactor was charged with 5 mL of CHO monomer and catalyst (10 mg), and then purged with several time with CO2. The mixture was heated to desired temperature. The polymerization was stopped after 4 hr and white polymer powder was separated by precipitation followed by filtration. The polymerization yield was determined by gravimetry. 2.4. Characterization X-ray diffraction (XRD) patterns of the catalysts were obtained with a PJNT2000 wide angle goniometer 185 using Cu Ka radiation at 40 kV and 30 mA. X-ray photoelectron spectroscopy (XPS) analysis of the catalysts was performed on an ESCALAB 250 induced electron emission spectrometer with Al Ka(1486.6. eV, 12 mA, 20 kV) X-ray sources. IR spectra of the catalysts were obtained in transmission mode using a React IR (Asi Applied

241

system). ! H and 13C NMR spectra of the polymers were performed on a Varian Gemini 2000 & HP5P with CDCI3 as a solvent. 3. RESULTS AND DISCUSSION 3.1. Catalyst characterization In the present study of catalyst preparation aqueous solution of ZnCl2, K3[Co(CN)6] and K4Fe(CN)6 were combined. The resulting precipitate was combined with 'BuOH as a major complexing agent (CA) and PTMEG as a co-complexing agent. The resulting catalyst is expected to have the general formula: Zn3[Co(CN)6]2 ZnCl2 yH2O z CA or Zn3[Co(CN)6]2 Zn2Fe(CN)6 ZnCl2 yH2O z CA Even if the insolubility of this complex makes the analysis difficult, the surface formulation of the catalyst could be investigated by XPS. Atomic identification in an insoluble solid is possible using XPS to measure the energy required to create a hole in a core electronic structure. The results of XPS analysis of the catalysts are summarized in Table 1. The binding energy (1023.7 eV) of Zn atom of ZnCl2 shifts towards lower value by 2.7 for MMC catalyst and 3.7 eV for DMC catalyst after complexation. These chemical shifts results from the coordination of oxygen atom to Zn by reacting ZnCl2 with K3[Co(CN)6] in the presence of 'BuOH and lgepal CA-520 as CAs. It is interesting to note that the O/Zn value indicating a presence of CA increases from 0.37 (DMC) to 1.03 (nano-sized DMC), demonstrating K4Fe(CN)6 enhances the coordination amount of CA. The type and the amount of oxygen coordinated to Zn play an important role in the ring opening polymerization of epoxide since oxygen atoms coordinated Zn ion are real active centers.6' Co 2P3 of K3[Co(CN)6] shifts from 781 eV to 778 for DMC and 779 for MMC and Fe 2P3 of K4Fe(CN)6 shifts from 707.2 eV to 706.7 for MMC, by forming complexes with oxygen coordinated Zn. IR spectrum of v(CN) band from 2133.4 cm"1 (K3[Co(CN)6]) to 2195.9 cm"1 (DMC catalyst). MMC catalyst showed two v(CN) bands at 2148.4 and 2192.4 cm"1. Note that the v(CN) bands of K4Fe(CN)6 and of free CN" are 2098 and 2080 cm"1, respectively. The v(CN) shift to higher frequencies demonstrates that the CN" ion acts as not only a a-donor by donating electrons to the Co but also an electron donor by chelating to Zn metal. Alteration of the crystal structure can be shown by X-ray diffraction patterns to be dependent on the different surfaces. The X-ray patterns for DMC prepared in the absence of the organic complexing agent resembled that for highly crystalline Zn hexacyanocobaltate hydrate.6' The X-ray patterns for DMC and MMC catalysts, which are synthesized in the presence of 'BuOH CA and PTMEG co-complexing agent, exhibit broad signals at dspacings of 5.75, 5.07, 3.59, 2.54, and 2.28 angstroms, ascribing to a cubic lattice structure of Zn3[Co(CN)6]. This means that the catalysts are substantially amorphous. The crystallinity of the catalyst may strongly influence the catalytic activity. Table 1. ESCA characterization of catalysts Zn2p3 Compound

Co2p3

Fe2p3

Ols

N Is

C Is

Cl 2p

O/Zn

B.E.

[AT]

B.E.

[AT]

B.E.

[AT]

B.E.

[AT]

B.E.

[AT]

B.E.

[AT]

B.E.

[AT]

[AT]

(eV)

%

(eV)

%

(eV)

%

(eV)

%

(eV)

%

(eV)

%

(eV)

%

%

ZnCl 2

1023.7

K3[Co(CN)6]

-

. 781

. -

K,Fe(CN) 6

.

.

-

.

-

.

-

. -

-

707.2

DMC

1020

19.2

778

8.8

MMC

1021

4.7

779

0.8

706.7

0.7

530

7.0

397

20.5

283

42.0

197

2.4

0.37

530

4.9

396

10.0

284

73.1

197

5.9

1.03

242

3.2. Alternating copolymerization of cyclohexene oxide with CO2 Semi-batch copolymerizations of CHO and CO2 using DMC and MMC catalysts have been carried out at the temperature range between 40 and 100 °C at a constant pressure (100 psi). Copolymerization of CHO and CO2 leads to the following general structure:

If the alternating copolymerization is perfect, only carbonate groups should be generated. However, homopolymerization of CHO results in ether linkage. Each linkage can be estimated exactly by assigning ! H NMR spectrum of methine hydride in cyclohexane backbone. Table 2 shows the results of copolymerization. All catalysts show higher activity comparing to the results obtained by using a homogeneous diethylzinc-based catalyst2'. DMC catalyst shows higher activity than MMC catalyst; however, DMC catalyst shows much lower carbonate content (f = 0.22) than MMC catalyst (f = 1.44 at 80 °C) possibly due to starvation of active sites for CO2. These results demonstrate that the incorporation of iron complex into the DMC catalyst gives a favorable condition to generate active sites for CO2. CO2 reactivity decreases sharply at high polymerization temperature (say 100 °C by MMC catalyst). Both catalysts produce copolymers showing very narrow MWD and medium molecular weight. Table 2. CHO/CO2 copolymerization results obtained by DMC and MMC catalysts TON TOF (hr 1 ) 202.0 50.5 80 DMC 90 188.2 DMC 47.05 40 MMC 15.9 4.0 60 47.5 MMC 11.9 80 102.3 25.6 MMC 90 MMC 142.6 35.6 212.4 MMC 100 53.1 ''[carbonatej/tether] ratio determined by H NMR spectrum. Catalyst

T P (°C)

Mn 3413 3441 3580 4180 4550 4190 3500

M w /M n

.02 .03 .09 .14 .22 .17 .01

f]) 0.22 0.08 0.79 1.04 1.44 1.00 0.05

ACNKOWLEDGEMENT This work was supported by grant No. R01-2003-000-10020-0 from the Basic Research Program of the Korea Science & Engineering Foundation. IK is also grateful to the Brain Korea 21 Project in 2003. REFERENCES 1. E. Chiellini, R. Solaro, Adv. Mater., 8, 305 (1996). 2. (a) D.J. Darensbourg, J.R. Wildeson, J.C. Yarbrough, J.H. Reibenspies, J. Am. Chem. Soc. 122 (2000) (b) ARokicki, W. Kuran, J. Macrmol. Sci. Rev. Macromol. Chem. C21 135 (1981) 3. S. Inoue, K. Matsuura, J. Polym. Sci., Polym. Lett., B7, 287 (1969). 4. (a) E. Tsuchida, M. Kasai, Markromol. Chem., 181,1613 (1980). (b) L. Chen, H. Chen, J. Lin, J. Macromol. Sci., Chem., A24, 253 (1987). 5. S. Motika, U.S. Patent 5,026,676, 1991. 6.1. Kim, J. -T. Ahn, C. - S . Ha, C. S. Yang, I. Park, Polymer, 44, 3417 (2003). 7. (a) R. J. Herold, US Patent 3,278,459, 1966. (b) J. Milgrom, US Patent 3,404,109, 1968. 8. J. Hofmann, P. Gupta, H. Pielartzic, EP 0892002, 1999.


243

Alternating Copolymerization of Carbon Dioxide and Epoxide [1]. Aluminum Schiff Base Complex - Quartenary Ammonium Salt Systems as Novel Initiators. Hiroshi Sugimoto, Hiroaki Ohtsuka, and Shohei Inoue Faculty of Engineering, Tokyo University of Science, 1-3 Kagurazaka, Shinjuku-ku, Tokyo 162-8601, Japan Alternating copolymerization of carbon dioxide and epoxide was successfully catalyzed by the aluminum Schiff base complex - Lewis basic additive systems to produce a copolymer. 1. INTRODUCTION Carbon dioxide (CO2) is considered as an important raw material for organic syntheses, because CO2 is inexpensive, nontoxic, nonflammable, and abundant in nature. However, its synthetic applications, especially in industry, have been rather limited so far, including the syntheses of urea, salicylic acid, and some carbonates. Polymer synthesis using CO2 as a monomer is of particular interest, because the polymer is expected to carry ester groups in main chain and exhibit remarkable properties. The first example of polymer synthesis from CO2 was discovered by us using epoxides as the comonomer.1 The catalyst was the equimolar reaction product of diethylzinc and water. With this catalyst, the alternating copolymerization proceeded even at room temperature under a CO2 pressure of 30 - 50 atm (441 - 735 psi). Since this discovery, a variety of catalyst systems has been developed mainly based on the combination of diethylzinc and protic compounds in order to obtain systems with enhanced catalytic activity. In relation to organozinc catalysts, organoaluminum catalysts have also been investigated, because organoaluminums are less expensive than organozincs, and because some aluminum compounds are known as good catalysts for the homopolymerization of epoxides. However, the aluminum-based systems examined so far were less effective for the copolymerization of CO2 and epoxide than zinc analogues, and gave only the products rich in ether linkages even

244

in the presence of CO2. Among them, a few systems realize the alternating copolymerization of CO2 and epoxide, where the common feature is that the effective systems were the combination

of aluminum

compounds and additives, such as triethylaluminum

-

triphenylphosphine and aluminum porphyrin - quaternary phosphonium salt systems.

The present paper focuses attention on the aluminum Schiff base complex (1) - various additive systems, for the purpose of understanding the scope of the alternating copolymerization of CO2 and cyclohexene oxide (CHO) by using the aluminum-based systems.

2. EXPERIMENTAL Copolymerization of CO2 and CHO A typical procedure for the copolymerization of CO2 and CHO by the aluminum Schiff base complex (1) - tetraethylammonium acetate (Et4NOAc) was as follows: Into a stainless steel autoclave (150 mL) containing a CH2C12 solution (1 mL) of 1 (0.1 mmol) and Et4NOAc (0.1 mmol) under dry nitrogen, a prescribed amount of CHO (25 mmol) was added with a syringe in the nitrogen stream. The copolymerization was started by pressurizing the catalyst solution by CO2 up to 50 atm (52 kg/cm2, 735 psi). After the mixture was stirred for 24 h at 80 °C, the autoclave was cooled, and excess CO2 was discharged, and then to which was added methanol to stop the copolymerization. A small portion of the products was subjected to *H NMR and IR analyses for determining the polymer structure and to GPC for estimating the average molecular weights. Copolymerization of CO2 and CHO by other aluminum complex - ammonium salt systems was carried out similarly to the above. The results were summarized in Table 1.

245 3. RESULTS AND DISCUSSION Copolymerization of CO2 and CHO, carried out at 80 °C under 50 atm of CO2 by using aluminum Schiff base co mplex (1; 0.1 mmol) as the catalyst in CH2CI2 (1 mL) in the presence of the equimolar amount of Et4NOAc at the initial mole ratio [CHO]0/[l]o of 250 (CHO; 2.5 g), proceeded readily to afford a polymer with Ma and MJMa of 10,000 and 1.5, respectively, in 24 h (Table 1, run 1). From the IR and H NMR spectra, the obtained polymer was characterized as the copolymer of CO2 and CHO with the alternating sequence, of which carbonate linkage value as estimated by *H NMR was 94 % (3.4 g; yield 94 %).

Among the series of aluminum Schiff base complexes (2 - 5), 2 and 3 with ferf-butyl groups on the phenyl groups, also brought about the copolymerization to give the copolymer with the high content of carbonate linkages (90 % in both cases) in good yield in the presence of Et4NOAc (Table 1, runs 2 and 3), although the Mns of the produced copolymers were lower than that of the alternating copolymer obtained by the 1 - Et4NOAc system. In contrast, the copolymerization initiated by 4 having aliphatic diamine unit in the ligand and 5 with monoSchiff-base ligand in combination with Et4NOAc produced the copolymers with lower CO2 content in lower yield (Table 1, runs 4 and 5).

When Et4NOAc salt was replaced by Et4NCl orEt4NBr, the CO2 content of the obtained copolymer was still high, while the Mn of the copolymer was lower and the appreciable amount of cyclic carbonate was produced (Table 1, run 6 and 7). Pyridine was also effective for the alternating copolymerization of CO2 and CHO by 1 to afford the polycarbonate with the highest content of carbonate linkage (97 %) in the present study in good yield (94 %) under the similar conditions, although the Mn of the copolymer was slightly lower than that of the copolymer obtained by using 1 - Et4NOAc system (Mn = 8,400, Mw/Mn = 1.5) (Tablel, run 8). On the other hand, when triphenylphosphine was used in

246

Table 1. Copolymerization of CO2 and CHO by the Aluminum Complex - Additive Systems.a reactionti conditions run catalyst

additive

Product yield

carbonate linkages1

(%)

(%) 10,000

1.5

4,700

1.6

MJMnQ

1

1

Et4NOAc

94

2

2

Et4NOAc

72

96 90

3

3

Et4NOAc

83

90

2,900

1.8

4

4

Et4NOAc

69

85

1,400

2.1

5

5

Et4NOAc

28

70

900

3.1

6

1

EUNC1

97

89

6,200

1.7

7

1

Et4NBr

64

89

4,100

2.5

8

1

Pyridine

94

97

8,400

1.6

9

1

Ph3P

36

56

3,000

1.7

10

1

-

36

2

4,300

7.3

a

Under 50 atm of CO2, [CHO]0/[additive]0/[l]o = 250/1/1, in CH2C12 ([l]o = 0.1 mmol/mL), at 80 °C, in 24 h b Estimated by 'H NMR. C Estimated by GPC based on standard polystyrenes.

combination with 1, the Mn and the CO2 content of the copolymer were decreased considerably (Table 1, run 9). The use of Lewis basic additive is essential for the alternating copolymerization of CO2 and CHO, since 1 alone gave a copolymer rich in polyether linkage in low yield (Table 1, run 10).

REFERENCES 1. S. Inoue, H. Koinuma, and T. Tsuruta, J. Polym. Sci., Polym. Lett. Ed., 7 (1969) 287. 2. H. Koinuma and M. Hirai, Makromol. Chem., 178 (1977) 1283. 3. T. Aida, M. Ishikawa, and S. Inoue, Macromolecules 19 (1986) 8.


247

Alternating Copolymerization of Carbon Dioxide and Epoxide [2]. The First Example of Polycarbonate Synthesis from 1-atm Carbon Dioxide by Manganese Porphyrin. Hiroshi Sugimoto, Hiromistu Ohshima, and Shohei Inoue Faculty of Engineering, Tokyo University of Science, 1-3 Kagurazaka, Shinjuku-ku, Tokyo 162-8601, Japan The first successful example of the formation of polycarbonate from 1-atm carbon dioxide and epoxide was achieved by the alternating copolymerization of carbon dioxide and cyclohexene oxide with (porphinato)manganese acetate under the 1-atm pressure of carbon dioxide to give a copolymer with alternating sequence. 1. INTRODUCTION Chemical fixation of carbon dioxide (CO2) is of much interest and importance from the viewpoints of energy resources as well as biological photosynthesis. Since we discovered that the Zn-complex catalyzed alternating copolymerization of CO2 and epoxide to afford a linear polycarbonate , various catalysts, including Zn, Al, Mg, Cd, Cr, and lanthanoid complexes, have been developed for the controlled alternating copolymerization. However, in many cases, a high CO2 pressure, a high reaction temperature, and a long reaction time were needed to realize the activity of the catalyst and/or the high CO2 content in the obtained copolymer. In

the course

of our studies on the alternating

Q ^ C rf\ s

copolymerization with transition-metal complexes, we have

t/~X'^' Hiin-N^3'

found that manganese porphyrin, [(TPP)Mn(OAc), or 1],

^^y== w 0 = 10 (w 0 = 2.4) V : PEO 2K -i-PFOMA 13K , 1 lwt% w = 5 (w ' ' ° ° = 3 ' 8 ' ' " : PEO 2K -*-PFOMA l3K , 5wt °' %> w » = 10 (wo« = 6) D : PEO 0 , 5K -6-PFOMA 81C; ° - 5 w t % - w " = 2 0 (WC = 12)

w0 = 20 and 10. Similar trends were recently reported for a cationic PFPE surfactant in CO2 [8]. As wo and thus droplet size is increased, the interface become less curved about water, hence the natural curvature - based phase separation occurs at higher temperature. For PEO-6-PFOMA surfactants, the cloud point pressure did not change appreciably upon adding water in these w0 ranges. REFERENCES 1. J. L. Kendall, D. A. Canelas, J. L. Young, and J.M. DeSimone, Chem. Rev., 99 (1999) 543. 2. J. M. DeSimone, Science, 297 (2002) 799 3. H. Ohde, F. Hunt, and C. M. Wai, Chem. Mater., 13 (2001) 4130 4. J. D. Holmes, D. C. Steytler, G. D. Rees, and B. H. Robinson, Langmuir, 14 (1998) 6371 5. J. Y. Heo, Y. S. Jeong, J. T. Kim, and K. T. Lim, J. Ind. Eng. Chem, 10 (2004) 389 6. K. T. Lim, M. Y. Lee, M. J. Moon, G. D. Lee, S. S. Hong, J. L. Dickson, and K. P. Johnston, Polymer, 43 (2002) 7043 7. H. S. Hwang, J. Y. Heo, Y. T. Jeong, S. -H. Jin, D. Cho, T. Chang, and K. T. Lim, Polymer, 44(2003)5153 8. C. T. Lee, Jr., P. A. Psathas, K. J. Zigler, K. P. Johnston, H. J. Dai, H. D. Cochran, Y. B. Melnichenko, and G. D. Wignall, J. Phys. Chem. B, 104 (2000) 11094


255

Aliphatic polycarbonate synthesis by alternating copolymerization of carbon dioxide with cyclohexene oxide using |3-diiminate zinc complex Sang Mook Kim, II Kim*, Dae Won Park, Chang-Sik Ha Division of Chemical Engineering, Pusan National University, Jangjeon-dong, Geumjeong-gu, Busan, 609-735, Korea; e-mail:[email protected] The copolymerizations of carbon dioxide with cyclohexene oxide using Zn or Mn complexes were investigated in this study. The well-defined homogenous Zn and Mn catalysts could be efficiently synthesized by using (3-diiminate and pyridine alcohol ligands. The Zn complexes showed much higher yield and CO2 incorporation than corresponding Mn complexes. The Zn complexes gave high polymers with very high carbonate linkage (> 90%) and narrow polydispersity (< 1.3), indicating single active sites. 1. INTRODUCTION One of profitable utilization of CO2 is to convert its resources to polymeric materials and use them in industrial fields. Carbon dioxide is an ideal synthetic feedstock since it is abundant, inexpensive, nontoxic, and nonflammable. ' Although it is estimated that nature uses CO2 to make over 200 billion tons of glucose by photosynthesis each year, synthetic chemists gave had little success in developing efficient catalytic process that exploit this attractive raw material. But, the past three decades witnessed the great progress in fixation of CO2 into aliphatic polycarbonates since the creative work of Inoue in 1969. Inoue's discovery has been the impetus of international efforts in obtaining more efficient catalysts for polycarbonate production. Kuran, Soga,3' Darensbourg,4' and Beckman et al.5' investigated various catalyst systems for CO2/cyclohexene oxide (CHO) copolymerization. Those catalysts are noncrystalline and extremely insoluble. As a result the exact composition of the active sites of the catalysts is hard to determine. Recently Coates developed p-diiminate Zn complexes for preparation of aliphatic polycarbonate.1' The P-diiminate Zn complex showed significantly higher activity than previous catalysts and produced aliphatic polycarbonate with very narrow polydispersity. These homogeneous single-site Zn complexes can give different activity and structure of the resulting copolymer according to P-diimine ligand structure. In this paper, the effect of the type of ligand and metal obtained by modifying the bulkiness of backbone and/or aryl ring on the activity of the catalyst and on the structure of the copolymer. Especially preliminary results obtained by the complexes using the pyridine alcohol ligand and by the Mn complexes are to be described. 2. EXPERIMENTAL 2.1. General procedure

256

All reactions with air and/or moisture sensitive compounds were carried out under dry nitrogen using standard Schlenk line techniques. ! H NMR spectra were measured on a Varian Gemini 2000 & HP5P with CDC13 as a solvent. Gel permeation chromatography (GPC) analyses were carried out using a Waters-400 spectrometer using polystyrene as a standard and dimethyl formamide as a solvent. All materials such as Zn(II) acetate [Zn(OCOCH3)2], Mn(II) acetate [Mn(OCOCH3)2], and n-butyl lithium were purchased from Aldrich and used without further purification. Cyclohexene oxide (CHO, Aldrich) was distilled for 6 hours over sodium metal and stored over Linde type 4 A molecular sieves. Tetrahydrofuran (THF) and methylene chloride were refluxed over sodium and store over 4 A molecular sieves under dry nitrogen. 2.2. Preparation of ligands and catalysts The p-diimine ligands and pyridine alcohol ligands7' were prepared according to literature procedures. The Zn and Mn complexes were prepared by similar procedure under inert atmosphere. To a solution of ligand (1.28 mmol) in THF (10 mL), n-BuLi (1.6M in hexane, 0.88 mL, 1.41 mmol) was added dropwise at 0 °C. After reacting the mixture for 5 min at 0 °C, the solution was cannulated to a solution of Zn acetate (or Mn acetate) (0.24 g, 1.41 mmol) dissolved in THF (10 mL). After stirring overnight at room temperature, the suspension was filtered and the clear solution was dried in vacuum. The pale yellow (Zn complex) or pale brown (Mn complex) solid was recrystallized from methylene chloride at low temperature. The resulting metal complexes were characterized by 'H NMR spectra. 2.3. Copolymerization of CO2 with cyclohexene oxide The copolymerizations were performed in a 10 mL high-pressure reactor equipped with a thermometer and a magnetic stirrer. The reactor was charged with catalyst (3.7 X 10"5 mol) and CHO (3.7 X 10"2 mol) in a dry box, and then transferred to a bath controlled at a desired temperature. The vessel was pressurized to 7.0 MPa with CO2 and allowed to stir. After 2 h of polymerization, the copolymer was dissolve in minimum amount of methylene chloride and precipitate from excess methanol. The product was then dried in vacuum. 3. RESULTS AND DISCUSSION 3.1. Synthesis of ligands and complexes The P-diimine ligands are readily synthesized by refluxing 2 equiv of the desired aniline with 1 equiv 2,4-pentanedione in acidified ethanol, yielding the corresponding hydrochlorides (Scheme 1). Neutralization with aqueous carbonate solution gives desired ligand in yields up to 80 % after recrystallization. The pyridine alcohol ligands were prepared by addition of 2-lithiopyridine to the appropriate ketones, followed by aqueous workup (Scheme 1). One of the advantages of these ligands is that the effects of electronic and steric modifications on catalysis can be readily proved by systematic ligand variation. Facile synthesis of pyridine alcohol ligand and commercial availability of various types of ketone must be a key advantage. The introduction of various ligands will make resulting metal complexes show different copolymerization behavior. Deprotonation of the ligands with 1 equiv. n-BuLi in THF at 0 °C and subsequent reaction with metal [Zn(II) or Mn(II)] acetate yielded the desired metal complexes that can be recrystallized from methylene chloride (Scheme 1). Scheme 2 shows structures of representative Zn(II) and Mn(II) acetate complexes utilized for CO2/CHO copolymerizations

257

in this study. These complexes could be recrystallized from methylene chloride in the case of (3-diimine Zn(II) complexes and from toluene/hexane mixture in the case of pyridine alcohol metal complexes.

Scheme 1.

Synthesis of P-diimine and pyridine alcohol ligands and metal [Zn(II) and

Mn(II)] acetate complexes

Scheme 2. Structure of metal complexes utilized for the copolymerization of this study: (a) metal complexes bearing p-diimine (BDI) ligand and (b) metal complexes bearing pyridine alcohol (PyA) ligand 3.2. Alternating copolymerization of CO2 with cyclohexene oxide Semi-batch copolymerizations of CHO and CO2 using Zn and Mn complexes have been carried out at the temperature range between 30 and 70 °C at a constant pressure of CO2 (7.0 MPa). Copolymerization of CHO and CO2 leads to the copolymers of the following general structure: If the alternating copolymerization is perfect, only carbonate groups should be generated. However, homopolymerization of CHO results in ether linkage. Each linkage can be estimated exactly by assigning 'H NMR spectrum of methine hydride in cyclohexane backbone. Zn complexes bearing both p-diimine ligand and pyridine alcohol ligand (Scheme

258

4) produced aliphatic polycarbonate which has very high carbonate contents ( ^ 94 %) as shown in Table 1. Recently reported homogeneous bis(phenoxide) Zn complexes (say A in Table 1) and heterogeneous fluorinated Zn catalyst (B in Table 1) suffered from low catalytic activities, low CO2 incorporation (requiring high C0 2 pressure) or long reaction time. However, Zn complexes bearing (3-diimine and pyridine alcohol ligand can produce aliphatic polycarbonate with high molecular weight and narrow PDI under mild conditions. The same results are reported earlier.1' As the bulkiness of ortho aryl substituents on p-diimine ligand decreases, the catalytic activity disappears due to either the bis-ligated complexes or the stabilized dimeric complexes or the stabilized dimeric complexes.1'However, (PyA)ZnOAc complex showed quite high activity (89.3 h"1) even if ortho position of aryl ring is not substituted with bulky group. Further studies are on going to clarify the effect of the bulkiness of ortho aryl substituents on pyridine alcohol ligands on catalytic activity and on CO2 reactivity. In general, Zn complexes showed significantly higher activity and much better effectiveness for CO2 incorporation than corresponding Mn complexes. Table 1. Results of CO2/CHO copolymerizations Catalyst

Temp.

fC)

Pressure Reaction time (h) (MPa)

Mn(xl0"3) (GPC)

Mw/Mn (GPC)

TOF (h-1)

linkage 2.4 A 38.0 80 91 55.1 69 4.5 93 24 Bb 6.4 100 17.0 137.8 9.0 96 2 1.21 21.0 50 7.0 247 (BDI)ZnOAc 38 2 2.34 50 1.2 (BDI)MnOAc 7.0 1.7 94 50 2 9.8 7.0 1.28 89.3 (PyA)ZnOAc 31 2 7.0 2.94 1.1 50 0.9 (PyA)MnOAc 4 b 'Data for (2,6-Ph2C6H3O)2Zn(Et2O)2, A. ' Data for HO2CCH=CHCO2(CH2)2C6Fi3/ZnO, B.5) a

Acknowledgement. This work was supported by grant No. R01-2003-000-10020-0 from the Basic Research Program of the Korea Science & Engineering Foundation. IK is also grateful to the Brain Korea 21 Project in 2003 and the Center for Ultramicrochemical Process Systems.

REFERENCES 1. M. Chen, E. B. Lobkovsky, and G. W. Coates, J. Am. Chem. Soc, 120, 11018 (1998). 2. P. Gorecki, W. Kuran, J. Poly. Sci., Polym. Lett. Ed., 23, 299 (1985). 3. K. Soga, E. Imai, I. Hartori, Polym. J., 13, 407 (1981). 4. D. J. Darensbourg, N. W. Stafford, T. Katsurao, J. Mol. Catal. A: Chemical, 104, LI (1995). 5. M. Super, E. Berluche, C. Costello, E. J. Beckman, Macromolecules, 30, 368 (1997). 6. J. Feldman, S. J. McLain, A. Parthasarathy, W. J. Marshall, J. C. Calabrese, S. D. Arthur, 16, 1514(1997). 7.1. Kim, Y. Nishihara, R. F. Jordan, Organometallics, 16, 3314 (1997).


259

Synthesis of poly(DOMA-co-AN) by addition of carbon dioxide to poly(GMA-co-AN) and the miscibility behavior of its blends with PEI Jin-Hee Hura, Dong-Ok Lima, Kyung-Hoon Kima, II Kim\ Hee-Chul Woo\ and Dae-Won Park a * a

Division of Chemical Engineering, Pusan National University, Pusan 609-735, Korea Department of Chemical Engineering, Pukyung National University, Pusan 608-739, Korea

Poly(glycidyl methacrylate-co-acrylonitrile)[poly(GMA-co-AN)] was converted to a poly[(2-oxo-l,3 dioxolane-4-yl) methyl methacrylate-co-acrylronitrile][poly(DOMA-co-AN)] by the polymer reaction with carbon dioxide using quaternary ammonium salt as a catalyst. Quaternary ammonium salts of lager alkyl group and more nucleophilic counter anions exhibited better catalytic activity in the addition of CO 2 to poly(GMA-co-AN). Poly(DOMAco-AN) formed clear films when blended with polyetherimide (PEI). DSC analysis also showed that poly(DOMA-co-AN) containing blends were miscible in the whole composition ranges. The glass transition temperature (Tg) of the polymer blends closely matched the Fox equation. 1. INTRODUCTION The reaction of carbon dioxide with oxiranes to produce cyclic carbonates has been of great interest as a useful method for its fixation by chemical processes [1]. The synthesis of cyclic carbonates from the reaction of carbon dioxide with oxiranes has been widely studied by using Lewis acids, transition-metal complexes, organometallic compounds, and alkali metal salts in the presence of crown ethers of quaternary ammonium salts under high pressure [2]. The reaction mechanism and catalytic effects for the synthesis of five-membered cyclic carbonates under atmospheric pressure using low molecular weight or polymer-supported quaternary onium salts were also reported [3-6]. Blends of poly(DOMA) bearing cyclic carbonate groups with some commercial polymers have also been reported for a more versatile application of the polymer [7]. In this study, we synthesized poly(glycidyl methacrylate-co-acrylonitrile)[poly(GMA-coAN) from glycidyl methacrylate (GMA) and acrylonitrile(AN). Then, CO2 was incorporated into the polymer to form poly[(2-oxo-l,3-dioxolane-4-yl) methyl methacrylate-coacrylonitrile][poly(DOMA-co-AN)]. Finally, blends of poly(DOMA-co-AN) with polyetherimide (PEI) were prepared to test their miscibility. 2. EXPERIMENTAL A radical copolymerization of GMA (0.18 mmol) with AN (0.02 mmol) [poly(GMA-coAN)] was prepared in dimethyl sulfoxide (DMSO, 300 mL) using 2,2'-azobisisobutyronitrile (AIBN, 0.27 g) as an initiator at 60 °C for 24 h under nitrogen atmosphere, then the solution was poured into distilled water to give a precipitate. The copolymer was recovered using an excess of methanol, and dried in vacuum at 30 "C for 12 h. The copolymer composition of poly(GMA-co-AN) was determined from the ratio of area in the copolymer using the 'HNMR spectrum. The ratio of area for the copolymer peak is 76.3:26.7 (GMA: AN). The synthesis of a copolymer of DOMA and AN [poly(DOMA-co-AN)] from poly(GMA-co-AN) and CO2 was carried out using quaternary ammonium salts. 0.5 mmol of catalyst was introduced to a 250 mL four-neck semi-batch reactor containing the mixture of 5

260

g of poly(GMA-co-AN) and 100 mL of DMSO, and the solution was heated up to a desired temperature (100 °C). Reaction was started by stirring the solution under a slow stream of CO 2 (10 mL/min), and continued for 8 h. The yield of CO2 addition to poly(GMA-co-AN) is defined as the number of unit of cyclic carbonate group in poly(DOMA-co-AN) divided by the number of unit of epoxide group in poly(GMA-co-AN). To prepare blend films, weighed amounts of poly(DOMA-co-AN) and PEI with given composition were cast from 10 wt% solution in DMF. The films were dried under vacuum for 3 days at room temperature. Glass transition temperatures (Tg) were measured using a differential scanning calorimetry (DSC, Perkin Elmer) calibrated with pure indium as a standard. 3. RESULTS AND DISCUSSION 3.1. Synthesis of poly(DOMA-co-AN) The conversion of epoxide ring in poly(GMA-co-AN) to the five-membered cyclic carbonate group in poly(DOMA-co-AN) could be identified by 'H-NMR and FTIR spectra. The characteristic peaks are as follows; 4.0-4.5 (-OCH2-, in side chain), 5.0-5.2 (-HCO-, in cyclic carbonate), 4.5-4.8 ppm (-OCH/r, in cyclic carbonate). The IR spectrum of the poly(DOMA-co-AN) exhibited an absorption peak at 1800 cm" (C=O of cyclic carbonate), whereas this peak was not observed on the IR spectrum for poly(GMA-co-AN). The catalytic activity of a quaternary ammonium salt usually depends on the corresponding catalyst cation and counter anion [8]. In order to understand the effects of the cation structure in the reaction of poly(GMA-co-AN) and CO2, quaternary ammonium chloride catalysts of different alkyl cations (TOAC, THAC, and TBAC) were used at 100 °C. The yields of CO2 addition with catalysts of different alkyl chain lengths are summarized in Table 1. The total number of units of epoxide group in the poly(GMA-cc-AN) was 591. The amount of CO2 added to poly(GMA-co-AN) was calculated from the difference of molecular weight of poly(DOMA-co-AN) and poly(GMA-co-AN). The yield of CO 2 addition increased in the order TBAC < THAC < TOAC. The quaternary salt catalyst having large alkyl group is easily dispersed in organic solvent. Bulky quaternary salt, having longer distance between cation and anion, is generally known to exhibit higher activity in activating anions [7]. This explains why they are more effective in nucleophilic attack of the anion to the oxirane ring of poly(GMA-co-AN). Figure 1 shows time variant yield of CO2 addition for different cations of quaternary ammonium salts. The addition of CO2 into poly (GMA-co-AN) increased with increasing alkyl chain length of the cation of quaternary ammonium salts. In order to understand the effects of anions in the reaction of poly(GMA-co-AN) and CO2, quaternary ammonium salt catalysts of different anion (Cl", Br", I") were used. Table 1 also shows that the addition rate of CO2 increased in the order I" < Br" < Cl". When different halide ions are used for the quaternary ammonium salts in the reaction of poly(GMA-co-AN) and CO2, the solvation of the anion is an important factor for the liquid phase reaction [8]. In an aprotic solvent like DMSO, stronger solvation may be expected for a soft anion like I" than for a hard anion like Cl". Therefore, the order of nucleophilicity will increase in the order I" < Br' < Cl", which accords with the order of the reactivity of quaternary ammonium halides. It can be suggested that high nucleophilicity of Cl" in an aprotic solvent enhances the attack of the anion to epoxide ring of poly(GMA-co-AN) [4]. Figure 2 shows time variant yield of CO2 addition for the quaternary ammonium salts of different anions. One can confirm the highest CO2 addition yield for TBAC. 3.2. Blends of poly(DOMA-co-AN) with PEI In order to examine the degree of miscibility of the poly(DOMA-co-AN) and PEI blend, optical clarity was first investigated. Table 2 shows that all the poly(DOMA-co-AN) with PEI blends formed clear films, which seems to mean that the blends are miscible over the whole concentration range. For the detailed study of the miscibility polymer blends are to measure the glass transition temperature of the blend component polymers. The blend of different poly(DOMA-co-AN) compositions with PEI exhibited single glass transition temperature

261 Table 1 .Effect of catalysts on the yield of CO 2 addition to poly(GMA-co-AN). Unit Mna Cat. Material Mw/Mn Amount6 Epoxide CO2 Poly(GMA-coAN)

-

110,000

2.23

-

23,000 1.36 133,000 TOAC 18,000 1.45 128,000 THAC Poly(DOMA16,000 1.41 126,000 TBAC CO-Pi.rs) 10,000 1.53 120,000 TBAB 1.44 118,000 TBAI 8,000 a Number average molecule weight. b The amount of CO 2 added to poly(GMA-co-AN).

CO2 addition (%)

591

-

-

69 182 228 364 410

522 409 363 227 181

88.3 69.2 61.4 38.4 30.6

Table 2. Optical clarity and glass transition temperatures of the blends of poly(DOMA-coAN) with PEL Optical Composition of TE(°C) clarity poly(DOMA-c6>-AI 400 nm) of a THF solution containing 0.1 mM [Re(dmb)(CO)3]2, Re-Re bond homolysis takes place to form the Re(dmb)(CO)3(THF) monomer, a dmb anion radical with absorptions at 467 and 496 nm. The monomer reversibly reforms the dimer in =50 minutes after photolysis. Formation of the dimer is surprisingly slow, and is second order in [Re(dmb)(CO)3(THF)] with kd = 40 5 M"1 s"1 in THF (27). Activation parameters for the dimerization reaction of Re(dmb)(CO)3(THF) in THF are AH* = 11 kcal mol"1 and A&= -14 cal mol"' deg"1. Flash photolysis studies of compounds containing the M-M bond shows that visible excitation induces homolysis of the M-M bond with formation of 17e" metal radicals such as Re' (CO)s and CpM' (CO)3 which then dimerize to reform the parent species with a rate constant of 109 M"1 s"1 in solvents such as cyclohexane, CH3CN and THF (29, 30). In our case, the species formed by the homolysis is not a 17e" metal radical, but rather an 18e" six-coordinate Re(dmb)(CO)3(THF) with a dmb anion radical. If the dimerization reaction that we observe proceeds via a preequilibrium to the metal-centered radical, eq 6, then the unfavorable preequilibrium is the dominant barrier to the reaction. Reaction 7 proceeds rapidly due to the unstable nature of the six-coordinate 19e" species. Provided AS6 ~ 0, the equilibrium constant (Kg) may be estimated from AH6 = D H*/2 and K6 = exp(-DHV2RT)=10"4. Re(dmb')(CO)3(THF) :==! Re'(dmb)(CO)3(THF) Re" (dmb)(CO)3(THF) -+ Re' (dmb)(CO)3 + THF 2Re" (dmb)(CO)3 -> [Re(dmb)(CO)3]2

(6) (7) (8)

274

kd = (K6)2 k8 = 40 M"' s"1

(9)

Thus the extremely slow dimerization of Re(dmb)(CO)3(THF) may arise from the unfavorable equilibrium K^ (eq. 6). Assuming eq 9, then the dimerization rate constant (k8) by the metal radical (eq 8) is ~ 4 x 109 M"1 s'1, consistent with other metal 17e"systems. When photolysis experiments are carried out under 760 torr of CO2, CO is produced. However when the light is turned off, the dimerization competes with CO2 reduction, indicating that Re(dmb)(CO)3(THF) reacts with CO2 very slowly (kCo2 < 0.1 M"1 s"1 in THF). The reaction with 13CO2 was monitored by 'H NMR and 13C NMR in dvDMF. As shown in Figure 1, [Re(dmb)(CO)3]2(CO2), Re(dmb)(CO)3(OC(O)OH) and [Re(dmb)(CO)3]2(OCO2) were observed, together with formation of CO in 50 to 100% yield based on [Re(dmb)(CO)3]2. The CO yield increases with Re concentration consistent with the highly reactive nature of Re(dmb)(CO)3S, and with longer photolysis Figure 1. 'H NMR spectrum of reaction products of times consistent with the existence of the Re I (dmb-)(CO) 3 with CO2 in d7-DMF: dimerization reaction. Neither [Re(dmbXCO)3]2(CO2), [Re(dmb)(CO)3]2(OCO2), O [Re(dmb)(CO)3]2, D Re(dmb)(CO)3(OC(O)OH), and Re(dmb)(CO)3(OCHO) nor H2 was detected. A dmb. The predicted monomeric species, Re(dmb)(COWCO2), in eq 3 has not been observed in our experiments. The involvement of the dinuclear species, (CO)3(dmb)Re-CO(O)-Re(dmb)(CO)3, as an intermediate in the CO2 reduction system is unique. Each reduced dmb ligand donates one electron to the bound CO2. This is in contrast with the CoL+ (L = 5,7,7,12,14,14-hexamethyl-l,4,8,ll-tetraazacyclotetradeca-4,ll-diene) system, where the Co(I) center provides two electrons to the bound CO2 forming the Co(III) carboxylate (31, 32). The disappearance of [Re(dmb)(CO)3]2(CO2) to form [Re(dmb)(CO)3]2(OCO2) and Re(dmb)(CO)3(OC(O)OH) together with CO is first order in [CO2] with the bimolecular rate constant being 9.7 x 10 4 M"1 sat 25°C in DMF. Visible-light irradiation ( > 380 nm) by 150 W Xenon lamp can accelerate the reaction -50 times. The observed catalytic reactions are summarized in the scheme below.

275

Figure 2. Observed reactions for CO2 reduction with Re(dmb)(C0)3X 4. CONCLUSIONS Re(dmb)(C0)3S, prepared either by reductive quenching of the excited states of /ac-Re(dmb)(CO)3(CH3CN)+ or by homolysis of [Re(dmb)(CO)3]2 can react with CO2 to form a dimeric intermediate containing the Re-C(O)O-Re moiety. In Re(dmb)(CO)3S, the extra electron is located on the ligand and the equilibrium constant between the ligand- and metal-centered radicals (Re(dmb')(CO)3S and Re'(dmb)(CO)3) is 10"4. Thus the dimerization reaction of the ligand -centered radical via the metal radical is very slow compared to typical metal radicals of organometallic complexes. The reaction with CO2 is also slow, since the binding of CO2 requires an electron-rich metal center(s). ACKNOWLEDGMENT We thank Dr. Norman Sutin, Dr. Carol Creutz, Prof. Osamu Ishitani, Dr. Kazuhide Koike, and Prof. Kazutetu Shinozaki for valuable discussions. This work was performed at Brookhaven National Laboratory, funded under contract DE-AC02-98CH10886 with the U.S. Department of Energy and supported by its Division of Chemical Sciences, Office of Basic Energy Sciences. SK acknowledges financial support from the Mombu-Kagaku-sho, Japan, as a fellow for research-in-abroad.

REFERENCES 1. J. Hawecker, J.-M. Lehn, and R. Ziessel, J. Chem. Soc, Chem. Commun., 536 (1983). 2. J. Hawecker, J.-M. Lehn, and R. Ziessel, Helv. Chim. Acta 69, 1990 (1986). 3. C. Kutal, M. A. Weber, G. Ferraudi, and D. Geiger, Organometallics 4, 2161 (1985). 4. H. Hori, F. P. A. Johnson, K. Koike, K. Takeuchi, T. Ibusuki, and O. Ishitani, J. Chem. Soc, Dalton Trans., 1019 (1997). 5. H. Hori, F. P. A. Johnson, K. Koike, O. Ishitani, and T. Ibusuki, J. Photochem. Photobiol, A 96, 171 (1996). 6. K. Koike, H. Hori, M. Ishizuka, J. R. Westwell, K. Takeuchi, T. Ibusuki, K. Enjouji, H. Konno, K. Sakamoto, and O. Ishitani, Organometallics 16, 5724 (1997). 7. K. Koike, N. Okoshi, H. Hori, K. Takeuchi, O. Ishitani, H. Tsubaki, I. P. Clark, M. W. George, F. P. A. Johnson, and J. J. Turner, J. Am. Chem. Soc. 124, 11448 (2002).

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8. H. Hori, J. Ishihara, K. Koike, K. Takeuchi, T. Ibusuki, and O. Ishitani, /. Photochem. Photobiol.A 120, 119(1999). 9. J. Hawecker, J.-M. Lehn, and R. Ziessel, J. Chem. Soc, Chem. Commun., 328 (1984). 10. B. P. Sullivan, C. M. Bolinger, D. Conrad, W. J. Vining, and T. J. Meyer, J. Chem. Soc. Chem. Common., 1414 (1985). 11. A. I. Breikss and H. D. Abruna, J. Electmanal. Chem. 201, 347 (1986). 12. T. R. O'Toole, B. P. Sullivan, M. R.-M. Bruce, L. D. Margerum, R. W. Murray, and T. J. Meyer, J. Electmanal. Chem. 259, 217 (1989). 13. A. Klein, C. Vogler, and W. Kaim, Organometallics 15, 236 (1996). 14. F. P. A. Johnson, M. W. W. George, F. Hartl, and J. J. Turner, Organometallics 15, 3374 (1996). 15. B. P. Sullivan, M. R. M. Bruce, T. R. O'Toole, C. M. Bolinger, E. Megehee, H. Thorp, and T. J. Meyer, in "Catalytic Activation of Carbon Dioxide" (W. M. Ayers, ed.), Vol. 363, p. 52. American Chemical Society, Washington, DC USA, 1988. 16. T. Scheiring, A. Klein, and W. Kaim, J. Chem. Soc. Perkin Trans. 2, 2569 (1997). 17. N. Sutin, C. Creutz, and E. Fujita, Comments Inorg. Chem. 19, 67 (1997). 18. D. H. Gibson, X. Yin, H. Y. He, and M. S. Mashuta, Organometallics 22, 337 (2003). 19. D. H. Gibson and X. Yin, J. Am. Chem. Soc. 120, 11200 (1998). 20. B. P. Sullivan and T. J. Meyer, /. Chem. Soc, Chem. Commun., 1224 (1984). 21. D. H. Gibson and X. Yin, J. Chem. Soc, Chem. Commun., 1411 (1999). 22. J. P. Casper and T. J. Meyer, J. Am. Chem. Soc. 87, 952 (1983). 23. M. Wrighton and D. L. Morse, J. Am. Chem. Soc. 96, 998 (1974). 24. J. A. Riddick, W. B. Bunger, and T. K. Sakano, "Organic Solvents, Physical Properties and Methods of Purification." Wiley, New York, 1986. 25. T. Hamada, B. S. Brunschwig, E. Eifuku, E. Fujita, M. Korner, S. Sasaki, R. van Eldik, and J. F. Wishart, J. Phys. Chem. A 103, 5645 (1999). 26. D. W. Thompson, J. F. Wishart, B. S. Brunschwig, andN. Sutin, J. Phys. Chem. A 105, 8117 (2001). 27. Y. Hayashi, S. Kita, B. S. Brunschwig, andE. Fujita, J. Am. Chem. Soc. 125, 11976 (2003). 28. T. R. O'Toole, B. P. Sullivan, and T. J. Meyer, J. Am. Chem. Soc. I l l , 5699 (1989). 29. T. J. Meyer and J. V. Caspar, Chem. Rev. 85, 187 (1985). 30. S. L. Scott, J. H. Espenson, andZ. Zhu, J. Am. Chem. Soc. 115, 1789 (1993). 31. E. Fujita, L. R. Furenlid, and M. W. Renner, J. Am. Chem. Soc. 119, 4549 (1997). 32. E. Fujita and R. van Eldik, Inorg. Chem. 37, 360 (1998).


277

Electrochemical Reduction of CO2 at Alloy Electrode in Methanol Satoshi Kaneco,a>* Hiroki Yamauchi,a Hideyuki Katsumata,a Tohru Suzuki*3 and Kiyohisa Ohtaa a

Department of Chemistry for Materials, Faculty of Engineering, Mie University, Tsu, Mie 514-8507, Japan

b

Environmental Preservation Center, Mie University, Tsu, Mie 514-8507 The electrochemical reduction of CO2 in methanol-based electrolyte at low temperature

was investigated on copper alloy electrodes. 53 % with advance electrode.

The best methane current efficiency was

This research can contribute to the large-scale

manufacturing of fuel gases. 1. INTRODUCTION The chemical fixation

of carbon dioxide (CO2), by radiochemical, chemical,

thermochemical, photochemical, electrochemical, and biochemical procedures, has been of significant interest form both fundamental and practical viewpoints [1]. The electrochemical method appears to be a very suitable method for the conversion and reduction of CO2.

In

the electrochemical reduction of CO2, in water, at most metal electrodes the major reaction products were carbon monoxide and formic acid [2].

However, only copper has proven a

suitable electrode for the formation of hydrocarbons such as methane and ethylene, which can be used as fuel gases [2,3].

Azuma et al. [2] investigated the electrochemical reduction

of CO2 at a Cu electrode in 50 mmol-dm"3 KHCO3 aqueous solution, at 293 K, and obtained methane, ethylene, and ethane with Faradaic efficiencies of 17.8, 12.7, and 0.039 %, respectively.

Moreover, at 273 K, the current efficiency was 24.7 % for methane, 6.5 %

for ethylene, and 0.015 % for ethane.

Kyriacou et al. [3] described the formation

efficiencies in the electrochemical reduction of CO2 on Cu, in 500 mmol-dm"3 NaHCO3 solution, at 298 K, as follows: 19 % for methane; 11 % for ethylene; etc. Recently, many investigators have actively studied the electrochemical reduction of CO2 using various metal electrodes in organic solvents, given that organic aprotic solvents * Corresponding author: Phone: +81-59-231-9427. Fax: +81-59-231-9442, 9471 or 94 27. E-mail: [email protected]

278

dissolve much more CO2 than water [4-6].

Reduced products containing carbon monoxide,

oxalic acid, and formic acid were produced by the electroreduction of CO2 in dimethyl sulfoxide, N,N-dimethyl formamide, propylene carbonate, and acetonitrile.

However, even

at a copper electrode, no hydrocarbons were obtained in these organic solvents [6]. Methanol is a better solvent of CO2 than water, particularly at low temperature.

The

solubility of CO2 in methanol is approximately five times that in water, at ambient temperature, and eight to fifteen times that in water, at temperatures below 273 K [7]. Therefore, methanol has been industrially used as a physical absorbent of CO2 in the Rectisol method, at low temperature [7]. process.

Currently, over 70 large-scale plants apply the Rectisol

In addition, acidic gases such as SOX and NOX can be easily removed with the

chemical adsorption method using water.

Due to these two properties of methanol, the

direct electrochemical reduction of CO2 in methanol-based electrolyte is an advantageous choice.

Thus, we have investigated the electrochemical reduction of CO2 on pure copper

electrodes by using methanol as the solvent, at 243 K [8-12]. the formation of methane and ethylene was observed.

In methanol-based catholyte,

However, there is little information

on the electrochemical reduction of CO2 at copper alloy electrodes in the methanol. This study deals mainly with the electrochemical reduction of CO2 at copper alloy electrodes in methanol at low temperature. 2. EXPERIMENTAL The apparatus and experimental conditions for the electrochemical reduction of CO2 are shown in Table 1.

The electrochemical reduction of CO2 was performed in a home-made,

divided H-type cell.

An Aldrich Nafion 117-type ion exchange membrane (0.18 mm

thickness) was used as the diaphragm.

The cathode potential was measured with respect to

a silver rod quasi-reference electrode (QRE).

Methanol (99 %, Nacalai Tesque, Inc., Japan)

was used as received without further purification.

