Reaction Kinetics and the Development and Operation of Catalytic Processes

Studies in Surface Science and Catalysis 133 REACTION KINETICSANDTHE DEVELOPMENTAND OPERATION OF CATALYTIC PROCESSES S...

Author: G.F. Froment | K.C. Waugh

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Studies in Surface Science and Catalysis 133 REACTION KINETICSANDTHE DEVELOPMENTAND OPERATION OF CATALYTIC PROCESSES

Studies in Surface Science and Catalysis 133 REACTION KINETICSANDTHE DEVELOPMENTAND OPERATION OF CATALYTIC PROCESSES

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S t u d i e s in S u r f a c e S c i e n c e a n d C a t a l y s i s Advisory Editors:

B. Delmon and J.T. Yates

Vol. 133

REACTION KINETICSANDTHE DEVELOPMENTAND OPERATION OF CATALYTIC PROCESSES Proceedings of the 3rd International Symposium, Oostende, Belgium, April 22-25, 2001

Edited by G.E Froment

TexasA&M University, Department of Chemical Engineering, College Station, TX 77843, USA

K.C.Waugh

UMIST, Department of Chemistry, Manchester, UK

I

2001 ELSEVIER Amsterdam - - London -- New York -- Oxford - - Paris -- Shannon - - Tokyo

E L S E V I E R S C I E N C E B.V. S a r a B u r g e r h a r t s t r a a t 25 P.O. B o x 2 1 1 , 1 0 0 0 AE A m s t e r d a m , T h e

Netherlands

92001 E l s e v i e r S c i e n c e B.V. A l l r i g h t s r e s e r v e d . This work is protected under copyright by Elsevier Science, and the following terms and conditions apply to its use: Photocopying Single photocopies of single chapters may be made for personal use as allowed by national copyright laws. Permission ofthe Publisher and payment of a fee is required for all other photocopying, including multiple or systematic copying, copying for advertising or promotional purposes, resale, and all forms of document delivery. Special rates are available for educational institutions that wish to make photocopies for non-profit educational classroom use. Permissions may be sought directly from Elsevier Science Global Rights Department, PO Box 800, Oxford OX5 1DX, UK; phone: (+44) 1865 843830, fax: (+44) 1865 853333, e-mail: [email protected] may also contact Global Rights directly through Elsevier's home page (http://www.elsevier.nl), by selecting 'Obtaining Permissions'. In the USA, users may clear permissions and make payments through the Copyright Clearance Center, Inc., 222 Rosewood Drive, Danvers, MA 01923, USA; phone: (+1) (978) 7508400, fax: (+1) (978) 7504744, and in the UK through the Copyright Licensing Agency Rapid Clearance Service (CLARCS),90 Tottenha m Court Road, LondonWl P0LP,UK; phone: (+44) 207 6315555; fax: (+44) 207 6315500. Other countries may have a local reprographic rights agency for payments. DerivativeWorks Tables of contents may be reproduced for internal circulation, but permission of Elsevier Science is required for external resale or distribution of such material. Permission of the Publisher is required for all other derivative works, including compilations and translations. Electronic Storage or Usage Permission ofthe Publisher is required to store or use electronically any material contained in this work, including any chapter or part of a chapter. Except as outlined above, no part of this work may be reproduced, stored in a retrieval system or transmitted in any form or by any means, electronic, mechanical, photocopying, recording or otherwise, without prior written permission of the Publisher. Address permissions requests to: Elsevier Science Global Rights Department, at the mail, fax and e-mail addresses noted above. Notice No responsibility is assumed by the Publisher for any injury and/or damage to persons or property as a matter of products liability, negligence or otherwise, or from any use or operation of any methods, products, instructions or ideas contained in the material herein. Becauseof rapid advances in the medical sciences, in particular, independent verification of diagnoses and drug dosages should be made.

First edition 2001 Library of Congress Cataloging-in-Publication Data Reaction kinetics and the development and operation of catalytic processes : proceedings of the 3rd international symposium, Oostende (Belgium), April 22-25,2001 / editors, G.F. Froment, K.C. Waugh. p. cm. -- (Studies in surface science and catalysis ; 133) Includes index. ISBN 0-444-50559-8 1. Catalysis--Congresses. 2. Chemical kinetics--Congresses. 3. Chemical processes--Congresses. I. Froment, Gilbert F. II. Waugh, K. C. III. Series. TP 156.C35 R28 200 l 660'.2995--dc21 2001023023 ISBN: 0-444-50559-8 I S S N : 0167-2991

O T h e p a p e r u s e d in t h i s p u b l i c a t i o n m e e t s t h e r e q u i r e m e n t s ( P e r m a n e n c e of Paper). P r i n t e d in T h e N e t h e r l a n d s .

of ANSl/NISO Z39.48-1992

Ontvn,8 Introduction

XlII

KEYNOTE LECTURES Adventures in catalytic nanospace: 'seeing' spillover in-situ for the first time

M. Bowker, R.A. Bennett, A. Dickinson, D. James, R.D. Smith and P. Stone

From first-principles to catalytic turnovers: ethylene hydrogenation over palladium

19

Development of microkinetic expressions by instationary methods

41

Self sustained oscillations over copper in the catalytic oxidation of methanol

57

Novel reactor configurations and modes of operation

71

New methodologies and reactors for catalytic process development

87

M. IVeurock, E. Hansen, D. Mei and P.S. Venkataraman

F.J. Keil

L B6ttger, B. Pettinger, Th. SchedeI-Aliedrig, A. Knop-Gericke and R. Schlo'gl D.Luss

M.P. Harold, P.L. Mills andJ.F. Nicole

KINETICS: STEADY STATE OPERATION Hysteresis kinetics of propene oxidation characterized by a Ag-Re supported membrane reactor

105

Kinetics of oxidative dehydrogenation of LPG to olefins on Dy-Li-CI-Zr-O catalyst

113

De Donder relations and the theory of reaction routes

123

A kinetic study of NO reduction by CO over a NiO/AI203 catalyst

131

M. Kobayashi, T. Kanno, J. Horiuchi, S. Hoshi, IV. Hattori and J. Togawa

M.L. Kaliya, O.V. Malinovskaya, M.V. Landau, M. Herskowitz and P.F. van den Oosterkamp I. Fishtik and R. Datta

T.N. Angelidis and M. Papapetrou

vi Reaction kinetics of the hydrodenitrogenation of methylcyclohexylamine over fluorinated NiMoS/AI203 catalysts

139

Use of the "dusty-gas" model for the analysis of ethylbenzene oxidehydrogenation process

147

Using low melting point alloy intrusion to quantify pore structure: Studies on an alumina catalyst support

155

L. Qu and R. Prins

D. Ardissone, A. Bachiller and J. Orejas

L. RuftTno, R. Mann, R.J. Oldman, S. Rigby and S. Allen

CATALYTIC PROCESSES Mass and heat transfer effects in the kinetic modelling of catalytic cracking

165

Kinetics of ethylene polymerisation over CrY zeolites

173

Investigations of the selective partial oxidation of methanol and the oxidative coupling of methane over copper catalysts

181

A new kinetic model for hydrodesulfurization of oil products

189

Low-temperature, carbon-catalyzed, solvent-washed, trickle-bed sulfuric acid process

195

Butane oxidation to maleic anhydride over a VPO catalyst following the riser regenerator approach

205

Effects of particle size and modified SAPO-34 on conversion of methanol to light olefins and dimethyl ether

211

P. Hagelberg, V. Alopaeus, K. Lipi~inen, J. Aittamaa and A.O.I. Krause Y. Zhang, X. Wang, M.A.N.D.A. Lemos, F. Lemos, R.T. Henriques and M.M. Marques