Water content in the pure methanol was

less than 0.01 % (confirmed by the Karl Fischer test).

Sodium hydroxide (96 %, Nacalai

Tesque, Inc.) was used as the ionophore in the methanol-based catholyte.

Monel (Cu: 33 %,

Ni+Co: 65 %, Fe: 2 %) and advance (Cu: 55-60 %, Ni+Co: 40-45 %, Mn: 240 nm would be consistent with not only an existence of tetrahedral coordinated Ti oxide species but also the presence of oligomeric and/or octahedral TiO character. As summarized in Table 1, the photocatalytic formation yield for CH4 on the TiSBA-15(0.25M HC1) increases two times larger than those of the TiSBA15(0.23M HC1) and the yield for CH3OH increases three times. These results clearly indicate that the increase of photocatalytic reaction yields strongly dependent on acidic synthetic condition is closely associated with a large modification of the energy level, resulting from the localized photo-excited state of the tetrahedral TiO species in mesoporous silica medium. As shown in these results, the acidic condition at the moment of TiSBA-15 synthesis plays a very important role to decide the environment of TiO species. The solubility of titania rise at very low pH less than 05, so aging in strong acidic solution leads to a highly dispersed tetrahedral TiO gels. This is the first clear indication that the acid strength in direct synthesis of TiSBA-15 control the dispersion of TiO gels in mesoporous silica medium. We believe that these highly dispersed tetrahedral TiO gels lead to a suppression of the nonradiative transfer of absorbed photon energy as well as to an enhancement of the activity of photoexcited Ti oxide species as suggested previously.7 Table 1. Photocatalytic reduction results of CO2 with H2O and band edge position for various porous titanosilicate photocatalysts CH3OH ,.,. .„. difference from yields of]aroducts photocatalysts . b valence band to (umol* mol -TiO^'-h 1 ) selectivity conduction band0 (eV) CH4 CH3OH (%) TiSBA-15 -„ 510 "C as shown Fig. 1 (b). Therefore, we can obtain kinetically controlled conditions and mass transfer limited condition in order to study the intrinsic

Fig. 1. The results of Arrhenius plot for dehydrogenation of ethylbenzene on VSb/Al catalyst. Symbols : ) Laboratory-scale reaction, (A) Bench-scale reaction. Reaction conditions: EB:CO2 ratio = 1:5, Time-on-stream = lh, W/F = 68.6g-cat h/mol, Catalyst loading : 50cc, Catalyst particle size : 4mm

336

Fig. 2. Yields of styrene and calculated H2O/H2 ratio as a function of CO2/EB ratio in the reaction. Symbols : ) Yields of styrene, (A) H2O/H2 ratio. Reaction conditions: Reaction Temperature (a) = 480 °C and (b) = 520 °C, W/F = 68.6g-cat h/mol, Time-on-stream = lh, Catalyst loading : 50cc, Catalyst particle size : 4mm reactions of dehydrogenation of ethylbenzene in bench-scale reaction. Fig. 2 shows the catalytic activity performance and calculated H2O/H2 ratio as a function of CO2/EB ratio in the reaction. The yield of styrene and ratio of calculated H2O/H2 based on reactant on EB to CO2 ratio from 1:0 to 1:10. In order to compare catalytic performance of intrinsically controlled reaction with that of reaction that was affected mass transfer limitation, the catalytic reactions as a function of CO2/EB ratio were carried out at different temperature, 480 °C and 520 °C, respectively. The temperatures were considered by the results of Fig. 1 (b). According to increase of EB to CO2 ratio, the yield of styrene and the ratio of calculated H2O/H2 were improved not only in intrinsic reaction regime but also in of mass transfer regime as shown in Fig. 2. However, the aspects of increase of calculated H2O/H2 ratio was different according to reaction temperature. As shown in Fig. 2 (a), H2O/H2 ratio in intrinsic reaction regime was formed plateau nearby CO2/EB = 5. On the other hand, in mass transfer regime, H2O/H2 ratio was gradually increased in proportion to the increase of CO2/EB ratio as shown in Fig. 2 (b). The formation of water as well as carbon monoxide implies that on the VSb/AkCh catalyst CO2 molecule dissociates CO and surface oxygen, which can abstract hydrogen from EB, and the results in the generation of water. Taking into account the formation of large amount of water, high activity for supported VSb/AkOs catalyst under CO2 carrier may be mainly ascribed to oxygen species dissociated from carbon dioxide molecule, thus due to the promotion of the oxidative process by carbon dioxide. Therefore, increase of H2O/H2 ratio according to increase the CO2 flow reveals that the concentration of CO2 on the oxygen deficiency was insufficient state. It means that mass transfer of CO2 strongly affect the intrinsic reaction. These effects of mass transfer were also observed in the results of catalyst particle size. Fig. 3 shows selectivity of styrene at 60% conversion as a function of catalyst particle size, and selectivity of styrene was decreased according to increase the

337

Fig. 3. Selectivity at 60% conversion as a function of particle size in the bench-scale reaction. Reaction conditions: 550°C, EB:CC>2 = 1:5, Time-on-stream = lh, W/F = 68.6g-cat h/mol, Catalyst loading = 50cc, Catalvst particle size = 4mm particle size of catalyst. This result illustrates that internal-diffusion strongly affect to the CO2-EBDH with increase of particle size. 4. CONCLUSION Antimony oxide was found to be an effective co-component of alumina supported vanadia oxide catalyst for the CO2-EBDH reaction to improve activity and especially time-on-stream stability. This beneficial effect of antimony can be explained with the increased dispersion of active VOx-component, enhanced redox properties of the systems and formation of a new mixed V-Sb oxide phase V1.1Sbo.9O4. In the bench-scale reaction, the reactions above 510°C were found to be in mass transfer limitation regime and intraparticle mass transfer resistances of CO2 were important for reaction performance. Moreover, enough concentration of CO2 on the catalyst surface needed to increase the yield as well as selectivity of styrene. ACKNOWLEGMENT This work was supported by the Korea Ministry of Science and Technology through the Research Center for Nanocatalysis (KN-0329) and Institutional Research Program. V.P. Vislovskiy gratefully acknowledge the financial support as a Brain Pool Program by KOFST. REFERENCES 1. F. Cavani and F. Trifiro, Appl. Catal. A, 1995, 133, 219.

338

2. V. P. Vislovskiy, J.-S. Chang, M.-S. Park and S.-E. Park, Catal. Comm., 2002, 3, 227. 3. J. S. Yoo, Catal. Today, 1998, 41, 409. 4.1. J. Doh, Y. I. Pae, and J. R. Sohn, J. Ind. Eng. Chem., 1999, 5, 161. 5. N. Mimura and M. Saito, Catal. Today, 1995, 121, 125. 6. Y. Sakurai, T. Suzaki, N. Ikenaga and T. Suzuki, Appl. Catal. A, 2000, 192, 281. 7. M.-S. Park, V. P. Vislovskiy, J.-S. Chang, Y.-G. Shul, J. S. Yoo and S.-E. Park, Catal. Today, 2003, 87, 205 8. J.-S. Chang, V. P. Vislovskiy, M.-S. Park, D.-Y. Hong, J. S. Yoo and S.-E. Park, Green Chem., 2003, 5 587.


339

Oxidative dehydrogenation of ethane with carbon dioxide over supported chromium oxide catalysts Min Jia, Do-Young Honga, Jong-San Chang3'*, Vladislav P. Vislovskiy3, Jin S. Yooa, and SangEon Parkb'* a

Catalysis Center for Molecular Engineering, Korea Research Institute of Chemical Technology (KRICT), P.O. Box 107, Yusung, Daejon 305-600, Korea

b

Department of Chemistry, Inha University, 253, Yonghyun-Dong, Nam-ku, Inchon 402-751,

Korea FAX: +82-42-860-7679. E-mail: [email protected], [email protected] Oxidative dehydrogenation of ethane with carbon dioxide as oxidant has been investigated over chromium oxide catalysts on ZrO2, CeO2 and Ce-ZrO2 supports. The modification of zirconia with ceria led to transformation of monoclinic zirconia phase to tetragonal one, enhancement of specific surface area, surface concentration of strong basic sites and substantially higher on-stream stability of CrOx/Ce-ZrO2 catalyst. 1. INTRODUCTION Currently, ethylene is produced by high temperature (> 800°C) processes including thermal dehydrogenation of ethane, which is cheap and abundant feedstock. Catalytic ethane dehydrogenation (EDH) also needs severe operating conditions for the achieving of economically attractive conversions and is endothermic and hence heat-deficient. However, oxidative dehydrogenation of ethane with molecular oxygen as an alternative option to produce ethylene is free from the equilibrium conversion restriction but it suffers from loss of selectivity for olefin due to over-oxidation. New approach to use carbon dioxide as mild oxidant for the EDH has been also pursued. Metal oxide catalysts (Ga2O3, CrOx, MnOx, VOX, CeO2, MoOx), both bulk and supported on several supports (oxidized diamond, SiO2, TiO2, A12O3, ZrO2, MCM-41, silicalite-2, H-ZSM-5) are reported recently for the CO2-EDH [1-3]. Among metal oxides, catalytic performance of chromium oxide has known to be highly dependent on the nature of the support [2]. hi this work, we report oxidative dehydrogenation of ethane over chromium oxide catalysts supported on ceria, zirconia and ceria-modified zirconia as supporting materials.

340

2. EXPERIMENTAL The zirconia (Z1O2) and ceria-zirconia Ceo.15Zro.85O2 (Ce-ZrO2) supports were from our previous study [4]. Cerium (CeC>2) and bulk chromium (CrOx) oxides were prepared by decomposition of corresponding nitrates. Supported catalysts with 5 wt.% CrO x based on Cr2O3 were synthesized by impregnation of supports with aqueous solutions of Cr(NO3)3 9H2O followed by drying and calcinations at 700 °C. Catalytic activity measurements were performed under the atmospheric pressure in a continuous flow system with a tubular fixed-bed quartz microreactor connected with on-line gas chromatograph. Quartz chips were placed above and below the catalyst bed to avoid homogeneous reactions in the reactor void volume. In the absence of the catalyst, the reactor completely filled with quartz chips showed negligible activity when tested in the conditions used in catalytic measurements (see the Table footnote). BET specific surface area

(SBET)

of the catalysts was measured from N2-adsorption

isotherms using a Micrometrics ASAP 2400 adsorption analyzer. Powder X-ray diffraction (XRD) analysis was performed with Ni-filtered CuKa radiation with a Rigaku 2155D6 X-ray diffractometer. X-ray photoelectron spectra (XPS) were obtained using a VG ESCALAB 210 spectrometer with aluminum anode. Coke amounts deposited on the used catalysts were measured using a thermal gravimetric analyzer (Setaram TG-DTA 92-12). A surface concentration of basic sites able to chemisorb CO2 was measured, after CO2-adsorption at room temperature in a conventional flow system, as amount of CO2 desorbed in a temperature interval 100-270 °C expressed as a TCD detected desorption curve area, a.u., per a sample surface area, m2. 3. RESULTS AND DISCUSSION Table 1 shows bulk and surface characteristics of supports, bulk chromium oxide and supported CrOx catalysts along with the initial yields of ethylene over them. Specific surface areas of both Ce-ZrC>2 support and CrOx/Ce-ZrO2 catalyst are essentially higher than those of zirconia and ceria as well as of CrOx materials on these single supports. Contrary to the initial ZrC>2, which was monoclinic, the combined Ce-ZrC>2 is found to form the solid solution of ceria in tetragonal zirconia phase [ 5]. No peaks of crystalline chromium oxide phase were found in the XRD patterns of all the supported CrOx catalysts, indicating that CrOx-component is well dispersed as amorphous oxide on the surface of supports. The observed XPS binding energy (BE) for Cr 2p3/2 near 576 eV points out the presence of Cr3+ species on Ce-ZrCh support in accordance with references [2,6]. However, higher BEs (577.6-577.8 eV) of Cr 2p3/2 on the surface of bulk CrOx and CrOx/ZrO2 samples indicate

341

Table 1. Catalytic and physicochemical properties of supports and bulk and supported CrOx systems SBET

Catalyst

2

(m /g)

XRD phase

Cr 2p3/2 (eV)a

Basic sites concentration, a.u./m2

Yield of (%)

Coke amount in used samples (wt.%)

CeO2

5

CeO2 (c)

Not detected

2.3

n.d.

ZrO2

15

ZrO2 (m)

-

0.87

1.7

n.d.

Ce-ZrO2

45

ZrO2 (t)

-

1.08

3.4

n.d.

CrOx

5

Cr2O3 (r)

577.6

2.17

12.5

n.d.

CrOx/CeO2

5

CeO2 (c)

n.d.

1.21

8.2

n.d.

CrOx/ZrO2

14

ZrO2 (m)

577.8

2.02

12.4

0.8b

CrOx/Ce-ZrO2

40

ZrO2 (t)

576.4

2.68

16.2

0.4c

Reaction conditions: catalyst weight = 0.2 g, T = 650°C, total flow rate = 30 ml, C 2 H 6 /CO 2 = 0.2 (molar ratio), time-on stream 30 min. a Binding energy of Cr 2p 3 / 2 level; b after 2 h on-stream;

c

after 6 h on-stream. Notation: c, cubic; m, monoclinic; t, tetragonal; r,

rhombohedral; n.d., not determined.

that the existence of Cr5+ or even Cr6+ species whose characteristic BE value is around 579 eV, cannot be excluded. All catalysts and supports studied have exhibited very high ethylene selectivity (> 97.5 %) with only small amounts of methane as a by-product. All supports were low active whereas bulk and supported chromium oxide displayed much higher initial yields of ethylene, which, however, were (except CrCVCeCh) rather similar (>12.5-16 %). These catalytic results show that the activity of the catalysts

studied

determined

by

is

mainly supported

chromium oxide species. At the same

time,

Ce-ZrC^-supported

CrO x system exhibited much more stable catalytic performance in the CO 2 -EDH compared to CrO x /ZrO 2 (Fig.

1). Strong deactivation of

CrOx/ZrO2 is believed to stem mainly from coke formation; coke Fig 1. Ethane conversions over CrOx/ZrO2 and CrOx/Ce-ZrO2 catalysts as a function of time-onstream under CO2 (solid symbols) and inert (hollow symbols) atmospheres in the ethane dehydrogenation

was

observed

amounts

on

in the

much

lower

surface

of

CrOx/Ce-ZrO2 even after longer

342

on-stream operation (Table 1). The presence of carbon dioxide causes clear promotional effect on stability of the CrOx/Ce-ZrO2 catalyst but does not influence the intrinsic catalytic performance of the CrOx/ZrO2 (Fig. 1). Better catalytic stability of the CrOx/Ce-ZrO2 system could be due to its higher CO2 adsorption capability related with surface basicity (Table 1). Modification of zirconia with ceria has already been reported to distinctly contribute to the enhancement in the CO2 adsorption [4]. The CCVdissociation and re-oxidation of partially reduced cerium oxide was observed [3,7]. Considering the inherent ability of cerium oxide in ceria-zirconia solid solution to store and release oxygen [8], it can be suggested that active surface oxygen species generated from CO2 can afford the enhanced catalytic stability in the CO2-EDH via the easier redox cycle between Cr +- Cr + species, hi addition, on the basis of our present and previous results for the dehydrogenation of ethylbenzene with CO2 [9], ceria in the CrOx/Ce-ZrO2 catalyst could be considered also as a structural stabilizer, keeping zirconia in tetragonal phase, which has higher specific surface area and is much less reactive in coke formation than monoclinic one of the unmodified CrOx/ZrO2 system. ACKNOWLEDGEMENT This work was supported by the Korea Ministry of Science and Technology through the Research Center for Nanocatalysis (KN-0329) and Institutional Research Program. M. Ji and V.P. Vislovskiy gratefully acknowledge the financial support as a Brain Pool Program by KOFST. REFERENCES 1. K. Nakagawa, M. Okamura, N. Ikenaga, T. Suzuki, T. Kobayashi, Chem. Commun., (1998) 1025. 2. S. Wang, K. Murata, T. Hayakawa, S. Hamakawa, K. Suzuki, Appl. Catal. A, 196 (2000) 1. 3. R.X. Valenzuela, G. Bueno, V. Cortes Corberan, Y. Xu, C. Chen, Catal. Today, 61 (2000) 43. 4. J. Noh, J.-S. Chang, J.-N. Park, K.Y. Lee, S.-E. Park, Appl. Organometal. Chem., 14 (2000)815. 5. H.S. Roh, K.W. Jun, and S.-E. Park, J. Ind. Eng. Chem., 9 (2003) 261. 6. B. Grzybowska, J. Sloczynski, R. Grabowski, K. Wcislo, A. Kozlowska, J. Stoch, J. Zielinski, J. Catal, 178 (1998) 687. 7. S. Sharma, S. Hilaire, J.M. Vohs, R.J. Gorte, H.-W. Jen, J. Catal, 190 (2000) 199 8. A. Trovarelli, Catal. Rev. Sci. Eng., 38 (1996) 439. 9. J.-N. Park, J. Noh, J.-S. Chang, S.-E. Park, Catal. Lett., 65 (2000) 75.


343

Oxidehydrogenation of Ethane with CO 2 over transition metal doped MCM-41 mesoporous catalysts Ying-Li Bi a , Vicente Cortes Corberan b , Hong Zhuang a and Kai-Ji Zhen d a

Department of Chemistry, Jilin University, Changchun 130023, P. R. China, e-mail: [email protected] b

Institute of Catalysis and Petroleumchemistry, CSIC, Marie Curie s/n, Cantoblanco, 28049 Madrid, Spain, e-mail: [email protected] Summary The catalytic behavior of transition metal-doped M-MCM-41 (M = Ni, Co, Cr) mesoporous materials for oxidative dehydrogenation of ethane (ODE) with CO2 was investigated. Crcontaining catalysts showed the best performance, being active already at 450 °C and reaching 51.2 % conversion with 94.5% ethene selectivity at 700 °C.

Keywords: Cr-MCM-41, ethane, oxidative dehydrogenation, mesoporous redox catalysts, use of CO 2 as oxidant Introduction Due to its relevance as one of the major responsible for the greenhouse effect, emission control and efficient utilization of carbon dioxide is attracting an increasing attention from academia and industry. Carbon dioxide can play a role as an oxidant, but its low reactivity requires using high temperatures. Ethene is the raw material for the production of valuable chemicals such as ethyl benzene, styrene, ethanol, acetaldehyde, acetic acid, etc. It is presently produced by the steam cracking process, but the recent increasing market demand could lead to the need of additional sources. Activation of ethane could be one of them, but requires very high temperatures. This drawback could make the use of carbon dioxide especially suitable as a selective oxidant for the oxidehydrogenation of ethane (ODE). Several catalytic systems have been reported to be effective for this reaction such as calcium-doped ceria [1] or supported and doped chromium oxide [2-4]. Most often used catalysts in literature are oxides or supported and doped oxides and there are catalyst lifetime problems. The introduction of transition metals in the framework of mesoporous MCM-type materials opens new ways to disperse and stabilize active oxidic species. However, catalytic research in the field of alkane oxidative activation on these materials has been almost exclusively devoted to V-containing catalysts, using molecular oxygen as the oxidant [5,6]. We report here, for the first time, the use of transition metal-doped M-MCM-41 mesoporous silicas as catalysts for the (ODE) with CO2. The effect of the homogeneous reaction, the effect of the transition metal dopant and the role of CO2 in the reaction are studied. The differences in catalytic behavior of various M-MCM-41 catalysts in the reaction are discussed on the basis of the nature of the transition metal M and the characterization results.

344

Experimental MCM-41 mesoporous materials doped with Cr, Co and Ni were prepared by direct hydrothermal synthesis using Ci6H33(CH3)3NBr (CTABr) ammonium salt as "supramolecular template": 1 g of this salt was dissolved in 26 ml H2O under stirring, the adequate amount of the metal nitrate to obtain between 5-14 at% contents of the dopant oxide was added into the solution, then 5 ml TEOS and 12 ml NH4OH were added and the mixture stirred for 4-5 h. The mixture was allowed to react for 24 h. The precipitate was filtered, dried and calcined in air at 550 C for 5 h. The calcined samples were pressed, crushed and sieved into granules of 0.25-0.42 size for catalytic tests. Low angle x-ray powder diffractograms (XRD) were obtained with a Rigaku D/max RB instrument at room temperature using the Cu Ka radiation. Surface areas were determined by the BET method using N2 as adsorbate. Elemental composition was determined by EPMA in an electronic microscope Philips SEM-515. Catalyst samples (0.5 g) were tested for the ODE with CO2 at 400-750 C in a fixed-bed tubular reactor, with on line GC analysis, using a mixture C2H6:CO2:He = 1:2:17 (total flow 6 1/h). These reaction conditions were selected to minimize the homogeneous gas phase reactions, based on previous studies [1]. Selectivity to ethene was calculated as ratio of formed ethene to converted ethane, in mol %. Results and discussion

Fig. 1. XRD patterns of Cr-MCM-41 catalyst: A) as prepared, B) after 19 h on stream, and C) after 41 h on stream.

Figure 1 shows representative XRD patterns of fresh and used Cr-MCM-41 samples. They show in all cases up to four of the characteristic peaks in the low range of 2 0 (< 10°), corresponding to the (100), (110), (200) and (210) reflections of a hexagonal cell analogous to the MCM-41-type mesoporous silicas. Surface areas and mesopores diameter depended on the transition metal cation introduced, ranging between 800 and 1200 m2/g, and 2.6 - 3.2 nm, respectively. All these results confirm the formation of the hexagonal mesoporous structure. Under the reaction conditions used, homogeneous reactions in the absence of catalyst (reactor filled with SiC bits) were detectable only above 700 °C, with an apparent activation energy (Ea) of near 100 kcal/mol. The homogeneous reaction of ethane and CO2 is actually a dehydrogenation of ethane, as the products were ethene, hydrogen and minor amounts of methane at the highest temperatures but no CO2 was converted. Cobalt-containing MCM-41s were little active, showing ODE activity only above

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Fig. 2.- ODE with CO2 on M-MCM-41 catalysts. Left: effect of temperature and transition metal: rhombs, 14 at% Cr-MCM-41, squares, 8 at% Co-MCM-41, circles, homogeneous reaction. Right: rate of formation of ethene and CO on Cr-MCM-41 catalyst. Reaction conditions in text. 600 °C (Fig. 2), with an Ea = 53 3 kcal/mol, being the main product ethene and minor amounts of CO, almost two orders of magnitude lower. This indicates that the reaction is mostly a non-oxidative dehydrogenation, without the participation of CO2. The best ethene selectivity (99 % at 35 % ethane conversion) was obtained at 700 °C with sample containing 3 at.% Co. The activity varied very little when increasing the Co contents from 3 to 8 at.% Co. On the contrary, the main product of the reaction over Ni-MCM-41 samples was CO, which formed already at 450 °C and reached 57 % yield (referred to ethane) at 600 °C, and minor amounts of methane: besides, little ethene was formed ( C2H4 + CO +H2O (1) and the unselective oxidation to CO: C2H6 + 5 CO2 -> 7 CO + 3 H2O (2) So, the observed formation of only CO might indicate either that only the unselective reaction proceeds or if ethene is formed, it is further transformed into coke, as often observed on Nicontaining catalysts. The blackening and deactivation of the samples points to the latter. Interestingly, Cr-MCM41 catalysts were also active already at 450 °C but in this case the main products were ethene and CO, and the Ea of ODE much lower (around 20 kcal/mol).

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Selectivity to ethene was high, always above 96 %. Their activity increased with the chromium contents: v.g., at 550 °C conversion was 2.5 % on the catalyst with 5 at% chromium and 14.2 % on the one with 14 at% chromium. Noteworthingly, in the presence of CO2 the Cr-MCM41 catalytic activity was stable for at least 25 h on stream at 600 °C, and the surface area of fresh Cr-MCM41 sample (835 m2/g) or XRD pattern (Fig. 1) showed no marked change after 41 h on stream. However, after testing at 725 °C, ethane conversion at 600 °C decreased from 25 to 11 %. Activity was not recovered upon treatment with CO2, but it was by treatment under oxygen (15 mol % in He) at 600 °C for 2 h. So, this deactivation could be due either to carbonaceous deposits or to over-reduction of the catalyst at the highest temperature (above 700 °C). The rate of formation of ethene and CO and their evolution with temperature is shown in Fig. 2 (right). At any temperature the rate of ethene formation was always higher than that of CO, but their relative difference decreased when temperature increased (the ratio between both rates decreased). This involves that both oxidative and anaerobic dehydrogenations were produced simultaneously, because if only the oxidative route were active, both rates should be either equal via the selective reaction (1) or that of CO higher via the unselective reaction (2). Additional tests of ethane dehydrogenation in the absence of CO2 (by substituting it for helium while keeping all the other reaction conditions equal), evidenced the dehydrogenating activity of the Cr-MCM-41 catalyst, but the ethane conversion was lower and the catalyst deactivated very fast. This seems to confirm the co-existence of the two reactions in the presence of CO2. Thus, one may conclude that, in the ODE reaction with CO2 over Cr-MCM-41, carbon dioxide plays a double role: it can dissociate on the catalyst surface to produce active oxygen species, as previously reported on ceria-based catalysts [1], and it inhibits the coke formation on the catalyst surface, thus keeping the stability of the catalytic dehydrogenating activity. The observed variation of the ratio between the rates of formation of ethene and CO involves that the contribution of the oxidative dehydrogenation with CO2 to the total conversion of ethane increases with temperature. Further studies are going on in our labs to identify the active species in both reactions. Conclusions Cr-MCM-41 catalysts are highly effective for the ODE with CO2 even at moderate temperatures, while other M-MCM-41 catalysts are either almost inactive (M = Co) or non selective (M = Ni). The role of CO2 in this reaction over Cr-MCM-41 is double: as an oxidant and as inhibitor of coke formation. Acknowledgement. This work was financed by Spanish AECI under project 2002CN002. References [1] R. X. Valenzuela, G. Bueno, V. Cortes Corberan, Y. Xu, C. Chen, Catal. Today, 61 (2000) 43. [2] B. Grzybowska, J. Slocynski, R. Grabowski, L. Kerommes, K. Wcislo, T. Bobinska, Applied Catal. A 209 (2001) 279. [3] M. Cherina, M. S. Rao, A.M. Hirt, I.E. Wachs, G. Deo, J. Catal., 211 (2002) 482. [4] X. Zhang, Y. Hue, Z. Gao; Catal. Lett. 83 (2002) 19. [5] B. Sulikowski, Z. Olejniczak, E. Wloch, J. Rakoczy, R.X. Valenzuela and V. Cortes Corberan, Applied Catal. A: General, 232 (2002) 189. [6] M. Jia, R. X. Valenzuela, P. Amoros, D. Beltran-Porter, J. El-Haskouri, M. D. Marcos and V. Cortes Corberan, Catal. Today, accepted for publication.


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Selective formation of styrene via oxidative dehydrogenation of 4vinylcyclohexene over Zr0 2 -supported iron oxide catalysts Jong-San Changa'*, Do-Young Honga, Yong-Ki Parka, and Sang-Eon Parkb>* a

Catalysis Center for Molecular Engineering, Korea Research Institute of Chemical Technology (KRICT), P.O. Box 107, Yusung, Daejon 305-600, Korea. b Department of Chemistry, Inha University, 253, Yonghyun-Dong, Nam-ku, Inchon 402-751, Korea. FAX: +82-42-860-7679. E-mail: jschang@pado,krict.re.kr, [email protected] Styrene has been selectively formed by oxidative dehydrogenation of 4—vinylcyclohexene with oxygen and/or carbon dioxide as oxidant over ZrO2-supported iron oxide catalysts. It showed high styrene selectivity (88.9%) with VCH conversion of 80.1% at 400°C as well as a low rate of deactivation, which is confirmed to maintain its activity for more than 50 h. It is found that the addition of carbon dioxide to molecular oxygen as oxidant in the feed stream leads to the enhancement of the selectivity to styrene. 1. INTRODUCTION Styrene is produced commercially through the dehydrogenation of ethylbenzene with a large excess of steam over KiO-promoted iron oxide catalysts. However, dehydrogenation of 4-vinylcyclohexene (VCH) available from 1,3-butadiene (BD) via Diels-Alder cyclodimerization to styrene could provide an attractive route for styrene production and is regarded as an interesting alternative to the classical ethylbenzene dehydrogenation process. Dow Chemical has first introduced this new process to produce styrene from butadiene via VCH [1]. In addition, Neumann et al. has reported oxidative dehydrogenation of VCH using carbon-supported heteropolyacid catalyst [2]. This catalyst could be operated by lower reaction temperature, but it tended to be quickly deactivated. Bulk iron-containing oxides are known to be favorable catalysts for dehydrogenation reactions of aromatic hydrocarbons and olefins [3]. Moreover, the use of a supported iron oxide catalyst can provide a successful remedy for overcoming the problems encountered with the industrially applied bulk catalysts used for the dehydrogenation of hydrocarbons, such as ethylbenzene [4,5]. Zirconia is a promising support to be applied for not only oxidation catalysts but also reforming catalysts because it has nice characteristics such as high stability under oxidizing and reducing atmospheres, mild acid-base pairs, and oxygen vacant sites [6-8]. In this work, we try to prepare iron oxide-promoted zirconia catalyst and apply it to the oxidative dehydrogenation of VCH with oxygen and/or carbon dioxide 2. EXPERIMENTAL According to our previous paper [9], ZrC^-supported iron oxide (Fe2O3/ZrO2) catalyst was prepared by deposition of aqueous suspension of iron(II) sulfate heptahydrate onto ZrC*2

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at 60°C and pH = 10-11, followed by drying at 110°C and calcination at 400°C for 4 h. The loading of iron oxide was adjusted to 2 wt%. For CaO-Fe2O3/ZrO2 catalyst, an aqueous calcium nitrate solution corresponding to 1 wt% CaO was further added to Fe2O3/ZrO2 by impregnation, dried at 110°C, and calcined at 400°C for 4 h. The oxidative dehydrogenation reaction of VCH was carried out in a conventional fixedbed micro reactor made of quartz in the temperature range of 300 - 600°C and 1 bar. The catalyst was placed on the reactor and then pretreated with He at 600°C for 1 h. The VCH corresponding to the liquid hourly space velocity (LHSV) of 1.0 - 3.0 h"1 was injected with a syringe pump and then vaporized for the mixing with carrier gases. Reaction products were analyzed by an on-line gas chromatograph (Hewlette-Packard model 5890) equipped with a flame ionization and a thermal conductivity detectors, which employed a capillary column (J&W, DB-WAX) and a packed column (Chrompack, Heysep Q). 3. RESULTS AND DISCUSSION Table 1 summarizes catalytic results of the zirconia-based catalysts in the oxidative dehydrogenation of VCH. During the reaction, main products are styrene and ethylbenzene. In addition, benzene and unidentified polyolefinic hydrocarbons are formed as by-products. It is noted that even ZrO2 itself shows significant activity in the presence of oxygen, for instance, X(VCH) = 26.1% and S(SM) = 78.2% at 400°C, LHSV = 3, and O2/VCH = 6 (molar ratio). Its catalytic activity is improved by the modification with Fe2O3 or CaO-Fe2O3. In all catalysts tested, the formation of carbon dioxide due to the combustion of VCH is almost negligible even upon using a large excess of oxygen to VCH, indicating that the zirconiabased catalysts have mild and selective dehydrogenation activity. However, zirconia has no activity to produce styrene in the absence of oxygen (data not shown). This result strongly suggests that dehydrogenation of VCH to styrene over ZrO2 proceeds favorably with oxidative pathway. In contrast, the dehydrogenation activity of the Fe2O3/ZrO2 catalyst is still survived in spite of much decrease in the activity, suggesting that the introduction of iron oxide onto ZrO2 induces its own activity for dehydrogenation of ethylbenzene like a commercial styrene monomer catalyst. The catalytic activity for further dehydrogenation to styrene is much facilitated by molecular oxygen. For dehydrogenation of VCH, two possible pathways for producing ethylbenzene and styrene, respectively, are postulated as scheme 1. hi this scheme, styrene may be formed directly without isomerization via two consecutive dehydrogenation steps, while ethylbenzene may be formed by isomerization from an exocylic to endocyclic diene, followed by dehydrogenation. Actually, the ZrO2-supported catalysts produce styrene from VCH even without oxygen at low temperature, different from ethylbenzene dehydrogenation. Considering the difficulty of ethylbenzene dehydrogenation, it is likely that the formation of Table 1. Oxidative dehydrogenation of VCH over zirconia and zirconia-supported catalysts. VCH Selectivity (%) Catalyst Conversion Styrene Ethylbenzene Benzene Others ZrO2 Fe2O3/ZrO2 CaO-Fe2O3/ZrO2

26.1 34.5 45.2

78.2 86.3 88.9

6.1 3.0 1.3

4.2 2.0 2.5

Reaction conditions: Temp. = 400°C, LHSV = 3 h"\ O2/VCH = 6, P(VCH) = 1 kPa.

11.5 8.7 7.4

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styrene is not from or through ethylbenzene as an intermediate. In other words, it is supposed that ethylbenzene and styrene may be formed in parallel competitive reaction pathways.

Fig. 1 illustrates catalytic results according to carrier gas towards the CaO-FeiCVZrCh catalyst to exhibit the highest activity. In this case, lower space velocity (LHSV = 1 h"1) was applied to the reaction. Under a helium carrier gas, ethylbenzene is predominantly formed instead of styrene as a desirable product. This means that the dehydrogenation processes of

Fig. 1. Dehydrogenation of VCH over CaO-Fe2(VZrO2 under different reaction environments. Reaction conditions: T = 400°C, P(VCH) = 1 kPa and LHSV = 1 h"1, Carrier gas/VCH = 6, CO2/O2 = 0.6 in the case of mixed carrier gases. VCH and ethylbenzene appears to be apparently similar each other, leading to the final product of styrene. However, they show quite different behaviors for dehydrogenation due to different reaction pathway. In the case of ethylbenzene, only the ethyl group attached to aromatic ring needs to be dehydrogenated. However, since the ethyl group of VCH has been already dehydrogenate only cyclohexene ring of VCH is further required for dehydrogenation. Therefore, the poor selectivity of styrene (13.2%) in nonoxidative dehydrogenation of VCH might be caused by different reactivities toward cyclic and branched hydrocarbons. This implies that the participation of oxidant is necessary to produce styrene. Here one can expect another way to achieve high styrene selectivity while suppressing the formation of ethylbenzene in VCH dehydrogenation. The role of oxidant is more important in the oxidative dehydrogenation of VCH as compared with that of ethylbenzene because it is operated at lower temperature (< 450°C). However, the selectivity of styrene under mild oxidizing atmosphere like CO2 is not so high, suggesting that the oxidizing ability of CO2 is not enough to abstract ring hydrogens of VCH due to low temperature operation. When the dehydrogenation of VCH is carried out in oxygen atmosphere instead of CO2, the selectivity of styrene increases from 42.9% to 82.8% while that of ethylbenzene decreases from 33.3% to

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4.8%. Furthermore, when CO2 is partially introduced together with oxygen the highest selectivity of styrene (88.9%) is obtained. The activity and selectivity to styrene in the presence of oxygen are maintained without the significant change for 50 h. The conversion is not much influenced by the reaction atmosphere different from the strong dependence of selectivity to styrene. Thus the same level of conversion is obtained at the same space velocity of VCH regardless of carrier gas. It is likely that the oxygen species activated on oxygen vacant sites of the catalyst play a key role in the oxidation of hydrogens abstracted from ring hydrocarbon of VCH, resulting in an increase of styrene yield. Zirconia is known to have oxygen vacant sites as well as adjacent acid-base pairs [8]. In some case, amphoteric properties of zirconia stemmed from acid-base pair sites could provide to some extent dehydrogenation activity of aromatic hydrocarbons, e.g. ethylbenzene to styrene [10]. However, these sites are not able to work well for the dehydrogenation of VCH at lower temperature (400°C). Instead, oxygen vacant sites are able to play a key role on the conversion of VCH to styrene or ethylbenzene through the formation of mobile oxygen species after taking up gaseous oxygen. The introduction of iron oxide onto ZrO2 support may accompany an increase in the number of dehydrogenation sites which are considered to be its oxygen deficient sites and base sites. Actually, surface basicity of the catalyst is confirmed to increase upon loading iron oxide on the ZrO2 support from CO2-TPD analysis [10]. The addition of CaO to the catalyst would be effective for inhibiting the catalyst deactivation by means of enhancing surface basicity of the catalyst and then suppressing carbon deposition. 4. CONCLUSION For the oxidative dehydrogenation of VCH over the zirconia-based catalysts, it was found that the ZrO2 itself has intrinsic dehydrogenation activity for the selective production of styrene with the help of molecular oxygen. When the zirconia is further modified with Fe2O3 and CaO, the highest styrene selectivity of 88.9% is obtained. Especially, the introduction of CaO in the catalyst is effective for inhibiting the catalyst deactivation due to suppressing carbon deposition and the dehydrogenation activity is maintained for more than 50 h. ACKNOWLEDGMENT This work was supported by the Korea Ministry of Science and Technology through the Research Center for Nanocatalysis (KN-0329) and Institutional Research Program. We thank Mr. Y.S. Choi for experimental assistance. REFERENCES 1. D.A. Hucul (by Dow Chemicals Co.), US Patent 5,336,822 (1994). 2. R. Neumann, and I. Dror, Appl. Catal. A: General, 72 (1998) 67. 3. A. Miyakoshi, A. Ueno, and M. Ichikawa, Applied Catalysis A: General, 16 (2001) 137. 4. P.G. Menon, Chem. Rev., 94 (1994) 1021. 5. L.A. Boot, A.J. van Dillen, J.W. Geus, and F.R van Buren, J. Catal, 163 (1996) 186. 6. H.S. Roh, K.W. Jun, and S.-E. Park, J. Ind. Eng. Chem., 9 (2003) 261. 7. H.S. Roh, K.W. Jun, and S.C. Baek, J. Ind. Eng. Chem., 9 (2003) 168. 8. T. Yamguchi, Catal. Today, 20 (1994) 199. 9. J.-S. Chang, M.S. Park, and S.-E. Park, Chem. Lett., (1997) 1023. 10. J. Noh, J.-S. Chang, J.-N. Park, K.Y. Lee, and S.-E. Park, Appl. Organomet. Chem., 14 (2000)815


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Catalytic Transformation of Cyclohexanol over Mixed Metal Oxides with and without CO2 V. R. Chumbhale, J. S. Kim, S. B. Lee and M. J. Choi* Advanced Chemical Technology Division-ENR, Korea Research Institute of Chemical Technology, P.O. Box 107, Daejeon 305-600, Korea The reaction of cyclohexanol with CO2 is studied over alkali doped ferric-molybdate catalyst to check the possibility of formation of carboxylated product. Prior to this the cyclohexanol transformation over binary oxides of Fe and Mo and modified Fe-Mo catalysts was extensively studied. It is observed that the nature of the catalyst and reaction parameters have profound effect on product distribution and selectivity to different products. 1. INTRODUCTION CO2 is a major source for the greenhouse effect leading possibly to a changes in climate due to its incremental rate of ca 1% per year mainly due to combustion of fossile fuels and deforestation [1-3]. Although, the content of CO2 in atmosphere is balanced partly due to photosynthesis conversion and its accumulation in seas, it is not substantially reduced by these natural activities. Hence there is a need to mitigate CO2 by making use of it's physical and chemical properties as it is one of the cheapest and most abundant carbon raw material in the world [4] The chemical property of CO 2 (CO 2 as a reactant) is used in the synthesis of urea, methanol salicylic acid and cyclic carbonates in addition to its application as an acid in water purification or in neutralization processes [5]. Although there are many chemical reactions studied by making use of CO2 due to its potential as a building block for carbon-carbon chains or as a competitive carbon source in chemistry, many reactions are involved with use of expensive organometallic complexes in stoichiometrical amount [6]. There are reports in literature that various useful products like fine chemicals, intermediates in pharmaceuticals and food industries can be produced by making use of CO2 [5]. Oxides, because of their ability to take part in the exchange of electron as well as in the exchange of protons or oxide ions are used as catalysts in both redox and acid base reactions. They constitute the active phase not only in the oxide catalysts in the strict sense, but also in the case of many metal catalysts, which in the conditions of catalytic reactions, are covered by a surface layer of a reactive oxide. The properties of oxide phases are the determining factor in the evolution of the catalytic activity and selectivity. Oxides of main group elements, used as acid-base catalysts in the form of solid solutions, oxysalts, oxyacids or bases in heterolytic reactions [7]. In the present work we have made an attempt to utilize CO2 towards the production of cyclohexanone carboxylic acid by making use of insitu generated hydrogen in catalytic dehydrogenation of cyclohexanol over mixed metal oxides. Prior to carboxylation, the cyclohexanol transformation over ferric-molybdate and modified ferric molybdate are studied. The product cyclohexanone (produced on cyclohexanol dehydrogenation) is used in the production of e-caprolactam from the cyclohexanone oxime route, s-caprolactam has become

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a commodity chemical with basically a single outlet - nylon 6. The cyclohexene (produced on dehydration of cyclohexanol) has got applications in the manufacture of adipic acid, maleic acid, hexahydrobenzoic acid [8]. The carboxylated product of cyclohexanone on esterification finds application as an intermediate for pharmaceuticals and other applications. 2. EXPERIMENTAL Preparation and Characterization of the Catalyst: The parent ferric-molybdate Fe2(MoC>4)3 was prepared by adopting procedure as reported elsewhere [9]. This was modified by treating with appropriate quantity of aqueous NaOH solution and evaporated to dryness at 90 °C (ca. 5g ferric-molybdate treated with 5.2ml of 10wt% aqueous NaOH solution resulted in 6wt% Na doping). The doped samples were further calcined in a muffle furnace at 500 °C for 5 h. On XRD investigation it was observed that there was formation of new inorganic phase NaFe(MoO4)2 [10]. The Reaction: Cyclohexanol transformation was carried out in a continuous fixed bed down flow tubular metal reactor (24cmxl0mm i.d) that was heated electrically. The catalyst (lg, mean particle size=1.42 mm; L/D=2.5; D/dp=5.63) was packed in the quartz sand. The temperatures were recorded by means of a thermocouple that was placed parallel with the catalyst bed. The catalyst was activated at a temperature of 500°C for 4h and was purged with N2 for 20 minutes prior to the reaction. The product gases were condensed in an ice-cold trap and analyzed by gas chromatography using HP-Innowax (30M length X0.25 mm i.d. having film thickness of 0.25um) on FID under programmed temperature. All reactions were carried out at atmospheric pressure. The conversion, selectivity and yield are defined as follows: Cyclohexanol conversion = 100 - % weight of cyclohexanol in the product Selectivity to specific product = wt. of product / cyclohexanol conversion xlOO Yield = Conversion x Selectivity STY= Product in g/ h / cc of the catalyst

(1) (2) (3) (4)

3. RESULTS AND DISCUSSION The process of cyclohexanol dehydration includes the complex of consequent-parallel reactions, such as dehydrogenation of cyclohexanol to cyclohexanone (the main reaction), aromatization of cyclohexanol to phenol, dehydration of cyclohexanol to cyclohexene, and condensation of cyclohexanone to cyclohexyldencyclohexanone (as a byproduct) [11]. The effect of temperature on conversion and selectivity over Fe-Mo parent catalyst is shown in Figure 1. It was observed that the main products were cyclohexene, cyclohexanone and small amount of phenol with other products. As the temperature was increased from 200°C to 300°C (atWHSV=3.6 h"1) the cyclohexanol conversion was linearly increased from about 16 to 98% with parallel rise in the % selectivity to cyclohexene which was in the range of 85- 98% due to dehydration at corresponding higher temperature. Lower temperature, (225°C) favoured aromatization to phenol due to dehydrogenation. The cyclohexene selectivity was prominent at 300°C. Beyond 300 °C (at 325°C), the cyclohexanol conversion and selectivity to cyclohexene was slightly decreased which is attributable to the by product formation [11, 12]. The impurities formation was higher at lower temperature (9-17 weight % in the range of 200-225 ° C which were reduced substantially at higher reaction temperature). However, there was no significant cyclohexanone formation noticed. The cyclohexene yield was highest at 300°C.The space-time yield was in the range of 3.31 to 4 g cc"1 h"1. Figure 2 and 3 describe

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the product distribution as a variation in space velocity at 300°C. From Figure 3 it can be stated that the space-time yield increased monotonously with increase in space velocity. The cyclohexanol conversion was in the range of 96-98% with cyclohexene yield in the range of 95-96%. No phenol formation was noticed due to cyclohexanol aromatization. Beyond 6.8 h"1, although there was an improved STY, the cyclohexanol conversion and cyclohexene yield was significantly reduced to 90 and 88% respectively. From these screening of the catalyst it was thought to be 300°C (reaction temperature) and space velocity of 6.8h"' as an optimum reaction parameters. When cyclohexanol transformation was carried out over alkali doped ferric-molybdate catalysts it was observed that the conversion of cyclohexanol was reduced with formation of cyclohexanone higher than the parent catalyst. The higher doping of alkali favoured the formation of cyclohexanone (Fig 4). It was found that the Arrhenius activation energy calculated from these data showed enhanced value due to new phase formation (not shown here). The possibility of formation of corresponding carboxylic acid was checked by passing CO2 with organic substrate at ambient pressure (Table 1). It is observed that under the set of experimental conditions studied, there was no significant formation of carboxylic acids over the binary oxides of Fe and Mo. The lower selectivity to cyclohexanone in the carboxylation of cyclohexanol over Fe-Mo-Na (6wt%) is probably could be due to lesser extent of CO2 activation on the newly phase formed by doping of alkali on parent catalyst.

Fig. 1. Effects of Temperature on Conv., Sel. and Yield.

Fig. 2. Effects of WHSV on Conv., Sel. and Yield.

Fig. 3. Effects of WHSV on STY.

Fig. 4. Effects of wt% Na doping on Conv. and Sel.

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Table 1: Data on Carboxylation reactions Catalyst

Reactant feed CNOL/CO2 (molar ratio=l:l)

Reaction parameters T: 300°C; WHSV: 7.3 h'1

Observations %X-CNOL: 29.7 Fe-Mo-Na(6wt%) %S-CH: 96.8 %S-CNONE: 2.84 %S-OTHER: 0.36 T: 250°C; %X-CNONE: 5.33 CNONE/CO2 Fe-Mo-Na(6wt%) (molar ratio=l:l) WHSV: 7.6h"' %S-CNOL: 99.81 %S-OTHER:0.19 %X-CNONE: 7.85 T: 300°C CNONE/CO2 Fe-Mo-Na(6wt%) (molar ratio=l:l) WHSV: 7.6 h'1 %S-CNOL: 95.85 %S-OTHER:4.15 CH/CO2 %X-CH: 1.41 T: 250°C (molar ratio=l:l) Fe-Mo %S-NONE: 80.85 WHSV: 4.4h"' %S-OTHER: 19.15 %X-CNOL: 79.6 T: 300°C CNOL/CO2 Fe-Mo (molar ratio=l:l) %S-CH: 93.2 WHSV: 7.3 h'1 %S-CNONE: 6.45 %S-OTHER: 0.35 CNOL: cyclohexanol, CNONE: cyclohexanone, CH: cyclohexene, %X: conversion, %S: selectivity

4. CONCLUSION Cyclohexanol transformation over ferric-molybdate catalyst yields cyclohexene as a major product. The optimum temperature at a space velocity of = 7 h"1 was found to be 300 °C. On doping aqueous alkali salt, the new inorganic phase is formed. The NaFe(MoO4)2 phase facilitated the formation of cyclohexanone in cyclohexanol transformation. The new inorganic phase formed enhanced activation energy. Carboxylation reactions studied at 250 and 300 °C (at ambient pressure) by passing CO2 at a substrate to CO2 molar ratio of 1 and using ferric-molybdate and alkali doped ferricmolybdate catalysts, did not reveal the formation of corresponding acids in the appreciable amounts. REFERENCES 1. W.C. Clarck, Ed.Carbon Dioixide Review: 1982, Clarendon Press: oxford, U.K., 303. 2. H. Oman, Chemtech, Feb (1988) 116. 3. J. Goudriaan, Eenergispectrum Nov.(1991) 268. 4. A. Behr, Carbon Dioxide activated by metal Complexes, VCH: Weinheim, Germany (1988). 5. G. Musie, M. Wei, B. Subramaniam and D.H. Busch, Coordination Chemistry Reviews 219-221 (2001) 789. 6. M.E. Volpin and I.S. Kolominikov, Pure and App. Chem., 33919730 567. 7. D.B. Dadyburjor, S.S. Jewur, E. Ruckenstein, Catal. Rev., 19 (1979)293. 8. The Merck Index 10th Edition, Merck & Co. Rahway, NJ,USA. (1983) P 2722. 9. V.R. Chumbhale and P.A. Awasarkar, Appl. Catal., 205 (2001) 109. 10.JCPDS Powder Diffraction Files for Inorganic Phases (international center for diffraction data) Swarthmore (Publishers), PA, U.S.A. 1989, P 307. 11. V.Z. Fridman and A.A. Davydov, J. Catal., 195(2000) 20. 12. F.M. Bautista, J.M. Campelo, A. Garcia, D. Luna, J.M. Marinas, R.A. Quiras, A.A. Romero, Appl. Catal., 243 (2003) 93.


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The Utility of Carbon Dioxide in Homogeneously-Catalyzed Organic Synthesis Philip G. Jessop* Department of Chemistry, Queen's University, Kingston, Ontario, Canada K7L 3N6 [email protected] While CO2 fixation is obviously of great interest, CO2 has increasing utility even in reactions which do not result in its fixation. This review briefly summarizes many of the techniques that involve CO2 without incorporating it into the product. An emphasis is placed on reactions involving homogeneous catalysis. CO2 as a reaction solvent Supercritical and liquid CO2 have been used as solvents for an enormous number of homogeneously-catalyzed reactions. The field was reviewed in 1999,1'2 so there is no need for a repeat of that information here. Important developments in the area since then include a greater emphasis on asymmetric catalysis,3 studies of the phase behaviour of reaction mixtures as a function of conversion,4 new developments in the search for more CC^-soluble catalysts or reagents, and the discovery of the virtues of insoluble catalysts. The latter two points need further explanation. Because many catalysts, especially charged complexes or those with several aromatic groups, have very poor solubility in CO2, researchers have adopted several strategies for developing CC^-soluble catalysts, including the use of trialkylphosphines instead of triarylphosphines,5'6 the attachment of fluorinated alkyl chains to the meta or para positions of arylphosphines,7 or the use of anions containing fluorinated groups.8 For situations in which arylphosphines are necessary, there have been several attempts to design a triarylphosphine which has high solubility in CO2 without incorporating fluoride groups (because of their expense and environmental persistence). Xiao's group recommended the use of an ester group in the para position of triphenylphosphine (eg. Ph^PCeHUpCC^Me), but solubility measurements by Beckman and Jessop10 showed that such phosphines are no more soluble than triphenylphosphine itself. Beckman and Jessop also found that tri(2-furyl)phosphine (1) and especially the peracetylated sugar-phosphine 2 below are much more soluble in SCCO2 than is triphenylphosphine.10 This exciting result should make it possible to use these ligands in the design of CO2-soluble homogeneous catalysis.