H.-J. W61k, G. Weinberg, G. Mestl and R. Schlo'gl

T. Mejdell, R. Myrstad, J. Morud, J.S. Rosvoll, P. Steiner and E.A. Blekkan

P.L. Silveston, M.A. Panthaky, K. Duval, R.V. Vladea, A. Lohi and R.R. Hudgins

St. Hess, H. Freund, M.A. Liauw and G. Emig

M.G. Abraha, X. Wu and R.G. Anthony

MOLECULAR MODELING From DFT calculations to dynamic Monte Carlo simulations. The reactivity of CHx on the Ru(0001) surface I.M. Qb/~'c~ F. Frechard, A.P.3. Jansen and R.A. van Santen

221

REACTORS : NON STEADY STATE OPERATION Non-steady state operation of trickle-bed reactors

E.G. Boelhouwer, H.W. Piepers and A.A.H. Drinkenburg

231

vii

Performance enhancement of a microchannel reactor under periodic operation

239

A dynamic study of steam-methane reforming

247

Model-based optimization of the periodic operation of the Fischer-Tropsch synthesis R.M. de Deugd, S.M. Ypma, F. Kapteijn, ~:.IV~.Meeuse,J.A. Moul/jn and P.J.T. Verhe/jen

255

Direct determination of periodic states of cyclically operated packed bed reactors

263

Dynamics and control of a radial-flow ammonia synthesis reactor

271

A.Rouge and A. Renken

M.A. EI-Bousifti and D.J. Gunn

T.L. van Noorden, S.IV/. Verduyn Lunel and A. Bliek

N.S. Schbib, M.N. Pedernera and D.O. Borio

MISCELLANEOUS Mass transfer through the gas-liquid interface at the presence of adsorbed active phospholipid monolayer

283

Oxygenates and olefins from catalytic partial oxidation of cyclohexane and n-hexane in single-gauze chemical reactors R.P. O'Connor and L.D. Schmidt

289

Mechanistic studies of hydroxyl radical-induced catalytic wet oxidation of dyehouse effluents at atmospheric pressure

297

Reaction pathways of photocatalytic oxidation of PCE in water with the TiO2film on glass

303

Liquid phase hydrogenation of naphthalene on Ni/AI203

309

Effect of NO and oxygen upon the deactivation of Cu-ZSM-5 in NO decomposition

317

L. Grado~ and T.R. Sosnowski

D.-K. Lee, D.-S. Kim and S.-C. Kim

D.-K, Lee and I.-C. Cho

P.A. Rautanen, M.S. Lylykangas, 3.R. Aittamaa and A.O.I. Krause

V.I. P~rvulescu, E. Segal, B. Delmon and P. Grange

KINETICS: TRANSIENTS Non steady-state production of hydrogen from natural gas: Experiments and modeling E. Odier, Y. Schuurman, H. Zanthoff, C. M//let and C. Mirodatos

327

viii

Dynamic study of methane interaction with active sites involved in the total oxidation of methane over Pd/AI203 catalyst

333

A non-stationary kinetics approach for the determination of the kinetic parameters of the protolytic cracking of methylcyclohexane

341

Kinetic modelling of transient NO reduction by CO in the presence of 02 over an automotive exhaust gas catalyst

349

Transient kinetics of lsNO-decomposition on Pt/AI203

357

Transient kinetics of the propene oxidation over silver catalysts

365

s. Fessi, A. Ghorbel, A.Rives and R. Hubaut

v. Fierro, J i . Duplan, J. Verstraete, K Schuurman and C. Mirodatos

J.M.A Harmsen, J.H.B.J. Hoebink and J. C. Schouten

A.R. Vaccaro, G. Mul and J.A. Moulijn

A. Zwijnenburg, IV~.Boer, G. Mul, IV~.Makkee and J.A. Moulijn

CATALYST

n9Xe NMR of absorbed xenon and IH NMR imaging" new methods to study the diffusion of gaseous hydrocarbons in a fixed bed of zeolite P. N'Gokoli-Kekele, Iv~.-,4. Springuel-Huet, 3.-L. Bonardet, I-M. Dereppeand3. Fraissard

375

Catalytic activity of carbon nanotubes and other carbon materials for oxidative dehydrogenation of ethylbenzene to stryrene

383

Heterogenous catalysis at high temperature for space applications

391

Catalyst design for reactions with synthesis gas

399

IV. Maksimova, G. Mestl and R. Schldgl M. Balat-Pichelin andJ.M. Badie

G.P. Valenga and E.S. Gongalves

REACTORS : STEADY STATE OPERATION

Enhancement of non isothermal autocatalytic reactions by intraparticle diffusion

411

An adsorptive reactor for operation of exothermic series reactions

419

M. Grzesik andJ. Skrzypek A.J. Kodde and A. Bliek

POSTERS CATALYSIS

Catalytic activity and characterization of quaternary ammonium(methylstyrene-co-styrene) resin in an organic solvent/alkaline solution H.-S. Wu and C.-S. Lee

429

Catalyst development for methanol and dimethyl ether production from blast furnace off gas

435

Heterogeneous photo- and thermal catalytic oxidation of CO: effects of metal deposition

445

Harmonization of the empirical information on the de-NOx catalysts for the comparison of the catalyst performance through reaction modeling

453

Synthesis of cumene (isopropylbenzene) from diisopropylbenzenes in the presence of benzene using triflic acid as catalyst at room temperature

459

J. Yagi, T. Akiyama and A. Muramatsu

I. Ozen and D. Uner

D. Liner and S. Kaya

M.C. AI-Kinany, S.H. AI-Khowaiter and F.H. AI-Malki

D I F F U S I O N A N D MASS T R A N S F E R

Diffusion analysis of cumene cracking over ZSM5 using a jetloop reactor

465

Simulation of the effect of mass transfer limitations in complex gas-liquid reactions

471

Evaluation of external diffusional effects in a microreactor for SRGO hydrotreating

477

A study of a phase transfer catalytic reaction between N-butyl bromide and sodium phenolate in an oscillatory baffled reactor

481

Kinetics of the main and side .reactions of the methanol oxidation over iron molybdates

489

P. Schwan, P.J. Henry and K.P. MSIler

M. Pisu, A. Cincotti, G. Cao and F. Pepe

L.C. CastaEeda-L6pez,F. Alonso-Mart[nez and J. Ancheyta-Ju~rez

B. Wilson, X. Ni and D.C. Sherrington

A.P. Vieira Soares, M. Farinha Portela and A. Kiennemann MECHANISMS

Mechanistic investigations of oxidation of purine and pyrimidine base components of nucleic acids by bromamine-B in aqueous alkaline medium: a kinetic approach

495

Activity-acidity relationships in solid acid catalysis a quantum chemical study

501

N. Vaz and Puttaswamy

X.. Wang, M.A.N.D.A. Lemos, F. Lemos and F. Ram6a Ribeiro

Catalytic reactions mechanisms with some nonlinear conservation laws E.S. Patmar and N.I. Koltsov

507

PROCESSES Modelling diffusion, cracking reactions and deactivation in FCC catalysts

509

Oxidehyrogenation of ethylbenzene to styrene on P-O-Ni-Mn/alumina catalysts

515

Hydrodechlorination of chlorobenzene-tetrachloroethylene mixtures over a Pd/AI203 catalyst E. L6pez, S. Ord6fiez, F.V. D/ezand H. Sastre

521

Kinetic study of the liquid-phase hydrogenation of 1,3-butadiene and n-butenes on a commercial Pd/AI203catalyst

527

F. L6pez-Isunza, N. Moreno-Montiel, R. Quintana-Sol6rzano, J.C. Mor~no-Mayorga and F. Hern~ndez-Beltr~n

D. Ardissone, A. Bachiller, M. Ponzi and J. Orejas

N.O. Ardiaca, S.P Bressa, J.A. Alves, O.M. Mart[nez and G.F. Barreto

KINETICS Kinetics and mechanism of ruthenium(III) and osmium(VIII) catalyzed oxidation of dopamine with bromamine-B in acid and alkaline media

535

Kinetic models for esterification of methacrylic acid using n-propanol and isopropanol

541

Kinetics models for esterification of levulinic acid with 2-ethylhexanol using different catalysts

547

Time resolution of the TAP reactor

553

A novel catalyst preparation and kinetic study on the dechlorination of chlorinated hydrocarbons

559

Kinetics of copolyesters of poly(butylene terephthalate), hydroquinone diacetate and terephthalic acid: a simple rate model for catalysed synthesis in melt

565

Catalytic oxidation of sulfite/bisulfite in a falling-film absorption column

575

Puttaswamy and N. Vaz

M. Grzesik, J. Skrzypek and M. Witczak

M. Grzesik and 7:. Gumula D. Wang and Z. Li

Y.S. Cho, J.-C. Park, B.-S. Choi, J. Moon and J. Yi

A. AI-Haddad and J. Mathew

M. Vorbach, R. Marr and M. Siebenhofer

Influence of fluorination on the kinetics of the HDN of o-toluidine over tungsten sulfide catalysts ex ammonium tetrathiotungstate

581

Kinetics of the hydrodenitrogenation of cyclohexylamine over suifided NiMo/y-AI203

587

H2-TPR-kinetics; case study on the reduction of a CrOJAl2Os catalyst ./.M. Kanervo and A.O.I. Krause

593

Kinetic study of methane combustion over Lao.gCeo.lCOO3

599

Theoretical determination of the kinetic parameters of a reaction intermediate by degeneration of the precursor process

605

An easy methodology for estimating kinetic constants in complex kinetic models

611

Experimental validation of a kinetic model for naphtha reforming

615

Kinetic analysis of enzymatic esterification of fatty acids and ethanol

619

Catalytic combustion of methane over Pd/y-AI203 in a monolithic reactor: kinetic study

625

Software functionality assessment for kinetic reaction model development, model discrimination, parameter estimation and design of experiments

631

-Authors' index

637

M. Sun and R. Pdns

F. Rota and R. Prins

R. Auer, L. Warnier and F.C. Thyrion

M. Elkhatib, C. Duriche, R. Metz, J.R. Vignalou and H. Delalu

3. Ancheyta-3u~rez and R. 5otelo-Boy~s

3. Ancheyta-3u~rez and E. Villafuerte-Madas

v.u. Miguel, G.C. Trubiano, G. Perez, D.O. Borio and A.F. Errazu

E. Ldpez, C. G[gola, D.O. Borio and V. Bucal~

R.3. Berger, 3. Hoorn, 3. Verstraete and3. W. Verwijs


...

Xlll

The 3rd International Symposium on "Reaction Kinetics and the Development and Operation of Catalytic Processes" will be a trendsetter. The Keynote Lectures, authored by top scientists, will cover a broad range of topics like fundamental aspects of surface chemistry, in particular dynamics and spillover, the modeling of reaction mechanisms, with special focus on the importance of transient experimentation and the application of kinetics in reactor design. Most of these topics are reflected in the communications, oral or poster. Fundamental and applied kinetic studies are well represented. More than half of these deal with transient kinetics, a new trend made possible by recent sophisticated experimental equipment and the awareness that transient experimentation provides more information and insight into the microphenomena occurring on the catalyst surface than steady state techniques. The trend is not limited to purely kinetic studies since the great majority of the papers dealing with reactors also focus on transients and even deliberate transient operation. It is to be expected that this trend will continue and amplify as the community is getting more aware of the predictive potential of fundamental kinetics when combined with detailed realistic modeling of the reactor operation. The response to our call for studies related to commercial processes is rewarding. We hope that this attitude will further develop. The field has reached a state at which commercial processes can be analyzed and modeled in their full complexity without resorting to the gross simplifications of the past and thus contribute to a more optimal and safe operation in today's more protected environment. G.F. Froment Department of Chemical Engineering Texas A & M University

K.C. Waugh Department of Chemistry UMIST


[GcturGr


Studies in Surface Science and Catalysis 133 G.F. Froment and K.C. Waugh (Editors) Published by Elsevier Science B.V., 2001

Adventures in Catalytic Nanospace: 'Seeing' spillover in-situ for the first time. Michael Bowker, Roger.A. Bennett 1, Amanda Dickinson, David James, Rupert D. Smith and Peter Stone Centre for Surface Science and Catalysis, Dept. Chemistry, Univ. Reading, Reading RG6 6AD. Email [email protected] 1

now at Physics Dept., University of Reading.

Abstract. The interaction between metal nanoparticles and a support may result in a changed reactivity pattern from what might be expected, for a variety of reasons. Here we give two examples of this, one where the active site involves the support, and one where transfer of material from the metal particle to the support occurs. In the first example, namely the photocatalytic reforming of methanol on Pd/TiO2 at ambient temperature, we believe that the reaction is determined by sites at the interface between the metal and the support. In the second example, and in a related vein, we have used STM to image Pd nanoparticles on a rutile TiO2 (110) surface, a model catalyst, and have identified spillover of oxygen from the Pd particles onto the surrounding TiO2 surface. This is witnessed by growth of new layers of TiO2 in the immediate vicinity of the nanoparticles, due to reaction between spillover oxygen and interstitial Ti3+ ions which migrate to the surface. We propose that this is a method of oxygen storage akin to that used in automobile catalysts.

Introduction.

The modem era of surface science began in the 1960's with the aim of simplifying complex heterogeneous systems by using well-defined, single crystal samples of known surface atomic structure, and utilising very simple adsorbates, often diatomics. In the new millenium surface science has advanced to studying more complex surfaces, with much more complex adsorbates, under a wide range of conditions, including high pressure and liquid phases. This is the proper situation for a subject which is as widely based as the term 'surface science' implies. One of the major areas of impact of surface science has been in the field of catalysis, as recently acknowledged by one of the most prestigious practitioners of catalysis in the 20th century, namely Michel Boudart [1], who says" .. the methods of surface science have permeated theory and practice of heterogeneous catalysis. Some key findings of surface science include the importance of surface reconstruction [2] and of the diffusion of weakly held precursors in surface reaction kinetics [3]. Much of this fundamentally important information has been obtained using single crystals, and it is imperative that this kind if work continues. However, it is also very important to recognise that such model surfaces are 2-D catalysts, whereas the industrial catalyst is usually a 3-D surface, as depicted in fig 1, and this may have different properties to a single crystal, especially when the nanoparticle is very small (sayl-2nm.). These differences are 1. If a particle is very small it is no longer a metallic conductor, because the conduction band splits into individual levels, the spacing of which is given approximately b y AE =EtCn Thus a lnm diameter particle (ca. 16 atoms) has energy levels split by approximately 0.3eV, or 30kJ mole 1, which makes electron excitation a significantly activated process.

Very small particles are not metallic

IMetai i

Edge sites

Support

Spillover, Reverse Spiliover

Electron transfer

Figure 1. Schematic model of a metal nanoparticulate surface, showing several of the ways in which it may differ from a bulk metal.

lOO t 80 A

"E 60 3

m 40

0

0

0

0

0

..,= (II

IZ:

20 0 0.001

oq 0 0.01

0.1 Wt. Loadingi%

1

10

Figure 2. The dependence of rate upon Pd loading for the photocatalytic reforming of methanol at 320K. Circles are data points, while squares and the fitted line are data from the theoretical model [7].