356 A CCh-insoluble catalyst can sometimes be an advantage instead of a disadvantage. Sellin and Cole-Hamilton" have reported that a liquid mixture of [Rri2(OAc)4] and P(OAr)3 (Ar = Ph or C6H4PC9H19) is visibly insoluble in scCO2, so that the hydroformylation of 1hexene could be performed in a CO2/H2/CO mixture in contact with a drop of the liquid catalyst. The reaction took place either at the surface of, or inside, the liquid catalyst. No significant quantities of the catalyst was leached out into the SCCO2 phase. Leitner also used a scCC>2-insoluble phosphine, but more specifically one containing a poly(ethylene glycol) chain. In Leitner's method, the catalyst bearing such ligands is dissolved in liquid substrate. After the reaction the product is extracted with SCCO2, leaving the CCh-insoluble catalyst behind in the vessel.12 CO2 as half of a biphasic solvent mixture Catalyst recovery is an important consideration in the industrialization of homogeneous catalysis. Obviously, recovery without catalyst destruction is preferred, but it is not the dominant method in industry.13 Biphasic catalysis is a method developed to address this need.14"16 In biphasic catalysis, the system is designed so that at the end of the reaction there will be two immiscible liquid (or fluid) phases, one of which contains the catalyst and the other contains the product. Typically one of these phases is the reaction solvent and retains the catalyst whereas the other is either immiscible product, immiscible excess starting material, or a second solvent added before or after the reaction. Product separation from the catalyst and re-use of the catalyst are both facile. In those cases where a second solvent is added, it is important to choose a solvent which will extract neither the reaction solvent nor the catalyst. Environmental considerations as well as concerns over cross-contamination of the two solvents suggest that traditional organic solvents should be avoided in such systems. In particular, researchers have recently suggested that use of SCCO2 as the product-bearing phase is advantageous because there are certain solvents of very low solubility in SCCO2 and because the recovery of the product from the scCO2 can be achieved simply by a reduction in the pressure or by extraction with water without pressure reduction (as practiced in caffeine extraction).17 Given that the product-bearing solvent is scCCh, there is still a selection of appropriate catalyst-bearing solvents, which obviously must not be miscible with SCCO2. Liquids that are not readily extracted by scCCh and could act as solvents for catalysis include water, ionic liquids, and liquid polymers. All of these combinations have been tested, and the first reports of each are cited here. The first example of catalysis using the SCCO2/H2O solvent pair was the hydrogenation of cinnamaldehyde reported by Arai's group (eq 1, tppts = P(C6H4mSO3Na)3).18 They obtained 99 % selectivity for the unsaturated alcohol if they used RuCl3 as the catalyst precursor and 100 % selectivity for the saturated alcohol if they used RI1CI3.

40 bar H2 40 , 2 h

(!)

The first example of catalyst recycling using an ionic liquid/CO2 solvent pair was the hydrogenation of a,(3-unsaturated carboxylic acids by Jessop's group (eq. 2).19 Other examples include the hydrogenation of CO2,20 hydroformylations,21 and a hydrovinylation.22 The last two of these were performed as continuous-flow processes.

357

(2)

The first example of the use of the liquid polymer/scCO2 solvent pair was reported very recently by the Jessop group.23 Wilkinson's catalyst, RhCl(PPri3)3, was used for the hydrogenation of styrene in liquid poly(ethylene glycol) (PEG) under SCCO2. The CO2 served as a solvent and as a modifier of the PEG, lowering its melting point enough that it could be liquid at the reaction temperature of 40 °C. The catalyst solution in PEG was used 5 times with no significant loss of activity. Higher molecular weight PEG fractions have the advantage of not being significantly extracted by SCCO2 but have higher viscosities and melting points. A molecular weight of 1,600 was found to be a reasonable compromise. The researchers are currently evaluating PEG, poly(propylene glycol), poly(methylphenylsiloxane), poly(tetrahydrofuran), and modified versions of these with altered end-groups, as media for various asymmetric homogeneously-catalyzed reactions. CO2 as a modifier of viscosity or gas solubility In situations where the solvent or reaction mixture is so viscous that mass transfer limitations exist, or if the solubility of reagent gases such as H2, CO or O2 are so low that reaction rate or selectivity is decreased, then CO2 can be used to great advantage. Fortunately, CO2, when dissolved in a liquid, decreases the viscosity of the liquid and increases the solubility of reagent gases.24 The CCh-induced changes in viscosity require pressures of typically 40-80 bar, so that the CO2 in the upper phase is either subcritical gaseous or just barely supercritical. A CC>2-in-organic liquid mixture is referred to as an "expanded" liquid, in reference to the volumetric expansion that is also observed. Changes in viscosity upon "expansion" by CO2 have been observed for ionic liquids25'26 and for liquid polymers.27 The drop in viscosity has already been used to advantage in the filtering of viscous used motor oil. 8 Modification of an ionic liquid solvent by CO2 expansion has been shown to increase the selectivity for linear aldehyde in a hydroformylation21 and to strongly affect the enantioselectivity of the hydrogenation of two unsaturated carboxylic acids, favourably for one and unfavourably for the other (eq. 2).29 Modifications of rates and selectivity by the dissolution of CO2 or related gases have also been observed in reactions in traditional organic liquid solvents.30"34 CO2 as a trigger for miscibility/solubility changes The application of CO2 gas over a pair of liquids can drastically change the mutual miscibility of the two liquids. For example, Brennecke has shown that CO2 gas pressure applied over a miscible pair of liquids can force them to become immiscible. Liquid/liquid pairs that exhibit this behaviour include ionic liquid/organic liquid35 and ionic liquid/water pairs.36 This technology has obvious applications in post-catalysis separations.

358

Similarly, some mutually immiscible solvents can be made miscible by the presence of CO2. Dissolution of CO2 into a fluorophobic organic solvent (i.e. one that is unable to dissolve highly fluorinated species) makes the solvent become fluorophilic. This phenomenon makes it possible to use dissolved CO2 as a trigger for the reversible coalescence of an organic liquid with a fluorous liquid (i.e. a perfluorinated liquid). For example, mixtures of an organic solvent and a fluorous liquid are often biphasic because of mutual immiscibility. In the technique known as fluorous biphasic catalysis.37'38 the biphasic mixture is rendered monophasic by raising the temperature. However, it is now known that the same liquid/liquid miscibility can be triggered by the application of CO2 gas to the system without any temperature increase.39 The required pressure of CO2 can be as low as 20 bar, depending on the choice of the two solvents. This represents an alternative to the application of heat in fluorous biphasic catalysis, for those reactions which may have poor selectivity at higher temperatures. Dissolution of CO2 into an organic solvent can also trigger the solubilization of a highly fluorinated solid such as a catalyst precursor. The slow release of the pressure causes precipitation or crystallization of the complex, often leading to crystals of sufficient quality for X-ray crystallography.40 Dissolution of CO2 into a solution that contains a non-fluorinated solute is more likely to cause precipitation rather than greater solubilization of that species. This can be of value in precipitating a catalyst while keeping a product dissolved.34 CO2 as a trigger for a change in a partition coefficient The partition coefficient of a solute between two phases can be changed drastically by the dissolution of CO2 into one or both phases. For example, if one of the phases is water, then the presence of pressurized CO2 causes the aqueous phase to become quite acidic, which could cause acidic solutes to partition more into the nonaqueous phase and cause basic solutes to partition more into the aqueous phase. This technique has yet to be applied to separations problems in catalysis. For another example, if the solute is partitioning between a fluorous (highly fluorinated) phase and a conventional organic solvent phase, the dissolution of CO2 into the latter will cause the solute to partition more into the organic phase from the fluorous phase. This was demonstrated by Jessop's and Eckert's groups, who used this phenomenon to make a catalyst that could reversibly switch from heterogeneous to homogeneous (Scheme 1). This switch is triggered by the application of a subcritical pressure of gaseous CO2. The fluorinated catalyst, which starts the experiment impregnated into the surface of fluorous silica, is extracted into the expanded organic phase, thus becoming a homogeneous catalyst. Upon release of the CO2 pressure, the catalyst partitions back again into the fluorous silica phase, becoming heterogeneous again.41'42 The technique has been tested on the hydrogenation of styrene, with successful catalyst recycling, and is now being tested on other reactions.

359

Scheme 1. The preparation and use of fluorous silica (R = perfluoropolyether) as a solid fluorous phase for the trapping of a fluorous homogeneous catalyst. The catalyst transfers to the liquid phase (cyclohexane) by the dissolution of CO2 in the cyclohexane. Upon release of the CO2 pressure, the fluorouscatalyst partitions/adsorbs back onto the fluorous silica.41'42 CO2 as a temporary protecting group Leitner's group has provided two examples of the use of SCCO2 as both solvent and a reagent for the creation of a temporary protecting group.43 In both cases, the reaction which creates the protecting group is carbamate salt formation from a secondary amine (eq. 3). The carbamate functional group does not interfere with homogeneously-catalyzed olefin metathesis (eq 4)44 or hydroformylation,43 reactions which would otherwise have not proceeded to the desired products.

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CO2 as an accelerator for solventless solid/solid or solid/gas reactions Solventless reactions are obviously environmentally superior to reactions in solution, but solventless reactions of solids react extremely slowly. The situation is more complicated because the melting point of a solventless reaction mixture is a function of conversion. Typically the melting point starts high, drops during the reaction (due to the presence of more compounds in the mixture) and then climbs again towards the completion of the reaction as the limiting reagent is used up. If the reaction temperature is lower than the melting point of the reactants or products, then the reaction mixture may be observed to melt after partial conversion and/or freeze as the reaction approaches completion. In such situations, the reaction can be accelerated by the addition of CO2 gas. Exposure of an organic solid to a pressure of gaseous CO2 causes a lowering of the melting point of the solid. Depending on the pressure and the nature of the solid compound, the melting point depression is typically 20-25 degrees.45 As a result, a solid-solid or solidgas reaction taking place at a temperature somewhat below the melting point of the solid (or solid/solid mixture) will proceed more rapidly in the presence of CO2 than in its absence. This is not an example of the effect of simple hydrostatic pressure; the same acceleration is not observed with a pressure of N2 for example. An example of an accelerated reaction is the hydrogenation of vinylnaphthalene (eq 5), which, when run for 30 min at 33 °C gave 0 % conversion in the absence of CO2 and 52 % at 56 bar of CO2.46 Note that this is subcritical gaseous CO2, not liquid or supercritical CO2. A similar effect but with a different reaction has been observed by Leitner.-47

(5) A solventless reaction in which the reaction mixture freezes as the reaction nears completion will most likely stall before complete conversion is obtained. CO2, by lowering the melting point of the reaction mixture, can allow the reaction to proceed further towards completion. For example, reaction 6 at 35 °C stalls at 90% and can not proceed further even if left for 25 h. In the presence of 60 bar CO2, however, the conversion reaches 97% within the first hour.46

(6) Conclusions Researchers in the area of homogeneous catalysis have invented several new techniques for using CO2, beyond the most obvious technique of using it as a reagent. Even the concept of using CO2 as a simple solvent is still evolving, although one might be tempted to consider it a maturing field. We now know that CO2 can serve as a reagent, a solvent, a solvent modifier, a miscibility trigger, a trigger of partitioning constant changes, a reversible protecting group, and an accelerant of solventless reactions. CO2 has proven itself to be an extremely versatile tool for the improvement of homogeneous catalysis.

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Acknowledgements The author acknowledges support from the Division of Chemical Sciences, Office of Basic Energy Sciences, U. S. Department of Energy (grant number DE-FG03-99ER14986) and from the Natural Sciences and Engineering Research Council, Canada. The author, Canada Research Chair in Green Chemistry, also acknowledges the support of the Canada Research Chairs program. References (1) Jessop, P. G.; Ikariya, T.; Noyori, R. Chem. Rev. 1999, 99, 475-493. (2) Jessop, P. G.; Leitner, W., Eds. Chemical Synthesis using Supercritical Fluids; VCH/Wiley: Weinheim, 1999. (3) Jessop, P. G. In Drug Delivery and Supercritical Fluid Technology; York, P., Kompella, U. B., Shekunov, B. Y., Eds.; Marcel Dekker, in press. (4) Ke, J.; Han, B. X.; George, M. W.; Yan, H. K.; Poliakoff, M. J. Am. Chem. Soc. 2001,725,3661-3670. (5) Jessop, P. G.; Ikariya, T.; Noyori, R. Nature 1994, 368, 231-233. (6) Bach, I.; Cole-Hamilton, D. J. Chemical Commun. 1998, 1463-1464. (7) Kainz, S.; Koch, D.; Baumann, W.; Leitner, W. Angew. Chem., Int. Ed. Engl. 1997, 36, 1628-1630. (8) Burk, M. J.; Feng, S.; Gross, M. F.; Tumas, W. J. Am. Chem. Soc. 1995,117, 8277-8278. (9) Hu, Y.; Chen, W.; Xu, L. J.; Xiao, J. L. Organometallics 2001, 20, 3206-3208. (10) Ablan, C. D.; Jessop, P. G.; Beckman, E. in preparation. (11) Sellin, M. F.; Cole-Hamilton, D. J. J. Chem. Soc.-Dalton Trans. 2000,11, 1681-1683. (12) Leitner, W. presented at 226th A CS National Meeting: New York, NY, 2003. (13) Falbe, J.; Bahrmann, H. J. Chem. Ed. 1984, 61, 961-965. (14) Cornils, B.; Herrmann, W. A., Eds. Aqueous-Phase Organometallic Catalysis; Wiley-VCH: Weinheim, 1998. (15) Cornils, B. Org. Proc. Res. Devel. 1998, 2, 121-127. (16) Keim, W. Green Chem. 2003, 5, 105-111. (17) McHugh, M.; Krukonis, V. Supercritical Fluid Extraction; 2nd ed.; Butterworth-Heinemann: Boston, 1994. (18) Bhanage, B. M.; Ikushima, Y.; Shirai, M.; Arai, M. Chem. Commun. 1999, 1277-1278. (19) Brown, R. A.; Pollet, P.; McKoon, E.; Eckert, C. A.; Liotta, C. L.; Jessop, P. G. J. Am. Chem. Soc. 2001,123, 1254-1255. (20) Liu, F. C ; Abrams, M. B.; Baker, R. T.; Tumas, W. Chem. Commun. 2001, 433-434. (21) Sellin, M. F.; Webb, P. B.; Cole-Hamilton, D. J. Chem. Commun. 2001, 781782. (22) Bosmann, A.; Francio, G.; Janssen, E.; Solinas, M.; Leitner, W.; Wasserscheid, P. Angew. Chem. Int. Ed. 2001, 40, 2697-2699. (23) Heldebrant, D. J.; Jessop, P. G. J. Am. Chem. Soc. 2003,125, 5600-5601. (24) Freitag, N. P.; Robinson, D. B. Fluid Phase Equilib. 1986, 31, 183-201. (25) Baker, S. N.; Baker, G. A.; Kane, M. A.; Bright, F. V. J. Phys. Chem. B 2001, 105, 9663-9668.

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(26) Lu, J.; Liotta, C. L.; Eckert, C. A. J. Phys. Chem. A. 2003,107, 3995-4000. (27) Gerhardt, L. J.; Manke, C. W.; Gulari, E. Journal of Polymer Science Part B: Polymer Physics 1997, 35, 523-534. (28) Sarrade, S.; Schrive, L.; Gourgouillon, D.; Rios, G. M. Separation and Purification Technology 2001, 25, 315-321. (29) Jessop, P. G.; Stanley, R.; Brown, R. A.; Eckert, C. A.; Liotta, C. L.; Ngo, T. T.; Pollet, P. Green Chem. 2003, 5, 123-128. (30) Combes, G. B.; Dehghani, F.; Lucien, F. P.; Dillow, A. K.; Foster, N. R. In Reaction Engineering for Pollution Prevention; Abraham, M. A., Hesketh, R. P., Eds.; Elsevier: Amsterdam, 2000, pp 173-181. (31) Thomas, C. A.; Bonilla, R. J.; Huang, Y.; Jessop, P. G. Can. J. Chem. 2001, 79, 719-724. (32) Musie, G.; Wei, M.; Subramaniam, B.; Busch, D. H. Coord. Chem. Rev. 2001, 219, 789-820. (33) Subramaniam, B.; Busch, D. H. In Carbon Dioxide Conversion and Utilization; Song, C , Gaffney, A. F., Fujimoto, K., Eds.; ACS: Washington, 2002, pp 364386. (34) Wei, M.; Musie, G. T.; Busch, D. H.; Subramaniam, B. J. Am. Chem. Soc. 2002,724,2513-2517. (35) Scurto, A. M.; Aki, S.; Brennecke, J. F. J. Am. Chem Soc. 2002,124, 1027610277. (36) Scurto, A. M.; Aki, S. N. V. K.; Brennecke, J. F. Chem. Commun. 2003, 572573. (37) Horvath, I. T.; Rabai, J. Science 1994, 266, 72-75. (38) Fish, R. H. Chem. Eur. J. 1999, 5, 1677-1680. (39) West, K. N.; Bush, D.; Hallett, J. P.; Brown, J. S.; Liotta, C. L.; Eckert, C. A. In Proceedings of the 2nd International Meeting on High Pressure Chemical Engineering; Brunner, G., Ed.: Hamburg, Germany, 2001. (40) Jessop, P. G.; Olmstead, M. M.; Ablan, C. D.; Grabenauer, M.; Sheppard, D.; Eckert, C. A.; Liotta, C. L. Inorg. Chem. 2002, 41, 3463-3468. (41) P. G. Jessop, C. A. Eckert, C. L. Liotta, U. S. Prov. Patent Appl. 30 May 2001, Int. Patent Appl. 30 May 2002, published 5 Dec 2002, WO 02/096550 Al. (42) Ablan, C. D.; Hallett, J. P.; Jones, R. S.; Eckert, C. A.; Liotta, C. L.; Jessop, P. G. Chem. Comm. 2003, 2972-2973. (43) Wittmann, K.; Wisniewski, W.; Mynott, R.; Leitner, W.; Kranemann, C. L.; Rische, T.; Eilbracht, P.; Kluwer, S.; Ernsting, J. M.; Elsevier, C. L. Chem. Eur. ./. 2001, 7, 4584-4589. (44) Filrstner, A.; Koch, D.; Langemann, K.; Leitner, W.; Six, C. Angew. Chem., Int. Ed. Engl. 1997, 36,2466-2469. (45) McHugh, M. A.; Yogan, T. J. J. Chem. Eng. Data 1984, 29, 112-115. (46) Jessop, P. G.; DeHaai, S.; Wynne, D. C ; Nakawatase, D. Chem. Commun. 2000, 693-694. (47) Francio, G.; Wittmann, K.; Leitner, W. J. Organomet. Chem. 2001, 621, 130142.


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Aldol reactions of propanal using MgO catalyst in supercritical CO2 Keitaro Matsui, Hajime Kawanami and Hiromichi Hayashi* Supercritical Fluid Research Center, National Institute of Advanced Industrial Science and Technology, 4-2-1 Nigatake, Miyagino-ku, Sendai, 983-8551, Japan. ABSTRACT Aldol reaction of propanal was carried out by using basic catalysts of MgO with different particle size under supercritical CO2 in the presence or absence of water. The obtained results showed that the catalytic activities of MgO depended on the order of catalyst size in the absence of water. In contrast, size dependence of the catalytic activity could not be observed in the reaction system, whereas the activity was improved by the addition of water. TEM and XRD analyses revealed that microscopic structure of MgO was converted to hydromagnesite after the reaction, since carbonic acid dissolved MgO during the reaction in the presence of water. The roles of MgO catalysts and carbonic acid in aldol reaction were discussed. 1. INTRODUCTION From the view point of greener processes, the design of organic synthetic methods using supercritical carbon dioxide (SCCO2) has become of much interest in recent years, because SCCO2 is considered as an environmentally benign and cheap medium [1-7]. It has several advantages for chemical syntheses such as relatively moderate critical condition (7c = 304.2 K, Pc = 7.38 MPa) and easily tunable physicochemical properties such as density and solubility for substrate by adjusting the pressure and temperature [7]. Though quite a large number of homogeneous catalytic reactions have been reported so far, it has suffered from serious disadvantages of the separation of reactant/product and of side reactions like decomposition, which could occur during the distillation after the reaction. Hence, solid heterogeneous catalyst characteristics, such as thermostability and easy catalyst/product separation, are advantageous compared with homogeneous catalysts for reuse.

g author:E-mail:[email protected]; Fax:+81-22-237-5215; Tel: +81-22-237-5214.

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The aldol reaction is one of the most important reactions to form C-C bonds [8]. Many attempts have been made to obtain aldol or enal products in high selectivity from unmodified aldehydes using heterogeneous catalysts such as alkaline oxides, alkaline earth oxides [9,10], zeolite [11], and hydrotalcites [12]. However, it is difficult to obtain the aldol or enal product in satisfactory selectivity, because of the high reactivity of aldehyde group. In our previous work, we studied the self-aldol reaction by using MgO as heterogeneous catalysts under SCCO2, and it was found that the addition of a small amount of water in the reaction system enable us to control the products selectivity by simple pressure manipulation [13]. Under compressed CO2, carbonic acid will be formed easily by the combination of added water and CO2, and it was thought that the dehydration of aldol compounds might be promoted by carbonic acid acting as Bransted acid. In contrast, the role of MgO catalysts in the reaction system under scCO2 was not clearly understood. Thus in this study, for the estimation of the characteristics of MgO during the reaction, self-aldol reaction of propanal was conducted by using MgO with different catalyst size, and catalytic properties and structural change of catalysts were investigated. 2. EXPERIMENTAL Four kinds of MgO fine powder (Ube Materials JRC-MGO-4 100A, 500A, 1000A and 2000A) purchased from The Catalysis Society of Japan were used as basic catalysts in the present study. BET surface area and surface basicity of the catalysts were determined by N2 adsorption/desorption at 77 K and CO2/TPD technique, respectively. The self-aldol condensation of propanal (Scheme 1) was conducted in a batchwise operation. For the catalytic reaction, propanal (0.72xl0"3dm3:10 mmol), MgO (50 mg:1.2 mmol) and water (0.1xl0"3dm3:5.6 mmol) were charged into a 25xl0"3dm3 stainless steel reactor, and the reactor was heated at 353 K. CO2 was subsequently charged into the reactor using a high-pressure liquid pump to the desired pressure(l IMPa). Supercritical state of the fluid phase at pressure of 11 MPa was confirmed by a high pressure view-cell observation. The reactions were started by stirring and continued for 6 h. After the reaction, reactor was cooled to 273 K, and the CO2 was released slowly. The yields and selectivities of products were obtained by GC and GC-MS. Microscopic structures of MgO catalysts after the reaction were observed with TEM and XRD.

Scheme 1 Self-aldol condensation of propanal.

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3. RESULTS AND DISCUSSION The properties of MgO catalysts are summarized in Table 1. The surface area and the basicity of the catalysts were increased with decreasing the particle size of MgO. Figure 1 shows the results of catalytic reaction under SCCO2 in the presence or absence of water. In the absence of water (Figure la), the conversion of la and selectivity of 3a were increased with decreasing the catalysts size, and catalyst 100A with the smallest particle size showed the highest activity in this reaction condition. As seen in Figure lb, the addition of water caused the improvement of reactivity, and the conversion of la and selectivity of 3a were increased twice as much as that in Figure la. In contrast, catalyst size did not effect to the results, the conversion of la or the selectivity of 3a. Table 1. The properties of MgO catalysts. CO2(m mol)/MgO(g)

Catalyst

Particle size (nm)

100A

14

120

1.94

500 A

60

28

0.36

1000A

105

16

0.19

2000 A

220

8

0.08

BET (rrfVg)

Fig. 1. The results of catalytic reaction in SCCO2 at 353 K for 6h under 11 MPa. a) in the absence of water, b) in the presence of O.lxlO"3 dm3 of water. Catalytic feature after the reaction were examined by TEM observation of the catalyst 100A used in the reaction in the presence or absence of water, and the TEM images were shown in Figure 2. As seen in Figure 2a, the morphology of the catalyst used in the absence of water was

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solid shape with about 10 nm in particle size, which was approximately the same as that of pristine MgO powder before catalytic reaction. On the other hands, the catalyst used in the presence of water had the indefinite structures with needle- or film-like morphologies which were about a few hundred nm in domain size. Further investigation of the catalysts features after the reaction were carried out by using XRD (Figure 3), and it was revealed that the chemical form of catalysts used under SCCO2 in the absence and the presence of water were MgO and hydromagnesite (Mg5(CO3)4(OH)2 4H2O:JCPDS 25-513), respectively.

Fig. 2. TEM images of MgO catalysts after catalytic reaction in SCCO2 at 353 K for 6h under 11 MPa. a) in the absence of water, b) in the presence of 0. lxl 0"3 dm3 of water.

Fig. 3. XRD profiles of MgO catalysts after catalytic reaction in SCCO2 at 353 K for 6h under 11 MPa. a) in the absence of water, b) in the presence of O.lxlO'3 dm3 of water.

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Generally the catalyst which has smaller particle size and larger surface area shows a higher catalytic activity. The trend in the reactivity of each catalysts seen in Figure la might be attributed to the strength of surface basicity, on the contrary, the catalytic activity of MgO seen in Figure lb disregards such a manner depending on the particle size. TEM and XRD results revealed that the catalysts used in the reaction system with a small amount of water suffer the change in its chemical component and morphology. The particle size of hydromagnesite derived from MgO was ten times lager than that of pristine MgO catalysts. Such structural change of catalysts might be caused by the dissolution of MgO with the carbonic acid during the catalytic reaction, and as a result, the original particle size of the catalyst which might affect the catalytic activity is invalidated. In contrast, the MgO catalysts used under SCCO2 in the absence of water maintain their particle size and morphology, and therefore the reactivity of catalysts well correlated with its particle size. As reported earlier [13], we proposed that the dehydration reaction of aldol compound might be promoted by carbonic acid as Bransted acid catalyst. To reveal the reason for the increase of the selectivity of 3a, catalytic activity of hydromagnesite were examined under scCO2 in the absence of water. The selectivity of 3a obtained by using hydromagnesite as catalyst was 12 %, which was lower than the selectivity in the case of MgO as seen in Figure la. Thus, this result indicates that hydromagnesite is not so active for the dehydration reaction as compared to MgO, and well supports our assumption that the promotion effect on the dehydration reaction in the presence of water might be caused by carbonic acid formed in the reaction system. Because water is more soluble under the supercritical conditions than the subcritical conditions, solvated water reacts with CO2 to generate carbonic acid which could promote the dehydration of 2a as Bransted acid. In the present results, the addition of water in the reaction system under SCCO2 increased not only the selectivity of 3a but also the conversion of la. Although the promotion effect on the conversion of la caused by the addition of water cannot yet be explained sufficiently, the reason for increasing the conversion of la is supposed as described below. At first, CO2 was dissolved in the water, and pH of water will be decreased with the formation of carbonic acid. As reported by Toews et al. [14], pH of aqueous solution is decreased to pH3 under the similar condition of the present study. Such an acidic solution dissolves MgO, and as a result, Mg2+ cation and especially OH" anion species with higher basicity as compared to MgO may be formed in the reactor. Consequently, we carefully propose that the conversion of la to 2a might be promoted by OH species derived from MgO in carbonic acid. The role of OH species on the surface of MgO catalysts in the aldol reaction was studied [15], and it was reported that OH species act as base and promote the aldol reaction. Thus, the reason for increasing the conversion of la under SCCO2 in the presence of water proposed here is not so curious. It is note worthy that the catalyst can be separated from the product as hydromagnesite after the reaction, even though MgO reacts like as homogenous catalyst under SCCO2.

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4. CONCLUSION We demonstrated the self-aldol reaction of propanal using basic catalysts of MgO with different particle size under SCCO2 in the presence or absence of a small amount of water, and the properties of catalysts were studied. The obtained results revealed that under SCCO2 with the small amount of water would affect the progress of the dehydration reaction of aldol derivatives, and the catalyst features are changed from MgO to hydromagnesite. Carbonic acid formed by the reaction of water and compressed CO2, which dissolves the MgO catalysts to form hydromagnesite during the reaction, as a result, size and chemical component of catalyst were drastically changed. The aldol condensation might be promoted by OH species derived from dissolved MgO with water, which act as base. Hence the effects of the difference in the particle size of catalysts used were not appeared in the presence of water. In contrast, notable change of catalyst feature could not be observed in MgO catalysts used under SCCO2 in the absence of water. In summary, this study has exhibited the advantage of addition of water to improve the catalytic activity of MgO for the aldol reaction under SCCO2. References 1. J. Hyde, W. Leitner and M. Poliakoff, in High Pressure Chemistry, R. Eldik and F.-G Klarner (eds), Wiley-VCH, Weinheim, 2002, p. 369. 2. S. Fujita, B. M. Banage, Y. Ikushima and M. Arai, Green Chem., 3 (2001) 87. 3. H. Kawanami and Y. Ikushima, Chem. Commun., (2000) 2089. 4. H. Kawanami and Y. Ikushima, Tetrahedron Lett., 43 (2002) 3841. 5. H. Kawanami and Y. Ikushima, J. Jpn. Petrol. Inst, 45 (2002) 321. 6. H. Kawanami, A. Sasaki, K. Matsui and Y. Ikushima, Chem. Commun., (2003), 896. 7. Y. Ikushima and M. Arai, Chemical Synthesis Using Supercritical Fluids, P. G Jessop and W. Leitner (eds.), Wiley-VCH, Weinheim, 1999. 8. W. Carruthers (ed.), Some Modern Methods of Organic Synthesis - 3rd edn., Cambridge Univ. Press., New York, 1986. 9. K. Tanabe, G Zhang and H. Hattori, Appl. Catal, 48 (1989) 63. 10. H. Tsuji, F. Yagi, H. Hattori and H. Kita, J. Catal., 148 (1994) 759. 11. E. Dumitriu, V. Hulea, I. Fechete, A. Auroux, F.-F. Lacaze and C. Guimon, Microporous Mesoporous Mater., 43 (2001) 341. 12. (a) Y. Anzai, M. Goto, A. Kodama and T. Hirose, Proc. 14th Symp. on Chem. Eng., Kyushu-Taejon/Chungnam, Taejon Univ., Korea, Dec. 1st 2001, p. 13; (b) M. Goto, Y. Anzai, A. Komada and M. Yoshida, Chem. Eng. Trans, 2 (2002) 85. 13. K. Matsui, H. Kawanami, Y. Ikushima and H. Hayashi, Chem. Commun, (2003) 2502. 14. K. L. Toews, R. M. Shroll, C. M. Wai andN. G. Smart, Anal. Chem, 67 (1995) 4040. 15. G Zhang, H. Hattori and K. Tanabe, Appl. Catal, 36 (1989) 189.


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Critical point and phase envelope calculations - some practical aspects based on CO2-mixtures H.A. Turunena, J.-P. Pokkib, R. Raudaskoskia, E. Turpeinena, E. Muurinena, K.I. Keskinenbc, and R.L. Keiskia a University of Oulu, Department of Process and Environmental Engineering, POB 4300, 90014 University of OULU, Finland. E-mail: [email protected] b Helsinki University of Technology, Dep. of Chem. Tech., POB 6100, 02015 HUT, Finland c Neste Engineering Oy, POB 310, 06101 PORVOO, Finland Critical point calculations for binary systems, CO2-hydrocarbon and CO2-methanol, were carried out by an indirect method, the phase envelope calculation method, using Flowbat process simulation software. The Soave-Redlich-Kwong equation of state was used. The calculated critical points of CO2-hydrocarbon mixtures were compared to the measured values collected from the literature. Most of the calculated values met rather well the measured ones. The phase envelope calculation for a binary system with an azeotrope, CO2-ethane, appeared to be a demanding task. The influence of supercritical conditions to mass transfer and reactor selection was also discussed. 1. INTRODUCTION The vapour liquid critical point (CP) of a fluid is a special point where the properties of liquid and vapour phases become similar. The CP of a pure component is a fixed value of temperature, pressure and molar volume. The CPs of pure components are tabulated for many common components in various reference books, articles and databases. For example, CP of carbon dioxide is Tc=304.19 K, pc=7.3815 MPa and vc=0.094 nrVkmol [1]. The CP of a mixture is strongly dependent on the components and the composition of the system. Only a small number of CPs of mixtures, mostly for binary mixtures, are published. The theory to calculate the CPs for mixtures was laid down by Gibbs in 1876 [2] about 50 years after the phenomenon itself was reported by Cagniard de La Tour in 1822 [3]. The development in the computer technology during the recent decades has given new possibilities to solve thermodynamic problems. Knowing the exact behaviour of fluids in different process conditions is a fundamental basis in designing safe and smoothly operating chemical processes. Reported advantages of the supercritical fluid (SCF) operation are enhancement of reaction kinetics and avoiding mass transfer limitations [4]. 2. EXPERIMENTAL The methods to calculate the CPs of mixtures can be divided into empirical and rigorous methods [5]. The empirical methods provide the pseudo CP of the mixture by relating the

370

critical properties of individual components with empirical composition-dependent mixing rules. The simplest empirical correlation for the critical temperature, Tc is Nc

T c > m = SziT c>i

(1)

i=l

where subscript m denotes mixture, i is the component, zj is the mole fraction and Nc is the total number of components. The same analogy can be applied also for the critical pressure and volume. A small set of common pseudocritical mixing rules can be found in the publication by Pokki et al. [6]. The rigorous methods can be divided into indirect and direct methods and they both need thermodynamic models. In the indirect methods the entire phase envelope is constructed, while the direct method provides the CP of a mixture directly [5]. The equation of state (EoS) describes the pVTz behaviour of mixtures in the fluid state. The "ideal gas law" and virial EoS are simple and famous EoS, but they are not capable to model critical phenomena and liquid phase properties. During the past century hundreds of modifications have been proposed to the cubic and higher degree EoS and their mixing rules in order to improve the accuracy and to extend the range of validity. Most of the equations have been developed empirically and numerable coefficients are determined by fitting all available experimental data using the criterion of minimum deviation between experimental and calculated data. [7] The methods to calculate the CPs of multicomponent mixtures based on EoS have been developed further by Peng and Robinson (1977), Heidemann and Khalil (1980), Assilineau et al. (1979), Michelsen (1980) [5,8] and Stockfleht and Dohrn (1998) [5]. The development of computers has made the rigorous methods more attractive compared to the empirical methods because of their better accuracy and consistency. The benefit of the empirical methods is their mathematical simplicity. [5] In this study the CP calculations by an indirect method, the phase envelope calculation method, for CO2-mixtures are presented. Also some aspects to the practical issues in the context of operation conditions and reactor selection have been presented, especially the use of supercritical fluids (SCF) is considered. 3. RESULTS AND DISCUSSION 3.1 Critical Point Calculation The Flowbat simulation program [9] has been used in the phase envelope calculations to obtain CPs by an indirect rigorous method. Figure 1 presents phase envelopes for the CO2methane mixtures in different concentrations. The equation of state used in simulation is Soave-Redlich-Kwong (SRK) [10], which has been proved to be suitable for these kinds of purposes in practice [5]. The binary interaction parameter used is 0.0933 and omega values for CO2 and CH4 are 0.239 and 0.012, respectively [9]. Crosses illustrate the calculated CPs of mixtures and the plus marks correspond to the measured CPs from the literature [11]. CPs are the particular points on the curve where vapour and liquid phases become indistinguishable, where the bubble-point and the dew-point lines meet. The area enveloped by the border-line curve represents a two-phase region in which both liquid and vapour are present in equilibrium. [12,13] The area above the CP (above critical temperature and pressure) is the region of supercritical fluid [13]. In the same figure the vapour pressure lines of pure components, starting from the triple-point and ending to the CP, are also presented. From Figure 1 it can easily be seen, that the CP as well as the shape of the phase envelope depend strongly on the mixture composition. In the CO2-methane mixtures CPs are higher

371

than CPs of pure components. The calculated CPs meet well the measured points in the three middle cases (err.2(88%) were found. The calculated CP meets well one of those points (err. PDI4) 2) 2) (wt%) (wt%) (bar) (his.) (g) 3 (|j,m) ' SI S2 S3 S4 S5 S6

3.0 4.5 6.0

1.0

10.0

70

350

48

6.0

2.0

10.0

65 70 75

350

24

1-2 2-5 1-5 3-4 1-4 2-3

61,000 70,000 80,000 51,000 35,000 31,000

3.0 3.1 2.5 2.1 2.5 2.6

1) poly(HDFDMA) 2) Weight percent on the monomer 3) Determined by FE-SEM 4) Determined by GPC

Fig. 2. SEM images of PS particles polymerized (SI) with 3.0g styrene, (S2) 4.5g styrene, (S3) 6.0g styrene, (S4) in 65 °C, (S5) in 70 °C and (S6) in 75 °C. 3.3 Copolymerization Table 2 shows the results of the dispersion copolymerization of poly(styrene-co-butyl acrylate) and poly(styrene-co-butyl methacrylate). Tg of copolymer is decreased as the butyl (meth)acrylate is added. Figure 2 shows the SEM images of three different polymers. Particle size of copolymer was about l-4um. And size distribution is broad than PMMA at similar reaction condition. At 20 wt% of «-butyl (meth)acrylate in copolymer, slightly aggregated

384

particle was obtained. So in this case, it is impossible to determine particle size. Table 2 Effect of comonomer ratio on St/BA2) Tg > entry (wt%) (°C) 99 S7 100/0 SBA1 85 95/5 SBA2 90/10 83 SBA3 60 80/20 3

l

dispersion copolymerization^ St/BMA 2 ) Mw4> PDI 4 ) entry (wt%) S7 1.8 22,000 10/0 SBMA1 1.5 15,000 95/5 SBMA2 1.6 18,000 90/10 SBMA3 1.4 80/20 16,000

Tx 3) g

(TC) 99 92 78 73

Mw4>

PDI 4 '

22,000 20,000 16,000 17,000

1.8 1.6 1.4 1.5

1) Reaction conditions: 6.0g comonomer, 0.24g AIBN, 0.6g poly(HDFDMA), 65 °C, 350bar and 48hours 2) St/B(M)A represents weight ratio of styrene to «-butyl (meth)acrylate 3) Determined by DSC 4) Determined by GPC

Fig. 3. SEM images of (S7) Polystyrene, (SBA1) Poly(styrene-co-butyl acrylate) (95/5) and (SBMA1) poly(styrene-co-butyl methacrylate) (95/5) particles. 4. CONCLUSION We could synthesize polystyrene, poly(styrene-co-butyl acrylate) and poly(styrene-comethacrylate) particles through dispersion polymerization in scCO2. These polymers were characterized by 'H-NMR, FT-IR, DSC, TGA, GPC and SEM. In general, when n-butyl acrylate or n-butyl methacrylate was added in polystyrene, Tg was decreased, molecular weight was slightly decreased and particle size was slightly increased. ACKNOWLEDGMENT This work was supported by the BK21 project of Ministry of Education and the Cleaner production program of KNCPC. REFERENCES 1. J. L. Kendal, D. A. Canelas, J. L. Young and J. M. Desimone, Chem. Rev., 99 (1999) 53. 2. M. A. McHugh and V. J. Krukonis, Supercritical Fluid Extraction, 2nd Ed., ButterworthHeinemann, Stoneham, MA, 1994. 3. E. D. Sudol, in: J.M. Asua (Ed.), Polymeric Dispersions: Principles and Applications, Kluwer Academic, Dordrecht, 1997. 4. E. J. Beckman, J. of supercritical Fluids, in print (2003). 5. S. Hong and M. Park, J. Korean Society for Imaging Science, 3 (1999) 40. 6. W. Bae, H. Lee, S. Kwon and H. Kim, Proceedings of the 6th International Symposium on Supercritical Fluid, Versailles, France, Tome 2 (2003) 1393.


385

Phenol hydrogenation over supported metal catalysts under supercritical carbon dioxide Chandrashekhar V. Rode, Uday D. Joshi, Takafiimi Sato, Osamu Sato, and Masayuki Shirai* Supercritical Fluid Research Center, National Institute of Advanced Industrial Science and Technology, Nigatake, Miyagino, Sendai 983-8551, Japan. E-mail: [email protected] Phenol hydrogenation was examined with several metal-loaded catalysts under supercritical carbon dioxide.

A charcoal-supported rhodium catalyst was highly active for the ring

hydrogenation.

The selectivity to cyclohexanone and cyclohexanol was controlled by

changing carbon dioxide and hydrogen pressures. 1. INTRODUCTION Hydrogenation reactions in liquid phase using supported metal catalysts have wide ranging applications in commodity, pharmaceutical and fine chemical production processes, however, reaction rates are not so high and the separation of pure products from solvents is critical.

Organic gas-liquid hydrogenations with solid catalysts under supercritical carbon

dioxide have several advantages; i) high reaction rates, ii) easy separation of catalyst and products, iii) environmentally benign processes without using toxic organic solvents. '

KA

oil (a mixture of cyclohexanone and cyclohexanol) is a raw material for adipic acid and s-caprolactam, which are important nylon intermediates, and commercially produced by the oxidation cyclohexane or the hydrogenation of phenol in chemical industry.

In this paper,

we investigated the catalytic hydrogenation of phenol with supported metal catalysts under supercritical carbon dioxide. 2. EXPERIMENTAL All catalysts used in this study were purchased from Wako Chemicals, Japan.

The

metal loadings of the charcoal-supported catalysts were 5 wt %.

The dispersion values of

metal particles were determined by a hydrogen adsorption method.

Hydrogenation reactions

were carried out in stainless-steel stirred batch reactors (50 ml capacity of liquid).

The

weighed amount of a supported metal catalyst (0.02 g) and phenol (0.02 mol) were placed in a

386

reactor and flushed with carbon dioxide.

After the reaction temperature was attained,

hydrogen and carbon dioxide were introduced into the reactor to the desired pressure.

After

the reaction was over, the pressure from the reactor was released slowly and contents were discharged, diluted with methanol and catalyst was separated by simple filtration. The liquid sample was then analyzed with GC-FID and GC-MS. 2 3. RESULTS AND DISCUSSION The catalyst screening results for the phenol hydrogenation under supercritical carbon dioxide are shown in Table 1. Charcoal-supported rhodium and ruthenium catalysts were highly active for the ring hydrogenation of phenol to cyclohexanone and cyclohexanol at 353 K under supercritical carbon dioxide conditions and the conversion of phenol for different catalysts were in the following order: Rh/C > Ru/C »

Pd/C > Pt/C. Hitzler et al. have

reported that ring hydrogenation of cresols over a commercial palladium catalyst proceeded under supercritical carbon dioxide at very high temperature (523 K).

However, maximum

selectivity to the corresponding cyclohexanol derivative was only 70 %, with formation of several other side products.

3

In our work, at very low temperature (353 K) rhodium was

found to be the most active for the ring hydrogenation of phenol, while for gas and liquid phase phenol hydrogenation at temperatures > 453 K palladium is reported to be a suitable catalyst for this hydrogenation. 4'5 Table 1 Catalyst screening for the hydrogenation of phenol' Catalyst

a

Dispersion(%)

Conversion (%)

Selectivity (%) Cyclohexanone

Cyclohexanol

5% Rh/C

12

35

68

32

5% Ru/C

25

11

69

31

5% Pd/C

3

0.9

91

9

5% Pt/C

8

0.6

80

20

Reaction temperature 353 K; reaction time 10 min.; hydrogen partial pressure 3 MPa; carbon

dioxide partial pressure 10 MPa; initialn amount of phenol 0.02 mol. Figure 1 shows the conversion profile of phenol hydrogenation over 5% Rh/C catalyst at 353 K.

Phenol hydrogenation is a successive reaction in which first phenol is hydrogenated

to cyclohexanone followed by its hydrogenation to cyclohexanol4 however, it is interesting to note that the selectivity to cyclohexanone and cyclohexanol did not depend on the conversion under supercritical carbon dioxide conditions, indicating that cyclohexanol would be formed

387

directly from phenol (Scheme 1).

Figure 1 The hydrogenation of phenol under hydrogen partial pressure 3.0 MPa, carbon dioxide partial pressure 10.0 MPa over a 5 wt% Rh/C catalyst (0.02 g). Initial phenol 0.02 mol; Reaction temperature 353 K; : phenol; : cyclohexanone; T : cyclohexanol.

Scheme 1 Reaction pathway of phenol hydrogenation. We examined the influence of the partial pressure of hydrogen on activity and selectivity for 5 wt % Rh/C catalyst (Figure 2).

For this purpose, separate hydrogenation experiments

were carried out at three different hydrogen pressures while carbon dioxide partial pressure was kept constant (10 MPa) in all these experiments.

It was found that both hydrogenation

activity and selectivity to cyclohexanol increased with increasing hydrogen pressure (at 10 MPa carbon dioxide), caused by an increase in the concentration of surface hydrogen with increasing hydrogen pressure. Cyclohexanol was observed at low phenol conversion under

388

high hydrogen pressure, indicating that cyclohexanol would be formed not only via the cyclohexanone intermediate but also directly from phenol.

At high concentration of surface

hydrogen, the hydrogenation of phenol to cyclohexanol would easily occur, resulting in higher selectivity to cyclohexanol.

Figure 2 Effect of hydrogen pressure on the hydrogenation of phenol over a 5 wt% Rh/C catalyst. Reaction temperature 353 K; reaction time 10 min.; CO2 10 MPa. : conversion; selectivity to A: cyclohexanone; T : cyclohexanol. We also examined the influence of the carbon dioxide pressure of 5 wt % Rh/C catalysts. The hydrogenation activity increased with increasing carbon dioxide pressure.

It is probable

that enhanced mass transfer contributes to the increased activity at higher carbon dioxide concentration. REFERENCES 1. M.GHitzler and M.Poliakoff, Chem. Commun. (1997) 1667. 2.

C.V.Rode, U.D.Joshi, O.Sato and M.Shirai, Chem. Commun, (2003) 1960.

3.

M.G, Hitzler, F.R.Smail, S.K.Ross, and M.Poliakoff, Org. Proc. Res. Dev, 2 (1998) 137.

4.

S. Narayanan, and K. Krishna, Appl. Catal. A, 174 (1998) 221.

5. N.Mahata and V.Visshwanathan, J. Mol. Catal., 120 (1997) 267.


389

Synthesis of Various Block Copolymers Containing Poly(L-lactide) in Supercritical Carbon Dioxide Ha Soo Hwang, Eun Ju Park, Yeon Tae Jeong, Hoon Heo, and Kwon Taek Lim* Division of Image Science and Engineering, Pukyong National University, Pusan 608-739, Korea. Poly(L-lactide) (PLLA) and its block copolymers were synthesized by a ring-opening polymerization of L-lactide in supercritical CO2 (scCC^) using 1-dodecanol, lH,lH-perfluoro1-decanol, monomethoxy polyethyleneglycol (mPEG-OH, Mw = 550), poly(dimethylsiloxane) mono hydroxy terminated (PDMS-OH, Mw = 4670), and polypropyleneglycol monobutylether (buPPG-OH, Mw = 1000) as initiators. The homopolymer and di-block copolymers could be prepared with fairly low molecular weight distribution (1.11 - 1.18) as analyzed by GPC. 1. INTRODUCTION Polylactides and its copolymer with PEG have received great interest in medical applications because of their excellent biodegradabilities and biocompatibilities [1-3]. Though polymers can be made either in bulk and solution, bulk polymerization is favorable for medical applications as it eliminates the use and potential residual of organic solvents. Recently ring-opening precipitation polymerization of n,L-lactide and glycolide was reported in non-toxic solvent of SCCO2 [4]. In addition to its use as an environmentally friendly solvent, CO2 could also be utilized to purify products via supercritical fluid extraction following polymerization. The present work aims at the synthesis of PLLA and its block copolymers via ring-opening polymerization of LLA in the presence of various CC>2-soluble initiators. PDMS and low molecular weight PEG and PPG are known to be soluble in CO2 [5]. This research may lead to a novel synthetic route to PLA based block copolymers in SCCO2, otherwise it is difficult to be prepared due to the bad solubility of polymers in CO2. In addition, fluorinated groups are a very interesting and useful class of materials for surface modifications due to their low surface energy [6].

* E-mail: [email protected] This work was supported by the program for cultivating graduate students in regional strategic industry from Korea Industrial Technology Foundation

390

2. EXPERIMENTAL 2.1 Materials L-Lactide, stannous 2-ethylhexanoate (Sn(Oct)2), 1-dodecanol, lH,lH-perfluoro-l-decanol, mPEG-OH (Mw = 550), PDMS-OH (Mw = 4670), and buPPG-OH (Mw = 1000) were purchased from Aldrich. Dry CO2, 99.99% purity, was purchased from Myungsin General GAS CO., LTD. All reagents were used as received. 2.2 Polymerization Pre-determined amount of L-Lactide, Sn(Oct)2, an initiator and teflon-coated stirred bar were introduced into a stainless steel view cell reactor (36ml) prior to pressurization. The reactor was pressurized by ISCO syringe pump (Model 260D) containing compressed CO2. Following pressurization, the reactor was heated to 70°C by immersing the reactor in a water bath in which temperature was controlled by the Thermo-Circulator (LABTECH®). The polymerization was conducted in supercritical CO2 at 3300 psi for 6 hrs. After polymerization, any unreacted monomers, Sn(Oct)2 and initiators were extracted with ~ 20 mL of liquid CO2 at the reaction pressure and ambient temperature. The CO2 was vented from the top of reactor as a gas. 2.3 Characterization Size exclusion chromatography (SEC) was carried out with a HP 1100 apparatus equipped with a set of three columns (104-103-102 A: polymer standards service) with CHCI3 as the eluent. Polystyrene samples were used as standards to construct the calibration curve. ! HNMR spectra of the homo-and di-block copolymers were obtained in CDCI3 using a JNMECP 400 (JEOL). 3. RESULTS AND DISCUSSION The ring-opening polymerization of L-lactide is performed by using hydroxy-terminated compounds as initiators and Sn(Oct)2 as a catalyst, thus CO2-soluble polymers with hydroxyl end groups can be used as the macroinitiator for synthesis of polylactide based block copolymers in SCCO2 (Scheme 1). The ring-opening polymerization of LLA was studied in a systematic way in SCCO2 at 70 °C under pressure range 207 to 228 bar. At the beginning of

391

Table 1. Ring-opening polymerization of L-lactide with various initiators in SCCO2 Polymer

LA/Initiator (mol/mol)

Initiator

PLLA

Mn (obs)a

Mn (obs)b

M w /M n b

8990

1.12

1H,1H-Perfluoro1-decanol 1-dodecanol

694 694

-

6910

1.11

PEG-6PLLA

MPEG-OH

42

5750

5420

1.15

PLfr

™"

694

6500

6000

1.18

POMS-OH

65

8000

7634

1.12

PLLA

J™f

Calculated from H-NMR analysis, determined by GPC. the polymerization, the reaction medium is completely transparent and the coexistence of two phases is not observed through the sapphire windows. Nevertheless, after a few minute, the solution turned cloudy due to precipitation of the growing PLLA chains. The conversions of the polymerizations were calculated to be about 50 - 60 % from weighing the products. The compositions of the polymers and their molecular weights were determined by ' H-NMR and GPC analysis, respectively. From Table 1, it appears that the molecular weight of PLLA synthesized from lH,lH-perfluoro-1-decanol initiator is higher than those from 1-dodecanol initiator. It may be suggested that highly CO2-philic fluoro alkyl chain provides some stabilization effect to the polymer in the early stage of polymerization. For the PEG-6-PLLA, there is an excellent agreement between theoretical and observed molecular weight. Thus the molecular weight of the resulting copolymers could be controlled by the monomer to initiator (LLA/mPEG-OH) ratio. In contrast, the molecular weight of PPG-6-PLLA and PDMS-6-PLLA was found to be lower than expected. The reason for the deviations is now under investigation.