2. The particle-support interface may affect the reactivity, and the following

effects are all enhanced in several ways at small particle size due to the increased number of atoms at the interface compared with those elsewhere in the particle, i)

The active site may be at the interface between the support and

the particle, an example of Which is given below, and is therefore likely to be very different from the behaviour of the bulk metal. ii)

Spillover may occur, either from or to the support, which will then affect the reaction kinetics

iii)

Electron transfer can result in charging of a small particle and an

alteration of the electronic structure For these reasons it is important to study supported nanoparticles in a rigorous way using the methods of surface science. A number of surface scientists have embarked on this path in recent years, and reviews of some of this work are available [4-6]. In this paper we will highlight some of the work at the University of Reading in the last two years which relates to catalysis by nanoparticles. In the following section we will highlight the importance of point 2i above, namely photocatalytic reforming of methanol at ambient temperature. We will then go onto describe some of the beginnings of surface science work on models of the catalyst used in that work using STM as the probe of surface structure. We can directly observe the spillover of active species from the metal to the support in such experiments.

Photocatalysis Involves Sites at the Interface between the Precious metal and Support.

We have recently reported on the photocatalytic reforming of methanol under anaerobic conditions using Pd/TiO2 catalysts [7-9]. The rate of this reaction has a very marked dependence on loading of Pd as shown in fig. 2. The rate is quite high at very low loadings of Pd, whereas TiO2 alone, or even TiO2 loaded with other non-precious metals, has a very low rate [8,9]. As loading increases, so there is a slow increase in the reaction rate, with a maximum at around 0.5 wt% Pd and a sharp drop in rate above that loading to

give zero initial rate at 2 wt%. This can be modelled on the assumption that only those sites at the interface are the active ones and these are given by the following equation Na = N2/3.(12n2V)1/3.{ 1-(cos'l[d/(3V/2Nn)l/3]/O)}

Here Na is the number of active sites, N is the number of particles, V is the volume of metal in the sample, d is the particle separation, and 0 is the angle between particles (30 degrees for a hexagonal array).This has the functional form of the real data. However the maximum should be at much higher loadings, which has lead us to conclude that the active region is at a perimeter related to the Pd particles, but remote from them. The details of this modelling will be given in a future publication, though the outline is in the earlier paper [7]. This modelling then leads us to a mechanism for the reaction as shown in figure 3, which schematically illustrates then course of the reaction. Essentially methanol dehydrogenation takes place on the Pd, which then becomes poisoned with CO in the absence of light, but which is relieved by light-induced oxidation of the CO by active oxygen species from the TiO2 support created by band gap excitation. Water is oxidised then on the support. In the next section we report beginnings of surface science work which is related to this, concerning the fabrication of model catalysts, their thermal stability and the first studies in the Reading group of their reactivity [ 10-13]. We have not yet studied photocatalytic reactions on these materials.

Fabrication and Properties of Model Catalysts A number of methods can be used to make model catalysts consisting of an array of nanoparticles on a flat, well-defined (usually single crystal) support, and these are outlined in a review by Campbell [4]. We have used both metal-vapour deposition (MVD) and metal- organic chemical vapour deposition (MOCVD) [13]; here only data for the former method will be presented. Fig 4 shows images for a layer deposited at 300K, and heated to several temperatures [ 10]. It is clear that sintering occurs at the highest temperatures, but at intermediate temperatures it appears that the main effect is one of wetting the surface - the particle diameter increases, but the particle height

@

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=

c.

9

Light > 3 . 2 e V

Ca)

CH3OH

oO

-.r

I

|

C]

,

CO2

v /OOoo

c |

i

(~

.....