6.0

5.0

4.0

3.0

2.0

1.0

0 6.0

5.0

4.0

3.0

2.0

1.0

0

(A) (B) Fig. 1. ' H-NMR spectrum of PEG-6-PLLA (A) and PDMS-6-PLLA in CDC13(B)

392

Fig. 2. GPC trace of block copolymers Typical ]H-NMR spectra of PEG-6-PLLA and PDMS-6-PLLA copolymers are shown in Fig. 1. The molecular weights of block copolymers can be calculated by comparing methine proton peak (5.2 ppm) of PLLA block with methylene protons (3.66 ppm) of PEG block and methyl protons (0.07 ppm) of PDMS block, respectively. Fig. 2 illustrates the GPC curves of the homo and di-block copolymers. The peaks of starting PDMS-OH and PEG-OH shifts toward the higher molecular weight copolymer after block copolymerization and the molecular weight distribution of the copolymer was analyzed to be narrow (the polydispersity is below 1.18), indicating that block copolymers were produced successfully. 4. CONCLUSION The ring-opening polymerization of L-lactide in scCO2 was investigated using various CO2soluble initiators and Sn(Oct)2 catalyst. Well-defined PLLA based block copolymers with narrow molecular weight distribution could be prepared by utilizing CO2-soluble macroinitiators such as mPEG-OH, buPPG-OH, and PDMS-OH. This polymerization technique may have the potential to form high purity PLLA based block copolymers for biomedical applications. REFERENCES 1. G. Schwach, J. Coudane, R. Engel, M. Vert, Polym. Bull., 32 (1994) 617 2. J. H. Jeong, D. W. Lim, D. K. Han, T. G. Park, Colloids Surf. B., 18 (2000) 371 3. C. S. Kweon, D. -W. Kim, H. -K. Cho, S. -T. Noh, J. Ind. Eng. Chem., 9 (2003) 146 4. D. D Hile, M. V. Pishko, Macromol. Rapid Commun., 20 (1999) 511 5. M. L. O'neill, Q. Cao, M. Fang, K. P. Johnston, S. P. Wilkinson, C. D. Smith, J. L. Kerschner, and S. H. Jureller, Ind. Eng. Chem. Res. 37 (1998) 3067 6. W. K. Lee, I. Losito, J. A, Gardella Jr., and W. L. Hicks Jr., Macromolecules, 34 (2001) 3000


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Vapour pressure of CO2 as well as phase envelopes and critical points for CO2 and CH 4 mixtures calculated by Aspen Plus and Flowbat simulation software R. Raudaskoski*, E. Turpeinen, H. Turunen, K. Rahkamaa-Tolonen and R.L. Keiski Department of Process and Environmental Engineering, P.O.Box 4300, 90014 University of Oulu, Finland. E-mail: [email protected] The calculation of vapour pressure of CO2 as well as phase envelopes and critical points with four different mixtures of CO2 and CH4 has been carried out with two process simulation software, Aspen Plus and Flowbat, using Soave-Redlich-Kwong (SRK) equation of state. A comparison has been made between calculated and measured values, collected from the literature. Calculated vapour pressures and most of the calculated phase envelopes and critical points are rather similar to the measured values. 1. INTRODUCTION CO2 is a major contributor to the greenhouse effect. Its content in the atmosphere is increasing at the rate of about 1 percent per year, from 250 ppm in the pre-industrial period to circa 400 ppm today. The increase in atmospheric CO2 content is mainly caused by the combustion of fossil fuels and deforestation. [1] The greenhouse effect causes e.g. global warming, a rise in sea level and an increase in the desert formation in the tropical zone [2]. In order to avoid the climate change, it is necessary to mitigate or control the CO2 concentration in the atmosphere. CO2 separation from flue gases and its conversion to other chemicals is one possible method of reducing CO2 content in the atmosphere. Recently, many kinds of laboratory experiments have been done to determine possible reactions and applicable reaction conditions for the CO2 chemical utilisation [2-6]. Reactions between CO2 and CH4 are especially attractive, because CH4 is also a greenhouse gas. Possible products from those reactions are e.g. carbon monoxide, hydrogen, methanol and acetic acid [2]. Unfortunately, laboratory experiments are very expensive and time-consuming ways to obtain new knowledge. Calculating thermodynamics and especially phase envelopes can help to find suitable reaction conditions for the laboratory experiments, which saves time and reduces expense. It is very important that the calculated values are correct. In this work, phase envelopes for different CO2 and CH4 mixtures calculated by two different process simulation software are shown. In addition, vapour pressure of CO2 and critical points for CO2 and CH4 mixtures are presented. Furthermore, the calculated values are compared to the measured values, collected from the literature [7, 8]. 2. EXPERIMENTAL Calculation of vapour pressure of CO2 as well as phase envelopes and critical points for different mixtures of CO2 and CH4 has been carried out with two different process simulation software; Aspen Plus (Version 11.1) and Flowbat. Aspen Plus is a commercial software,

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while Flowbat was created and developed mainly in a research environment, but it is also used in industry. In the calculation of vapour pressure of CO2, the SRK equation of state was used [9]. A comparison has been done between calculated and measured vapour pressures, collected from the literature [7]. Phase envelopes and critical points for the mixtures of CO2 and CH4 have been calculated with four different mixture ratios using also the SRK equation of state [9]. These calculated phase envelopes and critical points are compared to measured values, reported in the literature [8]. 3. RESULTS AND DISCUSSION 3.1. Vapour pressure of CO2 The vapour pressures of CO2 calculated by Aspen Plus and Flowbat process simulation software are shown in Figure 1, starting from the triple point and ending at the critical point. In the same Figure, the measured vapour pressure of CO2, collected from the literature [7], is presented. The calculated vapour pressures are quite similar to each other and to the measured vapour pressure.

Figure 1. Vapour pressure of CO2. 3.2. Phase envelopes and critical points for CO2 and CH4 mixtures Phase envelopes for four different mixtures of CO2 and CH4 calculated by Aspen Plus and Flowbat are shown in Figure 2. The area enclosed by the borderline curve represents a twophase region in which both liquid and vapour phases are in equilibrium. Above this region, all the fluid is in the liquid phase, whereas below this region, only vapour is present in the system. Usually, it is easier to carry out laboratory experiments in conditions where only one phase is present. The higher vapour-liquid equilibrium curve is the bubble point line and the lower one is the dew point line. As can be seen from in Figure 2, the dew point and the bubble point lines calculated by Aspen Plus and Flowbat simulation software are exactly the same. Comparing the bubble point and the dew point lines to the measured data, collected from the literature [8], indicates that the bubble point and dew point lines when 54.3%, 70.5% or 88% CO2 is in the mixture are quite similar to the measured data. When 18% CO2 is in the mixture the calculated and measured [8] lines have some differences near the critical region. The bubble point and the dew point lines meet at the critical point (marked on the 70.5% CO2 + 29.5% CH4 phase envelope calculated by Flowbat). At the critical point the liquid and vapour phases become similar. The mixture is in a supercritical state when its pressure is

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Figure 2. Phase envelopes for different mixtures of CO2 and CH4 calculated by Aspen Plus and Flowbat and one example of a critical point calculated by Flowbat. higher than the critical pressure and its temperature is higher than the critical temperature. The reported benefits when operating under the supercritical conditions are the enhancement of reaction kinetics and the avoidance of the mass transfer limitations [10]. Critical temperatures and pressures for different mixtures of CO2 and CH4 calculated by Aspen Plus and Flowbat have been indicated in Table 1. In the same Table the measured critical points for the same mixtures are presented. The accuracy of the calculated critical points compared to the measured ones have been defined using the following equation: A c c u r a c y = 100%*((xcaicuiated-XmeaSured)/xmeasured)

(I)

where x is the critical temperature or pressure. This equation gives the relative errors. Relative errors are presented in Table 2. The relative errors are quite small except for the critical pressures in the case of 18% CO2 and 82% CH4 calculated by Aspen Plus and Flowbat. Table 1. Critical temperatures (T^) and critical nressures flM for CO-> and Composition Literatiire [8] AspenLPIUS Pc(kPa) Po(kPa) Tc(K) T 0 (K) 8377 88% CO2 + 12% CH4 286.48 297.37 8248 8618 278.65 8653 70.5% CO2 + 29.5% CH4 273.71 261.44 54.3% CO2 + 45.7% CH4 256.48 8446 8653 18% CO, + 82% CH4 222.04 6791 214.39 6110 Table 2. Relative errors (%) of calculated critical points compared points Composition Aspen Plus Tc(%) Pc(%) 88% CO2 + 12% CH4 3.80 -1.54 70.5% CO2 + 29.5% CH4 1.80 0.41 54.3% CO2 + 45.7% CH4 1.93 2.45 18%CO 2 + 82%CH 4 -3.45 -10.03

CFL mixtures Flow/bat Pc (kPa) TC(K) 7998 294.85 278.71 8664 257.89 8628 213.77 6115

to the measured critical Flowbat T c (%) 2.92 1.83 0.55 -3.72

Pc(%) -4.52 0.53 2.15 -9.95

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4. CONCLUSION The primary scope of this work was to define the accuracy of phase envelopes for CO2 and CH4 mixtures calculated by two different process simulation software; Aspen Plus and Flowbat, using the SRK equation of state. Reactions between CO2 and CH4 are very relevant because they both are greenhouse gases. Phase envelopes of their mixtures can give useful information for experimental design, because already at research stage one should try to aim at doing the experiments under the a single phase operation conditions. From the point of view mass and heat transfer it is known that it is preferable to operate under a one phase region than under a two-phase region. However, before the information from the phase envelopes can be used, the validity of the data must be ensured. Possible experimental errors must also be taken into account. The calculated dew point and bubble point lines are quite similar to the measured data except in the phase envelopes of 18% CO2 and 82% CH4 near the critical region. In addition, the vapour pressure of CO2 as well as the critical points for four different mixtures of CO2 and CH4 have been presented. Vapour pressures calculated by Aspen Plus and Flowbat are quite similar to each other and to the measured vapour pressure. The relative errors of calculated critical points compared to measured critical points are quite small except for the critical pressures in the case of 18% CO2 and 82% CH4. Overall, calculated values, presented in this study, can be considered quite reliable. ACKNOWLEDGMENTS This work has been carried out with the financial contribution from the Graduate School in Chemical Engineering (GSCE) and the Finnish Foundation of Technology (TES). REFERENCES 1. X. Xiaoding and J.A. Moulijn, E & F, 10 (1996) 305. 2. M.M. Halmann and M. Steinberg, Greenhouse Gas Carbon Dioxide Mitigation: Science and Technology, Lewis Publishers, Florida, 1999. 3. S. Wang, K. Murata, T. Hayakawa, S. Hamakawa and K. Suzuki, Appl. Catal. A, 196 (2000) 1. 4. S. Wang and G.Q. (Max) Lu, E & F, 10 (1996) 896. 5. Y. Choi, K. Futagami, T. Fujitani and J. Nakamura, Appl. Catal. A, 208 (2001) 163. 6. B.J. Liaw and Y.Z. Chen, Appl. Catal. A, 206 (2001) 245. 7. D.R.Lide (edit.), CRC Handbook of Chemistry and Physics, 74th edition, CRC Press, Inc., Florida, 1993. 8. H.G. Donnelly and D.L. Katz, Industrial and Engineering Chemistry, 46 (1954) 511. 9. G. Soave, Chemical Engineering Science, 27(1972)1197. 10. G. Ertl., H. Knozinger and J. Weitkamp (Ed.), Handbook of Heterogeneous Catalysis, VCH, Weinheim, 1997.


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CO 2 Recovery Pilot Plant Kyoung Ryong Jang*, Jun Han Kim, Jae Goo Shim, Young Mo Ahn and Hee Moon Eum Radiation and Environment Group, Korea Electric Power Research Institute, 103-16 Munjidong, Yuseong-gu, Daejeon 305-380, South Korea

The pilot plant in this study was using the recycle separation process adopted the absorption and stripping. Recovery of CO2, one of the main parameters, was experimented at different temperatures and flow rates of flue gas and absorbent, Monoethanolamine(MEA). When the temperature of flue gas was increased by 10°C from 30 °C to 40 °C, CO2 recovery ratio was decreased by around 5%. When the flow rate of flue gas was increased to 690m3/hr, equivalent to 120% of design value of pilot plant, CO2 recovery ratio was decreased by around 15%. CO2 recovery ratio was maintained stably at both 15wt% and 25wt% of MEA concentration over 2.5 m /hr of MEA flow rate. 1. INTRODUCTION It is well known that global warming caused by increase of greenhouse gas like CO2 and CH4 in the air gives many ill effects to the earth. IPCC joint-established in 1988 by UN Environmental Plan(UNEP) and World Meteorology Organization(WMO) with the aim of preparing countermeasures for global climate change, predicted uprise of geo-surface temperature by 1.0-3.5 °C and of sea level by 15 ~ 95 cm in the report on Global Climate Change in July, 1997 [1]. To keep current energy system effectively and abate CO2, the unavoidable byproduct from industrial activities, it is essential to develop related technologies. Around the world, many researches to seek for the solution to reduce CO2 are under way, but yet such an effective method as not to affect industrial productivity has not developed. Of the CO2 abatement methods under research, it is known that chemical absorption process is suitable when CO2 concentration in flue gas is comparatively low(about 10% or so) as thermal power plants emit. CO2 separation process using chemical absorption

* Corresponding author. Tel. : +82-42-865-5480 ; Fax : +82-42-865-5504 E-mail address : [email protected]

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has some problems, for example, deterioration of absorbent, corrosion of equipment, and high energy consumption. So many researchers have tried to solve this obstacles developing new style of absorbents [2]. In this study with the long term goal of being a commercialization of CO2 absorptionseparation technology, a pilot plant with the capacity of 2 ton CCVday using chemical absorption process was constructed to separate CO2 from flue gas of a thermal power plant. MEA(Monoethanolamine) was used as an absorbent and evaluated of its absorption characteristics in a series of experiments giving some kinds of variations like temperature and flow rate of MEA, temperature and rate of flue gas, and MEA concentrations [3-4]. 2. EXPERIMENTAL 2.1 Experimental pilot plant The pilot plant used in the study was followed by the general recycle separation process adopted the absorption and stripping. And this unit was constructed near unit 5 of Seoul Thermal Power Plant in order to use the real flue gas (Fig. 1). Absorber is 18.8m high and 0.46m in diameter, and stripper is 16.7m high and 0.34m in diameter, of which capacity is equivalent to treat 2 ton CO2 per day. Describing the process, flue gas was injected to the bottom of the absorber by blower and then reacted with MEA, while treated flue gas was discharged to the existing stack through the top of absorber. MEA was injected to the top of the absorber to react with flue gas. MEA solution after reaction with CO2 (rich amine) was collected in the bottom of the absorber, and then pumped to the top of the stripper, where it was separated into MEA and CO2. Separated CO2 can be either converted into various high value-added compounds or transferred for storage. CO2 separated MEA solution (lean amine) passed through the heat exchanger and the filter, and recycled into the absorber to react with CO2 again. A certain amount of packing materials was filled in the absorber to allow CO2 and MEA to contact effectively as well as to increase the residence time. Stripper was also filled with packing materials so as to make easy separation as a washing zone. This was done by heating the solution at the reboiler located on the bottom of the stripper. Cooling equipment was installed in front of the absorber to maintain the proper experimental temperatures. Control system is installed for adjusting the main accessories and equipments, and whose remote indicators were enable to monitor and control through computer. 2.2 Experimental method MEA, the CO2 absorbent, its concentration was adjusted to the experimental condition by adding pure water. Flue gas used in the study was real and made from the thermal power plant. CO2 concentration of the flue gas streamed into the pilot plant was analyzed as a real

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time by ND-IR typed analyzer before entering the absorber, after reacting with MEA inside the absorber. CO2 recovery ratios were calculated from the differences. And also MEA solution before and after reaction were sampled and analyzed to check the loading values for lean and rich amine by the acid-base neutralization titration method.

Fig. 1. Schematic diagram of pilot plant.

3. RESULTS AND DISCUSSION 3.1. Effects of MEA temperature and flow rate Fig. 2 shows CO2 recovery ratios to MEA temperature and MEA flow rate at 15% wt MEA concentration. The lower was the temperature of MEA, the higher was the CO2 recovery ratio. It is highly estimated that it is caused that the reaction of CO2 with MEA is an exothermic process which the reaction was active at low temperature. CO2 loadings of lean and rich amine to flow rates of MEA were like Fig. 3. This shows the highest CO2 loading was around 0.66 at the flue rate of 2.0m3/hr. But the optimum of loading is needed to control less than 0.6 because it is known that corrosivity increase as loading goes higher[5-6]. CO2 loading means CO2 moles absorbed per 1 mole MEA. Generally CO2 loading of rich amine is the loading of residual MEA after reaction with CO2 in the absorber, and CO2 loading of lean amine is a value after MEA was separated from CO2 in the stripper. As for rich amine, MEA, which is the primary amine, reacts theoretically with 0.5-1.0 mole CO2 per 1 mole MEA, according to final products, either carbamate or bicarbonate[7-10].

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Fig. 2. CO2 recovery ratios on MEA flow rate and temperature.

Fig. 3. CO2 loading of lean and rich amine on MEA flow rate and temperature.

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3.2 Effects of flue gas temperature and flow rate Temperature and flow rate of flue gas are the fundamental parameters which should be obtained when designing and operating the pilot plant. So the effects of CO2 recovery ratio from those two parameters were checked. These experiments had been done at the gas volume between 570m3/hr and 690m3/hr. Fig. 4 shows the CO2 recovery ratios to flow rates of flue gas at the temperature of flue gas 40 °C. And this also does CO2 recovery ratios were more than 90% until flue gas flow rate went to 640m3/hr. And CO2 loadings didn't change a lot compared to the changes of CO2 recovery ratio.

Fig. 4. CO2 recovery ratio and loading on flow rate of flue gas at 40 °C.

Fig. 5 shows CO2 recovery ratios on flue gas flow rate and temperatures, 30°Cand 40 °C. In the above experiments about CO2 recovery ratio to the different MEA temperature, the change between the two items was not so big. But in this experiment, when the flue gas temperature was increased to 40 °C, CO2 recovery ratio was decreased very sharply, and when the flue gas flow rate was increased, CO2 recovery ratio was decreased also a lot. It is possible to assume that the quantity of flue gas flow rate gave bigger effects to the process more than MEA concentration did.

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Fig. 5. CO2 recovery ratio on flow rate of flue gas at 3 0 °C and 40 "C.

3.3. Effects of MEA concentration Absorbent has more economic advantages when it absorbs CO2 gas more at the similar operation conditions. In this experiment, CO2 recovery ratio and amine loading were examined at various MEA concentrations, 10wt%, 15wt%, and 25wt%. As in Fig. 6, 15wt% and 25wt% MEA concentration showed more than 90% of CO2 recovery ratio over 2.5m3/hr of MEA flow rate. At the 10wt% MEA, CO2 recovery ratio had much lower values compared to the other MEA concentrations, but this ratio was gotten over 95% especially when MEA flow rate was increased to far over 3.0 m /hr. CO2 loadings of lean amine and rich amine at various concentrations are presented in Fig. 7. MEA concentration of 10wt% and 15wt% showed similar pattern to the variation of loading. At 2.0 m3/hr of MEA flow rate, CO2 loading of rich amine went over 0.65 which regarded as higher value for the corrosion. And 25wt% MEA showed the most stable results in CO2 loading to flow rate.

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Fig. 6. CO2 recovery ratio on MEA flow rate and concentration.

Fig. 7. CO2 loading of lean- and rich-amine on MEA flow rate and concentration.

4. CONCLUSIONS Using a pilot plant for absorption and separation of CO2, a series of experiments with real flue

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gas from a thermal power plant were carried out and the conclusions are as follows. 1) Recovery of CO2 was experimented at different temperatures and flow rates of flue gas and absorbent, MEA. When the temperature of flue gas was increased by 10°C from 30 °C to 40 °C, CO2 recovery ratio was decreased by around 5%. When the flow rate of flue gas was increased to 690m3/hr, equivalent to 120% of design value of pilot plant, CO2 recovery ratios was decreased by around 15%. As for MEA, the absorbent, when its temperature was varied to 40 °C, 45 °C, and 50 °C, CO2 recovery ratios decreased in proportion to flow rate. This means that temperature and flow rate of flue gas, rather than those of MEA, give more effects to CO2 recovery ratios. 2) More than 95% of CO2 recovery ratio was maintained stably at both 15wt% and 25wt% of MEA concentration over 2.5 m3/hr of MEA flow rate. On the other hand, it didn't reach 95% until MEA flow rate became over 3.0 m3/hr, at 10wt% MEA concentration. As for CO2 loadings, all the concentrations didn't give good values at the standpoint of corrosion problem below 2.5m3/hr of MEA flow. And about the 25wt% of MEA concentration, lean amine had a too high CO2 loading which shows low CO2 absorption. So all things considered, the optimal MEA concentration and flow rate is 15wt% and 3.0m /hr, respectively, considering both CO2 recovery ratio and loading value. ACKNOWLEDGMENTS This work was supported by Korea Energy Management Corporation R&D headquarter, Grant 2000-C-CD02-P-01, and by Korea Electric Power Corporation. REFERENCES 1. IPCC, The Regional Impacts of Climate Change : An Assessment of Vulnerability, (1997). 2. Chakma, A. and Tontiwachwuthikul, P., Greenhouse Gas Control Technologies, (1999) 35. 3. Kohl, A. L. and Riesenfied, F. C , Gulf Publishing Co., Houston, (1979) 28. 4. Mimura, T., Simayoshi, H., Suda, T., Iijima, M. and Mituoka, S., Energy Convers. Mgmt, 38(1997)s57. 5. Veawab, A., Tontiwachwuthikul, P. and Bhole, S. D., Ind. Eng. Chem. Res., 36(1) (1997) 264. 6. Mimura, T., Shimojo, S., Kagaku Kogaku Ronbunshu, 21(3) (1995) 478. 7. Caplow, M., J. Am. Chem. Soc, 90 (1968) 6795. 8. Danckwerts, P. V., Chem. Eng. Sci., 34 (1979) 443. 9. Xu, S., Wang, Y. W., Otto, F. D., and Mather, A. E., Chem. Eng. Sci., 51 (1996) 841. 10. Sharma, M. M., Trans. Faraday Soc, 61 (1965) 681.


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A 2-stage PSA process for the recovery of CO2 from flue gas and its power consumption Soon-Haeng Cho, Jong-Ho Park, Hee-Tae Beum, Sang-Sup Han, and Jong-Nam Kim Korea Institute of Energy Research, 71-2, Jangdong, Yusungku, Taejon, 305-343, Korea FAX: +82-42-860-3021. E-mail: [email protected] A PSA process for the recovering of CO2 from flue gas is usually composed by 2-stage adsorption unit. At the first stage of the 2-stage PSA, CO2 is enriched to 40-60%, and then further concentrated to 99%, at the second stage. Performance of the 2-stage PSA process for CO2 recovery from flue gas is investigated experimentally and theoretically. Power consumption of a pilot scale process was experimentally measured and it ranged from 1.26 kWh/Nm3-CO2 to 1.52 kWh/Nm3-CO2 depending on the type of vacuum pump used. By a simulation study on the power consumption, it was shown that substantial reduction of power consumption could be obtained with a more simple process scheme. With the 2-stage PSA process composed of 2-bed for each stage, 99% CO2 is obtained with 80% recovery from the feed containing 10.5% CO2. Theoretical power consumption of the above process was estimated as 0.28 kWh/Nm3-CO2. 1. INTRODUCTION Carbon dioxide is considered as the main cause of global warming. For the sustainable development, efforts are being undertaken to mitigate the emission of CO2 to the atmosphere. Capture and storage of CO2 to the ocean or depleted oil field is considered as one of the possible options for the CO2 abatement. In order to make the option to be viable, the cost for the capture of CO2 from its source should be as small as possible. The pressure swing adsorption (PSA) process is a highly efficient gas separation process, and it can be applied for the removal of carbon dioxide from various gas mixtures [1]. The configuration of the PSA process for CO2 recovery varies according to the CO2 concentration of the feed. When the concentration of CO2 is higher than 20%, a 1-stage PSA process is enough to produce 99%CO2 with high recovery [2]. On the other hand, when the concentration of CO2 is lower than 15%, it is difficult to recover CO2 with high purity and recovery using the 1-stage PSA process. In this case, a 2-stage PSA process is applied [3]. In the first part of this study, the performance of a pilot-scale 2-stage PSA process that could produce CO2 of 99% at 13Nm3/hr is reported briefly. The power consumption of the pilot scale process was significantly affected by the operating conditions such as purge amount at the first stage. In the second part of this study, theoretical and experimental investigation on the performance of a 2-stage PSA process is presented. The experimental apparatus used in the second part is small in size compared to the pilot plant but all the components of the pilot plant are included in it. The power consumption of the 2-stage PSA process is analyzed with a mathematical model and a possible way of reducing the power consumption is proposed.

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2. EXPERIMENTAL A schematic diagram of the 2-stage PSA process is shown in figure 1. A scrubber, dryer, and 2-stage PSA process are three main parts of the process. The scrubber plays the role to remove dust and part of SO2 and to cool down the flue gas to ambient temperature. The dryer reduces the water content of the flue gas before it enters the main 2-stage PSA process. Dried flue gas is fed to the first stage PSA process, where carbon dioxide is concentrated to 40-60%. The cycle sequence of the first stage is the adsorption, pressure equalization, blowdown, low pressure purge, pressure equalization, and feed pressurization steps. A part of the effluent of the adsorption step is used to purge the adsorber in the low pressure purge step and the rest of the effluent is used to regenerate the dryer. The CO2 -rich stream obtained in the first stage PSA is further processed in the second stage PSA and CO2 is concentrated to over 99%. The CO2 flowing out of the adsorber during the adsorption step is recycled to the inlet of the scrubber. The cycle sequence of the second stage is almost the same as that of the first stage. The only difference is that the second stage is operated without the low pressure purge step. The cycle sequence of the process is given in Table 1. Feed gas for the experiment was the flue gas of a residential boiler. The concentration of CO2 in the flue gas was 10.5%.

Figure 1. Schematic diagram of the 2-stage PSA process. Table 1. Cycle sequence of 2 stage PSA process; a) a) _ Time(s) I 5 I 5 I 30 | 120 Bed A EQ FF 1 AD BedB | EQ I EV | Purge b) Time(s) I 10 I 10 I tAD Bed A EQ ~ FF | AD Bed B EQ EV

1st stage b) 2 nd stage _ ^ I 5 I 5 | 30 EQ ~ EV | EQ | FF | I

10 EQ EQ

I

10

I |

120 Purge AD

|

tAD EV

FF

|

AD

~

3. RESULTS and DISCUSSION Figure 2 is the photograph of the pilot plant for the recovery of CO2 from the flue gas of a power plant, which was installed in 1996. The process configuration of the pilot plant was slightly different from that shown in Fig 1. Major difference is the number of bed employed in the main PSA process. In the pilot plant, each stage of PSA process consisted of 3 beds. A

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later study revealed that a unit of 2 beds is sufficient for the recovery of CO2, which was applied for the modified process. The power consumption was about 1.26 kWh/Nm3-CO2 and the productivity of the process was 30.2 Nm3/ton/hr when the CO2 concentration of the feed was 12%. The high power consumption of the pilot plant was attributed to several factors; a) high purge flow rate, b) low recovery at the second stage. The higher the purge flow rate, the lower the CO2 concentration fed to the second stage. The second factor is closely related to the first factor. Since the CO2 concentration fed to the second stage was low, the recovery of the second stage was Figure 2. Photograph of CO2 recovery pilot about 45% for the production of 99%-CO2. 3 plant (Plant capacity: 110 Nm /hr)

Power consumption of the 1st stage PSA The effects of the low pressure purge step on the power consumption at the first stage was first analyzed by a mathematical model. It is the sum of the power consumptions of the blower and a vacuum pump. The power consumptions of the blower and vacuum pump were calculated from the following equation for adiabatic compression of gas. ( — *\

W^nRJ.JL-^y -I v

(1)

)

Figure 3 shows the effects of the low pressure purge step on the power consumption of the first stage with a fixed amount of feed. The specific power consumptions shown in Figure 3 are based on the total amount of CO2 desorbed in the evacuation and purge steps. It is seen in Figure 3 that the specific power consumption of the blower increases with the decrease of the purge flow rate. Considering that the amount of CO2 produced through the evacuation and purge steps, at the desorption steps, is F yco F R, the specific power consumption of the blower can be represented by the following equation.

wB=R/rf\~\{yCo2,F-R)

(2)

Where R is the recovery and P2 is the discharge pressure of the blower. Eq. (2) means that the power consumption of the blower is proportional to the inverse of the recovery. This explains the decrease of the power consumption with the increase of the purge flow rate. Contrary to the blower, the specific power consumption of the vacuum pump increases with the purge flow rate. Considering that the total amount of gas pumped during the desorption steps is F yCOi F Rl yco , the specific power consumption at the vacuum pump can be represented by the following equation.

Wy=RJrj{y\yco,

(3)

The discharge pressure P2 is constant, atmospheric pressure. Pl is a pressure between the

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adsorption pressure, PH, and the desorption pressure, PD, and it approaches the desorption pressure as the purge amount increases. jyco in Eq. (3) is the average CO2 concentration of desorption effluents, i.e., the evacuation and purge effluents. As the amounts of N2 and O2 in the desorption effluents increase, more and more energy is used to compress N2/O2. This is why the average concentration of CO2 is important. The average CO2 concentration of the desorption effluents is reduced with the purge flow rate above the purge flow rate about 1 L/min, which plays a role to increase the specific power consumption. Moreover, with the increase of the purge flow rate, the pressure, Pt , approaches the desorption pressure because more and more gas is obtained at lower pressure. This increases the compression ratio. Due to higher compression ratio and lower CO2 purity, the specific power consumption of the vacuum pump increases with the increase of the purge flow rate above the purge flow rate of 1 L/min. The situation is somewhat complicated below the purge flow rate of 1 L/min. Both of the CO2 purity and the compression ratio are increased with the Figure 3. Recovery, purity and power increase of the purge flow rate below the consumption of the 1st stage PSA process with the purge flow rate of 1 L/min. However, purge flow rate (Feed: 135 L/min, Desorption since the increase of the compression ratio pressure :50 mmHg, Wv: power consumption of is more significant than the increase of the vacuum pump, WB: power consumption at blower) purity, the specific power consumption of the vacuum pump is increased with the increase of the P/F ratio. Power consumption o/2-stage PSA Through the simulation study on the power consumption of the first stage it was revealed that the amount of purge and desorption pressure are important for the reduction of the overall power consumption. Here, the performance of the 2-stage PSA process is compared with the simulation results and the power consumption of the process is accessed by the simulation. Figure 4 shows the performance of the 2-stage PSA process with the adsorption step time of the second stage. The desorption pressures of the first and the second stage were 50 mmHg, 125 mmHg respectively. The low pressure purge flow rate was 7.8 L/min. Lines of figure 4 represent the simulation results. The power consumption shown in the figure is the sum of the power consumption of the blower and two vacuum pumps of the main PSA process. The recovery and purity of the process were fairly well predicted by the simulation. Small discrepancy between experiment and simulation was inevitable because simplifying assumption was introduced to the model. For example, volume averaged concentration of the 1st stage desorption steps and the 2nd stage adsorption effluent were used as concentration of the 2nd stage adsorption step feed and recycle stream to the 1st stage in the simulation, respectively. As the second stage adsorption step time increases, the recovery of the process decreased. This results from the larger loss of CO2 with the increase of the adsorption time. Contrary to the recovery, the purity of CO2 and the power consumption are increased with the 2nd stage

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adsorption step time. With the increase of the adsorption step time, the amount of carbon dioxide recycled to the first stage is increased, which increases the load of the first stage vacuum pump. The cumulative volume and concentration of CO2 flowing into or out of the column at each step during half cycle of the first stage were summarized in Table 2. The values shown in table 2 are the simulation results for the adsorption step time of 100s and 120s. It is seen that at longer adsorption step time of the 2nd stage more amount of CO2 is vented from the 2nd stage adsorption step, which is recycled to the first stage. In addition to this, the amount of gas desorbed from the 1st stage for the adsorption time of 120s is larger than that for the adsorption time of 100s. Because of this, vacuum pump of the first stage consumed more power when the adsorption step time of the 2nd stage is 120s. Except for the power consumption of the 1st stage vacuum pump, the power consumptions of blower and 2nd stage vacuum pump are the same regardless of the adsorption step time. According to the simulation, 0.28 kWh was consumed for the production of INm3 of Figure 4. Performance of 2-stage PSA process with the second stage adsorption step time 99%-CO2. (symbols: experiment, lines: simulations) Table 2. Cumulative volume and average concentration of CO2 flowing into or out of column at each step for two different 2nd stage adsorption step times. (Feed: 135 L/min, Purge: 7.8 L/min) 2 nd AD time Is1 feed 1st AD vent l s t EV+PU 2 nd Feed 2 nd AD Vent 2 nd EV Reco1st stage very 2 nd stage Power 1st W v Consu2 nd W v mption WB

120s Gco2 YcO2 (L) (%) 49.6 12.7 7.6 2.16 66.48 42.9 66.48 42.9 39.56 14.1 28.57 99.72 86.52 66.58 0.21 0.05 0.04

100s GcO2

(L) 41.2

Y

CO2 (%)

11.0 2.16 5.5 36.7 59.5 36.6 59.5 21.4 6.75 99.37 30.1 88.95 81.94 0.19 0.05 0.04

Table 3. Cumulative volume and average concentration of CO2 flowing into or out of column at each step for different 2nd stage desorption pressure. (Feed: 135 L/min, Purge: 4.7 L/min) 2 nd desorption pressure Is1 feed 1st AD vent 1 S 'EV+PU 2 nd feed a) 2 nd AD Vent 2 nd EV Reco1st stage very 2 nd stage Power 1st W v Consu- 2 nd W v mption WB

150mmHg OcO2

Y

CO2 (%)

(L) 10.4 38.7 2.01 6.9 63.4 32.8 63.4 32.6 17.1 3.8 28.89 99.16 84.71 88.08 0.17 0.05 0.04

100 mmHg Y °CO2 CO2 (%) (L) 37.7 10.1 6.7 1.97 63.2 31.9 63.2 31.9 2.3 11.1 29.5 98.94 84.6 92.44 0.17 0.06 0.04

Another variable of interest and importance is the desorption pressure of the second stage. The benefit of lower desorption pressure is that the recovery and productivity of the second stage could be increased due to the increase of the working capacity. However, lower desorption pressure might increase the power consumption due to high compression ratio. So, the effects of the desorption pressure of the second stage is not straightforward. The effects of the the desorption pressure of the second stage are investigated by simulation

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and the results are shown in Figure 5. As shown in Figure 5, when the desorption pressure was 150 mmHg, the recovery and power consumption for the production of 99% CO2 were 79% and 0.26 kWh/Nm3-CO2, respectively. However, when the desorption pressure was 100 mmHg, the recovery and the power consumption were 81% and 0.265 kWh/Nm3-CO2. That is, lower desorption pressure was favorable with respect to the process recovery but higher desorption pressure was favorable with respect to the power consumption. The cumulative volume and concentration of CO2 flowing into or out of the column at each step during half cycle of the first stage with the 2nd stage desorption pressure were summarized in 96 Table 3. While the recovery at the second stage for the desorption pressure of 100 mmHg was 92%, that for the desorption pressure of 150 mmHg was 88%. However, the power consumption of vacuum pump of the second stage was higher at lower desorption pressure due to higher compression ratio. But, because the difference in the power consumption Figure 5. Performance of 2-stage PSA process with the second stage desorption pressure with the second stage desorption pressure is small, it seems that lower desorption ( : 150mmHg, o : lOOmmHg) pressure giving high recovery is more favorable than higher desorption pressure. The power consumption of the modified process was solely estimated by the simulation Table 4. Ratios of desorption amount at each u s i n 8 i d e a l thermodynamic relation, i.e. Eq. (1). stage to product for the pilot plant and the Comparing the amounts of gas evacuated by two vacuum pumps in the modified and pilot modified process process will give some insight for the real Desorption amount/product power consumption of the modified process. 1sl stage 2nd stage The ratio of the desorption amount to product is given in Table 4. It is clearly seen that in the 1.7 Pilot plant 5.35 modified process the total amount of gas 2.03 1.0 Modified process evacuated by two vacuum pumps is 2.3 times less than that in the pilot plant process. This means that the modified process will consume less power. REFERENCES 1. 2. 3.

Sircar, S. (1988) Sep. Sci. Technol, 23, 519 Chue, K.T., Kim, J.N., Cho, S.H. and Yang, R.T. (1995) Ind. Eng. Chem. Res., 34, 591 Ishibashi, M., Ota, H., Akutsu, N. And Umeda, S., Tajika, M., Izumi, J., Yasutake, A., Kabata, T. and Kageyama, Y. (1996) Energy Corners. Mgmt, 1996, 37, 929-933


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Novel Nanoporous "Molecular Basket" Adsorbent for CO2 Capture Chunshan Song*, Xiaochun Xu, John M. Andresen, Bruce G. Miller and Alan W. Scaroni Clean Fuels and Catalysis Program, The Energy Institute, and Department of Energy & Geo-Environmental Engineering, Pennsylvania State University, 209 Academic Projects Building, University Park, PA 16802, USA A CO2 adsorbent with high adsorption capacity and high selectivity was prepared based on a novel "molecular basket" concept consisting of mesoporous molecular sieve of MCM-41 type and an immobilized branched polymer with CCb-capturing sites. The novel CO2 "molecular basket" adsorbent was characterized by X-ray powder diffraction (XRD). N2 adsorption/desorption and CO2 adsorption/desorption measurement. The "'molecular basket" adsorbent can selectively capture CO2 in a gas mixture and was also applied in the separation of CO2 from simulated flue gas and boiler flue gas. 1. INTRODUCTION The continued use of fossil fuels to provide clean and affordable energy supply depends on the technology development to reduce their negative environmental impact. On the one hand, the emissions of particulate matter and the oxides of nitrogen (NOx) and sulfur (SOx) need to be further reduced [1, 2]. On the other hand, considerable increase in global atmospheric CO2 concentration has caused serious concern for climate change and led to worldwide effort in research and development on the control of CO2 emissions [3, 4]. Capture and separation of CO2 from stationary sources is considered an important option for the control of CO2 emission. The capture/separation cost constitutes about threefourths of the total cost of the control of CO2 emissions, e.g., carbon sequestration [5]. It is therefore important to explore new cost-effective approaches for CO2 separation. A new concept called CO2 "molecular basket", has been proposed in our laboratory for developing a high-capacity, high-selective CO2 adsorbent to decrease the CO2 separation cost. The novel "CO2 molecular basket" is a solid adsorbent and can selectively capture CO2 from gas mixtures and "pack" CO2 in condensed form in nanoporous channels. In order to capture a large amount of CO2, the adsorbent needs to have large-pore channels filled with a CO2-capturing substance as the "basket". To make the "basket" CO2-selective, a substance with numerous CO2-affinity sites should be loaded into the pores of the support to increase the affinity between the adsorbent and CO2 and, therefore, to increase the CO 2 adsorption selectivity and CO2 adsorption capacity. In this paper, the preparation of the novel CO2 "molecular basket" adsorbent by using the large pore volume material of mesoporous molecular sieve MCM-41 and CO2-affinity substance of polyethylenimine (PEI) was studied. The separation of CO2 from simulated flue gas and boiler flue gas by using this novel "molecular basket" adsorbent are reported.

* Corresponding author, E-mail: csonaaDS.u.edu; Tel: 814-863-4466; Fax: 814-865-3248

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2. EXPERIMENTAL The CO2 "Molecular Basket" adsorbent was prepared by loading 50 wt% PE1 into the mesoporous molecular sieves MCM-41 (MCM-41-PEI-50) and was used as adsorbent in the adsorption separation of CO2 from simulated flue gas mixture and boiler flue gas. The "molecular basket" adsorbents were characterized by X-ray powder diffraction (XRD), N2 adsorption/ desorption as well as the CO2 adsorption/desorption performance. The details on the preparation and characterization of this adsorbent were published elsewhere [6, 7]. The adsorption separation of CO2 from simulated flue gas mixture and boiler flue gas was carried out in a flow adsorption separation system [8]. The simulated flue gas mixture contains 14.9% CO 2 , 4.25% O2 and 80.85% N 2 . In some experiments, the simulated flue gas mixture was mixed with 10% moisture. Typically, 2.0 g powder adsorbent was used for the adsorption separation. The adsorption separation was carried out at 75 °C and ambient pressure. The feed gas flow rate was 10 ml/min. After the adsorption, helium with a flow rate of 50 ml/min was used to purge the adsorbent bed to perform the desorption at 75 °C. The flow rate of the effluent gas was measured by a soap-film flowmeter. The concentration of the effluent gas was measured on-line using a SRI 8610 C Gas Chromatography (GC). The analysis was carried out every 5 minutes when the dry simulated flue gas mixture was used as adsorbate and every 15 minutes when the moist simulated flue gas mixture was used as adsorbate. Adsorption capacity in ml (STP) of adsorbate/g adsorbent and desorption capacity in percentage were used to evaluate the quality of the adsorbent. The adsorption and desorption capacities were calculated from the mass balance during the adsorption separation. The separation factor, a , j, was calculated from equation 1 as the ratio of the amount of gases adsorbed by the adsorbent. (nj/nj)adsorbed, over the ratio of the amount of gases fed into the adsorbent bed, (ni/nj)feed: \ni

'nj)adsorbed

,-.\

The coal-fired flue gas contains 12.5-12.8% CO2, -4.4% O 2 , 50 ppm CO, 420 ppm NOX, 420 ppm SO 2 , 6.2% H2O and 76-77% N 2 . Generally, 30 g adsorbent with particle size between 18 and 35 mesh was used. The beginning and the end of the adsorption column were filled with alumina (-170 g) to decrease the dead volume in the separation system. The adsorption separation was carried out at a feed flow rate of 5-6 1/min and temperature of 0 °C. After adsorption, gas was switched to helium at a flow rate of 5 1/min to perform the desorption at 0 °C. The flow rate of the effluent gas was measured by a rotameter. The concentrations of O2, CO, CO2, SO2 and NOX in the effluent gas were measured on-line using model NGA 2000 paramagnetic oxygen analyzer; model NGA 2000 nondispersive infrared CO analyzer; model NGA 2000 non-dispersive infrared CO 2 analyzer; model 890 ultraviolet SO2 analyzer; and model NGA 2000 chemiluminescence NOX analyzer, respectively. The analysis was carried out every 5-6 seconds. Since the alumina also adsorbed the gases, a blank separation test with the adsorption column only filled with the alumina (-210 g) was also carried out. Therefore, the adsorption/desorption capacity for the "molecular basket" adsorbent can be calculated by subtracting the adsorption/desorption capacity between the adsorption experiment and blank experiment. The adsorption/desorption capacity was calculated from mass balance during the adsorption separation. The separation factor was calculated from equation 1.

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3. RESULTS AND DISCUSSIONS 3.1 Preparation and characterization of CO2 "Molecular Basket" adsorbent The structure of MCM-41 before and after loading 50 wt.% PEI was characterized by XRD and the results are compared in Figure 1. The diffraction patterns of MCM-41 did not change after PEI was loaded, which indicated that the structure of MCM-41 was preserved. However, the intensity of the diffraction patterns of MCM-41 changed. The intensity of the diffraction patterns of MCM-41 decreased after PEI was loaded, which was caused by the pore filling effect [6, 7]. The pore structure analysis by nitrogen adsorption/desorption confirms that PEI was loaded into the pore channels of the MCM-41 support. Completely degassed MCM-41 shows a type IV isotherm. The surface area, pore volume and pore diameter were 1480 m2/g, 1.0 ml/g and 2.75 nm respectively. After loading the PEI, the mesoporous pores were completely filled with PEI, resulting in a type II isotherm and restricting the access of nitrogen into the pores at the liquid nitrogen temperature. The residual pore volume of the MCM-41-PEI-50 was only 0.011 ml/g, the surface area was estimated to be 4.2 m2/g and the average pore diameter was smaller than 0.4 nm. The CO2 adsorption and desorption performance of MCM-41 and MCM-41-PEI-50 was measured by TGA at 75 °C under pure CO2 atmosphere. The results are shown in Figure 2. Before PEI was loaded, the MCM-41 support alone showed a CO2 adsorption capacity of 8.6 mg/g adsorbent. The low adsorption capacity was caused by the weak physical interaction between CO2 and MCM-41 at relatively high temperature. In order to strengthen the interaction between CO2 and MCM-41, the branched polymeric substance PEI with numerous CC^-capturing sites was loaded into the channels of the MCM-41. After loading the PEI, the adsorption capacity increased substantially. The MCM-41-PEI50 showed a CO2 adsorption capacity of 112 mg/g adsorbent, which was much higher than that of the MCM-41 support and higher than that of the pure PEI (109 mg/g-PEl). The desorption was complete for both the MCM-41 support and the MCM-41-PEI-50. However, the desorption for pure PEI was slow and was not complete compared to the desorption time of the MCM-41-PEI adsorbents, which indicated that the "molecular basket" adsorbent facilitates the CO 2 desorption. The fast desorption of CO 2 from the "molecular basket" adsorbent can be explained by the high dispersion of PEI into the MCM-41 channels.

Figure 1 Comparison of the XRD patterns of MCM-41 and MCM-41-PEI-50.

Figure 2 Comparison of CO2 adsorption and desorption performance of MCM-41 and MCM-41-PEI-50 (on 1 g adsorbent). In order to evaluate the effect of the "molecular basket", the adsorption capacity

414

weighed on PEI in the "molecular basket" adsorbent was calculated with equation (2): PEI adsorption capacity (mg adsorbate/g-PEI) = [Adsorption capacity of the adsorbent -(MCM-41 weight percentage in the adsorbent x Adsorption capacity of pure MCM-41)]/ (PEI weight percentage in the adsorbent) (2) The adsorption capacity weighted on PEI for MCM-4l-PEI-50 was 215 mg/g-PEl. which is two times that of the pure PEI. This clearly shows that MCM-41 has a synergetic effect on the CO2 adsorption when the PEI was loaded into its porous structure. There are two possible reasons for the synergetic effect of MCM-41, i.e., its high surface area and its uniform mesoporous channel. When the PEI was loaded on the materials with high surface area, more CO2 affinity sites were exposed to the adsorbate and thus the adsorption capacity increased. However, when the PEI was coated on a high-surface-area silica gel (550 m2/g) with the PEI loading of 50 wt.%, the CO 2 adsorption capacity was only 156 mg/g-PEI [7], which was slightly higher than that of the bulk PEI and much lower than that of the "molecular basket" adsorbent with the same PEI loading. Therefore, only when the PEI was loaded into the mesoporous channels of MCM-41 did the "molecular adsorbent" show a highest synergetic effect on the adsorption of CO2. 3.2 Separation of CO2 from simulated flue gas Figure 3 shows the concentrations of CO2, O2 and N2 during the adsorption separation of CO2 from a simulated flue gas mixture at 75°C by the "molecular basket" adsorbent. At the beginning of the separation, CO 2 was completely adsorbed by the adsorbent and the concentration of CO2 was below the detection limit of the gas chromatograph, i.e. < 100 ppm. After 50 minutes of adsorption, CO2 began to break through and was detected in the effluent gas. After 120 minutes of adsorption, the adsorbent can only adsorb 5% of the CO2 from the adsorbate. The CO2 adsorption capacity was 37.5 ml (STP)/g adsorbent before breakthrough and was 45.4 ml (STP)/g adsorbent after 120 minutes of adsorption. The adsorption capacity is comparable with that previously measured from TGA at about 15% CO2 concentration [6]. The adsorption of O2 is much lower than that of the CO2 and the adsorption capacity of O2 is 0.07 ml (STP)/g adsorbent after 120 minutes adsorption. The CO2/O2 separation factor was calculated to be 185. The adsorbent hardly adsorbs any N 2 , with CO2/N2 separation ratio of > 1000. A similar experiment was carried out by using the MCM-41 as the adsorbent, which showed a CO2 adsorption capacity of 3.2 ml (STP)/g adsorbent; the separation factor for CO2/O2 and CO2/N2 was 3.3 and 2.9 respectively, which are much lower than those of the "molecular basket" adsorbent. These indicated that the high adsorption capacity and high selective CO2 adsorbent was prepared by using the "molecular basket" concept and by loading the PEI into the MCM-41 channels. 3.3 Separation of CO2 from moist simulated flue gas The influence of moisture on the adsorption separation of CO2 from simulated flue gas containing about 10% moisture by the novel "molecular basket" adsorbent (MCM-4l-PEI50) was investigated at 75 °C and ambient pressure. Figure 4 compares the CO2 breakthrough curve, where the amount of carbon dioxide was followed as the fraction of the CO2 concentration in the effluent gas from the adsorption column, C, over that of the CO2 concentration in the feed, Co, for the flue gas without moisture and with ~ 10% moisture. In the presence of moisture, the "molecular basket" adsorbent can still

415

Time (min)

Figure 3 Changes of gas concentration during the separation of CO2 from simulated flue gas. Operation condition: Weight of adsorbent: 2.0 g; Temperature: 75 °C; Feed flow rate: 10 ml/min. Temperature: 75 °C; Feed composition: 14.9% CO2, 4.25% O2 and 80.85% N2.

Figure 4 C O2 breakthrough curve with/without moisture in the simulated flue gas. Operation condition: Weight of adsorbent: 2.0 g; Temperature: 75 °C; Feed flow rate: 10 ml/min. Temperature: 75 °C. Dry feed composition: 14.9% CO2, 4.25% O2 and 80.85% N 2 ; Moist feed composition: 13.55% CO 2 , 3.86% O2, 72.72% N 2 and 9.87% H2O.

effectively adsorb CO2. The CO2 breakthrough time was 60 minutes, which is longer than that under dry flue gas conditions and indicated that the moisture had a promotion effect on the adsorption of CO2 by the "molecular basket" adsorbent. The CO 2 adsorption capacity increased from 45.4 ml (STP)/g adsorbent for dry flue gas to 65.0 ml (STP)/g adsorbent for moist flue gas. Meanwhile, the adsorption of O2 was inhibited at the moist condition. The separation selectivity for CO2/O2 was -180 at the dry flue gas condition and ~ 600 at the moist flue gas condition. The adsorption of N 2 was below the detection limit of gas chromatography. The "molecular basket" adsorbent also adsorbed significant amount of water. The water adsorption capacity was 59.0 ml (STP)/g adsorbent in the moist flue gas condition. However, the CO2 adsorption capacity increased 40% in the presence of water and the separation of water and CO2 is rather easy in reality. The desorption of CO2 and moisture were complete. 3.4 Separation of CO2 from coal-fired boiler flue gas The adsorption separation of CO2 from a coal-fired flue gas was investigated and the CO2 breakthrough curves are shown in Figure 5. (Note that the analysis of the effluent gas composition was carried out after removing the moisture in the gas mixture, the analyzed concentrations of O2, CO, CO2, SO2 and NOX were slightly higher than those in the real flue gas mixture.) Clearly, both the alumina and the "molecular basket" adsorbent can adsorb CO 2 . However, the adsorption performance of the "molecular basket" adsorbent was much better than that of the alumina. The lowest CO2 emission concentration was ~ 2.5 % for alumina and was below 0.1% for the "molecular basket" adsorbent. The CO2 adsorption capacity was 1.4 ml (STP)/g adsorbent for the alumina and 36 ml (STP)/g adsorbent for the "molecular basket" adsorbent. In addition, the "molecular basket" adsorbent showed a better selectivity. The "molecular basket" adsorbent did not adsorb O2. N2 and CO, while the CO2/O2 selectivity was 3.5 for alumina. The SO2 and NOX adsorption capacity for the "molecular basket"

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adsorbent were 0.11 ml (STP)/g adsorbent and 0.21 ml (STP)/g adsorbent, respectively. The separation selectivity for CO2/SO2 and CO2/NOX was 1.07 and 0.57 respectively. However, very little NOX and SO2 were adsorbed before CO2 breakthrough. The adsorption capacity for CO2, SO2 and NOX before CO2 breakthrough were 24 ml (STP)/g adsorbent, 0.0074 ml (STP)/g adsorbent and 0.028 ml (STP)/g adsorbent respectively. Therefore, the separation selectivity for CO2/SO2 and CO2/NOX were 10.7 and 2.86, respectively, before CO2 breakthrough. While the desorption of CO2 was complete, very little NOX and SO2 desorbed.

Figure 5 C0 2 breakthrough curve for coal-fired flue gas. Operation condition: Weight of adsorbent: 30 g; Feed composition: 12.5-12.8% CO2, 6.2% H2O, ~ 4.4% O2, 50 ppm CO, 420 ppm NOX, 420 ppm SOX and 76-77% N2; Feed flow rate: 6000ml/min; Temperature: .

4. CONCLUSIONS Novel CO2 "molecular basket" adsorbent with high adsorption capacity and high selectivity has been developed. The "basket" of MCM-41 channels displayed a synergetic effect on the adsorption of CO? by PEL The CO2 "molecular basket" adsorbent was successfully applied to the separation of CO2 from simulated flue gas and boiler flue gas. ACKNOWLEDGEMENTS

Financial support from U.S. Department of Defense (via an interagency agreement with U.S. Department of Energy) and the Commonwealth of Pennsylvania are highly appreciated (Cooperative Agreement No. DE-FC22-92PC92162). REFERENCE 1. J.J. Mooney, Annual Meeting of National Petrochemicals & Refiners Association, San Antonio. TX, March 26-28, 2000. 2. U.T. Turaga and C. Song, American Chemical Society Division of Petroleum Chemistry Preprints. 46(2001)275-279. 3. M. M. Maroto-Valer, C. Song and Y. Soong (Eds). Environmental Challenges and Greenhouse Gas Control for Fossil Fuel Utilization in the 21 st Century. Kluwer Academic/Plenum Publishers. New York, 2002, 447 pp. 4. C. Song, A. M. Gaffhey, K. Fujimoto (Eds). CO2 Conversion and Utilization. American Chemical Society (ACS), Washington DC, ACS Symp. Series, Vol. 809, 2002, 448 pp. 5. U.S. Department of Energy, Carbon Sequestration-Research and Development, 1999. http://www.fe.doe.gov/coal_power/sequestration/reports/ rd/index.html. 6. X.C. Xu, C. Song, J.M. Andresen, B.G. Miller and A.W. Scaroni, Energy & Fuels, 16 (2002) 1463-1469. 7. X.C. Xu, C. Song, J.M. Andresen, B.G. Miller and A.W. Scaroni, Microporous and Mesoporous Materials, 62 (2003) 29-45. 8. X.C. Xu, C. Song, J.M. Andresen, B.G. Miller and A.W. Scaroni, International Journal of Environmental Technology and Management, (2003) submitted for publication.


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Reversible Adsorption of Carbon Dioxide on Amine-Modified SBA-15 from Flue Gas Containing Water Vapor Norihito Hiyoshi, Katsunori Yogo* and Tatsuaki Yashima Research Institute of Innovative Technology for the Earth, 9-2 Kizugawadai, Kizu-cho Sorakugun, Kyoto 619-0292, Japan Aminosilane modified mesoporous silica as novel adsorbent for CO2 capture and separation from flue gases was prepared by grafting various aminosilanes; i.e., (3-aminopropyl) triethoxysilane (APS), N-(2-aminoethyl)-3-aminopropyltriethoxysilane (AEAPS) and (3trimethoxysilylpropyl)diethylenetriamine (TA), on mesoporous silica, SBA-15. Pore walls of SBA-15 were modified uniformly with these aminosilanes, and relatively high surface area and uniform pore size were retained. Adsorption capacities of CO2 in the presence of water vapor were compared with that in the absence of water by a flow method. It was found that adsorption capacities of aminosilane modified SBA-15 under wet condition were comparable to that under dry condition. In particular, the adsorption capacity of TA-anchored SBA-15 reached 1.28 mol-kg" in the presence of water vapor at 333 K. In addition, these adsorbents were completely regenerated by heating up to 423 K in He flow. 1. INTRODUCTION Increasing atmospheric concentration of CO2 due to fossil fuel combustion is a serious environmental problem. Recently, CO2 capture and sequestration has attracted considerable attention as one of the options to reduce CO2 emission. Various processes, such as liquid solvent absorption, membrane separation, and pressure (and/or temperature) swing adsorption (P(T)SA), have been proposed for separation and recovery of CO2 emitted by power plants, steel works, etc. [1]. However, the costs of CO2 separation from flue gases are accounted for approximately 70-80 % of total cost for CO2 sequestration. Therefore, it is important to develop a new efficient and energy-saving technique for CO2 separation. In addition, it is desired to downsize a plant for CO2 separation, because enormous amounts of gases must be treated. As for the conventional PSA or PTSA process using zeolite, a dehumidification process which consumes about 30 % of total energy is necessary, because water vapor is adsorbed more strongly than CO2 on zeolite surface. Therefore, development of a new adsorbent which is able to adsorb CO2 in the presence of water vapor is required to construct a simple and energy-saving process by elimination of the dehumidification process. Hayashi et al. showed that hydrated potassium carbonate supported on active carbon was able to absorb CO2 from flue gas containing water vapour [2]. Solid sorbents in which amines are supported on high surface area supports are also promising as sorbents for CO2 separation [3-6]. Leal et al. applied aminosilane modified silica to CO2 adsorbent [4]. Aminosilane modified silica has been extensively studied because of its widespread applications. However, it is difficult to modify micro pore wall of silica with aminosilanes

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molecules due to steric hindrance [7]. Therefore, mesoporous silicas such as M41S, FSM-16, and SBA-15 are suitable supports for surface modification with aminosilane on mesoporous silica, because they have large and uniform pores. Furthermore, higher loading of aminosilane than conventional silica gel is expected due to their high surface area. Recently, Xu et al. showed that polyethylenimine modified MCM-41 is an efficient CO2 adsorbent [5,6]. However, effect of water vapour on adsorption performance of polyethylenimine modified MCM-41 has not been studied. Among various mesoporous silicas, SBA-15 is a suitable material for applications in gases containing water vapor, due to its higher hydrothermal stability [8]. hi this study, we have prepared aminosilane modified mesoporous molecular sieve SBA-15 as a "water-tolerant adsorbents", which is possible to adsorb CO2 in the presence of water vapor, and applicability for PTSA was examined by CO2 adsorption-desorption measurement in a flow system. 2. EXPERIMENTAL 2.1. Preparation of adsorbents SBA-15 was prepared by a similar method reported by Zhao et al. [8]. 2-D hexagonal structure characteristic of SBA-15 was confirmed by X-ray diffraction pattern. The lattice constant of SBA-15 was estimated to be 11.3 nmbyXRD. Modification of SBA-15 was performed by grafting method using aminosilanes. Aminosilane (50 cm3) and SBA-15 (5.0 g) which was previously dried at 398 K for 6 h in air were refluxed in dehydrated toluene (250 cm3) (Wako Pure Chem. Ind. Ltd.) at 383 K for 24 h under Ar flow. The product was washed with toluene (200 cm3) and dried at 333 K over night. (3-aminopropyl)triethoxysilane (Aldrich), N-(2-aminoethyl)-3-aminopropyltriethoxysilane (Chisso Chem. Co.) and (3-trimethoxysilylpropyl)-diethylenetriamine (Gelest Inc.) were used as a grafting agent. The obtained samples were denoted as APS/SBA, AEAPS/SBA and TA/SBA, respectively. APS/SBA(i) was prepared by the impregnation with (3-aminopropyl)triethoxysilane into SBA-15. hi addition, a gelatinous solid (denoted as APG) was obtained by hydrolysis and condensation of (3-aminopropyl)triethoxysilane. 2.2. Characterization of adsorbents Adsorption-desorption isotherms of N2 were measured at 77 K by an N2 adsorption system (Auto Sorb 1, Quantachrome Corporation) after each sample was evacuated at 473 K for 3 h. The surface areas and the pore size distributions were calculated by the BET and BJH method, respectively. The amount of supported aminosilane was measured from TG curve. TG-DTA curves of aminosilane modified SBA-15 and APG were obtained with a thermal analysis system (TAS2000, Rigaku). The sample (ca. 10 mg) was heated from 303 K to 1173 K in air at a rate of 5 K-min"1. 2.3. CO2 Adsorption-Desorption Experiment Capacities for CO2 adsorption were determined by a flow method as follows. A 1.5 g of adsorbent placed in a Paylex tube with 13 mm in inner diameter was dried in He flow (30 cm3min"') at 423 K for 1 h. 12% H2O with He balance (total flow rate: 60 cm3-min"') was fed into the adsorbent at 333 K until the adsorbent was saturated with water, and then the gas introduced was switched to the mixture of 15% CO2, 12% H2O with N 2 balance (total flow rate: 30 cm3-min"'). A breakthrough curve of CO2 was obtained by the analysis of effluent gases. The effluent gases were analyzed by a gas chromatograph (GC-332, GL science Inc.) equipped with gaskuropack 55 column (2 m) and TCD detector. Measurement under dry

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condition was conducted using 15% CO2 with N2 balance (total flow rate: 30 cm3min"') without pre-adsorption of water vapor. The desorption experiment was performed by heating the saturated adsorbent from 333 K to 423K in He flow (total flow rate: 30 cm3min"') at a rate of 10 K-min"1. Desorption curves were obtained by the analysis of effluent gases. 3. RESULTS AND DISCUSSION 3.1. Characterization of the adsorbents The surface areas and pore volumes measured by N2 adsorption at 77 K are summarized in Table 1. Although the modification of SBA-15 with the aminosilanes resulted in decreasing the surface area, samples prepared by grafting method (APS/SBA, AEAPS/SBA and TA/SBA) maintained relatively high surface areas. On the other hand, the surface areas of APS/SBA(i) and APG were very low (2 m2-g"'). The pore volume measured by N 2 adsorption (VN2) are compared with the expected value (Vcai) calculated by eq. 1. Vcal ( c m V ) =

[VSBA x

(100 - w) /100] - [w / (100 x d)]

(1)

and w are pore volume of SBA-15 (cm3-g"') and aminosilane loading (wt%) determined by TG analysis, respectively, and d is the density (gem 3 ) of corresponding aminosilane gel obtained by hydrolysis and condensation of aminosilane. The pore volume measured by N 2 adsorption was very close to expected value for each adsorbent. This result indicates that SBA-15 was modified without pore plugging. Furthermore, although the pore diameter of APS/SBA, AEAPS/SBA and TA/SBA were lower than that of SBA-15 (6.0 nm), uniformity of the pore size distribution of SBA-15 was not changed after the modification with the aminosilanes. This suggests that the pore wall of SBA-15 was uniformly modified by grafting method. Surface coverage defined as the number of aminosilane molecules per 1 nm2 of SBA-15 surface is also summarized in Table 1. Surface coverage decreased in the order: APS/SBA > AEAPS/SBA > TA/SBA. This order is reasonable if molecular sizes of aminosilanes are considered. As shown in Table 1, amine content defined as the amount of nitrogen atom per 1 gram of adsorbent is increased in the order: APS/SBA < AEAPS/SBA < TA/SBA. It should be noted that these values are higher than that of conventional aminosilane modified silica [4,7]. VSBA

Table 1. Physical and chemical properties of adsorbent Amine contents5' Adsorbent SA" Coverage VN22) Vcalc.J) 3 3 2 /cm -g"' /molecules-nm" /N-mmol-g' /cm -g"' /m -g"' 1.06 0 0 SBA-15 910 0.53 2.7 APS/SBA 360 0.51 2.6 0.47 2.0 4.2 0.45 AEAPS/SBA 310 0.37 1.8 TA/SBA 0.36 5.1 240 2 0 6.0 12 APS/SBA(i) 2 9.0 0 APG 1) Surface area calculated by BET method. 2) Pore volume determined by N2 adsorption. 3) Pore volume expected from eq. 1. 4) The number of organosilane molecules per 1 nm2 of SBA-15 surface. 5) Defined as the number of nitrogen atom per 1 gram of the adsorbent.

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3.2. Adsorption-desorption properties The breakthrough curves of CO2 under wet condition are compared with that under dry condition (fig. 1), and the CO2 adsorption capacities obtained from the breakthrough curves are summarized in Table 2. As for the SBA-15 support alone, the relative concentration of CO2 immediately reached unity due to low adsorption capacity. On the other hand, considerable amounts of CO2 were adsorbed on APS/SBA, AEAPS/SBA and TA/SBA. It should be noted that these amine-modified SBA-15 gave almost the same adsorption capacity under the wet condition compared with dry condition. In particular, the adsorption capacity of TA/SBA reached 1.28 mol-kg" in the presence of water vapor. These results demonstrate that aminosilane modified SBA-15 is effective for CO2 adsorption not only under dry condition but also under wet condition.

Figure 1. Breakthrough curves of CO2 at 333 K. The mixture of 15% CO2, 12% H2O with N 2 balance or 15% CO2 with N 2 balance (total flow rate: 30 cn^-min"1) was fed into the adsorbent (1.5 g). Table 2. CO2 adsorption capacities of various adsorbents at 333 K 1 Adsorbent Adsorbed CCVmmol-j>"' (Amine Efficiency ') CO2( 15 %)-H2O( 12%)-N2 CO2(15%)-N2 SBA-15 0.05 0.04 (0.19) APS/SBA 0.52 0.56 (0.21) (0.21) AEAPS/SBA (0.23) 0.87 0.98 (0.22) TA/SBA 1.28 (0.25) 1.10 2.32 (0.39) 0.14 (0.01) APS/SBA(i) (0.30) 0.03 (0.00) 2.71 APG 1.23 Na-Y (Si/Al=2.7) 0.03 1) Amine efficiency = (adsorbed CO2 / amine content)Amine efficiencies of APS/SBA, AEAPS/SBA and TA/SBA, which are defined as eq. 2, are summarized in Table 2.

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Amine efficiency [-] = (adsorbed CO2 [mmol-g"1] / amine content [mmol-g"1])

(2)

The amine efficiencies of these adsorbents were similar; the adsorption capacity is proportional to the amine content of the adsorbents. Therefore, it was suggested that primary amine and secondary amine are involved in CO2 adsorption with similar efficiency. On the other hand, the adsorption properties of zeolites, APS/SBA(i) and APG were very different from aminosilane modified SBA-15. Adsorption of CO2 did not occur on zeolite Na-Y (Si/Al=2.7) on which water was pre-adsorbed by the same procedure. In the case of APS/SBA(i) and APG, sorption of CO2 was observed only under the wet condition. In spite of the low surface area, CO2 sorption capacity of APS/SBA(i) and APG were much higher than that of aminosilane modified SBA-15 under the wet condition. The internal amino group of APS/SBA(i) and APG would react with CO2 through swelling with water or hydrolysis of siloxane bonds. However, it is evident from the breakthrough curves that the rate of CO2 sorption on APS/SBA(i) and APG is much lower than these on APS/SBA, AEAPS/SBA and TA/SBA. This result indicate that amine-CO2 reaction on APS/SBA, AEAPS/SBA and TA/SBA proceeds immediately, since these amine compounds were anchored and distributed on the surface of mesoporous support. Measurements of the CO2 and H2O desorption curves were performed by heating the adsorbents up to 423 K in He flow after adsorption at 333 K. The typical desorption curves for TA/SBA and APG are shown in figure 2. The amounts of CO2 desorbed from the adsorbents were consistent with the amounts of adsorbed CO2. Therefore, it was confirmed that these adsorbents were completely regenerated by heating up to 423 K. Aminosilane modified SBA-15 prepared by grafting method was more rapidly regenerated than APG. For example, in the case of TA/SBA (Fig. 2A), 77% of adsorbed CO2 was desorbed before temperature reached 423 K, and then desorption was completed 7 min after temperature

Figure 2. Desorption curves of CO2 and H2O in He flow (30 cm3 min"1) from TA/SBA (A) and APG (B) after adsorption (333 K, CO2(15%)-H2O (12%)N2(balance)). reached 423 K. On the other hand, only 9% of adsorbed CO2 was desorbed from APG before temperature reached 423 K (Fig. 2B). At 423 K, desorption of CO2 from APG was

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accelerated as APG was dried, and then the rate of CO2 desorption was decreased gradually. It took about 2 h to complete desorption of CO2 from APG. These results indicate that amine compounds exposed to the surface are favourable for desorption of CO2. 4. CONCLUSION In this study, novel adsorbents of CO2 were prepared by grafting of aminosilane on SBA15 surface. It was found that CO2 adsorption and desorption proceeded immediately on these aminosilane modified SBA-15, since amine compounds were anchored and distributed on the pore wall of mesoporous support uniformly. Furthermore, aminosilane modified SBA-15 exhibited high adsorption capacities in the presence of water vapor. The results of CO2 adsorption/desorption experiments demonstrate that these adsorbents are effective for CO2 separation in the presence of water vapor and applicable to CO2 capture and separation from flue gasses containing water vapor. ACKNOWLEDGMENT This work was supported by the New Energy and Industrial Technology Development Organization (NEDO), Japan. REFERENCES 1.

H. Ohta, S. Umeda, M. Tajika, M. Nishimura, M. Yamada, A. Yasutake and J. Izumi, hit. J. of Global Energy Issues, 11 (1998), 203. 2. H. Hayashi, J. Taniuchi, N. Furuyashiki, S. Sugiyama, S. Hirano, N. Shigemoto and S. Hirano, Ind. Eng. Chem. Res., 37 (1998), 185. 3. S. Satyapal, T. Filburn, J. Trela and J. Strange, Energy & Fuels, 15 (2001), 250. 4. O. Leal, C. Bolivar, C. Ovalles, J. J. Garcia and Y. Espidel, Inorganica Chimica Acta, 240(1995), 183. 5. X. Xu, C. Song, J. M. Andresen, B. G. Miller and A. W. Scaroni, Energy Fuels, 16 (2002), 1463. 6. X. Xu, C. Song, J. M. Andresen, B. G. Miller and A. W. Scaroni, Microporous and Mesoporous materials, 62 (2003), 29. 7. E. F. Vansant, P. Van Der Voort and K. C. Vrancken, Stud. Surf. Sci. Catal., 93 (1995), 193. 8. D. Zhao, J. Feng, Q. Huo, N. Melosh, G. H. Fredrickson, B. F. Chmelka amd G. D. Stucky, Science, 279 (1998), 548.


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Hollow Fiber Membrane Contactor Hybrid System for CO 2 Recovery Kew-Ho Lee , Soon-Hwa Yeon, Bongkuk Sea and You-In Park Membranes and Separation Research Center, Korea Research Institute of Chemical Technology (KRICT), P.O. Box 107, Yusung, Taejon 305-600, Korea. As an alternative technology for CO2 recovery from the flue gas, membrane contactorstripper hybrid process was studied. Porous polytetrafluoroethylene(PTFE), polyvinylidenefluoride(PVDF) and polypropylene(PP) hollow fiber modules were used as membrane contactors and their performance was compared with a conventional packed column. The CO2 was removed by the hollow fiber membrane contactor using monoethanolamine (MEA) solution as the absorbent. PVDF module with asymmetric pore structure showed the most stable gas-liquid interface and the highest removal efficiency among the three different hollow fiber membrane modules used. Mass transfer resistance model for the CO2 absorption across the hollow fiber membrane was proposed and analyzed. In application of large-scale hollow fiber membrane contactor, the hybrid process was successfully operated and recovered CO2 with a lower cost. 1. INTRODUCTION The most common process for the removal of CO2 from gas streams is gas absorption using a conventional absorption column such as packed or plate columns. Although the traditional packed bed absorbers have been used in the chemical industry for decades, there are several disadvantages such as flooding, channeling and foaming, which leads to difficulty of mass transfer between gas and liquid. An alternative technology that overcomes these disadvantages ana also otters more interfacial area than conventional approaches is non-dispersive contactor via a microporous membrane [1, 2]. Using a suitable membrane configuration such as a hollow fiber, fluids to be contacted flow on opposite sides of the membrane and the gas-liquid interface forms at the mouth of each membrane pore. (See Fig. 1) This porous hollow fiber membrane contactor can be attractive alternative to packed towers for gas absorption process [3]. Membrane contactors offer many advantages over conventional absorption system. Ihe available contact area remain undisturbed even at high and low flow rates because the two fluid flows are independent. *Author to whom correspondence should be addressed. Phone: +82-42-860-7240, Fax: +82-42-861-4151, E-mail: [email protected]

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Membrane operations usually scale linearly, so that a predictable increase in capacity is achieved simply by adding membrane modules. Modular design also allows a membrane plant to operate over a wide range of capacities. Interfacial area is known and is constant, which allows performance to be predicted more easily than with conventional system. With packed columns the interfacial area per unit volume may be known, it is often difficult to determine the loading, that is, what fraction of the available interfacial area is actually used [4]. These advantages have been proved through many researches on gas(CC>2) absorption behavior using various kind of hollow fibers and absorbents [5-9] The success of using the membrane contactor process over conventional absorption processes will largely depend on the gas-liquid system and types of hollow fiber membranes used. In addition to the properties of hollow fibers, operating conditions will also play a major role in the overall membrane contactor performance. However, reports on effect of combination with the conventional process and properties of hollow fibers on CO2 removal efficiency are relatively rare. In principle, conventional absorption process can be replaced by a hollow fiber membrane contactor. Therefore, it is necessary to test the CO2 absorption capacity using various type of membrane modules combined with conventional stripper. In this study, hybrid process equipped with the membrane contactor instead of the packed tower as the absorber in the conventional process is applied for carbon dioxide removal. The combined process was operated continuously recycling absorbent through membrane contactor and thermal stripper. PTFE, PVDF and PP hollow fibers were used as the membrane contactor and their performance of CO2 absorption compared. Finally, we developed a large-scale membrane module for CO2 recovery and then evaluated the membrane contactor-stripper hybrid process from economical point of view. 2. EXPERIMENTAL 2.1 Absorption modules The properties of the hollow fibers used in this study were shown in Table 1. PTFE (Polytetrafluoroethylene, Sumitomo Co. Japan), PVDF (Polyvinylidenefluoride, KRICT Korea), PP (Polypropylene, Memtech Co. Australia) hollow fibers are used as membrane contactor. These hollow fibers are widely used as membrane contactor for gas absorption. Table 1. Dimensions of hollow fibers for membrane contactor Dimensions

PP

PTFE

PVDF

I.D.(iam) O.D.(nm) Pore size((im) Surface area(m2/m3)

250 550 0.25 2855

1000 1913 1.0 1340

830 1070 0.03 1488

Figure 2 shows SEM of the porous PVDF hollow fibers developed in the previous study. PVDF hollow fiber is asymmetric, while PTFE and PP hollow fibers are symmetric. Cylindrical module of 0.02m i.d. and 0.23m length was used as the hollow fiber membrane contactor. The packing density was 0.4 for PP and PVDF module and 0.64 for PTFE module. For comparison, a packed column was used as the conventional process. The 1/4" PTFE

425 rasching rings were used as packing material and the column was 0.027m i.d. and 0.2m length. In a large-scale test, hollow fiber membrane module of 0.07m i.d. and 0.52m length was used. Details of the hollow fibers for membrane contactor have been reported previously [10]. 2.2 Hybrid process The experimental apparatus for CO2 Fig. 2. SEM image of porous PVDF hollow recovery by hybrid process using fiber membrane. membrane contactor was shown in Fig. 3. A gas mixture containing 25 vol.% CO2 in balance of N2 was fed into the fiber lumen and was in countercurrent contact with a absorbent solution fed in the shell side of the modules. 5wt.% monoethanolamine (MEA, Junsei chemical co.) aqueous was used as CO2 absorbent. The carbon dioxide in the feed gas reacts chemically with the absorbent while the purified gas is vented to the atmosphere, and the absorbent enriched by CO2 is pumped from the contact tower to a heat exchanger. The rich absorbent is preheated in the lean/rich exchanger by the hot lean solution returning from the regenerator/stripper on its way back to the absorber. The rich absorbent solution enters the top of the stripper where it flows down through the vessel countercurrent to the stripping steam generated in the solution reboiler. Steam and absorbent vapours move up the regenerator column, condensing as CO2 is liberated and the absorbent solution is heated. Uncondensed steam and carbon dioxide leave the top of the regenerator and then enter the reflux condenser. The condensate is returned to the system while the carbon dioxide is removed to further processing. The lean absorbent solution is pumped from the bottom of the regenerator directly to the lean/rich exchanger. The absorbent leaves the lean/rich exchanger after giving up heat to the rich solution and then enters a cooler, where its temperature is further lowered before being returned to the absorber. The absorbent and feed gas temperatures were maintained as 313K and 298K, respectively. Fig. 3. Experimental set-up of membrane contactorstripper hybrid system for CO2 recovery. The outlet CO2 concentration of absorber was measured using the gas chromatography (Shimadzu GC-14B) equipped with thermal conductivity detector at an interval of 20 min. All data were obtained at steady state after sufficient operating time was allowed for system stabilisation. Ten samples were taken under the same operating conditions and the average value was calculated. The deviation between the sample values and the average value is less than 5%.

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3. RESULTS AND DISCUSSION 3.1 CO2 absorption rate For hollow fiber membrane contactor, CO2 absorption flux of PVDF hollow fiber was 2 and 3 times higher than that of PP and PTFE hollow fibers, respectively. For all hollow fibers, CO2 flux increased initially with gas velocity, the effect being more pronounced in the case of PTFE and PVDF, and the flux then tended to reach an asymptotic value at high gas velocity. This means that the absorption capacity of absorbent reach a maximum value. When CO2 flux is independent of gas velocity, mass transfer in the liquid phase will be important and control the flux if the membrane pores are non-wetted. It was known that the mass transfer resistance in the gas phase is negligible if the gas velocity is not very low. Figure 4 shows CO2 absorption rate per unit volume for absorber modules when CO2 removal efficiency is 90-95%. Comparing to packed column, membrane contactor module exhibited higher absorption rate and particularly, absorption capacity of PVDF hollow fiber module was 2 times higher than that of conventional packed column. It may thus be presumed that membrane contactor possess a large amount of available liquid-gas contact area accelerating mass transfer of CO2. Because of the independent flow path in the Fig. 4. CO2 absorption rate per unit volume hollow fiber membrane module without for absorber modules. flooding and channeling, the membrane contactors can be provide considerably higher CO2 absorption efficiency. 3.2 Mass transfer coefficient and resistance For the mass transfer across a porous gas-liquid membrane contactor, there exist the driving forces in the gas phase, membrane pores and the liquid phase [11]. The overall resistance to mass transfer is the sum of three individual resistance; gas phase, membrane and liquid phase. Thus, the overall liquid phase mass transfer coefficient, KL can be represent as (1/KL) = {VEk\) + (l/kuKRT/He)(d^do) +(l/kG)(VHe)(d1/d0) (1) where k°L, kM and kc are the liquid phase, membrane and gas phase mass transfer coefficients, respectively; d\ and d0 are inner and outer diameter of the fiber, respectively; He is the Henry's constant. E is the enhancement factor for the liquid-phase mass transfer coefficient due to chemical reaction [8]. The value of KL was determined experimentally using the following equation, KLa = Q/(APJHe) (2) where Q is the CO2 absorption rate per unit volume of the contactor, APtm is the log-mean partial pressure of CO2 calculated from the log-mean of the inlet and outlet CO2 concentration in the gas phase and, the average column shell side pressure. For each absorber module used in this study, the overall mass transfer rates, KLa determined from the equation (2) are listed in Table 2. Membrane contactors show higher mass transfer rate than packed column. In the membrane contactors, particularly, the PVDF hollow fiber module exhibited the highest mass transfer coefficient of 0.562m/s, which is three times the value typically obtained for

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conventional packed towers. Table 2. Overall mass transfer coefficients for CO2 absorption Absorber PVDF PTFE PP Tower

Gas(m/s) 0.133 0.091 0.360 0.016

Liquid(m/s) 0.0018 0.0035 0.0018 0.001

KLa(sl) 0.534 0.311 0.392 0.182

The membrane contactor increased the available gas-liquid contactor area and mass transfer coefficient, and thus the hybrid process showed a higher CO2 removal efficiency than the conventional absorption tower. PVDF module with smaller pore size showed the most stable gas-liquid interface and the highest removal efficiency among the hollow fiber membrane contactor modules. In the calculation of mass transfer resistance for the CO2 absorption, membrane resistance of PP and PTFE hollow fibers was dominated due to the penetration of the liquid into pores. PVDF module possessing the highest CO2 absorption capacity showed the smallest membrane resistance, and in this case, resistance in the liquid phase prevailed. 3.3 Large-scale module test A large-scale PVDF membrane module containing several thousand hollow fibers was designed and applied to the hybrid process for CO2 recovery from the flue gas. The membrane contactor module of 0.07m i.d. and 0.52m length was able to remove CO2 of 90% from the flue gas of 3.0Nm3/h, while conventional packed tower of 0.1m i.d. and 2m length could treat the flue gas of 4.5Nm3/h with CO2 removal efficiency of 90%. As shown in Fig. 5, therefore, the CO2 absorption rate per unit volume of the PVDF membrane contactor was 3 times higher than that of the packed column. In long-term stability test, the hybrid process was successfully operated for 30 days maintaining CO2 removal efficiency of above 90%. The economical evaluation through a recovery cost and electric power consumption showed that compared to conventional CO2 recovery processes [12], the membrane contactor

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hybrid process was able to recover CO2 with a lower cost as shown in Fig. 6. The electric power basic unit and recovery cost calculated from operating power consumption and module capital cost were 0.39kWh/Nm3-CO2 and 60.7$/ton-of-carbon, respectively. The economical hybrid process was due to higher CO2 removal efficiency and lower module cost than those of the conventional processes. 4. CONCLUSIONS Membrane contactor-stripper hybrid process was used to recover CO2 from the flue gas. Porous PTFE, PVDF, PP hollow fiber modules were used as the membrane contactor and compared with a conventional packed column. The CO2 was removed by the porous polymeric hollow fiber membrane contactor using monoethanolamine(MEA) solution as the absorbent. Then, the CO2 was recovered from the absorbent by the thermal stripping tower with steam and the stripped absorbent was recycled. The membrane contactor increased the available gas-liquid contactor area and mass transfer coefficient, and thus the hybrid process showed a higher CO2 removal efficiency than the conventional absorption tower. PVDF module with smaller pore size showed the most stable gas-liquid interface and the highest removal efficiency among the hollow fiber membrane contactor modules. The overall mass transfer rate, Kâ of the PVDF membrane contactor was found to be 3 times higher than that of the packed column. In the calculation of mass transfer resistance for the CO2 absorption, membrane resistance of PP and PTFE hollow fibers was dominated due to the penetration of the liquid into pores. PVDF module possessing the highest CO2 absorption capacity showed the smallest membrane resistance, and in this case, resistance in the liquid phase prevailed as 62% of total resistance. In addition, the membrane contactor hybrid process was successfully operated in a large-scale test with long-term stability and proved to be economic by evaluation through CO2 recovery cost and operating power consumption. References 1. A. Gabelman and S.-T. Hwang, J. Membrane Sci., 159 (1999) 61. 2. R.D. Noble and S.A. Stern, Membrane Separations Technology Principles and Applications, p.467 , Elsevier, Amsterdam, 1995. 3. S. Karoor and K.K. Sirkar, Ind. Eng. Chem. Res., 32 (1993) 674. 4. H. Kreulen, C. Smolers and W.P.M. van Swaaij, J. Membrane Sci., 78 (1993) 197. 5. O. Falk-Pesersen and H. Daninstorm, Energy Conversion & Management, 38 (1997) S81. 6. S. Nii and H. Takeuchi, Gas Sep. & Purif., 8 (1994) 107. 7. N. Nishikawa, M. Ishibashi, H. Ohta, N. Akutsu, H. Matsumoto, T. Kamata, H. Kitamura, Energy Conversion & Management, 36 (1995) 415. 8. H.A. Rangwala, J. Membrane Sci. 23 (1996) 229. 9. K. Li and W. K. Teo, Sep. and Purif. Tech., 13 (1998) 79. 10. Y.-I. Park, K.-S. Lee, S.-H. Yeon, B. Sea and K.-H. Lee, J. Ind. Eng. Chem., 9 (2003). 11. S.-H. Yeon, B. Sea, Y.-I. Park and K.-H. Lee, Sep. Sci. Tech., 38 (2003) 271. 12. Report on greenhouse gas reduction technology, KIPO, Seoul, 2002.


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Optimization of CO2 Absorption Process with MEA Solution Seok Kim, Hyung-Taek Kim and Byung-Chul Choi Dept. of Energy Studies, Ajou University, Wonchon-dong San-5, Paldal-gu, Suwon, Korea 442-749 FAX: +82-31-219-2969. E-mail: [email protected] The CO2 absorption with MEA (mono-ethanolamine) solution is simulated to evaluate the operating condition of the bench-scale, continuous CO2 absorption apparatus located in the Korea Institute of Energy Research. The main objective of this study is to determine the most energy-effective process in the regenerator through the simulation of various conditions of CO2 absorption. The flow rate of flue gases and MEA solution are varied from 35~80 1/min and 1.0-3.0 1/min. Regenerating process conditions of maximum CO2 separation efficiency on this simulation are determined with the change of reboiler heat duty and number of stage.. 1. INTRODUCTION Carbon dioxide is recognized as the major contributor to the greenhouse effect due to its abundance, and is thus a primary target for reduction. CO2 separation technologies by absorption with aqueous amine solution and by using membrane developed actively in these days.[4] Among them the gas absorption with a chemical reaction process using amine is considered to be the most cost effective and the best proven operability record. The main purpose of this study is to minimize the amount of energy required in desorption (commonly called regeneration) process through the simulation of various process concept of solvent absorption and to suggest the optimum operating condition to the actual bench-scale CO2 absorption experimental setup. Bench-scale, continuous CO2 absorption reactor (capacity =5 Nm3/hr) located in the Korea Institute of Energy Research is modeled and simulated to determine optimum operating condition. In the present study, simulations are focused on regenerating process. 2. EXPERIMENTAL Behavior of CO2 absorption as electrolyte solutions is described by using 'Electrolyte Non Random Two Liquid (ELEC-NRTL) Model' and 'AMINES PROPERTY' in AspenPlus™. For the simulation of absorber and regenerator, 'RadFrac Column' is used. During the simulation,

430

gas flow rates are typically varied from 35-80 1/min and liquid flow rates from 1.0-3.0 1/min. MEA solution concentration is fixed at 30 wt%. Flue gas flow rate is set at 50 1/min and the composition of flue gas is CO2=13.89%, N2=82.56% and 02=3.55%. Top stage pressure of absorber and regenerator are set at 1.05 atm. 2.1 Reaction chemistry The main reaction mechanism of MEA solution chemistry is as follows, which is coded in the simulation. 2H 2 o CO 2 +

H3 OH J

f

O
values as function of [HCO3']

Figure 4, Several empirical models prediction ofqcoz values as function of X

Cause of facts of the approximation of Haldane kinetic could be accepted in both of model plot of carbon dioxide fixation and cellular population, it was agreed that this model was suitable for growth and CO2 removal kinetic of A. cylindrica in small scale of reactor in series at low temperature and light intensity. Except the values in first reactor, comparison between experimental and calculated qcoi indicated a significant correlation between both of values during microbial growth (Figure 5). Similar to above explanation, in whole reactors, comparison result among experimental data

467

and calculation of ju, also indicated a deeply correlation between both of values during microbial growth.

Figure 5. Comparison result of calculation and observation values of qC02 and /i Change of CTRICUR values during cultivation at different temperature condition was presented in Figure 6. CO2 transferred rate (CTR) and CO2 uptake rate (CUR) was calculated by followed equations. CTR = gC02-X

(3)

CUR = a-— dt

(4)

Alpha value calculated from dry biomass carbon content was 1.74. CTRICUR values increased steeply at beginning step of cultivation and values decreased more slightly after maximum value was reached. Factually, [CTR/CUR]m, the maximum value of the first and fourth reactors revolved around 55. [CTR/CUR]m of second and third reactor tends to around 10 and the last reactor tends about 35. These are high certain value, as the same reason of the result of previous work [4]. hi case of this facts, Carbon dioxide Concentration Mechanism (CCM) have an important role. In this mechanism, each cell have a CO2 storage that was also called Q pool, is a kind of cell apparatus which kept of unusable CO2 in phenomenon of higher difference between CTR values and sum of CUR values, culture medium CO2 concentration and [HCO3"]. And similar with estimation of that previous work, we exactly estimated that, Q pool have responsibility to this experiment result of CTRICUR. Result of light energy utilities for microbial growth population of A. cylindrica in this experiment was concluded in Table 2. It was concluded that total energy utility (£x) or light energy utility efficiency (rj) for biomass production in multiple reactor series, factually, more efficient than results of microbial growth in single reactor (267 kJ/g and 0.0711 of EK and if) [7]. Except the result in first reactor, these were around 3.7 times more high in case of 77 value and of course, approximately 5 times more efficient in case of Ex value. Total energy utility

468

for carbon dioxide fixation (Ecf) in this experimental works, also higher than result of single reactor. Additionally, in case of both of experimental condition, no significant difference of the efficiency of light energy absorbed in microbial culture (?7Ab)-

Figure 6. Values of CTR/CUR N 1 2 3 4 5

Ex [kJ/g] 122 64.9 65.2 57.5 51.9

Ecf[kJ/g] 5.08 5.23 8.09 0.804 0.879

TlAb H

TIN

0.793 0.818 0.824 0.746 0.809

0.139 0.270 0.271 0.278 0.334

Based on 0.321 g/[dm3.h] of overall average carbon dioxide transfer rate (C77?av), multiple reactor series increased an ability of carbon dioxide removal system. The CTRm of multiple reactor series was approximately, two and half times of single reactor. Finally, refer to value of final cell concentration, multiple reactor series also increased biomass production and it was about three times. REFERENCES 1. 2. 3. 4. 5. 6. 7. 8.

Mustaqim D and K Ohtaguchi, Energy, 22 (1997) 353 Ohtaguchi K, Technology (suppl.), 7, (2000) 175 Post A F, R de Wit, L R Mur, J. Plankton Res., 7, (1985) 487 Wijanarko A and K Ohtaguchi, Comp. Biochem. Physiol., 134 (2003) S214, 2003 Utkilen H.C., J. Gen. MicrobioL, 128 (1982) 1849 Bailey J E and D F Ollis, Biochemical Engineering Fundamentals, New York, 1986 Schugerl K and K H Belgardt (eds), Bioreaction Engineering, Berlin, 2000 Ohtaguchi K and A Wijanarko, Technology, 8 (2002) 267


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Viability and activity of microorganisms in pressurized and supercritical carbon dioxide Baoquan Zhang*, Xiufeng Liu and Yiru Gan School of Chemical Engineering and Technology & State Key Laboratory of Ci Technology, Tianjin University, Tianjin 300072, China The viability and biological activity under various pH, exposure time and temperature in pressurized CO2, C2H4 and N2 were measured for Sacchawmyces cerevisiae and Absidia coerulea. A remarkable change of biological activity in CO2 fox Absidia coerulea appeared at near-critical region indicating the special effect of phase transition. The initial pH has a most profound influence on viability and biological activity in pressurized CO2 environment. It is possible to achieve high viability and biological activity in pressurized CO2 for both Sacchawmyces cerevisiae and Absidia coerulea. 1. INTRODUCTION Supercritical CO2 (SCCO2) possesses many advantages over organic solvents due to its low toxicity, non-flammability, high availability and environmentally benign nature. It has been widely used in food processing and preservation, alcohol production, and enzyme-catalyzed reactions [1-6]. When ordinary organic solvents are used in fermentation processes, the product inhibition or low yield of product extraction often leads to low process efficiency. In-situ removal of inhibitory products by supercritical fluid extraction, especially by SCCO2, can be a promising solution. The relevant research dates back to the late 1980s. The main contributions are listed in Table 1 [7-15]. Unlike most enzyme-catalyzed reactions, the addition of coenzymes is not necessary for the reaction catalyzed by whole-cell microorganisms. This is especially useful when the bioprocess undergoes a series of enzyme-catalyzed reactions or the reaction occurs on specific sites inside microbial cells such as the Clip-hydroxylation of Reichsterin's substance acetate (RSA) that needs C11 (3-hydroxylase and its coordination with coenzymes on cell membrane [6]. It is evident that most enzymes can keep active in SCCO2 [1, 2]. On the contrary, the cell-catalyzed reaction in SCCO2 is rarely studied. The results derived from the available literature are often contradictory [6-12, 15]. The objective of this contribution is to investigate the influence of various factors on the viability and biological activity of two microorganisms in pressurized CO2, C2H4 and N2 to find possible mechanism. * Corresponding author. Fax: +86-22-27405080. Email: [email protected]

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Table 1 Research on viability and activity of microorganisms in pressurized and supercritical fluids for fermentation or fermentation with in-situ extraction Pressure

Temperature

MPa 0.1-7.0 7.0

K 303 303

CO2, N2 CO2

Thabaultetal 1987 [7]

0.1-0.8 0-100

303 281-316

CO2, O2 CO2

Onken, 1990 [9] Isenschmid et al 1995 [10]

6.9 14.0 7.0

333 308 333

CO2, C2H6, N2 CO2 C2H6, C3H8, N 2

Knutsonetal 1999 [11] Fabreetal 1999 [12] Berberichetal2000[13]

Clostridium thermocellum

1.5-7.0

298, 333

C5H12, C6H]4... C2H6, C3H8, N2 C5H12, C6H14...

Berberichetal2000[14]

Geotrichum candidum Absidia coerulea Saccharomyces cerevisiae

10.0 0-8.0

308 301-312

CO2 CO2, C2H4, N2

Matsuda et al 2000 [15] Zhang et al, this study

Microorganisms Saccharomyces cerevisiae Saccharomyces cerevisiae Saccharomyces rouxii Pseudomonas fluorescens Candida utilis Kluyvermyces fragilis Saccharomyces cerevisiae Clostridium thermocellum Kluyvermyces marxianus Clostridium thermocellum

Gases

Sources

LTtalien et al 1989 F81

2. EXPERIMENTAL Saccharomyces cerevisiae (S. cerevisiae) and Absidia coerulea (A. coerulea) were used as model microorganisms. The influence of pressure, temperature, exposure time, and pH on the viability and activity of microbial cells was investigated. The experiment was conducted in a 0.5 L autoclave made of stainless steel (rated to 350 °C and 21.6 MPa). In all experiments, either pressurized or depressurized process took 30 min. A syringe pump combined with pressure reducing regulator and two back-pressure regulators was used to achieve the steady pressure change in the pressurized process. The pressure was monitored using a pressure transducer. Samples were collected after the pressure release. The experimental apparatus was shown in Fig. 1. The viability of S. cerevisiae was determined by the methylene blue method [12]. The result was the average value over six parallel counts by using a Thoma cell. For determination of the spore viability of A. coerulea, spores were first incubated on culture disc for 48 hours, and then the colonies were counted for the A. coerulea exposed in pressurized CO2 (or C2H4) and in atmospheric pressure. The viability is just the ratio of the above two colony numbers. The product P-hydrocortisone was analyzed by an Agilent 1100 HPLC. The 25cm long C18 column was used, and the absorption wavelength was chosen at 242nm. The culture medium and fermentation procedure for S. cerevisiae and A. coerulea were the same as previous literature reports [8, 16]. The optimal pH of A. coerulea and S. cerevisiae at atmospheric pressure was first measured. It is about 6.5 for A. coerulea and 5.0 for S. cerevisiae [6]. The specific activity of

471

A. coerulea pellets was calculated according to

Specific Activity = (XP/X0)xl

00%

where X represents the conversion of RSA after 24 hours transformation under atmospheric pressure and 28°C. Subscript P corresponds to the A. coerulea exposed in pressurized CO2 (or C2H4), and 0 to the A. coerulea cultured in atmospheric environment. Here, the specific activity indicates the effect of pressurized CO2 (or C2H4) on the catalytic ability of A. coerulea pellets.

Fig. 1. Flow sheet of experimental setups 1. CO2 cylinder, 2. syringe pump with cooling unit, 3. buffer, 4. two-way ball valve, 5. filter, 6. pressure reducing regulator, 7. pump, 8. thermostatic bath, 9. thermocouple, 10. autoclave, 11. jacketed heat exchanger, 12 electric motor, 13. pressure transducer, 14. display monitor, 15. back-pressure regulator, 19. tubing coil, 20. gas-liquid separator. 3. RESULTS AND DISCUSSION As shown in Fig. 2, the pressure of CO2 has a strong influence on viability for S. cerevisiae but not for A. coerulea spores. It is worth noting that the variation at near-critical region is most evident, implying that the influence of phase transformation of CO2 cannot be neglected. The influence of exposure time on the viability of S. cerevisiae becomes stronger above the critical point of CO2. This observation is in good agreement with literature reports [7, 8, 10]. Experimental results in Fig. 3 indicate the optimal initial pH (pH before the sterilization of cultural medium) for S. cerevisiae is going up to about 6 rather than 5 under pressurized CO2 environment. In comparison with pressurized nitrogen environment, a high viability can also be achieved in SCCO2 under an appropriate initial pH. Actually, higher initial pH value may compensate the acidic pH due to the increased solubility at pressurized CO2. So the increase of initial pH is necessary to make the system of pressurized CO2 operate at near the inherent optimal pH for S. cerevisiae. The same effect of pH on the viability of A. coerulea spores and

472

on RSA conversion in pressurized CO2 is also observed, as what is shown in Fig. 4. Here the optimal initial pH was about 8 instead of 6.5. Hence, the increase in viability could be achieved by adjusting the initial pH. By comparison, 85% viability of A. coerulea spores can be achieved in pressurized C2H4, which is independent upon initial pH. The specific activity of A. coerulea pellets in pressurized C2H4 can reach up to 75%. Besides, the optimal initial pH does not change much in pressurized C2H4 environment compared with that at atmospheric pressure. The high viability and biological activity can also be achieved in pressurized ethane, propane, and pentane and so on [11, 13, 14]. Even though the maximum biological activity at present is only about 20% of the A. coerulea pellets without exposure in pressurized CO2, it still possesses application potential due to the favorable characteristics and environmental benefits of SCCO2. It is worth mentioning here that the internal pH should play a vital role in cellular activity if the system is to be exposed in pressurized CO2 [3]. Therefore, RSA conversion in pressurized CO2 can be improved by regulating the diffusion of CO2 across cellular membrane.

Fig. 2. Variation of viability with pressure of CO2 at 306K and initial pH of 6.5. S. cerevisiae: 30 min ) and 60 min (O) exposure. A. coerulea spore: 60 min ) exposure, pc is the critical pressure of CO2.

Fig. 3. Effect of initial pH on the viability of Saccharomyces cerevisiae after 30 minutes exposure in pressurized nitrogen (O) or carbon dioxide (D)at 7.5MPa and 306.2K.

The effect of temperature on the biological activity of A. coerulea pellets was also investigated. The result is shown in Fig. 5. A distinct change appears at near-critical temperature. What has been observed for many solid- and enzyme-catalyzed reactions in the near-critical region is that the conversion gets improved in supercritical fluids [1, 2]. On the contrary, the observed outcome of this cell-catalyzed reaction is just opposite. It may be due to the low spore viability in SCCO2. As what will be demonstrated below, it may be caused by the hypha reduction of A. coerulea pellets on which the reaction occurs [6].

473

Fig. 4. Effect of initial pH on biological activity and spore viability for A. coerulea at 7.5MPa and 306.2K after 30 minutes exposure in CO2: O viability specific activity C2H4: viability specific activity.

Fig. 5. Variation of biological activity with temperature for Absidia coerulea pellets at initial pH of 6.5 after being exposed in 7.5MPa carbon dioxide for 30 minutes, where Tc is the critical temperature of carbon dioxide.

The micrograph illustrated in Fig. 6 indicates that the hypha of A. coerulea pellet is significantly reduced after being exposed in pressurized CO2. The reduction of hypha may be a major reason for the decrease in RSA conversion. However, the hypha does not change much after being exposed in pressurized C2H4.