~

o~

~++,

ooOO ~ & ~o

~~~,o::l

Figure 3. Proposed model for the photocatalytic reforming of methanol on Pd/TiO2 [7-9]. This involves dissociation of the methanol on the Pd surface to yield hydrogen and adsorbed CO (a). The CO poisons further reaction in the absence of light, but band gap excitation produces an active oxygen species which strips it off (b,c). The resulting vacancy on the Pd is filled by another methanol molecule, and the anion vacancy in the titania is filled by water splitting, both producing hydrogen (d,e). The cycle is complete at f, which is identical to b.

o

~

Figure 4. STM images of a Pd/TiOz layer deposited at 300K and heated to various temperatures. Significant sintering only begins above 700K. After annealing to high temperature (F) a range of particle types are produced, which clearly depend on the exact nature of the substrate close to them, and which exhlbit facetting.

rz)

10 decreases (fig. 5). Frenken has shown that these sintering processes occur with high fractional order power laws [14], as has generally been reported earlier for powdered materials [ 15]. Fig 4f shows that particles can facet, but apparently have strongly varying morphology, depending upon the local environment. Thus the Wulff construction approach [ 16] may be invalid here, because the nature of the metal-support interface may play a dominant role in dictating high temperature particle structure. In this way we can make nanoparticles which, at least in their size distribution have a strong similarity to those found in powdered, high surface area catalysts We can also control the size distribution by varying the loading on the surface and by controlling the annealing temperature.

Spillover, Encapsulation, SMSI and Oxygen Storage Seen for the First Time It is important that we understand something about the properties of the support TiO2 (110) crystal. The treatment in UHV to clean the sample (a major surface impurity initially is Ca) results in slight bulk reduction. We estimate that for the stoichiometric formula TiO2.• x is always less than 0.001 in the work reported here. However, this has a substantial effect on i) the bulk conductivity and ii) the surface structure. The evolution of surface structure with use is shown in fig 6 and shows a transformation from the expected (lxl) termination to (lx2), to (nx2), where n is large, for instance 12 [17]. This is often called the 'cross-linked (lx2)'. With even more bulk reduction crystallographic shear planes can terminate at the surface in long-range, straight line steps [18]. The (nx2) structures are relatively easily oxidised by annealing in low pressures of oxygen, but this oxidation occurs in very particular ways. It appears that, at least at elevated temperature, the dominant defect is not an anion vacancy, but is an interstitial Ti 3§ Part of the evidence for this is that, as the surface is heated at 673K in low pressures of oxygen, new layers of TiO2 grow at the surface, and they grow in a special way [19]. The (nx2) fills in with material to make the denser surface layer of the (lxl), then the (nx2) structure forms a new layer on top of this, then fills in to (lxl), and so on in a cyclic fashion, building up new layers on the surface. This behaviour can be viewed as a movie elsewhere [20]. Since this is the case we might expect to see novel effects when a nanoparticulate layer is exposed to oxygen, as indeed we do. Figure 7 shows the effect of dosing oxygen

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14 onto the model Pd/TiO2 (110) surface, where the particles are adsorbed onto the (nx2) cross-linked surface of TiO2 (110) [11,12]. A number of phenomena can be witnessed here:1) Spiilover of oxygen from the Pd to the support occurs (see fig.7C for instance) seen as a patch of additional TiO2 grown preferentially around the Pd particle. This layer grows in a slightly different way from that on the clean surface. Instead of oscillating between (nx2) cross-linked structures and (lxl) the region adjacent to the particle always grows in a (lxl) structure. The growth of these layers is faster here than in areas remote from the particle due to an enhanced rate of adsorption on the Pd compared with the titania, and an enhanced supply of oxygen to the TiO2 surface in the adjacent region. 2) Encapsulation and burial. The layers adjacent to the particle build up so much as to completely bury the particle after 7 layers have grown. After this the rate of titania growth slows considerably. 3) Oxygen storage. Clearly by this process oxygen is stored adjacent to the Pd particle; this is reminiscent of the use of CeO2 as an oxygen storage medium in car catalysts. Although CeO2 has a different bulk structure (fluorite vs. rutile), it also possesses a 3+/4+ redox couple like TiO2, which is the source of the oxygen storage capability. Of course, this storage ability is only the property of MO2.x, that is sub-stoichiometric oxides.

It is not likely that this phenomenon is akin to classical SMSI, since those are observed after reducing the catalyst. However, the SMSI has probably been seen, by Diebold et al. For Pt [21] and by us for Pd, as shown in fig 8. It appears that the surfaces of the Pd particles, after heating above 600K are decorated by a new structure, which is related to the titania [22]. We believe this to be a TiO layer by structural analysis and comparison with bulk structures. This layer severely modifies the adsorption capability of the surface [23].

Conclusions. Model catalysts can now be made routinely, and these have proved amenable to surface science studies of their properties. By the use of high

15

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Advances in computational chemistry and molecular simulation have also reached the stage whereby they can be Used to develop more advanced and robust kinetic models for catalytic systems. First-principle quantum chemical methods, for example, are being used to routinely calculate thermochemistry and kinetics for gas phase chemistry with accuracies on the order of 0.5-2 kcal/mol (26) The situation for the kinetics of heterogeneous reaction systems has also improved considerably but due to the complexity in these systems the accuracy is more on the order of 5-7 kcal/mol (27,28) In addition to the advances in raw computational power and more accurate quantum chemical methods, molecular simulation tools have also made great progress and can now begin to capture detailed quantum chemical results. In this paper, we discuss some of our recent results in coupling ab initio density functional quantum chemical methods along with dynamic Monte Carlo algorithms in order to simulate a detailed myriad of competing molecular transformations that govern the kinetics on metal surfaces. Stochastic models enable us to take full advantages of the atomic detail that is offered from the detailed analytical spectroscopies as well as the information from ab initio quantum mechanical modeling efforts since they can explicitly account for atomic surface structure. Rather than averaging the details over the entire surface by invoking deterministic methods, the Monte Carlo approach follows the position of each atom on the surface as a function of reaction conditions, As such, we can begin to map out the molecular transformations along with kinetic behavior. In particular, we will focus on the hydrogenation of olefinic intermediates. Ethylene is used as a model toward understanding the hydrogenation of complex olefin and aromatic intermediates (29) It is therefore relevant to various hydrocarbon processes. In addition, understanding ethylene hydrogenation kinetics is also important in the manufacturing of high purity ethylene feeds for polyolefin production. It is necessary to remove acetylene from the ethylene feeds for polyolefin

23 manufacturing because acetylene leads to a host of different processing problems. Ethylene feeds are first hydrotreated to selectively convert acetylenic intermediates to ethylene. The chemistry, however, is very sensitive to process conditions. Unselective routes lead to ethylene hydrogenation to form ethane. This will sharply decrease the overall selectivity, and can also lead to runaway reactors since ethylene hydrogenation is fairly exothermic. The mechanism for olefin hydrogenation has been debated for well over 30 years. The general chemistry is thought to involve a Horiuti-Polanyi (3~ like scheme. Hydrogen dissociatively adsorbs producing chemisorbed atomic hydrogen on the surface. Ethylene co-adsorbs in either di-~ or n-bound arrangement and then reacts with hydrogen to form the ethyl surface intermediate. The ethyl intermediate can subsequently react with a second surface hydrogen atom to form ethane or dissociate back to form ethylene and atomic hydrogen (29). Although the overall paths appear to be well established, the nature of the active site and the details of what controls the chemistry are still openly debated in the literature. A number of different schemes have been devised in order to appropriately fit the kinetic data (31,8,9,7). They include some of the following ideas: 1) hydrogen and ethylene compete for the same surface sites, 2) hydrogen and ethylene reversibly adsorb into separate islands and reaction occurs at the edges of these islands, and 