Fig. 6. Micrographs ofAbsidia coerulea pellets at initial pH of 6.5 and 306.2K The spore viability and biological activity of A. coerulea are very sensitive to CO2 pressure, pH and temperature, where the morphology of A. coerulea may also play an important role. RSA conversion in SCCO2 is not satisfactory at present stage. Based on the advantages and environmental benefits of SCCO2 over organic solvents in downstream separation, further

474

research is necessary to seek both piezophilic and piezotolerant microorganisms together with suitable transformation conditions. 4. CONCLUSIONS Experimental investigations on the viability and biological activity of S. cerevisiae and A. coerulea under various pH, exposure time and temperature in pressurized CO2, C2H4 and N2 have been performed. The viability and biological activity of A. coerulea in pressurized CO2 and C2H4 have not been previously reported yet. Pressure, temperature, exposure time, especially pH and phase transition have a profound influence on viability and biological activity in pressurized CO2 environment. Although the biological activity of A. coerulea in pressurized C2H4 is quite satisfactory, a better result in pressurized CO2 needs to be further sought due to its unique characteristics and environmental benefits. Acknowledgements: The authors acknowledge the financial support of NSFC under Grant No. 20076033. Baoquan Zhang thanks the Ministry of Education for a fellowship grant. REFERENCES 1. A.J. Mesiano, E.J. Beckman, A.J. Russell, Chem. Rev., 99 (1999) 623. 2. Z.Q. Zhu, Supercritical Fluids Technology: Principles and Applications, Chemical Industry Press, Beijing, 2000. 3. N.M. Dixon, D.B. Kell, J. Appl. Bacteriol., 67 (1989) 109. 4. E. Debs-Louka, N. Louka, G. Abraham, V. Chabot, K. Allaf, Appl. Environ. Microbiol., 65(1999)626. 5. S. Spilimbergo, N. Elvassore, A. Bertucco, J. Supercrit. Fluids, 22 (2002) 55. 6. X.F. Liu, Effects of Carbon Dioxide on Physiological Properties of Microorganisms, Dissertation, Tianjin University, Tianjin, 2002. 7. J. Thibault, A. LeDuy, F. Cote, Biotechnol. Bioeng., 30 (1987) 74. 8. Y. L'ltalien, J. Thibault, A. LeDuy, B-iotechnol. Bioeng, 33 (1989) 471. 9. U. Onken, Biotechnol. Bioeng, 35 (1990) 983. 10. A. Isenschmid, I.W. Marison, U. von Stockar, J. Biotechnol, 39 (1995) 229. 11. C.E. Fabre, J.-S. Condoret, A. Marty, Biotechnol. Bioeng, 64 (1999) 392. 12. B.L. Knutson, H.J. Strobel, S.E. Nokes, K.A. Dawson, J.A. Berberich, C.R. Jones, J. Supercrit. Fluids, 16 (1999) 149. 13. J.A. Berberich, B.L. Knutson, H.J. Strobel, S. Tarhan, S.E. Nokes, K.A. Dawson, Ind. Eng. Chem. Res, 39 (2000) 4500. 14. J.A. Berberich, B.L. Knutson, H.J. Strobel, S.E. Nokes, K.A. Dawson, S. Harhan, Biotechnol. Bioeng, 70 (2000) 491. 15. T. Matsuda, T. Harada, K. Nakamura, Chem. Commun, No. 15 (2000) 1367. 16. E. Brzezowska, J. Dmochowska-Gladysz, T. Kolek, J. Steroid Molec. Biol, 57 (1996) 357.


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Novel conversion of carbon dioxide to methanol catalyzed by sol-gel immobilized dehydrogenases Zhongyi Jiang *, Songwei Xu and Hong Wu State Key Lab. of C1 Chemistry and Chemical Technology, School of Chemical Engineering and Technology, Tianjin University, Tianjin 300072, China Effective utilization of carbon dioxide has attracted considerable research and development interests in recent years. Conversion of carbon dioxide into methanol has been regarded as a promising approach since it can recycle the greenhouse gas and produce valuable chemicals. Herein, we describe an enzymatic pathway of carbon dioxide utilization which employs sol-gel co-encapsulated formate dehydrogenase (FateDH), formaldehyde dehydrogenase (FaWDH) and alcohol dehydrogenase (ADH) as catalysts, NADH as an electron donor. The effects of reaction temperature, reaction pH, the amount of enzyme and NADH on the yields of methanol have been preliminarily investigated, the highest yield of methanol was up to 92.1%. The catalytic activity of ADH encapsulated by PEG-modified silica and the catalytic activity of pegylated ADH are also investigated. 1. INTRODUCTION Partial hydrogenation of CO2 by chemical approaches such as heterogeneous catalysis, electrocatalysis, photocatalysis with oxide-based catalysts has been extensively studied. These chemical approaches suffer from the following drawbacks including low efficiency, high cost, high energy consumption, and the need for special equipment adapted for high temperature or highly corrosive environments [1-3]. Herein, a new enzymatic approach was reported, CO2 was converted into methanol through three step reduction reactions catalyzed by three different kinds of dehydrogenases, formate dehydrogenase (FateDH), formaldehyde dehydrogenase (FaidDH) and alcohol dehydrogenase(ADH), immobilized by sol-gel entrapment [4-8]. Reduced nicotinamide adenine dinucleotide (NADH) was employed as the electron donor (Scheme 1) [9].

* Corresponding author. Fax: +86-22-27404066. Email: [email protected]

476

Scheme 1. Enzymatic conversion of carbon dioxide to methanol The enzymatic approach has several advantages: low temperature and pressure, high energy efficiency, high conversion rates, no need for special equipment adapted to high temperature or highly corrosive environments [10-14]. The effects of reaction temperature, reaction pH, the amount of enzyme and NADH on the yields of methanol have been preliminarily investigated, the highest yield of methanol was up to 92.1%. The chemical modification of sol-gel process and ADH on the catalytic activity of ADH is also investigated. 2. EXPERIMENTAL 2.1. Preparation of plain (typical) silica gel and enzymatic catalysis The initial sol was prepared by mixing TEOS with HC1 solution vigorously for a short period, and then NaOH aqueous solution was added dropwisely. The homogenous sol was formed rapidly. Enzyme solution in 0.1M phosphate buffers was added to l.OmL of the above sol to form gel in cuvette. The gels were aged for 72 hours and dialyzed for 48 hours. The gel was then transferred to a cuvette and immersed in 2.OmL NADH stock solution for 48 hours, which allowed NADH to diffuse into the gel completely .The concentrations of NADH ranged from 0.025M to 0.075M at pH 7.0-8.0. Then CO2 was bubbled into the reaction system within 8 hours for the maximum production of methanol at 25 "Cor 37°C.The pressure was maintained at 0.3MPa or 0.5MPa.The kinetic studies of free and immobilized dehydrogenases were carried out simultaneously. Table 1. Enzyme content in 1# and 2# gels Enzyme content 1# gel 2# gel

FateDH

FaidDH

ADH

7mg 14mg

2mg 4mg

2mg 4mg

2.2. Preparation of PEG modified silica gel and enzymatic catalysis The initial sol was prepared by mixing TEOS with HC1 and PEG 600 aqueous solution vigorously for a short period, and then NaOH aqueous solution was added dropwisely.. The homogenous sol was formed rapidly. ADH solution in 0.1 M phosphate buffers was added to 1 .OmL of the above sol to form gel in cuvette. The gels were aged for 72 hours and dialyzed for 48 hours. The gel was then transferred to a cuvette and immersed in 2.0mL NADH stock solution for 48 hours, which allowed NADH to diffuse into the gel completely. 2.3. Pegylation of ADH and preparation of silica gel

477

3.5mgADH and 10.5mg activated PEG was dissolved into a 0.1M phosphate buffer (pH7.0, 5mL) and the solution was stirred for 2h. The degree of modification was determined by TNBS method. The initial sol was prepared by mixing TEOS with HC1 solution vigorously for a short period, and then NaOH aqueous solution was added dropwisely.. The homogenous sol was formed rapidly. Pegylated ADH solution in 0.1M phosphate buffers was added to l.OmL of the above sol to form gel in cuvette. The gels were aged for 72 hours and dialyzed for 48 hours. The gel was then transferred to a cuvette and immersed in 2.0mL NADH stock solution for 48 hours, which allowed NADH to diffuse into the gel completely. The reaction of converting formaldehyde into methanol was employed to evaluate the catalytic activity of immobilized ADH. The reaction rates and Michaelis constants were fitted according to the Bio-Bio sequential reaction mechanism proposed by Deziel et al, 3. RESULTS AND DISCUSSION The yields of methanol were calculated based on the amount of NADH consumed. Yield of methanol (%)=3*moles of methanol produced/ moles of NADH initial amount* 100% The experimental results indicated that the optimum reaction conditions were around pH=7.0, T=37"C,the highest yield using 1* gel as catalyst reached 92.1%, and the highest yield of methanol catalyzed by free enzymes was 98.1%. The catalytic activity of immobilized enzyme was decreased mainly because of the minor three-dimensional structural change of enzyme and the extra diffusion resistance after immobilization [1]. Table 2. showed the typical results of immobilized enzymatic reaction. Table 2. Effects of experimental conditions on the enzymatic conversions of CO2 No.

Gel

NADH

Pressure

Yield

amount

/Mpa

/%

No.

Gel

pH

T/lC

(iniol

1

IS

7.0

2

2#

3

IS

4

2#

5

is

amount

/MPa

fy.

150

0.5

27

6

18

7.0

25

0.3

52.3

7

M

7.5

37

100

0.3

66.9

03

42.2

8

l»

1.0

37

100

0.3

49.5

03

35.1

>>

I"

S.O

25

50

0.3

68.8

0.3

30

10

I*

K.0

37

ISfJ

0J

32.0

0,3

7.0

37

100

7.0

37

150

7.0

37

150 100

Yield

92A

100

25

Pressure

,n:l"l

37

7.0

NADH

Both the free and immobilized dehydrogenases obeyed the Bio-Bio sequential reaction mechanism proposed by Deziel [15]. The catalytic activity of immobilized FateDH, immobilized Fa|dDH, and immobilized ADH were only 60-70%, about 60% and 40-50% of the free dehydrogenase counterparts. The increase of Km-based parameters reflected the affinity between the enzymes and the reactants decreased after immobilization, which was consistent with the results of conversion experiments. The TEM photographs showed that the dehydrogenases were uniformly dispersed in the

478

whole gels. And UV Spectra observed a minor shift of the maximum absorption wavelength. In addition, the even distribution of enzymes in silica gel matrix was also confirmed by measuring the catalytic activities of immobilized enzymes with different weights. The enzymatic reaction catalyzed by ADH in plain silica gel, ADH in PEG modified silica gel, pegylated ADH in plain silica gel obeyed Michaelis-Menten kinetics and the kinetic parameters were evaluated by Dalziel's double-substrate model (Fig.l).

Fig.l. Reaction rate constants fitted by Dalziel's double-substrate model. Kinetic constants for these immobilized ADH catalyzed reactions were shown in Fig. 1. The results indicated that peylation of ADH enhanced the affinity between the substrate and enzyme. The catalytic activity of ADH immobilized in the PEG-modified gel is higher than that in the plain gel. This may be due to: (1) enlargement of the pore size. From Fig.2, we can see the pore size of PEG -modified gel is 9.34nm while the pore size of plain gel(typical gel) is 7.76nm (PEG/TEOS molar ratio of 1%). The possible function mechanism of PEG in the formation of silica gel pore was shown in Fig.3. (2) The improved hydrophilicity and enhanced mechanical strength of silica gel.

Fig.2. The pore diameter distribution of the plain gel and PEG-modified gels

479

Our experiments presented in this study have also clearly demonstrated that PEG modification on ADH can significantly enhance the catalytic activity and stability of immobilized ADH. The relatively high catalytic activity and stability of modified ADH probably due to steric hinderance of protein unfolding of perhaps to perturbation of the bulk water structure in the region close to PEG. In addition, interaction between enzyme and its surrounding environment may lead to a conformational change in the tertiary structure. While large molecule PEG as a protection material between enzyme and environment may reduce the probability of enzyme conformation change.

Fig.3

The proposed function mechanism of PEG in the formation of silica gel pore

The stability of alcohol toleration of both native and SC-PEG modified ADH were studied by measuring the activities of the free enzymes in the enzymatic reaction of converting formaldehyde into methanol. Modified ADH in reaction after being immersed in 1% alcohol solution for 2 minutes resulted in 50% of original modified ADH specific activity, while native ADH after being immersed in identical alcohol solution retained 36% specific activity. The results showed that chemical modification in solution was an effective means of stabilizing ADH in organic solutions. Attaching a long-chain PEG molecule to the surface of the enzyme molecule may make the protein less flexible. And the long PEG-chain may be ascribed to a highly hydrogen-bonded structure around the enzyme. Thus, The relatively high stability of modified ADH probably due to favorable confinement of protein unfolding . In addition, our experimental data (not shown here) indicated that modification on gel and modification on ADH both resulted in enhanced immobilized ADH stability. This may be explained that the active sites were protected by the PEG molecules with branched structure. The good biocompatibility of PEG also provided ADH with a favorable microenvironment. 4. CONCULSIONS The immobilized dehydrogeases was efficiently entrapped in silica gels and the enzymatic conversion of CO2 to methanol was proved as the novel and feasible way for CO2 utilization. The enlargement of pore size by hydrophilic PEG and modification ADH by the activated PEG can both improve enzymatic conversion process. In addition, since coenzyme NADH is very expensive, a practically regeneration method must be introduced. Presently, a novel chemo-enzyme approach to regenerate NADH is being

480

investigated in our group. REFERENCES 1. G.Chiericato Jr., C.R.Arama, C.Casado, I.Cuadrado, H.D.Abrufia, Inorganica Chimica Acta., 300-302 (2000) 32. 2. Kuwabata, S.; Tsuda, R., Yoneyama, H. J., Am Chem. Soc. 116 (1994) 5437. 3. Aresta, M.; Quaranta, E.; Liberio, R., Trtrahedron, 54 (1998) 8841. 4. Gill, I.,Ballesteros, A. Trends in Biotechnology, 18(2000) 282. 5. Wei Y, Xu J., Feng Q.,.Dong H., Lin M., Mater. Lett., 44(2000) 6. 6. Smith K., Silvernail N. J., Rodgers K. R., Elgren T. E., Castro M., Parker R. M., J. Am. Chem. Soc, 124(2002) 4257. 7. Badjic J. D., Kostic N. M , J. Phys. Chem. B , , 104 (2000) 11081. 8. Livage J., Coradin T., Roux C , J. Phys. Condens. Matter , 13 (2001) 673. 9. Jiang Z.Y., Wu H., Xu S.W.et al, Chinese Journal of Catalysis, 23 (2002) 162. 10. Agarwal P.K., Webb S.P., Hammes-Schiffer S., J. Am. Chem. Soc, 122 (2000) 4803. 11. Antoniou D., Schwartz S.D., J.Phys. Chem. B, 105 (2001) 5553-58. 12. Kim k.S., Kim D., Lee J.Y., Tarakeshwar P., Oh K.S., Biochemistry, 41(2002) 5300. 13. Gulotta M., Deng H., Dyer R.B., Callender R. H., Biochemistry, 41 (2002) 3353. 14. Billeter S. R., Webb S. P., Agarwal P. K., Irodanov T., Hammes-Schiffer S., J. Am. Chem., Soc, 123(2001) 11262. 15. Pocker,Y, Chemico-Biological Interactions, 383(2001)130-132.


481

Combined Thermoneutral Processes for CO2 Emission Avoidance and Fuel Saying in the Metallurgical and Lime Industries M. Halmanna and A. Steinfeldb a

Weizmann Institute of Science, Department of Environmental Sciences and Energy Research, 76100 Rehovot, Israel b

ETH — Swiss Federal Institute of Technology, Department of Mechanical and Process Engineering, 8092 Zurich, Switzerland

By combining the endothermic reduction of iron oxide, or the endothermic calcination of limestone, with the exothermic partial oxidation of natural gas, overall thermo-neutral processes can be designed for co-producing syngas and iron or lime, in which the CO2 released by the endothermic reactions is utilized in situ by the reforming reactions, resulting in both 21-24% fuel saving and 75-86% CO2 emission avoidance relative to their conventional separate production. 1. INTRODUCTION Worldwide annual CO2 emissions from the iron, lime and cement, syngas, and zinc industries amount to 1.2, 0.70, 0.30, and 0.07 billion tons, together contributing about 8% to the worldwide anthropogenic CO2 emissions [1]. Substantial reduction of these CO2 emissions is not economically feasible by existing methods [2]. while sequestration disposal considerably raises the price of these important but low cost commodities [3]. We recently reported on a thermochemical process that combines the endothermic reduction of the metal oxides Fe2O3 and ZnO (Eqs. 1 and 2) with the endothermic H2O/CO2 reforming of CH4 (Eqs. 3 and 4), and with the exothermic partial oxidation of CH4 (Eq. 5), resulting in overall thermo-neutral reactions [4-6]. Fe2O3 +3CH4 = 2Fe+3CO + 6H2 ZnO + CH4 = Zn + CO + 2H2 CH4 + H2O = 3H2 + CO CH4 + CO2 = 2H2 + 2CO CH4 + I/2O2 = 2H2 + CO

AH°298K = 719 kJ mol"1 (1) AH°298K = 313 kJ mol"1 (2) AH°298 = 206 kJ mol"1 (3) AH°298 = 247 kJ mol"1 (4) AH°298 = -38 kJ mol"1 (5)

The conditions for thermo-neutrality were determined by thermochemical equilibrium calculations, and the reactions were confirmed experimentally at 1400 K by thermogravimetry and gas chromatography [5]. The co-production of metals and syngas predicts both fuel saving and considerably decreased CO2 emissions by comparison with their conventional separate production. Similar advantages are predicted for the thermo-neutral co-production of

* Corresponding author: M. Halmann. E-mail: [email protected]

482

lime and syngas from limestone, methane, and oxygen. By combining the endothermic CO?releasing calcination reaction (Eq. 6), CaCO3 = CaO + CO2

AH°298 = 178 kJmol"1

(6)

with the above H2O/CO2 reforming (Eqs. 3 and 4) and partial oxidation of CH4 (Eq. 5), an overall thermo-neutral process can be designed for co-producing lime and syngas, with both fuel saving and avoidance of CO2 release [7]. In the present study, the economic and environmental implications of the above processes are examined. 2. CO-PRODUCTION OF IRON AND SYNGAS 2.1. Economics The thermo-neutral co-production of Fe and syngas with natural gas (NG) and oxygen, when reactants are fed at ambient temperature, and products obtained in chemical equilibrium composition at 1400K and 1 bar, can be represented by [5], Fe2O3 + 4.87CH4 + 3.51O2 = 2Fe + 5.74H2 + 3.70CO + 3.99H2O + 1.17CO2 (7) Syngas is here assumed to be further processed to methanol. A preliminary economic assessment for such a plant is presented in Table 1, based on an assumed operation of 350 days/yr. The economics appear to be favorable. If syngas is destined for the production of hydrogen for ammonia synthesis, air could be applied instead of oxygen, thus further improving the economics. The current worldwide annual production of methanol is 30.4 million tons (metric tons) [8]. Thus, the expected production of 0.4 million tons methanol in the proposed plant would add ~1.3% to the current world supply. 2.2. C02-Emission Avoidance The plant, designed for an annual production of 0.50 million tons of iron, would according to Eq. (7) lead to the co-production of 42.5 million kmol (952xlO6 m3) of syngas, and the concomitant release of 0.34 million tons CO2. The conventional separate production of iron and syngas results in the emission of 2.20 ton CCVton Fe, and of 0.0015 ton CCVm3 syngas, respectively [2], By such separate production of the above amounts of iron and syngas, the amounts of emitted CO2 would be 1.10 million and 1.43 million tons, respectively, or a total of 2.53 million tons. The extent of CCVemission avoidance by the co-production would thus be 86%. 2.3. Fuel saving In the co-production of iron and syngas (for conversion to methanol) according to Table 1, the total annual fuel requirement for producing 0.50 million tons iron and 0.408 million tons methanol would be 19.5 million GJ. The separate production of Fe in efficient modern plants by the direct reduction process (DRI, operated below the melting point of iron) requires 12.8 GJ per ton of iron [9]. In this case, the production of 0.50 million tons Fe would require 6.4 million GJ. The total fuel consumption of methanol synthesis by conventional steam-methane reforming is 44.5 GJ/ton methanol [10a]. For the production of 0.408 million ton methanol, the total fuel requirement would be 18.16 million GJ. Thus, the separate conventional production of the above amounts of Fe and methanol would require 24.6 million GJ. The fuel saving by the proposed co-production vs. the separate production is therefore 20.6 %.

483 Table 1. Economic assessment for Fe and methanol production, based on the thermo-neutral thermochemical equilibrium achieved at HOOK and 1 bar by the reaction Fe 2 O 3 + 4.87CH 4 + 3.51O2 = 2Fe + 5.74H2 + 3.70CO +3.99H 2 O + 1.17CO2, followed by methanol synthesis. Ton = metric ton.

Design Parameters Annual Fe2O3 feed (ton/yr)

0.72x10"

Annual NG feed (GJ/yr)

19.5xlO6

Annual NG feed (mmbtu/yr)

18.5xlO6

Annual O 2 feed (ton/yr)

0.50xl0 6

Annual Fe production (ton/yr)

0.50xl0 6

Annual methanol production (ton/yr)'

0.408xl0 6

Annual CO 2 release (ton/yr)

0.341xl0 6

Capital Costs (million U.S.$) Equipment and facility2 Interest during construction (10% of investment) Startup expenses and working capital

100 10 10 120

TOTAL

Annual Costs (million U.S.S) Capital cost (15% of above TOTAL) Operation and maintenance (2% of TOTAL)

18.0

2.4

Fe 2 O 3 cost ($23/ton)3

16.6

NGcost ($4.80/mmbtu)4

88.8

O2 cost ($40/ton)

5

20.0 145.8

TOTAL

Annual Sales (million U.S.S) Iron ($130/ton)6 Methanol ($230/ton)

65.0 7

TOTAL

98.9 163.9

1

Assuming 90% yield of conversion of syngas to methanol.

2

Including the reactor, desulfurization, heat recovery, shift reactor, PSA, methanol synthesis reactor, and other related equipment and facility.

3

Price of iron ore at U.S. mines in 2002. See http://minerals.er.usgs.gov/minerals/pub/ commodity.

4

September 2003. Source: International Herald Tribune.

5

See Basye, L. and Swaminathan, S. "Hydrogen Production Costs - A Survey", 1997; Report by SENTECH, Inc. for DOE/GO/101-778.

6 7

Price of pig iron in 2002 in Australia. See: www.doir.wa.gov.au/investments/pdfs/DRI.pdf September 2003. See www.methanex.com/methanol.

484

Table 2. Economic assessment for CaO and methanol production, based on the thermo-neutral thermochemical equilibrium at 1200 K and 1 bar by the reaction CaCO 3 + 3CH4 + 2.55O2 + 3.2H2O = CaO + 4.59H2 + 2.31CO + 4.61H 2 O + 1.69CO2, followed by methanol synthesis. Design Parameters Annual CaCO 3 feed (ton/yr)

0.035xl0 6

Annual NG feed (GJ/yr)

O.935xlO6

Annual NG feed (mmbtu/yr)

0.89xl0 6

Annual O 2 feed (ton/yr)

0.0286xl0 6

Annual CaO production (ton/yr)

0.020xl0 6

Annual methanol production (ton/yr)'

0.0232xl0 6

Annual CO 2 release (ton/yr)

0.026xl0 6

Capital Cost (million U.S.S) Equipment and facility2

2.0

Interest during construction (10% of facility investment

0.2

Start-up expenses & working capital (10%)

0.2

TOTAL

2.4

Annual Cost (million U.S.S) Capital cost (15% of Total)

0.36

Operation and maintenance (2% of Total)

0.05

CaCO 3 ($8.65/ton)

3

0.30

NG cost ($4.80/mmbtu)4

4.27

O2 cost ($40/ton)5

1.14

TOTAL

6.12

Annual Sales (million U.S.S) CaO ($55/ton)3 Methanol ($230/ton)

1.10 6

TOTAL

5.34 6.44

1

Assume 90% yield in conversion of syngas to methanol.

2

Including the lime kiln, PSA, methanol synthesis reactor, and other related equipment and facility.

3

July - December 2003. See: www.ppc.co.za.

4

September 2003. Source: International Herald Tribune.

5

See Basye, L. and Swaminathan, S. "Hydrogen Production Costs", 1997; Report DOE/GO/101-778.

6

September 2003. See www.methanex.com/methanol.

485

3. CO-PRODUCTION OF LIME AND SYNGAS 3.1. Economics The calcination of limestone to lime is the main endothermic step in the lime and cement industries. The CO2 release is derived both from the combustion of fossil fuel for process heat, and from the decomposition reaction itself (Eq. 6). Based on the thermo-neutral thermochemical equilibrium obtained at 1200 K and 1 bar by the reaction [7], CaCO3 + 3CH4 + 2.55O2 + 3.2H2O = CaO + 4.59H2 + 2.31CO + 4.61H2O + 1.69CO2

(8)

and assuming syngas is further processed to methanol via the partial water-gas shift of CO to H 2 , the proposed co-production of lime and methanol should be economically favorable (see Table 2). The expected annual production of 0.0232xl06 tons methanol would add 0.08% to the current world supply. Alternatively, if syngas is further processed to hydrogen via the complete water-gas shift of CO to H2, the predicted H2 production is 4.35xlO6 kg/yr (2.17xlO6 kmol/yr) and, at a reported price of $4.50/kg H2 [11], its annual sale would amount to $19.6xlO6. However, in this case it would be necessary to add the cost of liquifaction of H2. 3.2. CO2 Emission Avoidance By the co-production of 0.020xl06 ton CaO and 2.415xlO6 kmol syngas (or 54.12xlO6 m3) (see Table 2), 0.0262x106 ton CO2 are released. The separate conventional production of lime results in the emission of 1.1 ton CO2/ton CaO [12a]. The production of 0.020x106 ton CaO would thus release 0.022x106 ton CO2. Similarly, the separate conventional production of syngas via methane-steam reforming involves the emission of 0.0015 ton CO2/m3 syngas. Therefore, the conventional production of 54.12x106 m3 syngas would release 0.0812xl06 ton CO2. The total amount of CO2 released by the conventional separate production of these amounts of lime and syngas would be 0.103xl06 ton. The predicted CO2 emission avoidance by the co-production of lime and methanol, relative to the conventional separate production is thus 75%. 3.3. Fuel saving By comparison with the conventional production of CaO in modern lime kilns, in which the fuel requirements are 3.5 GJ per ton of CaO produced [12b], and the separate production of syngas for methanol, in which the total fuel requirements are 44.5 GJ/ton of methanol [10a], the predicted fuel saving by co-production of 0.020x106 ton CaO and 0.026x106 ton syngas amounts to 23.5%. 4. DISCUSSION The proposed processes for the co-production of syngas (for its conversion to methanol or hydrogen) with either iron or lime are based on thermo-neutral reactions in thermochemical equilibria. The main reactions discussed above, i.e. the reduction of Fe2O3 and the calcination of CaCO3, are well-established major industrial processes, which readily proceed to completion at the temperatures used in the above calculations. In the proposed reactions, the released CO2 is consumed in situ by the CO2 reforming reaction. The partial oxidation of a fuel, such as natural gas, by oxygen or air has been reported to operate both with and without added catalysts [10b]. By comparison with the conventional separate production of syngas and of either Fe2O3 reduction or CaCO3 calcination, the proposed co-production predicts fuel

486

saving, substantial CO2 emission avoidance, and attractive economics. As noted above (Sections 2.1 and 3.1), the co-production of methanol with iron or lime in the proposed plants would add only —1.3% or 0.08% to the current world supply of methanol, which presumably could be absorbed by the existing demand. Many other uses for syngas, such as the production of hydrogen, ammonia, and various synthetic fuels should be considered as well. Further work will be required to determine the changes needed in the conventional processes for achieving the proposed co-generation of syngas and iron or lime, and for matching different company cultures of the petrochemical, metallurgical, and lime and cement industries. REFERENCES 1. IPCC, 2001. Intergovernmental Panel on Climate Change, Cambridge University Press. 2. A. Steinfeld and G. Thompson, Energy, 19 (1994) 1077. 3. M. Halmann and M. Steinberg, "Greenhouse Gas Carbon Dioxide Mitigation: Science and Technology", Lewis Publ., Boca Raton, Florida, 1999. 4. M. Halmann, A. Frei, and A. Steinfeld, 6th Internal Conf. On Carbon Dioxide Utilization, Breckenridge, CO., USA, Sept. 2001; Abstract PO-16. 5. M. Halmann, A. Frei, and A. Steinfeld, Energy, 27 (2002) 1069. 6. M. Halmann and A. Steinfeld, in "Greenhouse Gas Control Technologies", J. Gale and K. Kaya (eds)., Elsevier Science Ltd., 2003, p. 1013. 7. M. Halmann and A. Steinfeld, Energy and Fuels, 17 (2003) 774. 8. See www.methanol.org. 9. (a) M. Ottow et al., in Ullmann's Encyclopedia of Industrial Chemistry, Vol. A14, VCH Publ., 1989, p. 516. (b) See: www.doir.wa.gov.au/investments/pdfs/DRI.pdf. 10. H. Hiller et al., in Ullmann's Encyclopedia of Industrial Chemistry, Vol. A12, VCH Publ., 1989; (a) p. 289; (b) p. 183. 11. www.eere.energv.gov/hvdrogenandfuelcells/pdfs/gronich garbak_ee_tech_val.pdf. 12. (a) A. Thoma, Ecological Consideration of the Solar Production of Lime, (in German), Semesterarbeit, ETH-Zurich, 2002. (b) See: www.qualical.com.


487

Greenhouse Gas Hydrates in the Ocean S.M. Masutani University of Hawai'i, Hawai'i Natural Energy Institute 1680 East-West Road, POST 109, Honolulu, HI 96822 USA; [email protected] CH4 and CO2 hydrates occur in the deep ocean and sea floor sediment. CH4 hydrates are found in enormous quantities along all the continental margins; CO2 hydrates appear to be a transient anomaly. The magnitude of the carbon pool locked in the CH4 hydrate reservoir is not known precisely, but estimates range from 103 to 106 Gt C. Opening of this reservoir, either purposefully to extract the energy or inadvertently, will release the sequestered carbon into the environment, increasing radiative forcing and exacerbating the ongoing global warming trend. Data from the historical record and recent modeling results suggest that hydrate destabilization and CH4 outgassing may have catalyzed a number of runaway global warming episodes in the past. A careful reassessment of the role of CH4 in various climate change scenarios therefore appears warranted as well as an evaluation the vulnerability of the hydrate reservoir to unintentional destabilization. Commercial applications for CH4 and CO2 hydrates are discussed briefly in this paper and related research opportunities are identified. 1. INTRODUCTION Radiative forcing due to the accumulation of greenhouse gases in the atmosphere may profoundly impact global climate. The historical record indicates a clear correlation between global temperatures and atmospheric concentrations of carbon dioxide (CO2) and methane (CH4) [1], two primary greenhouse gases. Since the beginning of the industrial revolution, human activities have resulted in more than a doubling of CH4 levels and about a 30% increase in CO2 [1,2], with corresponding direct radiative forcing (i.e., change in net radiative flux) of approximately 0.5 and 1.6 W/m2, respectively. Although the global warming potential of CH4 per molecule over a 100 year time horizon is about 21 times that of CO2, recent attention has been focused on CO2 due to the enormous volume of anthropogenic emissions of this gas into the atmosphere, largely from the oxidation of fossil fuels, and its order-of-magnitude longer lifetime (i.e., approximately 120 years compared to about 10 years for CH4). Under most scenarios, it has been estimated that over the next century, anthropogenic CO2 will account for 61% of the total greenhouse gas radiative forcing, or more than four times the contribution of CH4 [1]. The presumed lesser role of methane in precipitating changes to the present global climate is being reassessed as the result of the emergence of information about the huge quantities of CH4 currently trapped in hydrates in ocean sediments and permafrost and new field and modeling studies that appear to link the extreme global warming episode known as the Paleocene-Eocene. Thermal Maximum (PETM), that occurred 55.5 million years ago to a large and sustained eruption of methane gas from destabilized methane hydrates [3,4]. These new data have lent support to the hypothesis that the now-recognized massive methane hydrate reservoir could catalyze a number of postulated runaway global warming scenarios.

488

Opinion remains divided, however, as a number of earlier studies have downplayed the significance of hydrate destabilization on climate change relative to CO2 emissions [5-7]. 2. MARINE HYDRATES AND CLIMATE CHANGE Hydrates are crystalline solids comprising water molecules linked by hydrogen bonds in a tight polyhedral cage structure. Guest molecules, which can include a host of species such as N2, CO2, CI2, halocarbons, and various paraffin hydrocarbons, reside in the interstices of this lattice. The temperature at which solidification of hydrates occurs can be above the freezing point of water and increases with pressure. Both CH4 and CO2 hydrates occur naturally in the deep ocean and seafloor sediments [6,8,9], where high pressures and low temperatures prevail. Natural CO2 hydrates appear to be something of a transient anomaly, while CH4 hydrates exist in huge quantities at many locales on the world's continental margins. This difference probably reflects the disparity in sources of concentrated CO2 and CH4 in the deep ocean and also the stability characteristics of the two hydrates in that environment. CH4 and CO2 form body-centered cubic structure I hydrates in proportions of slightly less than 6 molecules of H2O per molecule of CH4 or CO2. Since the solubility of these non-polar gases in water or sea water is of the order of parts per thousand or less, hydrate formation will not occur unless conditions exist that provide a means to exceed the solubility limit and allow sufficient quantities of the guest molecule to contact the water phase. While the mechanism(s) of in situ hydrate formation remains unknown, several models postulate either hydrate formation on the interface between free gas volumes (or gas adsorbed by solid sediment) and water; or in a single, supersaturated liquid phase [5]. Methane and carbon dioxide in natural hydrates can originate from either biological or thermal sources. Thermogenic gases evolve over long periods of time when organic matter is buried and transported by geological events deep into the earth, where it is chemically converted at high temperatures and pressures. These gases penetrate into the seafloor sediment layer through channels and faults and continue to migrate upward along the geothermal gradient into progressively colder zones. Biogenic CH4 and CO2 are products of microbial metabolism. CH4 is generated by methanogenic bacteria (or, more generally, a consortium of microbes), a major division of the archea kingdom, primarily via anaerobic acetate dissimilation and CO2 reduction. While there is considerable uncertainty about the size of the biogenic methane source, it has been estimated that between 10 and 10 5 g CH4 are produced annually by methanogenic bacteria worldwide. The portion of this inventory originating from ocean sediments is unknown, although IPCC estimates of the annual methane flux from the oceans into the atmosphere stand at approximately 1013 g [1]. Significant oxidation of CH4 probably occurs as it moves upward from the sediment through the water column; the amount reaching the atmosphere is a small fraction of the biogenic methane released from the seafloor. Recent studies have begun to reveal the magnitude of the marine methane hydrate reservoir. Estimates of the total volume of methane and other low molecular weight hydrocarbons locked in hydrate deposits worldwide vary widely, from about 105 trillion standard cubic feet (TCF) to 2.7 x 108 TCF (i.e., 2.8 x 1015 to 7.6 x 1018 standard cubic meters, or approximately 103 to 106 gigatonnes of carbon). This broad range reflects the difficulty in extrapolating from the limited pool of field assay data. Even at the lower end of this range, however, it is clear that methane hydrates in ocean sediments (and permafrost) represent a significant potential future energy resource, provided that practicable and environmentally acceptable recovery techniques can be devised. To put things in perspective, in 2001, world energy consumption,

489

including fossil fuels, nuclear, and renewables, reached 404 Quads, or 1.18 x 1014 kWh = 4.26 x 1017 kJ. The total higher heating value (evaluated at 25°C) of 105 TCF methane (i.e., the smallest estimated amount of CH4 stored in hydrates) is 1 x 1020 kJ, or about 240 times the current annual world energy use. On the other hand, recovery and oxidation of the CH4 sequestered in hydrates for energy production could generate emissions in excess of 103 gigatonnes of carbon (in the form of CO2), exacerbating the global warming problem. Figure 1 presents some preliminary results of a simple box diffusion model [10,11] to assess the effects of methane hydrate utilization on the atmospheric CO2 inventory and associated impacts on averaged ocean mixed layer temperatures. The model calculations consider two scenarios in which a carbon pool from methane hydrates of equal magnitude to all remaining recoverable coal, oil, and natural gas is oxidized over two different time periods and product CO2 is released into the atmosphere. A worst-case, "fast consumption" scenario, which presumes a continuation of the historical trend of exponential growth in fossil energy use has been assumed; adoption of policies to control greenhouse gas emissions, would reduce peak levels of CO2 significantly from the values shown. Although box diffusion models do not consider many important factors included in more complex General Circulation Models, they have frequently been applied to perform parametric studies and to gain general insight into the effect of different fossil fuel consumption scenarios and sequestration strategies on atmospheric CO2 inventories over time. The results shown in Figure 1 underscore the potential peril associated with liberating the carbon stored in the methane hydrate pool either purposefully, or inadvertently. Over the past 15 years, CO2 sequestration in the ocean, underground, or via biological pathways has been investigated as a means to reduce atmospheric emissions of this greenhouse gas. A number of these studies have explored CO2 hydrates as a vehicle to accomplish this goal [12,13]. With regard to direct ocean sequestration, it was originally hoped that discharged CO2 could be stably bound in the solid hydrate phase, resulting in longterm storage with minimal impacts to the deep ocean environment. Research has indicated, however, that CO2 hydrates decompose readily when exposed to sea water that is not

Fig. 1. Box diffusion model predictions of the impact of CH4 hydrate utilization for energy on atmospheric CO2 concentrations and ocean mixed layer temperature perturbation. Solid line is the baseline no-hydrate utilization case; dashed and dotted lines correspond to two slightly different hydrate consumption rate scenarios with identical total carbon release.

490

saturated with dissolved CO2 [14,15]. It has been suggested that dissolution of CO2 hydrates may be prevented or forestalled if they are buried in sea floor sediment. This concept has led to proposals to mine CH4 hydrates for energy and replace them in the sediment with CO2 (or mixed gas) hydrates formed from CO2 separated from the products of oxidation of the recovered methane [16]. While this would represent a no- or low-greenhouse-gas-emissions energy system, there are significant technical and economic questions that must be addressed to assess whether this concept of methane hydrate extraction and replacement is feasible. Furthermore, the impact of wholesale replacement of CH4 with CO2 on the microbial assemblage and biogeochemical processes occurring in the sea floor sediment (e.g., acidification of pore water) needs to be elucidated to ensure that benefits outweigh risks. Exploitation of the methane hydrate reservoir for energy will increase radiative forcing and exacerbate the global warming trend unless environmentally acceptable and practicable means to sequester the liberated carbon from the atmosphere are developed and applied in parallel. Even if the CH4 hydrate carbon pool is not intentionally broached, hydrate destabilization and methane release could occur inadvertently as a consequence of global wanning (note, however, that hydrate destabilization is endothermic and, therefore, self-modulating: destabilization results in local cooling which limits additional hydrate decomposition; the heat of dissociation of CH4 hydrate is approximately 55 x 103 kJ/kmol CH4 released, which means that liberation of 1 kg CH4 from the hydrate phase will result in a 1°C decrease in temperature of almost 1 m3 of sea water). A portion of the global hydrate reserves exists at the limit of stability. There is evidence that hydrates in sea floor sediment may undergo a continuous series of formation and decomposition events as methane migrates upward and the hydrate stability zone shifts due to progressive deposition and burial of detritus from the water column [5]. Modest warming of the deep ocean water masses (or, with regard to hydrates in permafrost, the land) could destabilize some of the hydrates and trigger a release of CH4. There is a growing body of evidence from the historical record that supports the hypothesis that massive emissions of methane from destabilized hydrates and high latitude gas fields precipitated rapid climate changes at the end of the last glaciation (ca. 10,500 and 13,000 years ago) [17,18]. Further back in time, the 100,000 year long episode of extreme global warming known as the PETM, in which the average global temperature increased by about 7°C, may have been a consequence of marine CH4 hydrate destabilization [19,20]. Falling sea levels, tectonic events, and deep ocean water warming due to a switch in deep water formation from high to low latitudes, have been postulated as the reason behind this destabilization [21,22]. Sediment cores data and other information are being analyzed to understand the genesis and consequences of the PETM. Recent modeling studies have shed new light on the atmospheric chemistry and radiative forcing associated with large and/or prolonged emissions of CH4 from decomposing methane hydrates [3]. Those results indicate that the emission of about 1,500 gigatonnes carbon as CH4 could explain the PETM climate perturbations. Furthermore, it appears that large releases of CH4 into the atmosphere depletes the tropospheric OH chemical sink for this species, increasing its lifetime significantly and providing for the formation of stratospheric water vapor (another greenhouse gas) as the unreacted methane penetrates to higher altitudes where it is oxidized. Elevated levels of CH4 and stratospheric water vapor were found to account for more of the peak radiative forcing than increases in atmospheric CO2 arising from oxidation of the methane [3]. While debate continues on whether massive methane emissions trigger rapid global climate changes or are simply a consequence that adds to a list of maladies associated with a global wanning trend due to natural or anthropogenic causes, the historical evidence indicates that the implications are significant and that they cannot be ignored. The integrity of the vast

491

greenhouse gas methane pool that exists as marine and permafrost hydrates clearly warrants concern and additional study. 3. HYDRATE APPLICATIONS AND RESEARCH OPPORTUNITIES A number of practical applications for hydrates have been explored over the years and include: (1) desalination; (2) gas storage/transport; (3) cold energy storage; (4) contaminant separation or chemical recovery; and (5) bioprocess engineering and biotechnology [5]. These applications exploit the ability of hydrates to exclude non-hydrate formers from the clathrate cage-effectively filtering an aqueous solution (applications 1,4, and 5); the relatively high density of trapped guest molecules (application 2); or the large heat of fusion (between 0.3 and 0.5 kJ/gm hydrate; heat of fusion of ice is 0.33 kJ/gm ice) of the solid (application 3). Halocarbons are probably better suited for desalination, although CH4 and CO2 hydrates were investigated for this purpose in the past and patents exist for CO2 and CH4 hydrate desalination processes. Hydrate desalination R&D was pursued aggressively between the early 1960's until the late 1970's. Activity dropped off thereafter and effort shifted to reverse osmosis and similar technologies. Recently, there has been a resurgence of interest in hydrate desalination, for example, the work at the Naval Research Laboratory (using methane mixed gas hydrates) [23] and the operation of a pilot facility in Hawaii employing R141b. The Hawaii facility was able to reduce total dissolved solids (TDS) in sea water from 34,500 ppm to less than 500 ppm. The operators indicated that, based on small scale laboratory tests, improvements in the hydrate slurry separation and washing processes could lower TDS to less than 50 ppm. Their economic analysis suggested that a commercial-scale facility (13,600 m3/day) could produce fresh water at a competitive cost of U.S.$2 per 1,000 gallons (3.8 m3). The bulk transport of CH4 and natural gas mixtures as solid hydrates has received attention periodically over the past half century. Although the energy density of CH4 hydrates is less than half that of LNG (6.9 x 106 kJ/m3 and 1.6 x 107 kJ/m3, respectively), they can be formed and stored at considerably higher temperatures. The water remaining after recovery of the fuel gas is an attractive byproduct for specialized transport scenarios such as space travel. Hydrate formation for desalination or fuel transport requires energy for pressurization and refrigeration. Some of this energy may be recovered and applied to the process when the hydrate is eventually destabilized to release the gas or fresh water, or used for off-process purposes such as air conditioning. Dedicated cold energy storage systems have been proposed [24] which use off-peak nighttime electricity to generate hydrates that are applied for refrigeration and air conditioning use during the day. Although the enthalpies of fusion of hydrates and ice are similar, hydrates can be formed at higher temperatures than ice which may offer certain cycle performance benefits. As in the case of desalination, halocarbons probably are a better choice than CH4 or CO2 for cold energy hydrate storage applications. While specialized circumstances exist where CO2 or CH4 hydrates may be employed for chemical or contaminant removal and recovery (e.g., de-watering), or bioprocessing, other hydrates are probably more relevant to these areas and are expected to constitute the focus of associated R&D efforts. Although greenhouse gas hydrates have the potential to be utilized commercially for a variety of practical ends, none of the applications discussed above have any direct impact on climate change. CH4 hydrates could, however, profoundly influence global climate if they are exploited for energy generation. CO2 or mixed gas hydrates, on the other hand, may provide a means to help stabilize radiative forcing by sequestering carbon in ocean sediments. Large scale recovery of CH4 for energy from marine methane hydrate deposits will not be realized for some years to come-major national R&D programs in Japan, India, and the U.S.

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notwithstanding-due to significant technical and economic challenges and uncertainties. For one thing, the major energy companies currently have limited interest and economic incentive to develop the methane hydrate resource, given that conventional oil, gas, and coal reserves will continue to be viable for another at least several decades [25]. Moreover, in situ destabilization of the hydrates (since removal of large volumes of sea floor material is probably not feasible nor environmentally acceptable) and collection of the released gas is not trivial. The hydrate resource may be distributed widely as fine crystals dispersed in heavy clastic sediments with problematic flow characteristics. Devising an effective gas recovery system will be a primary engineering research priority over the near-term as will be determining the associated impacts on the marine environment. Methane hydrates and methane seeps could be utilized within a 5 to 10 year time horizon for certain specialized subsea power applications. There is considerable interest by the U.S. Navy to develop CH4 fuel cells for a host of underwater missions ranging from long-term, low-power, monitoring and surveillance instruments moored on the deep ocean floor to propulsion systems for extended deployment AUVs (autonomous underwater vehicles) and other submersibles, as well as underwater fuel processing/storage/re-supply stations for these submersibles. The Naval Research Laboratory and the University of Hawaii have been collaborating on subsea power development for the past three years as part of a research program supported by the U.S. Office of Naval Research [26]. Figure 2 shows the concept for a biological fuel cell [27] employing syntrophic colonies of sulfate reducing and methanotrophic bacteria and a photograph of cells currently being tested in our laboratories. The prototype cells use graphite electrodes and contain sediment collected from hydrate fields on the Blake Ridge, Cascadia Margin, or the Gulf of Mexico, and synthetic sea water. The anode is immersed in the sediment and the cathode is placed in the overlying sea water. Parameters being investigated in this low-pressure facility and in a laboratory deep ocean simulator (i.e., pressure chamber) include circuit impedance, electrode placement, and CH4 and dissolved O2 levels. Molecular biological analyses and culruring of the collected sediments, as well as biofilm studies, are also being performed. These low output fuel cells are intended to serve as a long-life replacement for the battery packs that currently power sea floor-mounted instruments. The target power level is about 1 W. Utilization of methane hydrates for submersible propulsion and other high power level uses provides a host of research opportunities. While some solid oxide fuel cells (SOFC) have the capability to thermochemically convert CH4 or utilize it directly, the high operating temperatures of SOFCs may present a number of problems for deep subsea deployment. CH4 may alternatively need to be catalytically converted to H2, or consumed to provide power to electrolyze water to O2 and H2, for use in low temperature Proton Exchange Membrane (PEM) fuel cells. Development of high-output subsea power systems therefore demands R&D on moderate-volume methane extraction from hydrates in the sediment and recovery of O2 dissolved in the water column; identification, adaptation, and optimization of suitable power generation modules; and possibly in situ H2 production and storage. CO2 hydrate storage in deep marine sediments may reduce the impact of broaching the methane hydrate reservoir for energy. Research is needed to determine if this concept is technically, economically, and environmentally viable. Two primary issues are to definitively resolve the question of whether the CO2 hydrate will remain intact (over a reasonable period of time) or decompose in the sediment as it does in the unsaturated water column, and to understand the changes in microbial diversity and sediment biogeochemistry arising from large scale CO2 hydrate sequestration. Another point of interest is the possibility of utilizing CO2 mixed gas hydrates, instead of pure CO2 hydrates, to enhance stability.

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Fig. 2.

Methane biological subsea fuel cell concept and photograph.