3) hydrogen is adsorbed on two different sites- one is competitive with ethylene adsorption while the other site is not. There is little evidence, however, to support these ideas other than they provide a better fit the kinetic data. Fitting of the kinetic data alone, however, does not say very much about the mechanism. Westerterp (32) for example, recently showed that over eight different models could be readily correlated with the kinetic results for hydrogenating ethylene and acetylene mixtures. Rather than start with a preconceived notion about the mechanism, we have carried out a comprehensive series of first principle calculations in order to determine the intrinsic activation barriers, heats of adsorption and overall reaction energies for various speculate routes. Transition state theory as well as classical statistical mechanics are used to subsequently establish the Arrhenius factors and thus provide the intrinsic rate constants. One of the critical issues in using first-principles data involves the great difference between the ideal reaction environment that is chosen for the model and the actual environment under operating conditions. Most of the quantum chemical results that have been published over the past decade have examined the zero or low coverage limit in order to avoid introducing coverage effects. In order to model the active surface under reaction conditions, however, we need to explore the effects of the reaction environment. In particular, we focus here on the interactions between species on the surface. We explore these interactions in detail to determine their effects on adsorption and surface reactivity. First principle calculations enable us to easily calculate the interaction between two species coadsorbed to the surface. As we begin to think about more realistic systems, it very quickly becomes apparent that this essentially becomes impossible to do for more than two species because of the infinite number of ways in which adsorbates can assemble on a surface. For example, figure 3 depicts a single snapshot from kinetic Monte Carlo simulation that we performed for NO decomposition over Rh.

24

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We have chosen a smaller unit cell size and also the chemistry in this system is quite limited. Despite these simplifications, this system still contains hundreds to thousands of different configurations for just this single snapshot in time. As we begin to look at the myriad of events that can occur, it becomes clear that it would be simply impossible for us to carry out all of these configurations directly from first-principle calculations. Instead, we use set of first-principle calculations to map both the nearest and next-nearest neighbor interactions. This information is then used to parametrize a semiempirical model which can be called internally within the simulation. The Monte Carlo simulation therefore, builds up the database of calculated rate constants, calculates (through model equations) the influence of the

environment and subsequently simulates the outcome of different reaction steps. The apparent rate constants are subsequently used to calculate a reaction probability distribution, ri/Zri which is then used in the simulation to establish which events occur as we proceed in time. In addition to modeling coverage effects, the Monte Carlo algorithm also enables us to track individual sites, i.e. atop, bridge, hollow sites. It is well known from the surface science literature that while one site may prove to be strongest adsorption site, it may or may not be active site that drives the kinetics. We therefore intrinsically allow for site-specificity in the model.

II. Developing The Intrinsic Kinetic Data Base from Ab Initio Methods We have performed a number of cluster and periodic slab quantum chemical calculations to determine the intrinsic energetics for sequence of adsorption and reaction steps that control ethylene hydrogenation and decomposition. The methods are briefly described below. The (33-36) details can be found in the following papers:

A) OFT Calculations Large nineteen atom palladium clusters were used to model the Pd(111) surface (34). All calculations were performed using spin-polarized Density Functional Theory (DFT) calculations with nonlocal gradient corrections for the correlation and exchange energies. The local exchange correlation potential was modeled using the Vosko-Wilk-Nusair (VWN) potential along with nonlocal gradient corrections from Becke (37) and Perdew (38)for the exchange and correlation terms respectively. DFT-optimized double zeta gaussian type orbital basis sets which include

25 spin polarization were employed for all atoms. (39,40) The spin states were fully optimized for all calculations. The Pd cluster was held fixed at the metal positions of the ideal surface. The adsorbates however were allowed to fully optimize.

B. Periodic Pd(l l l) Calculations To examine higher coverage situations and the effect of surface relaxation, periodic DFT calculations were performed. The results were found to be quite consistent with calculations on the large Pd19 clusters (34) which helps to confirm the cluster results. The Pd(111) surface was described by constructing a supercell which is translated along the lattice vectors that lie in the surface plane to describe ( . ~ x.~/~) R30 ~ and (2x2) surfaces. The surface is defined here by three layers of Pd atoms. A vacuum layer of 10 A was placed above the adsorbate surface to avoid any electronic interactions between the slabs. In previous calculations we examined the effect of the number of metal layers on the binding energies for hydrogen on Pd(111) and found that the energy changes were less than 1 kcal/mol as the slab thickness was increased to three layers and beyond. In addition, the structural changes which occur in going beyond three layers were found to be negligible. Three layers of Pd, therefore, appeared to provide a computationally efficient and tractable model of the Pd(111) surface (41). The Kohn-Sham (KS) equations were solved using a plane wave basis set with a maximum kinetic energy of 40 Rydberg. Eighteen special k-points (Bloch vectors) were used in the description of the first Brillouin zone. The special k-points were chosen based on symmetry by the method developed by Chadi-Cohen. All calculations were performed self consistently by using non-local exchange and correlation potentials in the form of the Perdew-Wang 91 Generalized Gradient Corrections. Scalar relativistic corrections were described through the use of norm-conserving pseudopotentials. The first two palladium layers were allowed to relax in an effort to understand the effects of surface relaxation. The bottom layer was held fixed at the bulk structure of Pd(111). Adsorbates were then spaced in either a ( ~ x,~/~) R30 ~ or a (2x2) surface structure. The ( , ~ x,~/3) R30 ~ structure was found to lead to repulsive interactions between neighboring ethyl intermediates. The (2x2) structure was therefore used to help eliminate any through-space repulsive interactions. The geometries of the first two layers of the surface along with the adsorbates structure were completely optimized. (34)

C. The Catalytic Potential Energy Surface The low coverage adsorption energies for the relevant intermediates are shown in Table 1. Ethylene binds in both the n and di-cy adsorption states. At the low coverage, the di CH20 + H2 ARH~ +84 k J / m o l (dehydrogenation) > CH20 + H 2 0 A R H ~ = -159 k J / m o l (2) CH3OH + Y2 02 (oxidehydrogenation) (3) CH3OH + 3/2 02 > CO2 + 2 H20 ARH~ = -674 k J / m o l (total oxidation I) (4) CH20 + 02 > C02 + H20 ARH~ = -519 k J / m o l (total oxidation II) The reactions 1 a n d 2 lead to the desired partial oxidation p r o d u c t formaldehyde, w h e r e a s the total oxidation reactions p a t h w a y s 3 a n d 4 yield the m a i n b y - p r o d u c t s C02 a n d H20. The total oxidation is of specific relevance for the entire reaction s y s t e m due to the h e a t evolved from this highly exothermic reaction.

58

In the surface science based studies on the methanol oxidation over single crystalline Cu(110), Wachs and Madix [1, 2] have shown that the active state of the copper surface for the methanol oxidation was a partially oxidized copper surface exhibiting nucleophilic oxygen ad-atoms. The oxygen activated the surface for methanol adsorption and removed hydrogen released b y water formation on the surface via a low-energy reaction pathway. From investigations of the active phase of a bulk metalliC copper catalyst u n d e r reaction conditions of the partial methanol oxidation by m e a n s of in situ XAS at the oxygen K-edge and copper L2-,L3-edges it was concluded that the partial oxidation of methanol to formaldehyde is catalysed by a copper plus oxygen phase where oxygen atoms probe defects of the copper lattice, which represent the catalytically active sites [3, 4, 6]. Self-organised, periodic p h e n o m e n a have been found in several heterogeneous catalytic systems [7]. Such systems exist in a state far from thermodynamic equilibrium and exhibit at constant adjusted parameters of temperature, pressure and feed gas flow self-sustaining catalytic oscillations. Oscillations in the methanol-oxygen-copper-system were observed for the first time by H. Werner et al. [8]. They observed with stoichiometric methanol/oxygen mixtures (2:1) over Cu chips in a tube reactor at 660 K rate oscillations with a cycle duration of ca. 3.5 min. Characteristic were sharp maxima (several seconds) of the methanol conversion leading to maxima in the formaldehyde and the CO2 productions, followed by a decline of the methanol conversion and the CO2 partial pressure whilst the formaldehyde partial pressure showed a slow decrease to the next period. Oscillations were also found in other reactions, catalysed by copper or copper oxides, e. g. town gas combustion over supported Cu [9]; selective oxidation of propene to acrolein over Cu oxides [10, 11]; hydrogenation of PhNO2 (supported Cu [12]) and CO oxidation, catalysed by supported CuO [131. It is the purpose of the present work to investigate copper surfaces with regard to t h e self-organised periodic phenomena in methanol oxidation and to corresponding near surface and bulk properties. Therefore, oxygen/methanol pre-treated copper is studied u n d e r the conditions of methanol oxidation at atmospheric pressure in situ with m a s s spectrometry (MS), Raman spectrometry and visual video records in order to couple the oscillating variations of the catalytic conversion with macroscopic catalyst (near-) surface and bulk properties. The near surface related electronic structure of a model catalyst is studied by using core-level (XPS) and valence band (UPS) photoemission spectroscopy u n d e r clean UHV conditions in order to characterise the oxygen segregation process and copper sub-oxide formation. Both the photoemission study in this work and the EXAFS investigation [5] point to the importance of the participation of copper bulk oxygen species in the formation of the catalytic active surface. Such a participation was predicted from Werner et al. [8] and verified for the silver catalyst [14, 15] also active in the partial oxidation of methanol.

59

2.Experimental E x p e r i m e n t s with a copper (111) single crystal, cleaned by s t a n d a r d p r o c e d u r e s at T = 573 K in UHV were carried out in a modified Leybold LHS 12 MCD s y s t e m equipped with facilities for XPS, UPS a n d ISS. The base p r e s s u r e in the a n a l y s e r c h a m b e r was 1 x 10 -10 mbar. 120 W of Mg K non m o n o c h r o m a t i s e d radiation (1253.6 eV) a n d a fixed a n a l y s e r p a s s energy of 108 eV c o r r e s p o n d i n g to a resolution of 1.0 eV were u s e d to obtain XPS data. HeI s p e c t r a were a c q u i r e d u s i n g a n a l y s e r p a s s energy of 6 eV. Surface compositions were calculated from O i s a n d Cu 2p p e a k a r e a s after s u b t r a c t i o n of a Shirley-type b a c k g r o u n d u s i n g the cross sections from Ref. [16]. The Cu(111) s a m p l e was pre-oxidised with p u r e oxygen (Linde, purity: 99.9999 %) (30 min, T = 623 K, po2 = 1 mbar). Thereafter, the oxidised sample was step by step treated with m e t h a n o l (up to 2 5 0 0 0 0 L, at a p r e s s u r e of 1" 10 -4 m b a r , Merck, purity: 99.9 %, "for spectroscopy") a n d finally s p u t t e r e d by He + ions In situ m a s s s p e c t r o m e t r y investigations of the c o p p e r / o x y g e n / m e t h a n o l s y s t e m d u r i n g rate oscillations were carried o u t u s i n g as catalyst a polycrystalline copper foil (Goodfellow, purity: 99,99 %, surface size: 67 m m 2, thickness: 0.1 mm). To clean the fresh Cu s a m p l e s they were cyclical treated at high t e m p e r a t u r e s (720 - 770 K) with 02 a n d H2 prior the m e t h a n o l / o x y g e n t r e a t m e n t . The e x p e r i m e n t s were c o n d u c t e d in a microreactor equipped with a calibrated q u a d r u p o l e m a s s s p e c t r o m e t e r (Balzers QMS 200) coupled to the reactor outlet via a capillary (RESTEK Silicosteel i m p r e g n a t e d , i n n e r diameter: 0.25 m m , length: 2 m). The working p r e s s u r e in the reactor was 1 bar. The t e m p e r a t u r e of the sample (copper foil) was held c o n s t a n t l y by a n electronic t e m p e r a t u r e controller (the thermo couple was s i t u a t e d between the sample a n d a ceramic h e a t plate). The feed gases were supplied by Linde a n d h a d the following purities: 02, 99.995 %; He: 99.9999 %. Helium w a s s a t u r a t e d with m e t h a n o l (purity: 99.9 %, Merck, "for spectroscopy").The well mixed feed gases were a d m i t t e d t h r o u g h t h e r m a l m a s s flow controllers in the reactor. In situ R a m a n s p e c t r a of the copper s a m p l e were recorded d u r i n g reaction by a R a m a n s p e c t r o m e t e r from Dilor (Labram II). The s p e c t r o m e t e r integrated He-Ne laser was u s e d to obtain excitation radiation employing a laser power of 25 mW. The laser light was focussed to the sample t h r o u g h the q u a r t z window of the reactor by u s i n g a microscope. Video records were t a k e n from the sample with a CCD PAL colour video c a m e r a (SONY XC-999P).

60

Cu(111)

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The pre-oxidation of the Cu(111) sample resulted according to XPS analysis in a mixture of C u O / C u 2 0 . The total methanol treatment shifted the stoichiometry from Cui.20 to Cu3.20. The stoichiometry changed strongly for 0 L to 15000 L (1L= l x 1 0 -6 m b a r applied for ls) m e t h a n o l exposures (Cui.20 to

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Cu2.20). Whereas the stoichiometry Cu2.1- 2.20 remained a l m o s t u n c h a n g e d in the range from 15000 L to 100000 L m e t h a n o l t r e a t m e n t , the stoichiometry c h a n g e d more strongly u p to Cu3.20 for t r e a t m e n t s of 100000 L to 2 5 0 0 0 0 L m e t h a n o l . Valence b a n d spectra are p r e s e n t e d as difference d a t a between successive t r e a t m e n t steps to highlight the spectral changes. The Heldifference spectra were fitted [17] in the range from -0.5 to 4.9 eV BE by a linear combination of Cu-, Cu20- a n d CuO- reference spectra (Fig. 1). The

61 t r e a t m e n t u p to 15000 L led to a reduction of CuO to Cu20 (Fig. 1, C- B). The HeI difference s p e c t r u m for 50000 L - 15000 L m e t h a n o l t r e a t m e n t indicated despite a unmodified stoichiometry of Cu2.20, the r e d u c t i o n of CuO to a Cu20 like species a n d also a decrease of the a b u n d a n c e of metallic copper. Also the difference s p e c t r u m 150000 L - 100000 L m e t h a n o l shows similar changes. The difference s p e c t r u m of the i n t e r m e d i a t e m e t h a n o l t r e a t m e n t 5 0 0 0 0 L to 100000 L ( E - D ) i n d i c a t e s s u r p r i s i n g l y inverted changes, i. e. the d e c r e a s e of Cu20 features a n d the increase of a diffuse spectral c o m p o n e n t a n d of metallic Cu features. The HeI-difference s p e c t r u m of the last m e t h a n o l t r e a t m e n t (250000 L - 150000 L methanol) revealed a reduction of Cu20 to metallic Cu. However, a suitable fit of this difference s p e c t r u m is obtained only with a linear combination of Cu, Cu20 a n d the diffuse additional species. The r e d u c t i o n of Cu20 to Cu is inhibited, w h e r e a s this is not the case for the initial reduction of CuO to Cu20. The not inhibited CuO reduction is indicated by the s h a r p decline of stoichiometry from 0 L to 15000 L m e t h a n o l t r e a t m e n t (Fig. 1, A - C). This indicates t h a t CuO is highly unlikely species to exist u n d e r all practical reaction conditions applied here. A r e t a r d e d r e d u c t i o n of Cu20 by m e t h a n o l / o x y g e n m i x t u r e s was already observed by H. W e r n e r et al. [8]. They ascribed this finding to a possible generation of a copper sub-oxide t r a n s i e n t state on the way of the Cu20 reduction to Cu. The HeI difference s p e c t r a reveal also c h a n g e s of the chemical s t r u c t u r e for the t r e a t m e n t steps between 15000 L (C) a n d 100000 L (E) methanol. The s p e c t r a were fitted by linear c o m b i n a t i o n s of reference spectra of Cu a n d the known, stable copper oxides CuO a n d Cu20. It was n e c e s s a r y to include a n additional diffuse spectral c o m p o n e n t well compatible with a sub-oxide containing oxygen a t o m s in a w e a k chemical bonding to copper in order to fit the s p e c t r a of progressing reduction. The surface composition c h a n g e d d u r i n g reduction in the intermediate compositional range alternatingly between Cu20 --~ CuxO + Cu a n d CuxO --~ Cu20 + Cu. The fit r e s u l t s with the linear c o m b i n a t i o n of Cu, Cu20 a n d CuO show significant differences c o m p a r e d to the experimental d a t a at a r o u n d 6 eV binding energy. Emission at this binding energy is typical for oxygen species [18, 19]. This difference is particularly high after the m e t h a n o l deposition in the range of 15000 L to 150000 L methanol. The emission at this energy position c h a n g e s less t h a n the emission at lower binding energy. Hence the intensity of this emission is not exclusively linked with the Cu 3d b a n d from the spectral features of the k n o w n copper c o m p o u n d s Cu, Cu20 a n d CuO. This is an i n d e p e n d e n t hint to oxygen species not b o u n d in the k n o w n copper oxide p h a s e s of CuO or Cu20, with w e a k i n t e r a c t i o n s with copper, typical for e.g. a d s o r b e d oxygen species [18] or copper s u b oxide which can be i m p o r t a n t for the catalytic oxidation of m e t h a n o l [3,8]. 3.2.Self Sustained Oscillations-General