Finally, as noted in the preceding section, the role of methane emissions from CH4 hydrates in catalyzing or exacerbating runaway global warming deserves careful study along with assessing the vulnerability of the hydrate reservoir to unintentional destabilization under the ongoing warming trend. 4. SUMMARY The oceans contain vast quantities of greenhouse gas CH4 and CO2. The oceans currently serve as a net sink for atmospheric CO2 and as a source of CH4. Hydrates of these two species exist in the deep ocean and sea floor sediment. While pure CO2 hydrates appear to be a transient anomaly, CH4 hydrates are relatively stable and are found in large quantities along all the continental margins. The magnitude of the carbon pool locked in the methane hydrate reservoir is not known precisely, but current estimates range from 103 to 106 gigatonnes C. The energy stored in methane hydrates is believed to exceed by a significant margin the energy in all known conventional fossil fuels (i.e, coal, oil, and natural gas). Commercial exploitation of methane hydrates for energy will broach the heretofore sequestered carbon pool, increasing radiative forcing and exacerbating the ongoing global warming trend. Industrial production of CH4 from hydrate reservoirs is of considerable interest to nations lacking reserves of other fossil fuels and whose energy security and financial imperatives are different from the U.S. and the multinational oil companies. Significant technical and economic hurdles remain, however, that will probably delay large scale recovery of marine methane hydrates for at least the better part of a generation. The hydrate carbon pool also stands at risk of being inadvertently opened via destabilization due to changes in global climate, e.g., warming of deep ocean water. The historical record provides evidence that hydrate destabilization and CH4 outgassing may have catalyzed, or at least strongly reinforced, a number of past runaway global warming episodes. New field data and recent modeling results about the PETM indicate that it may be worthwhile to reassess the role of methane-and, specifically, methane emissions from marine and permafrost hydrates- in various climate change scenarios. While CH4 is recognized as a major con-

494 tributor to global warming, it currently is assigned a secondary role to CO2. The vulnerability of the hydrate reservoir to unintentional destabilization also needs to be carefully evaluated. Various commercial applications for CH4 and CO2 hydrates were discussed briefly and include desalination; gas storage and transport; cold energy storage; contaminant separation or chemical recovery; and bioprocess engineering and biotechnology. Although presenting interesting R&D opportunities, none of these applications provide a means to bind and isolate CH4 and CO2 or to significantly displace other fossil fuels or greenhouse gases and, hence, are not relevant to mitigating climate change. Significant research also is needed effectively and responsibly to utilize methane hydrates for energy. This will, however, negatively impact the global warming problem unless strategies to sequester the released carbon, such as storage of replacement CO2 or mixed gas hydrates in the sediment, can be developed. REFERENCES 1. IPCC, Climate Change 2001: The Scientific Basis, Cambridge Univeristy Press, New York, 2001. 2. Etheridge, D.M., L.P. Steele, R.J. Francey, & R.L. Langerfields, J. Geophys. Res., 103(D13), 15,979-15,993, 1998. 3. Schmidt, G.A. & D.T. Shindell, Paleoceanography, 18(1), 1004, 2003. 4. Dickens, G.R., J.R. O'Neill, D.K. Rea, & R.M. Owen, Paleoceanography, 10, 965-971, 1995. 5. Sloan, E.D., Jr., Clathrate Hydrates of Natural Gases, 2nd Ed., Marcel Dekker, New York, 1998. 6. Kvenvolden, K.A., G.D. Ginsberg, & V.A. Soloviev, Geo-Marine Letters, 13(1), 323, 1993. 7. Harvey, L.D.D. & Z. Huang, J. Geophys. Res., 100(D2), 2905-2926, 1995. 8. Sakai, H., T. Gamo, E.S. Kim, M. Tsutsumi, T. Tanaka, J. Ishibashi, H. Wakita, M., Yamamoto, & T. Oomori, Science, 248, 1093-1096, 1990. 9. Kvenvolden, K.A., Natural Gas Hydrate in Oceanic and Permafrost Environments (M.D. Max, ed.), Chapter 2, Kluwer Academic Publishers, Dordrecht, 2000. 10. Nihous, G.C., S.M. Masutani, L.A. Vega, & C M . Kinoshita, Climatic Change, 27, 225-244, 1994. 11. Nihous, G.C. & S.M. Masutani, Ocean Storage ofCO2: Ocean Circulation (W. Omerod, ed.), pp. 35-52, IEA Greenhouse Gas R&D Programme, Cheltenham, 1996. 12. Aya, I., K. Yamane, & R. Kojima, GHGT5 Greenhouse Gas Control Technologies (D. Williams, B. Durie, P. McMullan, C. Paulson, & A. Smith, eds.), 423-428, CSIRO, Collingwood, 2001. 13. Takano, S., A. Yamasaki, K. Ogasawara, F. Kiyono, M. Fujii, & Y. Yanagisawa, GHGT6 Greenhouse Gas Control Technologies (J. Gale & Y. Kaya, eds.), 843-848, Elsevier, Oxford, 2003. 14. Aya, I., Direct Ocean Disposal of Carbon Dioxide (N. Handa & T. Ohsumi, eds.), pp. 239-252, TERRAPUB, Tokyo, 1995. 15. Teng, H., C M . Kinoshita, & S.M. Masutani, Chem. Eng. Set, 50(4), 559-564, 1995. 16. Ohgaki, K., Proc. Deep Sea and CO2 2000, Ministry of Transport, Government of Japan, www.nmri.go.jp/co2/procj.htin, 2000. 17. Nisbet, E.G., Nature, 347, 23, 1990. 18. Nisbet, E.G., Can. J. Earth Sci., 27, 148-157, 1990. 19. Sloan, L.C., J.C.G. Walker, T.C. Moore, D.K. Rea, & J.C. Zachos, Nature, 357, 320-322, 1992. 20. Bains, S, R.M. Corfield, & R.D. Norris, Science, 285, 724-727, 1999. 21. Bice, K.L. & J. Marotzke, Paleoceanography, 17(2), 1018, 2002. 22. Haq, B.U., Natural Gas Hydrate in Oceanic and Permafrost Environments (M.D. Max, ed.), Chapter 11, Kluwer Academic Publishers, Dordrecht, 2000. 23. Max. M.D. & R E . Pellenberg, U.S. Patent 5,873,262, 1999. 24. Mori, Y.H. & T. Mori, AIChE Journal, 35, 1227, 1989. 25. Hawa'i Natural Energy Institute, First Workshop of the International Committee on Methane Hydrates, Final Report to the U.S. Office of Naval Research-International Field Office, 2001. 26. Masutani, S.M. & R.B. Coffin, Proc. 5lh International Symposium on CO2 Fixation and Efficient Utilization of Energy, pp. 78-82, Tokyo Institute of Technology, Tokyo, Japan, 2002. 27. DeLong, E.F. & P. Chandler, Nat. Biotechnol, 20, 788-789, 2002.


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Replacement of Methane Hydrate by Carbon Dioxide: 13C NMR Study for Studying a Limit to the Degree of Substitution Huen Lee al , Youngwon Seoa, Yu-Taek Seoa, Do Youn Kima, Igor L. Moudrakovskib, and John A. Ripmeester Department of Chemical and Biomolecular Engineering, Korea Advanced Institute of Science and Technology (KAIST), 373-1, Guseong-dong, Yuseong-gu, Daejeon 305-701, Korea Steacie Institute for Molecular Sciences, National Research Council Canada, 100 Sussex Drive, Ottawa, Ontario, Canada, K1A0R6 E-mail: [email protected] The results presented here show that, at least in a laboratory setting, the replacement of CH4 by CO2 is favored from the points of view of equilibrium thermodynamics. Pure and mixed-phase hydrates were prepared in porous silica gel in order to observe the guest distribution in hydrate cages by ' C NMR spectroscopy. The resulting ratio values for CH4 in small and large cages strongly suggested that at least 64% of methane should be recoverable from a hydrate of composition CH4«6.05H2O. 1. INTRODUCTION The discovery of vast hydrocarbon resources in hydrate form on the continental margins and in permafrost regions has sparked an interest in the possible recovery of relatively clean-burning methane gas. At the same time, it has been proposed that CO2, obtained as a by-product of combustion processes, possibly could be stored either in the deep ocean as liquid, or as solid hydrate, in order to reduce the release greenhouse gas into the atmosphere. As such, the conversion of methane hydrate to CO2 hydrate with the net recovery of methane seems quite attractive. The limiting equilibrium composition of the mixed hydrate then become of primary interest in assessing this potential technology. In this contribution we address this issue, which have been considered so far only from a macroscopic point of view, by following the distribution of guests over different cages using solid-state NMR methods. 2. EXPERIMENTAL For composition studies, pure and mixed-phase hydrates were prepared in porous silica gel (Aldrich, pore size 15 nm) in order to facilitate the rapid equilibration of hydrates. 13C NMR spectra were recorded at 200 K by placing the hydrate samples in a 4 mm o.d. Zr-rotor that was loaded into the variable temperature (VT) probe on a Bruker 400 MHz solid-state NMR spectrometer. All spectra were recorded at a Larmor frequency of 100.6 MHz with magic angle spinning (MAS) at about 2-4 kHz under high-power proton decoupling (HPDEC). A pulse length of 2 (is and pulse repetition delay of 20 s with radio-frequency field strengths of 50 kHz corresponding to 90° pulse of 5 (is duration were used. The downfield carbon resonance peak of adamantine, assigned a chemical shift of 38.3 ppm at 300 K, was used as

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an external chemical shift reference. For the mixed hydrate samples of CO2 compositions over 80 mol %, 13CH4 gas was used to obtain higher intensity CH4 peaks. Kinetic studies of the reaction of powdered solids ice or hydrate, particle size 5 - 5 0 /an) with gases were carried out in a high pressure NMR cell connected to a gas handling system. For CO2, the reactions were followed by measuring the integrated C powder pattern, characteristic of the guest in the large cage of str. I hydrate, as a function of time. The variation of methane content in the hydrate with time was obtained from the integrated [ H NMR signal. The absolute amount of product was obtained by comparing the signal with that of known samples. 3. RESULTS AND DISCUSSION In order to observe favourable exchange between CO2 and methane hydrate, there must be preferential partitioning of CO2 and CH4 between the gas and the hydrate solid phases. The only way this can arise is if CO2 has a preference for the large cage (51262) in the hydrate, as the large cages outnumber the smaller by a factor of three. This premise was investigated by examining the distribution of methane over the two cage sites by MAS 13C NMR for hydrate samples prepared from gas mixtures (Fig. 1).

Figure 1. 13C MAS NMR spectra of an si hydrate prepared using CH4/CO2 mixtures with compositions a s indicated. One can see that for pure methane hydrate L-CH*I S-Cff< is ~ 1.26 so that the small cage is occupied to a smaller degree than the larger cage, as reported before for synthetic and natural methane hydrates. With increasing CO2 in the gas mixture, the ratio declines steadily to a value that shows that now fewer than half as many large as small cages are filled with methane. This behavior can be explained in terms of the molecular sizes of CO2 and CH4. The size of CO2 is almost the same as the cavity diameter of the small 512 cage in si hydrate,

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the molecular diameter to cavity diameter ratio being 1.00 for CO2 and 0.855 for CH4. For CO2 hydrate, NMR and diffraction techniques have indicated that the 512 cages have smaller fractional occupancies than the large cages. Hydration numbers of- 6.0 for methane and 6.2 for CO2 hydrate are consistent with fractional occupancies of the small cages 0S of ~ 0.8 and ~ 0.7, respectively. Thus, when CO2 competes with CH4 in occupying the small cage of si 0 10 hydrate, CO2 is a relatively poorer guest, so that the occupancy ratio for CH4 ( L-CH2-rich conditions; the ratio of the flow rate of CO2 to that of water is around the hydrate stoichiometry. In such conditions, liquid CO2 drops formed in the pre-mixing unit would be divided into smaller drops in the static mixer. Upon division, new hydrate film would be formed quickly on the surface of the liquid CO2 drops. At the same time, the shear stress by the water flow would shed the hydrate film from the surface intermittently.

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Such a shedding of CO2 hydrate from the interface of liquid CO2 and water has been reported in a stirring vessel as well as in a tunnel tube type reactor under flow conditions [20]. Due to the higher number density of liquid CO2 drops and the slow velocity of water flow, the shed hydrate films and remained liquid CO2 drops would be agglomerated each other to form a hydrate chunk. For the water rich conditions, division of liquid CO2 drops and shedding of the surface hydrate film would occur in the static mixer as shown in Figure 4-b. However, no agglomeration would occur among the shed hydrate films and liquid CO2 drops due to higher velocity of water flow and lower number density of liquid CO2 drops. Instead, dispersed flow of liquid CO2 drops and tiny hydrate particles would be observed in the water-rich conditions, as shown in Figure 3-c. Thus, the observed results of the hydrate formation in the static mixer could be explained based on the following mechanisms; division of the liquid CO2 drops and shedding of CO2 hydrate film in the mixer, and agglomeration of the liquid CO2 drops and the shed hydrates. 3.2. Energy consumption of the static mixer Energy consumption is one of the main concerns for the application of the static mixer to the hydrate formation process. In this section, estimation of the energy consumption was conducted based on the present experimental conditions. Assuming the temperature conditions and pressure conditions of the fluids were prepared for the hydrate formation, we focused on the energy consumption for the hydrate formation in the static mixer; the energy consumption for the pumping of the fluid in the mixer. The pumping energy for the fluid into a pipe, £, [W/kg-total flow] under a turbulent flow condition is given by Eq. (1),

QAP Sy=^—

Vp

AP =

(1),

xp

where AP is the pressure drop along the pipe and Q [m3/sec] is the flow rate, and V is the inner volume of the pipe, and r[s] is the residence time of the fluid in the pipe. For the Kenics type static mixer, the pressure drop can be given by the following equation;

where fSM [-] is the friction coefficient of the mixer, D is the inner diameter of the mixer, L is the mixer length, p [kg/m3] is the density of the fluid, and u is the average fluid velocity [m/s]. The friction coefficient of the Kenics-type static mixer is given by the equations as an empirical function of Reynolds number, Re [19]. The energy consumption for the pumping into the static mixer was calculated assuming the following condition based on the present experimental results for the case of dispersed hydrate formation; water flow rate = 2,980 mL/min, CO2 flow rate = 93.4 mL/min. It is assumed that the fluid properties of two-phase flow could be replaced by those of water. The pressure drop for this case is, AP = 14.1 kPa, and the energy consumption is, &, = 45.3 W/kg-total flow, equivalent to the energy consumption per unit mass of CO2 flow = 1.66 kW7kg-CC>2 flow. Since the residence time of liquid CO2 in the static mixer is 0.31 s in this case, the energy consumption per 1 kg of

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CO2 is 0.52 kJ/kg-CC>2 assuming the conversion to hydrate in the mixer is 100 %. This result is about half of the case for the stirring vessel, of which the energy consumption for the hydrate formation (agitation energy) is about 1.1 kJ/kg-CO2 [14] based on a laboratory-scale experiments. From the above estimation, the application of the static mixer to the hydrate formation would have a large advantage in terms of energy consumption over the stirring vessel reactor, and a continuous hydrate formation process could be constructed by using the Kenics-type static mixer as a hydrate formation device with lower energy consumption. 4. CONCLUSIONS Various patterns of continuous CO2 hydrate formation was observed by the Kenics-type motionless mixer, and the formation patterns depended on the flow velocities of liquid CO2 and water. The process energy consumption was estimated to be much lower than that for a process by a stirring vessel type reactor. Thus, the use of motionless mixer is potentially a viable option for hydrate formation process in various applications. REFERENCES 1. T. Takaoki, T. Iwasaki, Y. Katoh, T. Arai, K. Horiguchi, Proc. 4th Int. Conf. Gas Hydrates (2002) 982 2. Y. Nakajima, T. Takaoki, K. Ohgaki, S. Ota, Proc. 4th Int. Conf. Gas Hydrates (2002) 987 3. S. Ota, H. Uetani, H. Kawano, Proc.4th Int. Conf. Gas Hydrates (2002) 991 4. K. Ohgaki, K. Takano, M. Moritoki, Kagaku Kogaku Ronbunshu (in Japanese), 20 (1994) 121 5. A. A. Khokhar, J. S. Gudmundsson, E. D. Sloan, Fluid Phase Equilibria 150-151 (1998) 383 6. J. -H. Yoon, H. Lee, AIChE J., 43 (1997) 1884 7. S. -P. Kang, H. Lee, Environ. Sci. Technol, 34 (2000) 4397 8. Y. Seo, H. Lee, Environ. Sci. Technol., 35 (2001) 3386 9. A. L. Ballard, E. D. Sloan, Proc. 4th Int. Conf. Gas Hydrates, Vol. 2 (2002) 1007 10. T. Austvik, K. P. Lziken, Energy Conv. Mgmt., 33 (1992) 659 11. T. Austvik, K. P. Laken, Energy Conv. Mgmt, 34 (1993) 1081 12. A. Saji, H. Toshihara, H. Sakai, T. Tani, T. Kamata, Energy Conv. Mgmt., 33 (1992) 643 13. A. Saji, H. Noda, Y. Takamura, T. Tani, T. Tanaka, H. Kitamura, T. Kamata, Energy Conv. Mgmt, 36 (1996) 493 14. A. Yamasaki, M. Wakatsuki, H. Teng, Y. Yamagisawa, K. Yamada, Energy, 25 (2000) 85 15. W. G. Knox, M. Hess, G. E. Jones, H. B. Smith, Chem. Eng. Prog, 57 (1961) 66 16. H. Kubota, K. Shimizu, Y. Tanaka, T. J. Makita, J. Chem. Eng. Jpn, 17 (1984) 423 17. J. Javanmardi, M. Moshfeghian, App. Therm. Eng, 23 (2003) 845 18. N. Harnby, M. F. Edwards, A. W. Nienow, Mixing in the Process Industries (2nd ed.), Butterworth-Heinemann (1992) 19. Noritake Co. Ltd, Basic Technology of Static Mixers (in Japanese), Technical Report, No.20 (2000) 20. H. Teng, A. Yamasaki, Energy&Fuels, 13 (1999) 624


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A New Type Separation Process of Condensable Greenhouse Gases by the Formation of Clathrate Hydrates K. Shiojiri1, T. Okano1, Y. Yanagisawa1, M. Fujii2, A. Yamasaki3, H. Tajima3, and F. Kiyono3 School of Frontier Sciences, The University of Tokyo, Tokyo, 113-8656, Japan. 2

Research Center for Material Cycles and Waste Management, NIES, Tsukuba, 305-8506, Japan. Institute of Environmental Management Technology, AIST, Tsukuba, 305-8569, Japan. A new concept of a gas separation process by the formation of clathrate hydrates in porous

membranes was proposed, and the concept was examined through the separation experiments for mixtures of greenhouse gases with nitrogen. The separation principle was based on the partition of the components in the mixture between the gaseous phase and the hydrate phase; a component that forms hydrate under milder conditions will be enriched in the hydrate phase. High selectivity can be expected for the separation of greenhouse gases from nitrogen because most of greenhouse gases such as Freons, SF6, and CO2 would form hydrate under much milder conditions than nitrogen. In this study, a novel separation process by using hydrate formation/dissociation in the pores of porous membranes was developed to realize a continuous separation process. Separation performances of model mixtures of HFC-134a and nitrogen were investigated experimentally by using various porous materials such as Vycor glass and porous alumina membranes, and relatively high selectivity of HFC-134a against nitrogen was observed under proper conditions within the hydrate formation region. 1. INTRODUCTION Clathrate hydrate is an inclusion compound of which the cage-like structure formed by hydrogen-bonded water molecules can hold gas molecules in it. In general, clathrate hydrates are thermodynamically stable under low-temperature and high-pressure conditions. A variety of applications of gas hydrates has been studied recently, especially in the field of gas storage and transportation, disposal of CO2, and gas separation[1]. " rnailto:[email protected]

508

Gas separation process by using hydrates is based on selective partition of the components in the mixture between the gaseous phase and the hydrate phase. Low-temperature and highpressure conditions are necessary for the hydrate formation in general. However, since stability of the hydrates depends on the shape and the size of guest molecules, interaction between guest and host molecules, thermodynamic conditions for the hydrate formation would differ in wide ranges of pressure and temperature depending on the guest molecules. Thus, for a mixture of components with large difference in the hydrate formation conditions, it can be expected that the component with milder conditions for the hydrate formation could be enriched in the hydrate phase against the components with severer conditions. In addition, the process could be environmentally benign because only water is concerned in the separation process itself. Applications of this principle have been examined for several gaseous mixtures through the phase equilibrium studies. From the practical point of view, the separation process is assumed to be conducted in the following three steps; hydrate formation, separation of the solid hydrate from the feed gas and water, and recovery of the enriched gas by dissociation of the hydrate. Such a multi-step procedure would make the separation process unnecessarily complex, which could reduce the process efficiency due to the energy requirement for the stirring to promote the hydrate formation, heat losses in the formation and dissociation steps, and leaking-off of the enriched gases in the solid-fluid separation step. To resolve the above difficulties, we propose a new type of separation process by using pores in the porous membranes for the hydrate formation, and the concept was applied for the separation of greenhouse gases from nitrogen or air for the recovery of such gases to mitigate global warming. To examine the practicability of the concept, separation performance was investigated for the mixtures of HFC-134a and nitrogen by using a laboratory-scale experimental setup. Mixtures of HFC-134a and nitrogen were selected as a model mixture not only because of a large difference in the hydrate formation conditions also because effective separation of HFC-134a from nitrogen is essential for preventing emission of this high-potential global warming gas into the atmosphere. 2. CONCEPT OF THE NEW TYPE SEPARATION PROCESS The concept of the proposed separation process was schematically illustrated in Figure 1. A gaseous mixture is fed into one side of the porous membrane (feed side

Figure 1 Schematic illustration of the concept of

hereafter), while the other side of the

the proposed separation process

509

membrane is filled with water (permeate side hereafter). With a proper choice of the pressure and temperature conditions, the feed side of the membrane could be maintained in the hydrate formation region, and the permeate side is kept atmospheric pressure, that is, the hydrate dissociation region at this temperature. Porous ceramic membranes would be appropriate for this purpose because a mechanical rigidity is required to maintain such a pressure drop. Water will be penetrated into the pores of the porous membrane by the surface tension, and will be contacted with the feed gas at a certain location in the pore. Since the feed side is in the hydrate formation region, hydrate will be formed at the water-gas interface. The component with milder criteria for hydrate formation will be enriched in the hydrate phase, and the other components will be remained in the gas phase; separation takes place. Once hydrate is formed, further formation of the hydrate will be prevented due to the large mass transfer resistance in the hydrate film. Instead, the hydrate film will be dissociated at the water side of the hydrate film. The hydrate dissociation will result in releasing the guest molecules captured in the cage-like structure. The released component will be transported to the permeate side of the porous membrane due to the concentration difference along the pore, and taken out as a product. The dissociation of the hydrate will then cause re-contact of water and the gaseous phase, and hydrate film will be reformed at the interface, which will eventually be dissociated at the water side. If this formation/dissociation processes could occur continuously with a short interval, a steady-state separation and transport process could be realized. 3. RESULTS AND DISCUSSIONS 3.1. HYDRATE FORMATION EQUILIBRIUM FOR THE MIXTURE OF HFC-134a AND NITROGEN Since the selectivity of the separation process is mainly determined by the partition of the species between the hydrate phase and the water phase, the phase equilibrium is of primary importance to evaluate the potential separation

performance.

Since no literature data are available for the present system of mixture of nitrogen and HFC-134a, measurements of the hydrate phase equilibrium in the bulk were conducted. A high-pressure vessel of Pyrex glass with magnetic stirrer was used for the measurements. The formation of the hydrate was confirmed by the

Figure 2 Mole fraction of HFC-134a in the hydrate phase plotted versus th at in the vapor phase. Total pressure = 5 bar,temperature = 277 K

510

direct observations in the vessel. In Figure 2, mole fraction of HFC-134a in the gas collected after dissociation of the hydrate phase was plotted against the mole fraction in the gaseous phase in equilibrium. The compositions in the hydrate were determined by a gas chromatography after separation of the hydrate from the fluid phase. For any composition of the gaseous phase studied, the corresponding mole fraction of HFC-134a in the hydrate phase (except water) is more than 90 %. 3.2 EFFECT OF PORE DIAMETER ON THE HYDRATE EQUILIBRIUM It is known that the hydrate formation equilibrium would be significantly affected by the surface effect in the pores of porous media. Since porous media with various pore size distributions were used in this study, effect of the pore size on the hydrate formation conditions should be evaluated. Figure 3 shows the results of theoretical calculations on the three phase equilibrium lines of hydrate-water (liquid)-vapor in the pore of various pore diameters. The calculations were conducted based on the Langmuir model of hydrate[2] with modification for the surface effects by the Kelvin equation.

Figure 3 Hydrate-Liquid Water-Vapor three-phase equilibrium line fo r HFC-134a (left) and nitrogen (right) in pores with various sizes. The basic equation was given in Eq. (1). Mw

RT0

+ k

—RT

dpF

h

— ^ - — d T - y v m ln 1+ > Cmjf. = RT1 „ { j )

-cosO rRT

(1) where ju is the chemical potential, Ah is the difference of enthalpy, Cmj is the Langmuir constant, a is the surface tension and 9 is the contact angle. Figure 3 indicates that the hydrate equilibrium pressure shifted to the higher region with a decrease in the pore size for a given temperature. However, the effect is almost negligible for the pores greater than 0.1 um.

511

3.3 SEPARATION EXPERIMENTS BY HYDRATE FORMATION IN POROUS MEDIA Schematic

drawing

of

the

experimental apparatus for the gas separation was shown in Figure 4. A

hollow

membrane

tube

of

was mounted

porous in a

permeation cell of stainless steel. A gaseous mixture was fed into the outer side of the porous membrane (feed side, hereafter) from the cylinders of nitrogen and HFC134a

after

mixing.

The

composition of the feed gas was controlled

by

the

mass

Figure 4 Schematic drawing for the experimental

flow

controllers, and the pressure was controlled by the back-pressure regulator. Water was fed into the inner side of the membrane (permeate side hereafter), and the permeated gas was collected by the water substitution method. The pressure of the permeate side was maintained at the atmospheric pressure. The composition of the permeated gas was determined by a gas chromatographic method (Shimadzu GC-8A), and the permeation rate was measured by a volumetric method. The porous membranes used in this study were as follows, Porous Vycor glass; mean pore diameter = 4 nm Porous alumina membranes; mean pore diameter = 0.1 um and 1 um Separation performances the alumina membrane with 0.1 um of the mean pore diameter were shown in Figs. 5 and 6. In Fig. 5, mole fraction of HFC134a in the permeate side was plotted versus that in the feed side. Selective permeations

of

HFC-134a

were

observed; HFC-134a was enriched to be more than 95 % in the permeate side for the feed with the mole fraction of

Figure 5 Relationship between the mole fraction in the permeate side and that in the feed side. Porous alumina membrane (average pore size = 0.1 |im), temperature = 277 K, pressure = 4 bar.

512

HFC-134a at 0.7 ~ 0.9. Figure 6 shows the corresponding permeation flux of HFC-134a as a function of the mole fraction of HFC134a in the feed. The permeation flux slightly depends on the feed

gas

composition,

presumably because a constant permeation rate of HFC-134a through the hydrate film in the pore. In some cases, however, no separation was observed in the permeate side. At the present stage, no clear criterion was found between the cases with rather good separation

and

without

selective permeation

separation.

No

of HFC-134a

was

observed in all the cases studied under high

Figure 6 Permeation flux of HFC-134a through the porous alumina membrane with average pore diameter = 0.1 urn.

temperature conditions out of the hydrate formation region (~298 K). Therefore, it can be speculated that the hydrate formation in the pores would be concerned in the selective permeation of HFC-134a in the pores. The cases with no selectivity would occur because sufficient hydrate was not formed in most of the pores due to insufficient nucleation for the hydrate formation. Similar results were obtained for Vycor glass membrane: selective permeation of HFC-134a, but the permeation fluxes are smaller in two order of magnitude, presumably due to smaller pore diameter. No separation was observed for the alumina membrane with 1.3 um of the mean pore diameter. This is so because the pore is too large to be filled with the hydrate film. 4. CONCLUSIONS A new type gas separation process by using hydrate in pores of porous membranes was proposed, and the feasibility of the process was examined experimentally for the mixtures of HFC-134a and nitrogen. It was observed that HFC-134a was enriched by the proposed separation process, indicating the new concept could be applicable for practical processes. REFERENCES 1. Okano, T., A. Yamasaki, F. Kiyono, M. Fujii, Y. Yanagisawa, Proceedings of the ICGH4, p. 1012 (2002) 2. Clarke M. A., M. Pooladi-Darvish P. R. Bishnoi, Ind. Eng. Chem. Res., 38, 2485 (1999) 3. Liang, D., K. Guo, R. Wang, S. Fan, Fluid Phase Equilibria, 187-188, 61 (2001)


513

The OACE (Ocean Abyssal Carbon Experiment) Research Project R. Kojima a *, I. Ayaa, K. Yamane a , S. Namie b , Y. Nakajima b , H. Shirotab, P. G. Brewerc, E. T. Peltzerc, P. M. Haugand, T. Johannessen d, and R. G. J. Bellerbyd a

National Maritime Research Institute, Osaka Branch, 3-5-10 Amanogahara, Katano, Osaka

576-0034, JAPAN b

National Maritime Research Institute, 6-38-1 Shinkawa, Mitaka, Tokyo 181-0004, JAPAN

c

Monterey Bay Aquarium Research Institute, 7700 Sandholdt Road, Moss Landing, CA

95039, U.S.A. d

University of Bergen, Allegaten 70, N-5007 Bergen, NORWAY The storage of liquid CO2 on the ocean floor, one of the promising measures to mitigate

global warming, requires >3500 meters depth for a gravitationally stable storage, breakthrough technology and a reasonable cost to realize, although it has many advantages such as a sequestration term longer than 2000 years. The National Maritime Research Institute (NMRI) conducted several joint field experiments with the Monterey Bay Aquarium Research Institute (MBARI) since 1999 under the auspice of the New Energy and Industrial Technology Development Organization (NEDO), and proposed the improved CO2 Sending Method for Ocean Storage (COSMOS), in which CO2 is released into 200 meters depth as a slurry mass (mixture of dry ice and cold liquid CO2). Since 2002, under the NEDO Grant, the NMRI started a new international joint research, OACE, Ocean Abyssal Carbon Experiment, with the MBARI and the University of Bergen (UoB), in order to accumulate the basic data on the long-term stability of stored CO2 and the environmental effects around a storage site.

1. INTRODUCTION The greenhouse effect, especially the one caused by CO2, has attracted much attention as a worldwide problem. Sequestration of CO2 in the ocean, as a promising measure to mitigate global wanning, was proposed earlier than a decade ago[1> 2]. CO2 hydrate plays an

Author to whom correspondence should be addressed. E-mail: [email protected]

514

important role in the evaluation of these measures, because CO2 reacts with seawater to form CO2 hydrate in the low temperature and high pressure conditions in the ocean deeper than 500 to 900 meters. Dissolution of liquid CO2 into seawater accompanied by CO2 clathrate-hydrate formation is anticipated during CO2 sequestration into the deep-sea ocean.[31 Sequestration of CO2 in the ocean is one of the strategies being considered for realizing the prevention of global warming. The OACE research project is designed to answer some of these questions. The NMRI conducted Fig.l. The concept of COSMOS.

several joint field CO2 release experiments with

MBARI since 1999 under the auspice of the NEDO, and proposed the improved COSMOS, in which CO2 is released below 200 meters depth as a mixture of dry ice and cold liquid CO2 (slurry mass)^. NMRI started a new international joint research, OACE with MBARI and UoB with a NEDO grant in 2002. The objective role of OACE is to acquire accurate data on the long-term stability of stored CO2 and its environmental effects around a storage site. 2. SITE OF IN-SITU EXPERIMENT As shown in Figure 2, the Monterey bay, looking like an ear is about 150 km south of San Francisco. At around the root of ear, there is the famous sightseeing city, Monterey. A very large and very steep oceanic canyon, Monterey Canyon, approaches Moss Landing from the Pacific basin. It is possible to conduct one-day cruise for a mid-depth ocean investigation from Moss Landing. It can be said that MBARI is located at an ideal site to conduct an in situ CO2 experiment'' . Moreover a large methane-burning power plant stands in front of MBARI across the state road. The shipping distance is very short if the CO2 recovered from this power plant would be used for CO2 Fig.2.

The

location

of

ocean storage in the future.

Monterey Bay about 150 km south of San Francisco. Author to whom correspondence should be addressed. E-mail: [email protected]

515

3. EXPERIMENT Ship/ROV availability in FY 2002 dictated that the experiment be carried out at a depth where liquid CO2 is buoyant in sea water, and thus an "upside down" experiment was created in which we were able to study physical forcing of a CO2 layer exposed to ocean currents. Since the CO2 saturated boundary layer is denser than seawater, a sinking plume should be observed as a low pH signal advected away from the source. It is this plume that is of concern for possible biological effects, and for knowledge of CO2 stability. In practice there are several unknowns, including details of boundary layer dynamics, and possible effects of slow hydration kinetics of CO2 at low temperatures. And while in situ glass electrodes are typically used for pH detection, dye techniques offer the potential for more sophisticated plume visualization and analysis. The initial task of the OACE project in 2003 will be to make use of the data successfully collected in the first field experiment, carried out in Monterey Bay on the research vessel "Point Lobos" and remotely operated vehicle "Ventana" as shown in Figure 3, respectively.

Fig. 3. Research Vessel Pt. Lobos (left side) and Remotely Operated Vehicle (ROV) Ventana (right side).

In this experiment, the density of seawater at the depth is greater than that of liquid CO2. Therefore the apparatus is equipped with an acrylic square box that is fixed upside-down on the top of apparatus. 4 sensors for pH were placed around the apparatus with different heights. A beacon-based guidance system, temperature sensors and recording CTD system are attached. Figure 4 shows the installation of the experimental apparatus on the shallow bottom of ocean as placed by the robot arms. This experiment was carried out with 20 liters of liquid CO2, at a depth where the liquid is strongly buoyant, and well within the hydrate forming Fig. 4. The image of the 1st experiment.

516

regime. The experiment was carried out for technical reasons: to field test sensors, to observe the fluid dynamics of liquid CO2 exposed to varying oceanic flow, to investigate techniques of pH sensitive dye injection, and to compare with hydrate formation characteristic in laboratory pressure vessels versus the natural ocean.

MEASUREMENT OF pH The study was carried out at 684m depth in Monterey Bay. A frame ~ 1.5 m high holding a square, transparent box for containing CO2 was constructed, and deployed by the ROV Ventana on the sea floor. Approximately 20 liters of liquid CO2 was dispensed into this from a steel accumulator, in two separate ROV dives. Two pH electrodes attached to a recording CTD unit were fastened to this frame, one electrode being close to the source. The liquid CO2 was injected into the acrylic box through a plastic tube. The surface of the liquid CO2 was quickly covered with hydrate, and a layer of CO2 covered with hydrate lay thick on the top of the box. If this stage of injection is continued, the box is only filled with hydrate and this is different from the expected storage formation. So, in order to avoid hydration formation during injection, the CO2 injection nozzle is inserted into the liquid CO2 droplets. As shown in Figure 5, the use of this injection prevents the liquid CO2from being exposed to seawater directly, and the liquid CO2 is stored at the top of CO2 container box as a large mass and is covered with hydrate at the interface between liquid CO2 and seawater. Figure 6 shows the record of pH we obtained adjacent to the CO2 container box, 2.0 m distant from it, along with conductivity (salinity), temperature, pressure and velocity of tidal flow the experimental site at 684 meters depth.

Fig. 5.

pH electrode attached to a recording system (left), and injection of the liquid CO2 into the acrylic

box through a plastic tube (right).

517

4. RESULTS AND DISCUSSION From these results we can clearly see the effects of tidal forcing in the periodicity of the signal, and the strong dilution recorded at the 1.5 m distant electrodes. We have also gained our first experience with system noise, and can calculate realistic detection limits etc. Local velocities were unexpectedly high (up to 40 cm/sec) and we were able to observe strong ripples on the CO2 surface, providing valuable guidance for a larger scale fluid dynamics study. The unexpectedly rapid mixing limits our interpretation of these early results, but clearly shows their potential. In this experiment, our research group examined so-called memory effect of hydrate formation, which is thought to influence much on the interaction between stored CO2 and sediment. The deep seawater and surface water showed no differences in the rate of hydrate formation. Some surface seawater was sampled to clarify the condition that the surface seawater does not show the

memory

effect

through

laboratory experiments. After the field experiment, using a 60 MPa high-pressure tank at the NMRI, a simulation experiment was carried out and measured the pH changes followed by the dissolution of stored

CO2.9'10

Moreover

the

sophisticated device to measure the thickness of thin CO2 hydrate membrane was improved. The second task of this project will be to prepare, both practically and theoretically, for 2003

field

experiment.

This

experiment will be of larger scale, and will take place at depths > 3500 meters, where the density of Fig. 6. Collected data of CO2 dissolution experiment at Monterey Bay about 684 m depths.

liquid CO2 exceeds that of CO2

518

rich seawater. The research questions now turn to the physical behavior of the liquid pool, and the dense boundary layer of low pH water above it. If ocean storage of fossil fuel CO2 is indeed seriously contemplated, then the ability to predict the behavior and fate of the material exposed to the normal range of ocean physical forcing is essential. The project will need to acquire the battery, thruster, and control technologies (MBARI), further test and develop the essential pH sensors to carry out first order simulations of the system so that correct scaling can be achieved (UoB),t7] and to study the interaction of sediment with liquid CO2 in laboratory high-pressure vessels (NMRI). ACKNOWLEDGEMENT

This work was supported by an International Research Grants from the NEDO, the Showa Shell Sekiyu Foundation for Promotion of Environmental Research and Japan Science and Technology Corporation of Cooperative System for Supporting Priority Research. The MBARI support was provided by the David and Lucile Packard Foundation, and by the U. S. Dept. of Energy Ocean Carbon Sequestration Program. The authors express their gratitude to Jim Barry and Chris Lovera of MBARI for preparation of data. The author wishes to express his sincere gratitude to the all members of above two international joint research teams.

REFERENCES 1. 2. 3. 4. 5. 6. 7.

Shindo, Y., Fujioka, Y., Ozaki, M., Takeuchi, K., Komiyama, H. (1993). In: Int. Symp. on CO2 Fixation & Efficient Utilization of Energy, Tokyo Institute of Technology(Eds), pp. 307-314. Nihous, G.C., Masutani, S.M., Vega, L.A., Kinishita, CM. (1996). Energy Corners. Mgmt., 37, 1039. Brewer, P.G., Friederich, G, Peltzer, E.T., Orr, F.M. Jr. (1999). Science, 284, 943. I. Aya, K. Yamane and K. Shiozaki, "Proposal of Self Sinking CO2 Sending System: COSMOS," Greenhouse Gas Control Technologies, Pergamon, Oxford, pp.269-274, 1999. P.G. Brewer, E.D. Peltzer, G. Friederich, I. Aya and K. Yamane, "Experiments on Ocean Sequestration of Fossil Fuel CO2: pH Measurements and Hydrate Formation," Marine Chemistry 72, pp.83-93. 2000. M.N. Tamburri, E.D. Peltzer, G. Friederich, I. Aya, K. Yamane and P.G. Brewer, "A Field Study of the Effect of CO2 Ocean Disposal on Marine Deep-sea Animals," ibid. 72,pp.95-101,2000. I. Fer and P.M. Haugan, "On the Fate of a Pur-posefully Disposed CO2 Lake in the Deep Ocean," Proc. 6th Int. Conf. on Greenhouse Gas Control Technologies, Kyoto, Japan, 2002.


519

Addition of carbon dioxide to phenylglycidyl ether using quaternary ammonium salts grafted on modified MCM-41 support Dae-Won Parka *, Dong-Hwan Shina, Byung-Soo Yua, Kyung-Hoon Kima, Sang-Wook Park". II Kima, and Jae-Wook Leeb a

Division of Chemical Engineering, Pusan National University, Kumjeong-ku, Busan 609735, Korea. E-mail : dwpark(5)pusan.ac.kr(D.-W.Park)

b

Department of Chemical Engineering, Sogang University, Mapo-ku, Seoul 121-742, Korea

The synthesis of phenoxymethyl ethylene carbonate(PMEC) from carbon dioxide and phenyl glycidyl ether(PGE) was investigated using chloro-propyl-methyl modified-MCM-41 supported quaternary ammonium salt catalysts. The catalytic activity was affected by the composition of organic modifier and the structure of trialkyl amine. The reaction was considered as pseudo first-order and the apparent activation energy for the reaction was estimated as 15.2 kJ/mol. 1. INTRODUCTION Most published works on resin-bound quaternary onium salts use styrene-divinylbenzene resins and related resins, taking advantage of the large amount of technology available on these resins due to their use as ion-exchange resin support]!]. However, the polymer-supported catalysts have yet some drawbacks such as diffusion limitations, high cost, tendancy to swell, mechanical and thermal instability. Since the discovery of ordered mesoporous silica[2,3], there have been continuous efforts to improve their stability and catalytic performance. Recently, the incorporation of 3-chloroprophyl and methyl substituents on MCM-41 was reported[4]. The chemical fixation of carbon dioxide has received much attention from the view of an environmental problem. An attractive strategy to deal with this situation is converting CO2 into valuable substances. Nishikubo et al.[5] reported the synthesis of cyclic carbonates by using insoluble polystyrene-bound quaternary salt catalysts. In our previous work [6], we reported the synthesis of 4-(phenoxymethyl)-l,3-dioxane-2-one (PMEC) from CO2 and phenyl glycidyl ether (PGE) using monomeric quaternary salt catalysts. PMEC is a good example of cyclic carbonates having high polarity and high reactivity for introduction of functional groups[7]. In the present study, the same reaction was performed using Cl-propyl-methyl-MCM-41 supported quaternary ammonium salt catalysts. Our research has focused on the effects of the concentration of the organic modifier and the structure of quaternary ammonium salts on the conversion of PGE in the synthesis of PMEC. A kinetic study on the reaction with the modified MCM-41 supported catalyst was also carried out. 2. EXPERIMENTAL Organically modified MCM-41 materials were synthesized according to a previous report[4] using a mixture of 3-chloropropyl-dimethoxymethyl silane (C1PDMMS, Shin-etsu) and tetraethyl orthosilicate (TEOS, TCI) as Si sources. Cetyltrimethylammonium bromide (CTMABr, TCI) was used as a template and tetramethylammonium hydroxide (25% aqueous TMAOH, Aldrich) was used to provide alkalinity of the medium. A typical molar ratio of the various constituents in the synthesis mixture was: 4.0 TEOS : 1.0 C1PDMMS : 1.25 CTMABr : 2.5 TMAOH : 300 H2O. To

520

the suspension of 1.2 g Cl-propyl-methyl-MCM-41, a solution of 4 g tributylamine dissolved in 20 mL benzene was added under stirring. After stirring overnight, ether was added and the product was filtered, washed with ether, and dried in vacuum to yield 1.5 g of the supported quaternary salt catalyst. The amount of quaternary ammonium salt bound to MCM-41 support was analyzed by an elemental analyzer (Varian EL). The morphology of the MCM-41 catalyst was analyzed by scanning electron micrograph (SEM, Hitachi S-4200). The syntheses of PMEC from PGE and CO2 using Cl-propyl-methyl-MCM-41 supported catalysts were carried out in a semi-batch reactor with a continuous supply of carbon dioxide. 3. RESULTS AND DISCUSSION 3.1. Preparation and characterization of Cl-propyl-methyl-MCM-41 XRD pattern of Cl-propyl-methyl-MCM-41 showed distinct peaks of MCM-41, the 100, 110. 200, and 210 reflections, indicating hexagonal ordering[2,3]. BET analyses revealed that the support has 1037 m2/g of surface area, 1.08 cnrVg of total pore volume, and 4.18 nm of average pore diameter. Table 1. The composition of Cl-propyl-methyl-MCM-41 support and the amount of immobilized quaternary ammonium salt. Catalyst TEOS(mol)

ClPDMMS(mol)

Tertiary Amine

"Attached amount (mmol N+/g-catalyst)

C-0

b

b

Tributyl

0.69

C-l

4

0.5

Tributyl

0.89

C-2

4

1.0

Tributyl

1.61

C-3

4

1.5

Tributyl

1.72

C-4

4

2.0

Tributyl

1.78

C-5

4

1.5

Triethyl

1.83

C-6

4

1.5

Trihexyl

1.54

C-7 4 1.5 Trioctyl a : Measured by elemental analysis b : C-0 is a silica-gel supported quaternary ammonium chloride catalyst

1.42

Table 1 shows the composition of the Cl-propyl-methyl-MCM-41 support and the amount of immobilized quaternary ammonium salt. All the Cl-propyl-methyl-MCM-41 supported catalysts showed higher amount of the quaternary salt than silica-gel supported one (C-0) did. The attached amount of ammonium salt increased as the amount of C1PDMMS increased. Among different trialkyl amines used for the quaternization of the Cl-propyl-methyl-MCM-41 support, triethyl amine showed the maximum attached amount of the quaternary salt group probably due to its small size, which facilitates its approach to the chlorine position of the C1PDMMS. 3.2. Addition of CO2 to phenyl-glycidyl ether The addition reaction of PGE and CO2 at 90°C was performed with the catalysts of equal amount of attached quaternary ammonium salts. Therefore, 0.3 mmol of pendant benzyltrialkyl ammonium chloride residue was used with 30 mmol of PGE. Fig. 1 shows time variant conversion of PGE with different types of catalysts. Tetrabutylammonium chloride salt attached to Cl-propyl-methyl-MCM-41 support(C-3) showed higher PGE conversion than that supported onto functionalized silica-gel. The silica-gel had specific surface area of 706 m2/g, total pore volume of 0.34 cm /g, and average pore diameter of 1.93 nm. Mesopores of MCM-41 might

521

Fig. 1. Time variant conversion of PGE for supported quaternary ammonium catalysts.

Fig. 3. Effect of amine structure on the conversion of PGE.

Fig. 2. Effect of TEOS/CIPDMMS ratio on the conversion of PGE.

Fig. 4. First-order plots of ln([PGE]o/[PGE]) vs. time at various reaction temperatures.

enhance the approach of PGE to the active catalytic site of the quaternary ammonium salt. However, C-3 showed slightly lower PGE conversion than a soluble tetrabutyl ammonium chloride catalyst at the same reaction condition. Fig. 2 shows PGE conversion after 6 h for the Cl-propyl-methyl-MCM-41 supported catalysts with different molar ratio of TEOS to C1PDMMS. Catalyst C-3 with TEOS/CIPDMMS of 4/1.5 showed the maximum PGE conversion. As shown in Table 1, however, C-4 catalyst with TEOS/CIPDMMS of 4/2 contained the highest amount of the quaternary salt. Therefore, one can suggest that too much crowded active sites on the modified MCM-41 support seemed to inhibit the diffusional approach of the PGE to the active sites due to the steric hindrance. Fig. 3 shows the effect of cation structure of the quaternary salt on the conversion of PGE. The catalyst prepared with tributylamine showed the highest conversion of PGE. In our previous work [6] on the same reaction with unsupported liquid quaternary ammonium salt catalysts, the order of catalytic activity for the salt was tetraoctyl- > tetrahexyl- > tetrabutyl- > tetraethyl-ammonium salt. Bulky quaternary salts, having longer distances between cations and anions, are generally known to exhibit higher activity in activating anions. However, the immobilized catalyst from trihexyl or trioctyl amines showed lower catalytic activity than that of tributylamine. This result indicates that the steric hindrance of bulky alkyl structure in the Cl-propyl-methyl-MCM-41 supported catalyst inhibits the approach of PGE to the supported active sites. To test the stability of the polymer-immobilized catalyst, 4 successive experimental runs were

522 carried out with the C-3 catalyst for the reaction of PGE and CO2 in DMF at 90°C. Each experimental run continued 6 h and then the used catalyst was separated and washed for the next run. The PGE conversion for the MCM-41supported quaternary ammonium salt catalyst was 47.5%, 46.1%, 44.7%, 42.5%, for the 1st, 2nd, 3rd and 4™ run, respectively. The catalyst can be reused for at least 4 successive runs with a decrease of its initial activity of less than 10%. Surface analysis of the used catalysts by energy dispersive spectroscopy (EDS, Jeol JXA-8600) confirmed the maintenance of chlorine groups bound to the polymer support after 4lh run. 3.3. Kinetic Studies in a Semi-Batch Reactor For the addition reaction of CCb with PGE using the Cl-propyl-methyl-MCM-41 supported catalysts, the following pseudo first-order rate equation can be applied since the catalyst concentration does not change during the reaction: ln([PGE],,/[PGE]) = k t where [PGE]o is the initial concentration of PGE. From the slope of the linear plot between In ([PGE]«/[PGE]) vs. time, one can estimate the pseudo first-order rate constant k. The addition of CO2 with PGE was carried out by using 1 mol% of C-3 catalyst in DMF. As shown in Fig. 4, the experimental data fit well to the pseudo first-order kinetics. The pseudo first-order rate constant k for 60, 70, 80, 90, and 100°C was 0.083, 0.097, 0.117, 0.123, and 0.149 h"', respectively. From the Arrhenius plot, the apparent activation energy for the CO2 addition reaction was estimated as 15.2 kJ/mol. 4. CONCLUSIONS Carbon dioxide can be effectively added to the epoxide ring of phenyl glycidyl ether (PGE) to produce the corresponding 5-membered cyclic carbonate. The Cl-propyl-methyl-MCM-41 supported quaternary ammonium salts showed good catalytic activity even at the atmospheric pressure of carbon dioxide. The catalytic activity highly depends on the composition of the organic modifier, the structure of trialkylamine, and the type of support. The catalyst can be reused for several times and it can have high possibility of easy recovery. The addition reaction of CO2 to PGE can be considered as pseudo-first order with respect to the concentration of PGE, and the apparent activation energy was estimated as 15.2 kJ/mol. ACKNOLEDGEMENT This work was supported by the Applied Rheology Center, and by Brain Busan 21 and Brain Korea 21 Program.

REFERENCES 1. C. M. Starks, C. L. Liotta and M. Halpern, "Phase Transfer Catalysis", Chapman & Hall, New York (1994). 2. C.T. Kresge, M.E. Leonowicz, W.J. Roth, J.C. Vartuli and J. S.Beck, Nature 359 (1992) 710. 3. S. Inagaki, Y. Fukushima and K. Kuroda, J. Chem. Soc. Chem. Commun. (1993) 680. 4. A. Bhaumik and T. Tatsumi, Catal. Lett. 66 (2000) 181. 5. T. Nishikubo, A. Kameyama, I.Yamashita, M. Tomoi and W. Fukuda, J. Polym. Sci. A . Polym. Chem. 31 (1993) 939. 6. D. W. Park, J. Y. Moon, J. G. Yang , S. M. Jung, J. K. Lee and C. S. Ha, Stud. Surf. Sci. Catal. 114(1998)403. 7. N. Kihara, N. Hara and T. Endo, J. Org. Chem. 58(1993) 6189.


523

Chemical absorption of carbon dioxide with NaOH in non-Newtonian w/o emulsion Sang-Wook Park 3 , Byoung-Sik Choi3, Tae-Young Kim3, and Jae-Wook Leeb a

Division of Chemical Engineering, Pusan National University, Pusan 609-735, Korea

b

Department of Chemical Engineering, Sogang University, Seoul 121-742, Korea

Carbon dioxide was absorbed into water-in-oil (w/o) emulsion composed of aqueous NaOH droplets as a dispersed phase and benzene solutions of polybutene (PB) and polyisobutlene (PIB) as a continuous phase in a flat-stirred vessel to investigate the effect of non-Newtonian rheological behavior on the rate of chemical absorption of CO2, where the reaction between CO2 and NaOH in the aqueous phase was assumed to be a pseudo-first-order reaction. It was expressed that PIB with elastic property made the rate of chemical absorption of CO2 accelerated by comparison of kL in the non-Newtonian liquid with that in the Newtonian liquid. 1. INTRODUCTION If the system is water-in-oil (w/o) emulsion such that the dispersed phase is the aqueous-solution containing reactant, and the continuous phase is the organic solvent having larger solubility of gas than water, then the specific rate of absorption may be enhanced because of larger solubility and chemical reaction. A qualitative explanation of this phenomenon has been given by various authors [1-3]: small droplets of a liquid immiscible with the continuous liquid phase absorb the gas in the hydrodynamic mass-transfer film, after which desorption of the gas takes place in the gas-poor bulk of the liquid. It is considered worthwhile to investigate the effect of non-Newtonian rheological behavior on the rate of chemical absorption in w/o emulsion, where a reaction between CO: and reactant occurs in the dispersed phase. In this study, the chemical absorption mechanism of CO2 into w/o emulsion composed of aqueous NaOH solution and benzene solution of PB and PIB is presented, and the measured absorption rates of CO2 are compared with those obtained from the model based on the penetration theory with chemical reaction. The volumetric mass transfer coefficient obtained from the previous work[4] is used to estimate the enhancement factor due to chemical reaction. 2. THEORY In case of absorption of CO2(A) into w/o emulsion with benzene solution of PB and PIB-aqueous NaOH(B) solution as shown in Figure 1, the mathematical model is developed to describe the absorption of CO2 into the continuous benzene phase through the gas-liquid interface under unsteady-state and transfer into the dispersed aqueous droplets through the

524

liquid-liquid interface under steady-state, where the chemical reaction of CO2 occurs.

Fig 1. Chemical absorption path of gas(A) into w/o emulsion. The conservation equations of CO2 transferred into the dispersed aqueous phase are given as D ,

A

( ^

{ dr2

+

^ ) - k

2

c

A

c

B

DeBf^+2CM=vk2cAcB { dr

2

(1)

r dr J

(2)

r dr J

Boundary conditions to be imposed are r = R; c A = c A = H A C A , ^ = 0 dr

(3)

r = 0;

(4) '

^

=^ dr

=0 dr

The conservation equation for the dissolved gas in the continuous phase with its volume fraction of e at unsteady-state is written as D

A ^ dz~

=

^ T dt

+ (1

-8)k^cB«Ef

(5)

Boundary and initial conditions are given as z-0,

t>0; CA=CA,

(6)

z>0, t = 0; CA = 0

(7)

z = oo, t >0; CA = 0

(8)

525 3. EXPERIMENTAL The w/o type emulsion made from benzene solution of PB(Mw=680) and PIB(Mw= 1000000) and water was made by the same procedure as those reported elsewhere [4] by adding Tween 80 (Aldrich Chem. Co.) and Arlacel 83 (Aldrich Chem. Co.) as surfactant, by using a homogenizer (Fisher Scientific Co.) in the range of agitation speed of 1500-10,000 rev/min, and the mean size of aqueous droplets was measured by Image Analyzer (Leitz TAS Plus Co.), and its value was 4x10° m. The rheological properties of w/o emulsion were measured by the parallel disk type rheometer(Ares, Rheometrics, U.S.A.) of the diameter of 0.05 m and the gap of 0.001 m. The w/o emulsion was composed of aqueous NaOH droplets as a dispersed phase and non-Newtonian viscoelastic benzene solutions of PB and PIB as a continuous phase. The absorption rates of CO2 were measured along the procedure similar to those reported elsewhere (Park, et al., 2002) at 25 °C and an atmospheric pressure. 4. RESULTS AND DISCUSSION The mass transfer coefficient, k^ of CO2 in COêmulsion system without NaOH in the aqueous droplets was estimated by using the empirical equation [5] correlating the relationship between k| a and the experimental variables in the w/o emulsion of non-Newtonian liquid as follows.

k I , a dVD eff =6.348xl0- 9 (d 2 Np/J Z536 (^/ f xJ 2397 (l

+

2461.3D e r 274 (a/aJ- 0039

(9)

where De is Deborah number defined as the ratio of the characteristic material time to the characteristic process time in the rheological behaviour. To observe the effect of viscoelasticity on the mass transfer coefficient of CO2. k| . in w/o emulsion of H2O/(PB/PIB/Bz) system, the effect of viscosity on kL was observed. The values of kL were estimated using Eq. (9), and Fig. 2 show typically plots of kL against \i at PB of 30 wt %. As shown in Fig. 2 kL decreases with increasing the viscosity. Note that mass transfer coefficient of a solute in the liquid phase is in inverse proportional to the viscosity of the liquid phase, because the diffusivity of the solute is in inverse proportional to the viscosity.

Fig. 2. Effect of viscosity on the mass transfer coefficient for w/o emulsion with PB of 30 wt% and PIB of 0.1-1 wt%.

526

To observe the effect of concentration of NaOH on the rate of chemical absorption, the rate of chemical absorption was measured according to the change of NaOH concentration. Fig. 3 shows typically a plot of RA against NaOH concentration at PB of 30 wt% and PIB of 1 wt%.

Fig. 3. Effect of NaOH concentration on the rate of chemical absorption into the w/o emulsion with PB of 30 wt% and PIB of 1 wt%.

As shown in Fig. 3, RA increases with increasing NaOH concentration. The rates of chemical absorption into w/o emulsion with the non-Newtonian and Newtonian liquid as the continuous phase are shown as a solid and dotted line in Fig. 3, respectively. As shown in Fig. 3, the value of the dotted line smaller than that of the solid line. This means that the elastic due to PIB also accelerates the rate of chemical absorption. ACKNOWLEDGEMENTS This work was supported by the Basic Research Program of the Korea Science and Engineering Foundation(KOSEF) though the Applied Rheology Center (ARC), Brain Korea 21 Project in 2003, and Brain Busan 21 Program.

REFFERENCES 1. Bruining, W. J., G.E.H. Joosten, A.A.C.M. Beenackers, and H. Hofman, Chem. Eng. Sci., 41(7) (1986) 1873. 2. Linek, V., and P. A. Benes, Chem. Eng. Sci, 31 (1976) 1037. 3. Mehra, A. A. Pandit, and M. M. Sharma, Chem. Eng. Sci, 43(4), (1988) 913. 4. Park, S. W. H. B. Cho, I. J. Sohn, and H. Kumazawa, Sep. Sci. Technol, 37(3), (2002) 639. 5. Park, S. W, I. J. Sohn, S. G. Sohn, and H. Kumazawa, Sep. Sci. Technol, 38(6) (2003) 1361.


527

CO2 Absorption and Regeneration using Na and K Based Sorbents Soo Chool Lee a , Bo Yun Choi a , Soo Jae Lee a, Suk Yong Jung a , Chong Kul Ryu b, and Jae Chang Kim a ' * a

Department of Chemical Engineering, Kyungpook National University, Daegu, 702-701,

Korea. b

Korea Institute of Energy Research, Daejeon, 305-380, Korea.

Carbon dioxide was chemically absorbed through the reaction K2CO3 + H2O + CO2 — 2KHCO3 to form potassium hydrogencarbonate over alkali metal supported on various sorbents in the presence of moisture. The amount of carbon dioxide captured depended on the dispersion of alkali metal and pretreatment conditions. The sorbents with low loadings and high dispersions were activated easily at low temperature and low humidity (60 °C and 6.5% relative humidity). But the sorbents with high loadings and the physical mixture of alkali metal and support needed a higher temperature and relative humidity (90 °C and 75 % relative humidity). So far the absorption capacities of various sorbents were proportional to the amounts of alkali metal loaded if the alkali metals were fully activated in the severe pretreatment conditions. 1. INTRODUCTION Carbon dioxide is a greenhouse gas that is released to the environment during the usage of fossil fuels. Because of these carbon dioxide emissions, global climate warming may occur to be disastrous to the environment. Carbon dioxide can be removed from flue gas and waste gas streams by various methods such as membrane separation, absorption with a solvent, and adsorption using molecular sieves.1 But these technologies are costly and an energy consuming operation. One of the improved techniques for the removal of carbon dioxide is the chemical absorption of carbon dioxide with solid sorbents.2> 3 The use of solid sorbents containing alkali metal carbonates such as Na2CO3 and K2CO3 has been previously reported in the literature.3'4 In this work, sorbents containing alkali metals deposited on the substrates with high surface area such as activated carbon, alumina and various zeolites and physical mixture of alkali metal and support were investigated for the removal of carbon dioxide in the presence of moisture. The effect of pretreatment conditions like relative humidity and temperature on the activation of these sorbents was studied.

* Corresponding author. Tel. +82-53-950-5622; Fax: +82-53-950-6615; E-mail: [email protected]

528 2. EXPERIMENTAL Sorbents prepared by two methods were used in this study. One sorbent was prepared by a physical mixing method and the other sorbents were prepared by impregnation of alkali metal carbonates like sodium and potassium carbonate on the various supports such as activated carbon, alumina and various zeolites. The carbon dioxide absorption and regeneration was carried out in a fixed bed reactor under atmospheric pressure. Before the CO2 absorption experiments, moist nitrogen with 6.5-75% relative humidity at 40-90 °C was supplied to all sorbents with flow rate 40 mL/min for 2 hours. The absorption and regeneration temperatures were fixed at 60 °C and 150 °C, respectively. Operations were carried out using a gas mixture with 1 vol% of CO2 and 1.3 vol% of H2O. The outlet gases from the reactor were automatically analyzed by a thermal conductivity detector (TCD; Young In Scientific Co.) equipped with an auto sampler (Valco Instruments Co. Inc.). The content of potassium and sodium of sorbents was analyzed by atomic absorption spectrometer (AAS; Shimadzu Co.). 3. RESULTS AND DISCUSSION Table 1 shows carbon dioxide capture capacities of various sorbents prepared by impregnation of alkali metal carbonates. The carbon dioxide capture capacities of K2CO3based alumina and activated carbon (AC) were 3.8% (0.038 g COi/g sorbent) and 4.1%, respectively. These sorbents showed higher carbon dioxide capture capacity than those of other supports. Carbon dioxide capture capacities of sodium based sorbents were less than those of the sorbents promoted by Table 1 potassium carbonate. In particular, CO2 capture capacities of sorbents incorporated s o d i u m carbonate supported on the with alkali metal carbonates at 60 °C and 1 atm alumina showed very small carbon dioxide capture capacity. The Impregnant CO2 regeneration of spent sorbents was Sorbent icontent (wt%) capacity (%) carried out by heating in the moisture 23.0 4.1 K2CO3/AC condition at 150 °C under atmospheric pressure. Except for K2CO3 and K2CO3/MCM-48 21.6 1.4 Na CO3-based alumina, most of alkali 2 21.2 K2CO3/NaY 2.3 metal based sorbents were completely 18.5 3.8 K2CO3/AI2O3 regenerated and transformed to their original phases, regardless of supports 21.3 1.8 K2CO3/Zeolite P and alkali metals. It has been shown K2CO3/USY 20.3 1.7 that the carbon dioxide capture capacity Na2CO3/AC 20.5 1.8 of sorbent supported on AI2O3 decreases after regeneration due to the formation Na 2 C0 3 /Al 2 0 3 19.5 0.7 of an inactive phase of potassium aluminum carbonate as reported in the previous paper.2'5 Figure 1 shows carbon dioxide breakthrough curves of K2CO3-based sorbents with different K2CO3 loading at 60 °C after pretreatment of sorbents with 1.3 vol% of moist nitrogen (6.5% relative humidity) at 60 °C. As shown in figure 1, the CO 2 capture capacities did not increase proportionally with the amount of potassium carbonate loaded on the support. As shown in Table 2, the carbon dioxide capture capacities of sorbents increased from 1.5% to 5.7% when the loading was increased to 33 wt%. But 50 wt% K2CO3 /AC showed the

529

Fig. 1. Carbon dioxide breakthrough curves of the various activated carbon sorbents with different K2CO3 loadings at 60 °C after pretreatment with 6.5% relative humidity at 60 °C for 2 hours.

Fig. 2. Carbon dioxide breakthrough curves of the various activated carbon sorbents with different K2CO3 loadings at 60 °C after pretreatment with 75% relative humidity at 90 °C for 2 hours.

carbon dioxide uptake of 5.4% which was almost the same as that of 33 wt% K2CO3 /AC. 50 wt% K2CO3 +AC prepared by physical mixing method showed the carbon dioxide uptake of 0.2%. These results indicate that the carbon dioxide capture capacity can be affected by the dispersion of alkali metal carbonate and activation processes before absorption. To confirm the effect of their dispersion and activation of sorbent, K^CC^-based sorbents were treated at different activation conditions. Figure 2 shows breakthrough curves of carbon dioxide of sorbents treated at a severe activation condition (75 % relative humidity at 90 °C instead of 6.5% relative humidity at 60 °C). As mentioned in figure 1, K.2CO3-based sorbent prepared by physical mixing showed very small carbon dioxide capture capacity, but after treated at a severe pretreatment condition, that of the sorbent increased to 11.3%. As shown in figure 2, the capacities of all alkali metal based sorbents increased with the amount of potassium carbonate loaded on the supports in contrast to the results shown in figure 1. Table 2 is the summary of the carbon dioxide capture capacity represented by the units of g CO2/ g sorbent and g CCV g K2CO3 after different pretreatments. As stated at the very beginning of this section, pretreatment conditions were 6.5% relative humidity at 60 °C and 75% relative humidity at 90 °C for 2hours. As shown in the last column of the table, the capture capacities per gram of K2CO3 were almost same for all the sorbents indicating that the capacity was increased linearly with loading after the samples were fully activated. It is clear that the enhanced carbon dioxide capture capacities are observed when the absorption is carried out after pretreatment with excess water at 60-90 °C. The reason for the enhanced carbon dioxide capture capacity after severe pretreatment is not clear, but it is thought that the absorption occurs through the active hydrate-intermediates and that the formation of the intermediate is affected by the dispersions and pretreatment conditions.

530

Table 2 The summary of the carbon dioxide capture capacity of alkali carbonate sorbents calculated from the breakthrough curves shown in figure 1 and 2 60 °C -6.5% relative humidity(2h) a

9 wt% K2CO3/AC 23 wt% K 2 CO 3 /AC a 33 wt% K2CO3/AC a 50 wt% K2CO3/AC a 50 wt% K2CO3+AC b a

g CO2/g sorbent 1.5% 4.1% 5.7% 5.4% 02%

90 °C -75% relative humidity(2h)

g CO2/g K£C h

g CO 2 /g sorbent

gCOz/gKzCQ,

17.0%

2.1%

23.0%

17.9%

5.1%

21.7%

172%

7.3%

222%

10.1%

10.8%

21.5%

0.3%

11.3%

22.5%

b

impregnation, physical mixing

4. CONCLUSION The carbon dioxide capture capacities of various sorbents prepared by adding alkali metals to various supports such as activated carbon, alumina and zeolites were investigated in the fixed-bed reactor. When the absorption was carried out after pretreatment with excess water at 60-90 °C, the enhanced carbon dioxide capture capacity was observed. The carbon dioxide capture capacity depended on the nature of the supports, the amount of alkali metal carbonates loaded and activation processes of sorbents. But if the sorbents were fully activated in severe pretreatment conditions, the capacities per gram of K2CO3 were almost same. The K^CCh-based activated carbon was clarified as the leading sorbent for CO2 removal. ACKNOWLEDGEMENT This research was supported by a grant (DAI-101) from Carbon Dioxide Reduction & Sequestration Research Center, one of the 21 st Century Frontier Programs funded by the Ministry of Science and Technology of Korean government.

REFERENCES 1. Okunev, A. G., Sharonov, V. E., Aristov, Yu. I., and Parmon, V.N., React. Kinet. Catal. Lett, 71, No. 2(2000)355. 2. Sharonov, V. E., Tyshchishchin, E. A., Moroz, E. M, Okunev, A. G, and Aristov, Yu. I., Russian J. Appl. Chem., 74, No. 3 (2001) 409. 3. Hirano, S., Shigemoto, N., Yamada, S., and Hayashi, H., Bull. Chem. Soc. Jpn., 68 (1995) 1030. 4. Hayashi, H., Taniuchi, J., Furuyashiki, N., Sugiyama, S., Hirano, S., Shigemoto, N. and Nonaka, T., Ind. Eng. Chem. Res., 37 (1998) 185. 5. Okunev, A. G., Sharonov, V. E., Gubar, A. V., Danilova, I. G., Paukshtis, E. A., Moroz, E. M., Kriger, T. A., Malakhov, V. V., and Aistov, Yu. I., Russian Chem. Bull., 52, No. 2 (2003) 359.


531

Preparation of 6FDA-based polyimide membranes for CO2 gas separation Kwang-Je Kim*, Won-Wook So and Sang-Jin Moon Advanced Chemical Technology Division, Korea Research Institute of Chemical Technology (KRICT), P.O. Box 107, Yuseong, Daejon 305-600, Korea. FAX: +82-42-860-7590. E-mail: [email protected] Polyimides for CO2 membrane separation were prepared from the thermal imidization reaction of 2,2-bis(3,4-dicarboxyphenyl)hexafluoropropane dianhydride(6FDA) with different diamines. The solubility of 6FDA-based polyimides to organic solvents played an important role in preparation of a dense membrane or a composite one. The composite membranes were available for 6FDA-based polyimides to be solube in 2-methoxyethanol or/and alcohols including methanol, ethanol, and butanol. CO2 separation performances for 6FDA-based dense membranes were investigated and compared with those for the composite membranes. The CO2 permeability and CO2/N2 selectivity of 6FDA-TrMPD polyimide dense membrane with different molecular weights were measured. 1. INTRODUCTION Global warming is a potential threat causing widespread concern. Removal and concentration of CO2 from the CO2 emission sources such as power stations, steelworks, and chemical industries are an important subject. Membrane processes are useful for efficiently separating CO2 from the combustion gas [1, 2]. 6FDA-based polyimides have attracted much attention as membrane materials for gas separation because of their high gas permselectivity and excellent physical properties [3, 4]. Polyimides are prepared by the imidization of polyamic acids obtained from a dianhydride and a diamine. In the preparation of polyimides, 6FDA is well known as an anhydride component of polyimide membranes. Diamine moieties can play an important role in the solubility of a polymer to organic solvents and in the control of membrane separation performances. In this study, 6FDA-based polyimide membranes having different diamine moieties were prepared. Diamines consisted of moieties with polar groups of a carboxyl or hydroxyl, and with different molecular weights of 2,4,6-trimethyl-l,3phenylenediamine(TrMPD). CO2 gas permeation properties for the membranes were investigated. 2. EXPERIMENTAL As materials for polyimide synthesis, 6FDA, 2,2-bis(3-amino-4hydroxyphenyl)hexafluoropropane(BAPAF), 2,4-diaminophenol dihydrochloride(DAP), 3,5diaminobenzoic acid(DABA), 3,3-dihydroxy-4,4-diaminobiphenyl(HAB), and TrMPD were used. The polyimides were prepared by polycondensation with equimolar amounts of the dianhydride and diamines followed by a cyclodehydration. Solubility test for the 6FDA-based polyimides was done using alcohol, glycol ether, and aprotic solvent. Gas permeabilities of 6FDA-based polyimides for CO2, N2 and CH4 were measured at 25°C as described in the

532

literatures [5, 6]. 3. RESULTS AND DISCUSSION A Solubility of 6FDA-based polyimides prepared in this study depended on organic solvents. 6FDA-BAPAF polyimide was dissolved in all solvent used, i.e. alcohol (methanol, ethanol, butanol), and glycol ether (2-methoxyethanol, 2-ethoxyethanol, 2-butoxy ethanol) as listed in Table 1. 6FDA-DAP and 6FDA-DABA polyimides were soluble in only 2methoxyethanol. All 6FDA-based polyimides including 6FDA-HAB and 6FDA-TrMPD polyimides were dissolved in aprotic solvents such as l-methyl-2-pyrrolidinone, dimethylacetamide, and dimethylformamide. The selection of 6FDA-based polyimide to be dissolved in alcohol and glycol ether can lead to the preparation of a dense membrane and a composite one using commercially available poly(sulfone) or poly(ether sulfone) as a support layer. It is possible for alcohol and glycol ether to be used as coating solvents because they Table 1. Solubility of 6FDA-based polyimides to organic solvents at 25°C. Polyimides

ME"

EE

BE

s

S

S

S

IS

IS

S

PS

IS

IS

IS

s

IS

IS

IS

IS

IS

IS

IS

IS

IS

IS

IS

IS

IS

IS

MeOH

EtOH

BuOH

6FDA-BAPAF

S

S

6FDA-DAP

IS

6FDA-DABA

IS

6FDA-HAB 6FDA-TrMPD

S: soluble; IS: insoluble; PS: partially soluble a

ME: 2-methoxyethanol; EE: 2-ethoxyethanol; BE: 2-butoxyethanol

Table 2. Comparison of separation properties for 6FDA-based polyimide membranes. Membrane type

Flow rate of CO2 gas 7 Barrer 6FDA-BAPAF(PI1) Dense 6 Barrer 6FDA-DAP(PI2) Dense 5 Barrer 6FDA-DABA(PI3) Dense 6FDA-TrMPD(PI4) Dense 100 Barrer 105 Barrer 6FDA-TrMPD(PI5) Dense 68 Barrer 6FDA-TrMPD(PI6) Dense 6 Barrer Dense 6FDA-3BDAF 30 Barrer Dense 6FDA-IPDA Dense 8 Barrer 6FDA-DAFO Asymmetric 5 GPU 6FDA-APPS 25 GPU 6FDA-BAPAF(PI1) Composite 6FDA-DAP(PI2) Composite 39 GPU 26 GPU 6FDA-DABA(PI3) Composite Barrer=10" 10 cm 3 (STP)- cm- cm"2- s"1' cmHg"1 GPU=l(r 6 cm 3 (STP)- cm'2- s 1 ' cmHg 1 Polyimide

a CO2/N2 18 27 23 5 11 5 27 23 22 26 20 29 28

a CO2/CH4 48 86 53 48 43 60 37 23 78 47

Ref. this work this work this work this work this work this work

HI 121 [31 [4] this work this work this work

do not hurt a support layer, which is not dissolved in the organic solvents. Polyimide membranes of 6FDA-BAPAF, 6FDA-DAP, and 6FDA-DABA showed the high

533

separation selectivity of up to 27 and up to 86 for CO2/N2 and CO2/CH4, respectively, and low CO2 permeability of 5 ~ 7 Barrers while 6FDA-TrMPD polymer exhibited a relatively low selectivity and high CO2 permeability in Table 2. CO2/N2 selectivities for the dense membranes of 6FDA-BAPAF, 6FDA-DAP, and 6FDA-DABA polyimides were similar to those for the composite membranes. The fluxes of the composite membranes were obtained without sacrificing a selectivity for CO2 gas separation. In Fig. 1. CO2 permeability for 6FDA-based polyimides decreased in order of 6FDA-BAPAF>6FDA-DAP>6FDA-DABA with increasing the inverse free volume fraction of the membranes, and CO2 diffusivity coefficient for the membranes decreased in order of 6FDA-DAP>6FDA-BAPAF>6FDADABA.

Fig. 1. Relationship between permeability & diffusion coefficient of CO2 for 6FDA-based polyimide dense membranes and inverse free volume fraction of the membranes.

Fig. 2. Plot of CO2 permeability and CO2/N2 selectivity over 6FDA-TrMPD polyimide dense membranes with different molecular weights. Molecular weight (Mw) PI4 = 180,000; PI5 = 102,000; PI6 = 53,000. The permeability and selectivity of 6FDA-TrMPD dense membranes with different

534

molecular weights were presented in Fig. 2. The 6FDA-TrMPD polyimide membrane with Mw = 102,000 showed a good separation performance. In the case of 6FDA-TrMPD dense membrane with low molecular weight of 53,000, its CO2 permeability and CO2/N2 selectivity were relatively low value due to reduction in the fee volume of the membrane. Despite higher CO2 permeability of 6FDA-TYMPD polyimide membrane, which was in agreement with the literature[7], the polymer membrane could not develop into a composite membrane because of limitation of solubility to coating solvent as listed in Table 1. Although 6FDA-HAB polyimide having two hydroxyl groups was a polar group-containing polymer, it was insoluble in alcohol and glycol ether. It indicates that a solubility of 6FDA-based polyimide to organic solvents is an important factor in preparation of a dense membrane or/and a composite one. 4. CONCLUSIONS 6FDA-based polyimide membranes having a polar group of hydroxyl or carboxyl in a diamine group could be applied to CO2 gas separation. Among the polymer membranes prepared in this study, 6FDA-BAPAF, 6FDA-DAP, and 6FDA-DABA polymer, which were soluble in 2-methoxyethanol, could be used in preparation of a dense membrane or/and a composite one. A solubility of 6FDA-based polyimides depended on organic solvents and was crucial for the preparation of different membrane types. The 6FDA-BAPAF polyamide dense membrane showed the highest CO2 permeability among above three polyimides, and CO2 diffusivity coefficient decreased with increasing inverse free volume fraction of the membranes. The 6FDA-TrMPD polyimide membrane with a proper molecular weight showed a good separation performances. ACKNOWDGEMENT This research(paper) was performed for the Carbon Dioxide Reduction & Sequestration Center, on of 21 st Century Frontier R&D Programs funded by the Ministry of Science and Technology of Korea. REFERENCES [1] B. Sea. Y.I. Park, K.H. Lee,/. Ind. Eng. Chem., 8 (2002) 290. [2] J.W. Rhim, J.R. Kim, Y.I. Park, K.H. Lee, J. Ind. Eng. Chem., 7 (2001) 299. [3] K. Tanaka, Y. Osada, H. Kita, K.I. Okamoto, J. Polym. Sci. Polym. Phys., 33 (1995) 1907. [4] H. Kawakami, M.Mikawa, S. Nagaoka, Macromolecules, 31 (1998) 6636. [5] S.H. Park, K.J. Kim, W.W. So, S.J. Moon, S.B. Lee, Macromol. Res., 11 (2003) 157. [6] K.J. Kim, S.H. Park, W.W. So, D.J. Ahn, S.J. Moon, J. Membr. Sci, 211 (2003) 41. [7] K. Tanaka, M. Okano, H. Toshino, H. Kita, and Okamoto, /. Polym. Sci., Polym. Phys,. 30 (1992) 907.


535

Chemical absorption of carbon dioxide into toluene solution of oxirane Sang-Wook Park3, Dae-Won Park, Tae-Young Kim, and Jae-Wook Leeb a

Division of Chemical Engineering, Pusan National University, Pusan 609-735, Korea

b

Department of Chemical Engineering, Sogang University, Seoul 121-742, Korea

Carbon dioxide was absorbed into the toluene solution phenyl glycidyl ether(PGE) as a oxirane and tricaprylylmethylammonium chloride (Aliquat 336, QX) as a catalyst using a semi-batch flat-stirred absorber at 85°C and latm to obtain a five-membered cyclic carbonate, phenoxy methyl ethylene carbonate. The reaction mechanism of oxirane - CO? reaction divided into three steps was used to obtain the reaction kinetics using the mass balance equations. The three steps are (i) a reversible reaction of PGE and QX to form complex, C\, (ii) reaction of Ci and carbon dioxide to form complex,C2, and (iii) dissociation reversible reaction of C2 to form QX and five-membered cyclic carbonate. The reaction rate constant in the reaction (ii) was obtained by the numerical solution of the mass balance equation using the measured molar flux of CO2 and the liquid-side mass transfer coefficient of CO2 at given concentration of PGE and QX. 1. INTRODUCTION Recently, the chemistry of carbon dioxide has received much attention[l], and its reaction with oxiranes leading to five-membered cyclic carbonate(oxirane-CO2 reaction) is well-known among many examples.[2,3] These carbonates can be used as aprotic polar solvent and sources for polymer synthesis.[4] In the oxirane-COb reaction, high pressure (5-50 atm) of CO2 has been thought to be necessary.[2] The oxirane-CCh reactions under atmospheric pressure have been reported[4] only recently. Many organic and inorganic compounds including ammines, phosphines, quaternary ammonium salts, and alkali methal salts are known to catalyze the oxirane-CO2 reaction. [3] Most purpose of these papers have been to show the reaction mechanism, the pseudo-first-order reaction rate constant with respect to the concentration of oxirane, and the catalyst dependence of its conversion. In the mass transfer accompanied by a chemical reaction, the diffusion may have an effect on the reaction kinetics. It is considered worthwhile to investigate the reaction kinetics of the gas-liquid heterogeneous reaction such as the oxirane-CO2 reaction. In this study, a chemical absorption mechanism of carbon dioxide into the toluene solution of PGE and Aliquat 336 was presented, from which the reaction rate constant of reaction with CO2 was obtained using the measured molar flux and liquid-side mass transfer coefficient of CO2 at 85°C and 1 atm. 2. THEORY

536 The overall reaction between oxirane and CO2 at atmospheric pressure of CO2 is presented as follows. O

/°\ y \

CO2(A) o

X p

Cat(QX)

R^(B)

> ^ (1) R The reaction path of reaction (1) considered in this study is assumed to be occurred as the following steps,[5] (1 B + QX -»==* C, ) A + C, - ^ >

C2

('»)

C, *=* QX + C (iii) At the steady state in the liquid film, the mass balances of the components in reaction (i)-(iii) and the boundary conditions are shown as follows, D.^-r.-O dz" + r ,.o

DC^4J.,,-,,.O

(4)

D r a

c

'

i»_ri dz 2

'

'

Dc,!O&l „,.,,.„ z = 0;[A] = [A] i; dB/dx = dC, /dx = dCJdx = dC/dx = 0 z = 5;[A] = 0,[B] = [B] 0 ,[C,] - [C,] 0 ,[C 2 ] = [C,] 0 ,[C] = [C]o where, r,-k1[B][QX]-k;[c1] r 2 =k 2 [A][C,]

(i

> (7) (8)

(y) (10)

r 3 =k 3 [C,]-k;[QX][C,] (11) Because it is impossible to get the solution of equations(2)-(6), the approximated solution shown as follows is used. Assuming that the reversible reactions (i) and (iii) are at equilibrium, the concentration of component Q at the initial time is obtained from the following equation, [C,] e =K e [B] f [QX] f (12) Where, Ke is the equilibrium constant, [B]f and [QX]f are the fed concentration of B and QX, respectively. The dimensionless mass balance equations (4) and (6) and the boundary conditions (7) and (8) are rearranged as follows. d2a w 7 (13) dx" d2c H ; (14) dx~

q

537

dc x = O;a = l , — = 0 dx x = l ; a = 0,c = l

(15) (16)

Where HA = Hatta number defined as ^k 2 [C,] c D A /k, , q = D c l [C,] e /(D A [A] ; ), x = z/8, a=[A]/[A];and c = [C,]/[C,] e . The reaction rate constant in reaction (ii), k2, is obtained using the concentration profile of CO2 in the liquid film, which is solved by the numerical solution of the differential equations (13)-(14) using the measured molar flux of CO2. 3. EXPERIMENTAL Absorption experiments were carried out in a semi-batch flat-stirred agitated vessel constructed of pyrex glass of 0.075 m inside diameter and of 0.13 m in height. Four equally spaced vertical baffles, each one-tenth of the vessel diameter in width, were attached to the internal wall of the vessel. A straight impeller with 0.034 m in length, 0.017 m in width and 0.005 m in thickness was used as the liquid phase agitator and located at the middle position of the liquid phase of 0.3 dm3. The absorption rates of CO? were obtained from the difference of the flow rate of CO2 between inlet and outlet of the vessel at 85°C. The molar flux of CO2 absorbed into the liquid was obtained from the measured rate of absorption. 4. RESULTS AND DISCUSSION Assuming that the reactions (i) and (ii) are at the reaction equilibrium, the concentration profiles of components A and d can be obtained by the numerical analysis of the equations (13) and (14). The estimated values concentration of component A and Ci are shown in Fig.l under the typical reaction condition such as [B]= 0.5 kmol/m3, [QX]= 0.05 kmol/m1, T= 85°C. As shown in Fig.l, the concentration of CO2 decreases and that of complex, Ci increases as the depth of the liquid increases. Using the measured molar flux of CO2 according to the change of fed concentrations of component, B and QX, the reaction rate constant of the reaction (ii), which can be obtained under the boundary condition such as [A]=0 at z=o, is 0.6 rrrVkmol s at 85°C.

Fig.l. Concentration profile of A and C r ([B]o=0.5 kmol/m1, [QX]o=0.05 kmol/m3, T=85"C)

538

The enhancement factor (|3) by the chemical reaction, which is defined as the ratio of absorption rate with chemical reaction to that without chemical reaction, is obtained from the number solution of equations (13) and (14) by the 4the order Runge-Kutta method.

Fig. 2. Effect of HA on |3 with q as parameter.

ACKNOWLEDGEMENTS This work was supported by KOSEF through ARC, the Brain Korea 21 project in 2003, and Brain Busan 21 program. REFERENCES 1. Inoue, S., In organic and bioorganic chemistry of carbon dioxide ; Inoue, S., N. Yamazaki, Eds, Kodansha Ltd, Tokyo, 1982. 2. Peppel, W.J, Preparation and properties of the alkylene carbonates, Ind. Eng. Chem., 50(5), 767-770 (1958). 3. Kihara, N , Hara, N , and Endo, T, Catalytic activity of various salts in the reaction of 2,3-epoxypropyl phenyl ether and carbon dioxide under atmospheric pressure,./. Org. Chem., 58, 6198-6202 (1993). 4. Rokicki, G, Cyclic dicarbonates as new monomers for the synthesis of poly(hydroxy ether)s, Makromol. Chem. 186, 331-337 (1985). 5. Aida, T. and S. Inoue, Activation of carbon dioxide with aluminum porphyrin and reaction with epoxide. J. Am. Chem. Soc, 105, 1304-1309 (1983).


539

A Study about CO 2 Absorption Process Design for Thermal Power Plant Flue Gas Sung-Youl Parka , Byoung-Moo Mina, Jong-Sup Leea and Hee-Moon Eumb a

Combustion Gas Clean-up Technology Center, Korea Institute of Energy Research, 71-2,

Jang-Dong, YuSeong-Gu, DaeJeon, Korea Radiation & Environment Group, Korea Electric Power Research Institute, 103-16, MoonJiDong, YuSeong-Gu, DaeJeon, Korea As an effort to develop CO2 removal technology by chemical absorption in packed column, continuous CO2 absorption facility using 30 wt% of monoethanolamine(MEA) aqueous solution was operated under various operating conditions like different packing materials and flow rates of flue gas and absorbent. Through the experiments, basic design data for CO2 absorption process, consideration factors for the selection of packing material and optimal flow rates of flue gas and absorbent, could be obtained. 1. INTRODUCTION Carbon dioxide is widely recognized as a major component of greenhouse gas contributing to global warming and produced in large quantity from various industrial sources, including fossil fuel firing electric power generation, steel production, chemical and petrochemical manufacturing, cement production, etc. Due to the increased CO2 mediated global warming problems, strong and adverse affects on the development of a large number of industries in the near future are expected. As a means of mitigating the global warming, removal of CO2 from industrial flue gas is considered important. The means of CO2 removal include absorption by chemical solvents, physical absorption, cryogenic separation, membrane separation and etc. Among these methods, CO2 absorption by alkanol amine(MEA) aqueous solution has been considered as the most efficient way and various research activities were conducted by this method and most commercial processes for the bulk removal of CO2 from * Corresponding author. Fax : +82-42-860-3691 E-mail address: [email protected]

540 gaseous streams involve the use of amines[l-3]. Traditionally, only random packings are used as the gas-liquid contacting media inside the absorption and regeneration towers. It has been suggested that using high efficiency column packings in these towers could substantially improve the efficiency of the gas treating process which reduce its capital cost. Hydrodynamics and mass transfer characteristics, including flooding capacities, gas and liquid mass transfer coefficients, interfacial area and liquid hold up are essential for evaluating the effectiveness of the tower packings and also important for the reliable design and operation of the CO2 absorption processes[4]. The purpose of this study was to find out the major operating parameters of packed column which can be used for the design of absorption process for CO2 removal from fossil fuel firing power plant flue gas. 2. EXPERIMENTAL 2.1. CO 2 absorption test facility Continuous CO2 absorption test facility is operated with same configuration as actual processes but has different capacities of CO2 absorption. As shown in Fig. 1, the facility is consists of oil boiler, flue gas supply pump, flue gas storage tank, four absorbers(packed column), two strippers, two heat exchangers, two coolers and three condensers. Flue gas(CC>2 13% , O2 3.8%) generated from the combustion of diesel oil is cooled, filtered, stored and then supplied to four absorbers by supply pump. The adjustments of flue

Fig. 1. Schematic diagram of CO2 absorption test facility.

541 gas flow rates and absorbent flow rates are achieved by the manipulation of flue gas supply valve installed at absorbers and rotation speed of absorbent circulation pump, respectively. Table 1 Properties of packing materials Pall Ring 10

Pall Ring 16

IMTP #25

Pro-Pak

Packing density(Kg/m Specific surface area(m2/m3)

410 950

396

336

362

188 171

1220

Void fraction(%)

93

95

97

96

Packing Material 3)

CO2 rich absorbent is heated to 95°C by heat exchanger and preheater and then supplied to stripper for absorbent regeneration. Regenerated absorbent is recycled to absorbers and the stripped CO2 pass through condenser and emitted to atmosphere. 2.2. Materials Pall Ring 10, Pall Ring 16, IMTP#25(Hanbal Masstech Co.) and Pro-Pak(Ace Glass Co.) were used for packing material and properties are shown in Table 1. As CO2 absorbent, 30 wt% of MEA aqueous solution was used. Weight concentration was controlled by the dilution of MEA and distilled water. After several times of cleaning by distilled water, the prepared aqueous MEA solution was supplied to the facility and circulated for the homogeneity of absorbent. 2.3. Analysis CO2 removal efficiencies were calculated from the CO2 concentrations of inlet and outlet of absorber. Humidity and fine particles of all gas sample were removed by the sample conditioner and then the samples were supplied to on-line CO2 analyzer, gas chromatograph (HP6890). CO2 concentrations were measured by a gas chromatograph equipped with packed column(Carboxen 1000) and thermal conductivity detector(TCD). 3. RESULTS AND DISCUSSION In this study, operating parameters affecting CO2 removal efficiency of randomly packed absorber with 30 wt% of MEA aqueous solution were measured. L/G(Liquid/Gas) ratio of CO2 absorption test facility is dependent on the flow rates of absorbent circulation and flue gas. The range of L/G ratio was 10-88 kg-mole liquid/kg-mole gas under the absorbent flow rates of 1-3 L/min. L/G ratio required to achieve above 90% of CO2 removal efficiency was 18-38 kg-mole liquid/kg-mole gas under different absorbent flow rates. Among properties of packing material, specific surface area is the major factor of CO2 absorption efficiency for

542

Fig. 2. The effect of L/G ratio on CO2 removal efficiencies.

Fig. 3. The effect of flue gas flow rate on CO, removal efficiencies.

packed column. Pro-Pak showed highest CO2 removal efficiency compared to other packing materials. Pro-Pak has the largest specific surface area and it means that the larger specific surface area gives larger gas-liquid contact area and consequently gives higher CO2 absorption efficiency. The experimental results are shown in Fig. 2 and 3. 4. CONCLUSION The results of experiment showed that 18-38 kg-mole liquid/kg-mole gas is required to achieve above 90% of CO2 removal efficiency under different absorbent flow rates and ProPak showed highest CO2 removal efficiency. Through this study, the basic design data for CO2 absorption in packed column could be obtained and it was expected that continuous absorption facility can be used for the test of durability and degradation of absorbent, CO2 absorption capacity and metal corrosion in chemical absorption of CO2. REFERENCES 1. Sheng H. Lin, Ching T. Shyu, Performance characteristics and modeling of carbon dioxide absorption by amines in a packed column, Waste Management, 19 (1999) 255-262. 2. T.Mimura, T. Suda, I. Iwaki, A. Honda, H. Kmazawa, Kinetics of reaction between carbon dioxide and sterically hindered amines for carbon dioxide recovery from power plant flue gases, Chem. Eng. Comm, 170 (1998) 245-260. 3. Sartori G, Savage D. W., Sterically hindered amines for CO2 removal from gases, Ind. Eng. Chem. Fundam., 22 (1983) 239-249. 4. Paitoon T, Research and development activities on high efficiency separation process technologies for carbon dioxide removal from industrial sources at university of Regina, Canada, Ener. Conser. Manage., 37 (1996) 935-940.


543

Membrane/PSA Hybrid Process for Carbon Dioxide Recovery at Low Concentration Jong-Ki Jeon1, Son-Ki Him2, Young-Kwon Park3*, Joo Sik Kim3, Jong-In Dong3, Seungdo Kim4, Ji Man Kim5, Seung-Soo Kim6, Kyung-Seun Yoo7 'Dept. of Chem. Eng., DongYang Univ., YoungJou, 750-711, Korea, 2Dept. of Chem. and Biomolecular Eng., KAIST, Taejon, 305-701, Korea, 3Faculty of Environmental Engineering, University of Seoul, Seoul 130-740, Korea, 4Dept. of Environ. System Eng., Hallym Univ., Chuncheon, 200-702, Korea, 5Dept. of Mol. Sci. & Tech., Ajou Univ., Suwon 442-749, Korea, 6Dept. of Environ. Eng., Dong Hae Univ., Dong Hae, 240-713, Korea, 7Dept. of Environ. Eng., Kwangwoon Univ., Seoul, 139-701, Korea. A hybrid system combining pressure swing adsorption (PSA) and gas membrane was investigated to recover low concentration CO2. At the first stage, a membrane unit concentrates the feed stream, and at the following stage, a PSA produces the purified gas rich in carbon dioxide. The membrane unit increases the concentration of the carbon dioxide from 15 to 40% by tube side feeding. The flow rate of permeate records more than 100 Nl/min, which is sufficient to fulfill the PSA capacity. The PSA cycle consists of pressurization, adsorption, blowdown, rinse, and desorption steps. The cycle configuration is determined to accomplish the purity of 99%. A higher heating rinse than normal rinse by 10°C resulted in 10% augments in recovery and productivity. 1. INTRODUCTION Many technologies to reduce the release of carbon dioxide from flue gases have been investigated. Evaluation of process economics is one of the most important factors in these fields. Capacity and discharge conditions at operational sites, however, play an important role in determining the process economics. Flue gases from power plants, iron foundries, and incinerators are estimated to represent 20-30% of total carbon dioxide emission, and the typical carbon dioxide concentration is less than 15% in mole base. The relatively low concentration adversely influences the economics of directly producing high purity carbon dioxide from flue gases compared with conventional approaches to pure carbon dioxide production.

* To whom correspondence should be addressed. FAX: +82-2-2244-2245. E-mail: [email protected].

544 The PSA and membrane processes are attractive for separation of carbon dioxide. The plant is simple, hence capital and operating costs are low at medium scale [1-3]. While PSA is suitable for producing higher purities (>99%), membrane process is very efficient for bulk separation. In case of PSA, it is necessary to add pre-concentration stage to produce higher purity of carbon dioxide than 99%. A PSA unit, however, requires a feed concentration greater than 25%. If a membrane process would be used, a serial configuration of modules with high permeation pressure is necessary. With the present generation of membrane processes, the PSA system still maintains an advantage in the high purity region while membranes become clearly advantageous when product purity requirements are less severe [2]. Therefore, it is reasonable to expect that an optimized gas separation process combining membrane and PSA would improve purity and/or recovery of carbon dioxide to compare with those of standalone systems. Although, several hybrid systems have been proposed in the literature for various applications [4-7], the study of hybrid system for CO2 recovery from flue gas has not been investigated yet. This study presents the performance of hybrid system for CO2 recovery consisting of PSA and membrane elements. 2. EXPERIMENTAL Membrane unit is composed of four hollow-fiber modules, of which fibers are made up with asymmetric polysulfone. To augment permeation capacity and diminish pressure drop, the four modules are arranged in parallel. Effective membrane area is about 40m . The PSA unit consists of three packed columns. Adsorbent is zeolite molecular sieve 13X manufactured by Grace Davison. The adsorbed amount of each gas is measured by the static volumetric method. Considering the rate and purity of product from membrane unit, the operating conditions of PSA are determined so that the product carbon dioxide concentration is not less than 99% by computer simulations. A simple cycle is employed; pressurization, adsorption, blowdown, rinse, and desorption. The cycle time is about 900 sec, and the bed length and diameter are 1.8 and 0.3m, respectively. Heating rinse is also applied to test whether or not it contributes to higher production rate as a result of increase in desorption. The feed gas with 15% carbon dioxide is introduced into membrane modules at a flow rate of 400 to 500 Nl/min and permeation pressure of 5 to 7 atm. To examine the effects of feeding conditions, the feed stream is provided to shell and tube sides, respectively. The permeating stream, concentrated to 38 -40% carbon dioxide, is then fed to PSA unit at adsorption step at a flow rate of 100 to 120 Nl/min. The part of product is employed as rinse gas enriches the purity of desorption gas. Fig. 1 shows the schematic diagram of hybrid system. 3. RESULTS AND DISCUSSION Fig. 2 shows the performance of membrane unit with respect to the pressure difference between shell and tube sides. The feed direction is applied to tube side as well as shell side. The purity and

545 recovery of carbon dioxide in permeate stream confirms that the feeding into tube side should be preferred. While the feeding strategy causes the enormous pressure drop in tube side, it completely precludes the concentration polarization, which is unavoidable in shell-side feeding. The membrane unit by tube side feeding allows the production rate of carbon dioxide to be 110 Nl/min with its purity up to 40% and recovery up to 60%. The results state that the membrane unit satisfies as a pretreatment or enrichment process for PSA. The production rate from membrane to PSA unit is 110 Nl/min, which is higher by 10% than predicted. The purity is evaluated to be around 40%, which shows a little variation with operating conditions. The performance of membrane used in this study remained constant during the 150 days.

Fig. 1. Schematic diagram of hybrid process for CO2 recovery

Fig. 2. Permeation efficiency for pressure differences with 15% carbon dioxide

The PSA operating on the prescribed condition allows production rate of 20 Nl/min with the purity of carbon dioxide higher than 99.5%, and its recovery of 45 to 50%. The low recovery of PSA unit is ascribed to the discharge of rinsing gas, and it is expected that some modification, for example, recycling it to membrane process, easily yield improved recovery.

546

Further efforts are recommended for concentrating the purged gas in a rinse step whose carbon dioxide concentration is estimated to be 30 to 80%. In the case of carbon dioxide PSA, the gas is usually purged with relatively low pressure close to atmospheric pressure. The high concentration, however, gives an impetus on membrane separation introducing vacuum in recovery side. In order to investigate the heating effect in rinse step, rinse gas was heated at a higher temperature by 10°C than that of normal rinse. Fig. 3 shows temperature variation inside column with time. The temperature distribution is relatively constant for the heating rinse, which seems to be due to the mitigation of a rapid temperature variation in top of column where CO 2 adsorption/desorption occurred well. The heating rinse also led to 10% augments in recovery and productivity, possibly thanks to ease of desorption resulting from increased volumetric rinse rate and temperature rise in the column. However, direct heating of adsorptive bed is not suitable due to the adverse effect that made it difficult for CO2 to adsorb on bed at adsorption step.

Fig. 3. Temperature profiles for the case of normal rinse and heating rinse

REFERENCES 1. D.M. Ruthven, Principles of Adsorption and Adsorption Processes, Wiley, New York (1984). 2. D.M. Ruthven, S. Farooq and K.S. Knaebel, Pressure Swing Adsorption, VCH, New York (1994). 3. R. T. Yang, Gas Separation by Adsorption Processes, Butterworths, Boston (1987). 4. M.B. Rao and S. Sircar, J. Membrane Sci., 110 (1996) 109. 5. T. Naheiri, K.A. Ludwig, M. Anand, M.B. Rao and S. Sircar, Sep. Sci. Tech., 32 (1997) 1589. 6. X. Feng, C.Y. Pan, J. Ivory and D. Ghosh, Chem. Eng. Sci., 53 (1998). 1689. 7.1.A.A.C. Esteves and J.P.B. Mota, Desalination, 148 (2002) 275.


547

Effect of Isotherm Selection on Performance Prediction of CO2 PSA Process Jong-Ki Jeon1, Son-Ki Ihm2, Young-Kwon Park3*, Joo Sik Kim3, Shin Do Kim3, Seungdo Kim4, Ji Man Kim5, Seung-Soo Kim6, Kyung-Seun Yoo7 'Dept. of Chem. Eng., DongYang Univ., YoungJou, 750-711, Korea,

2

Dept. of Chem. and

3

Biomolecular Eng., KAIST, Taejon, 305-701, Korea, Faculty of Environmental Engineering, University of Seoul, Seoul 130-740, Korea, 4Dept. of Environ. System Eng., Hallym Univ., Chuncheon, 200-702, Korea, 5Dept. of Mol. Sci. & Tech., Ajou Univ., Suwon 442-749, Korea, 6Dept. of Environ. Eng., Dong Hae Univ., Dong Hae, 240-713, Korea, 7Dept. of Environ. Eng., Kwangwoon Univ., Seoul, 139-701, Korea. Two adsorption isotherm models based on coupled Langmuir and ideal adsorbed solution theory (IAST) were applied to simulate CO2 PSA. CO2 adsorption amount on zeolite 13X was not precisely reflected by coupled Langmuir model, but by IAST. IAST is highly recommended for quantitative estimation of the productivity, recovery ratio, and purity in CO2 PSA using zeolite 13X. 1. INTRODUCTION The global warming due to CO2 accumulation in the atmosphere draws much attention. While several processes including chemical conversion, liquefaction, membrane, pressure swing adsorption (PSA), and hybrids of them have been suggested to treat CO2 effectively, PSA is considered to be a useful method to separate and recover CO2 from flue gas [1-3]. Numerous investigations in the field of PSA have been dedicated to analysis of recovery of weak adsorbates, especially hydrogen. In spite of process analogy to weak-adsorbate PSA, the strong-adsorbate carbon dioxide PSA is distinguished from the former, in view of significant adsorptive amount and relatively low-pressure operation with vacuum recovery. These distinctions prevent naive extensions of weak-adsorbate PSA to strongadsorbate ones. In previous work, a non-isothermal Linear Driving Force (LDF) model with a proper dimensionless formulation and a robust numerical algorithm was suggested [4,5]. To simulate CO2 PSA process from flue gas accurately, it is necessary to select a proper adsorption isotherm model. While most of simulation results were performed by coupled (or extended) Langmuir isotherm assuming ideal adsorption of gaseous mixture, its degree of discrepancy to CO2 PSA has not been tried [6-8]. To simulate CO2 behavior well, a relevant isotherm model must To whom correspondence should be addressed. FAX: +82-2-2244-2245 E-mail: catalica@,uos.ac.kr

548 be introduced in no time. In this study, two isotherm models based on coupled Langmuir and ideal adsorbed solution theory (IAST) were taken into consideration. The results of simulation were compared with experimental ones in view of productivity, purity, and recovery ratio. 2. EXPERIMENTAL The pilot CO2 PSA consists of subsidiary two PSA units. The first PSA serves concentration of feed stream, of which carbon dioxide fraction is usually less than 0.15. The second PSA carries out the main purification. In order to enhance recovery and process performance, a residual stream from the second PSA is returned to the first PSA. The PSA unit consists of three packed columns. Adsorbent is zeolite molecular sieve 13X manufactured by Grace Davison. The amount for each gas adsorbed is measured by the static volumetric method. A simple cycle is employed; pressurization, adsorption, blowdown, rinse, and desorption. The cycle time is about 660 sec, and the bed length and diameter are 1.8 and 0.3m, respectively.

3. MODEL 3.1 Adsorption isotherm model for binary system Langmuir parameters obtained by pure component was used for prediction of binary system. The coupled Langmuir isotherm for binary system is as follows [1-3].

The IAST is based on thermodynamic theory, and only pure component isotherms are needed to predict mixture equilibrium [1-3]. The basic equation of IAST in analogy with Raoult's law for vaporliquid equilibrium is written for a binary gas mixture of CO2 (component 1) and N2 (component 2): fVi = Pi(*. T)xi (2)

*(l-:yi) = *iiOr.r>(l-xi) 0) where Xi and yt are the mole fractions of CO2 in the solid and gas phases respectively, and Pf and P2° are the equilibrium pressures of pure gases at the spreading pressure. From the Gibbs adsorption equation, the spreading pressure is written in integration form:

Using the Langmuir isotherm for qf and from the equality of the same spreading pressures at equilibrium state, i.e. jti = n2, we obtain:

Q?ln(l + b1p

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