Rate oscillations were observed in the range of m e t h a n o l to oxygen ratios 1.0... 6.0 a n d a t e m p e r a t u r e range --640 ... 800 K. Cycle times r a n g e d from

62 ~1 to 5 min d e p e n d i n g on the m e t h a n o l / o x y g e n ratio a n d the s a m p l e t e m p e r a t u r e . An increase of t e m p e r a t u r e led to the d e c r e a s e of the periods a n d of the a m p l i t u d e of oscillations. With increasing 02 c o n c e n t r a t i o n s the periods b e c a m e longer. The copper catalyst consisted u n d e r the conditions of the p r e s e n t e x p e r i m e n t s mainly of copper oxide. At higher m e t h a n o l oxygen ratios, the rate oscillations were coupled with visible oscillations of the oxide state of the copper foils. At lower m e t h a n o l to oxygen ratios the rate oscillations also proceeded with optically p e r m a n e n t l y oxidised copper. In a g r e e m e n t with the surface analytical d a t a the oxidised surface u n d e r in-situ conditions m a i n l y consists of Cu20 according to R a m a n s p e c t r a recorded from the copper s a m p l e d u r i n g the oscillations. The formation of Cu20 w a s a p p a r e n t in m o s t s p e c t r a from the p e a k at ~220 cm -1 [20, 21, 22]. However, formation of additional CuO or metallic Cu c a n not be excluded b e c a u s e Cu metal is generally R a m a n inactive a n d CuO is k n o w n to be invisible by a laser excitation with 632 n m [23]. F r o m the c a r b o n b a l a n c e of the catalytic d a t a it o c c u r r e d t h a t no o t h e r species t h a n the deep oxidation p r o d u c t s a n d formaldehyde were formed in detectable a m o u n t s . Neither formic acid nor a n a d s o r b a t e which would deactivate the surface ( c a r b o n a c e o u s deposits) were observed.

3.3. Activationlpre-treatment of the copper samples A m e t h a n o l - o x y g e n p r e - t r e a t m e n t of the c o p p e r - c a t a l y s t s at high temperatures is n e c e s s a r y to obtain catalytic oscillations in the c o p p e r / o x y g e n / m e t h a n o l - s y s t e m . For the p r e - t r e a t m e n t the s a m p l e c a n be either treated with m e t h a n o l - o x y g e n m i x t u r e s or a l t e r n a t i n g with oxygen a n d m e t h a n o l at t e m p e r a t u r e s above 730 K. In a m e t h a n o l - o x y g e n - h e l i u m flow (methanol to oxygen ratio around 2:1, T > 730 K) the c o p p e r / o x y g e n / m e t h a n o l s y s t e m always oscillated within a few m i n u t e s after a d j u s t m e n t of the p a r a m e t e r s . After the p r e - t r e a t m e n t the s y s t e m oscillated also at lower s a m p l e t e m p e r a t u r e s down to N 650 K in m e t h a n o l oxygen-helium mixtures. An activation can also be achieved by a n e x t e n d e d t r e a t m e n t of several h o u r s at lower t e m p e r a t u r e s (N 670 K) in a m e t h a n o l oxygen mixture. After this low t e m p e r a t u r e t r e a t m e n t long r e d u c e d p h a s e s of the Cu sample u p to 30 s were observed, w h e r e a s after the high t e m p e r a t u r e p r e - t r e a t m e n t the r e d u c e d state h a s also at low sample t e m p e r a t u r e s a very short lifetime of ~0.2 s. Different cleaning p r o c e d u r e s of the copper s a m p l e s like polishing with d i a m o n d paste, cyclic oxidation (oxygen) a n d r e d u c t i o n (hydrogen) at t e m p e r a t u r e s u p to 770 K as well as rinsing of the s a m p l e s with different solvents like methanol, acetone a n d w a t e r h a d no positive influence on obtaining the oscillations.

3.4. Variation of the oxygen-methanol proportion At c o n s t a n t sample t e m p e r a t u r e T = 720 K the m e t h a n o l - o x y g e n flow ratio in the gas feed was gradually d e c r e a s e d from 1.6 to 1.1, respectively

63 from 7.9 vol-% to 12.2 vol-% 02. The m e t h a n o l c o n t e n t was held c o n s t a n t at ~12.8 vol-%. In general, the cycle times b e c a m e longer with d e c r e a s i n g m e t h a n o l to oxygen ratios (Fig. 2). At m e t h a n o l - o x y g e n ratios from 1.6 to 1.3 (phase I, Fig. 3) the sample showed a n oscillating oxidation state (see video pictures, at a m e t h a n o l oxygen flow ratio of 1.2 (phase II, Fig. 3) the sample h a d been p e r m a n e n t l y oxidised while catalytic oscillations still sustained. Below a m e t h a n o l - o x y g e n ratio of 1.1, no rate oscillations were observed.

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1.4 1.2 Methanol/02 (Educts)

FIG. 2. Rate oscillations with a methanol/O2/He flow over Cu: The m e t h a n o l conversion a n d the cycle time related to the m e t h a n o l : oxygen ratio. P h a s e I: oxide state oscillations of the Cu sample, p h a s e II: sample p e r m a n e n t l y oxidised.

1

Typical for the visible reduction process of the s a m p l e surface is a reduction front travelling over the entire surface a n d leading to a p p a r e n t l y metallic copper (video pictures). The lifetime of the r e d u c e d state is very short (N1 s). Shortly after the metallic state, a progressing surface re-oxidation could be followed by eye as a s u c c e s s i o n of various colours due to light interference in thin Cu20 films of increasing thickness. The m a x i m a l surface layer t h i c k n e s s c a n be e s t i m a t e d by the interference colours [24, 25] of at least 300 m o n o l a y e r s Cu20. After surface reduction of the sample the m e t h a n o l conversion increases strongly, i.e. the catalytic activity of the metallic sample is s u b s t a n t i a l l y larger t h a n FIG. 3. Volume fractions t h a t of the thick oxidised state. After surface a n d m e t h a n o l conversion (calculated from mass .5 .... . spectra). M e t h a n o l / O 2 / H e [Phase I I ', IP"~e Ill flow over a copper sample 60 4 (Cu foil), Tsample = 720 K. ~o ~ The volume fractions of 3 CO2 a n d formaldehyde as 40 N 0 well as the methanol 2 conversion exhibit rate ~~ ~ oscillations. The para1 20 " C m e t e r s in p h a s e I and p h a s e II differ slightly in 0 ,i . ,I , 50 55 60 65 the m e t h a n o l to oxygen Time ! min ratios: m e t h a n o l : o x y g e n reduction the CO2 yield increase m u c h = 1 : 1 . 3 / 1 . 2 (phase I/II), stronger t h a n the formaldehyde yield - the ~77 % He, ptotal 1 bar. =

64 selectivity to formaldehyde decrease significantly. As shown in Fig. 3 the amplitudes of the CO2 production (total oxidation activity) are substantially larger t h a n the amplitudes of the formaldehyde production. A methanol to oxygen ratio of 1.2 giving rise to a permanently oxidised sample induced only weak rate oscillations in agreement with the low activity of bulk oxide for catalytic action. Under the reaction conditions also the steam reforming of methanol (CH3OH + H20 > CO2 + 3 H2 ARH~ +62 kJ/mol) is possible, because water (ca. 6 vol-%) exist in the reactor. This reaction is catalysed by copper. It is well known from literature [26] that only metallic copper is the active site for this reaction channel. Since the H2 volume fraction does not strongly increase with the metallic episodes we conclude t h a t steam reforming did not take place indicating that the reduction was incomplete giving a further hint to the existence of a sub-oxide with a p p a r e n t metallic lustre. 3.5. Transition from an oscillating surface oxide state to a permanently oxidised sample during gas phase oscillation

In the following the transition of the visible oscillating oxide-metal state ("phase I") to a permanently oxidised state ("phase II") of the catalyst surface (see Figure 3) will be discussed. ' i i i I i i

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FIG. 4. Line profile comparison of the methanol conversion (a), the CO2 / formaldehyde volume fraction (b) and the formaldehyde selectivity (c) (calculated from MS) of one period of phase I (oxidation state oscillations) and phase II (permanently oxidized sample). (parameters cf. Fig. 3)

65 This transition was observed after the decreasing of the m e t h a n o l to oxygen ratio from 1.3 to 1.2. P r o n o u n c e d gas p h a s e oscillations were found in both states of the surface (cf. m a s s spectra, Fig. 3). This indicates t h a t the oscillating macroscopic oxide to metal transition is not n e c e s s a r y to s u s t a i n the gas p h a s e oscillations. Thus it is of interest to know if the gas p h a s e oscillation characteristics are identical for p h a s e I a n d p h a s e II. A c o m p a r i s o n of the line profiles (calculated from m a s s spectra) of the m e t h a n o l conversion (Fig. 4a) a n d the yields in CO2 a n d in formaldehyde (Fig. 4b) a n d of the formaldehyde selectivity (Fig. 4c) show similar line profiles whenever the sample is visibly oxidised. The differences between p h a s e I a n d p h a s e II can be solely explained by a modification of the catalyst activity due to the sample reduction in p h a s e I leading to an extra feature s u p e r i m p o s e d over the f u n d a m e n t a l rate variation. A strong d e p e n d e n c e of the cycle time on the m e t h a n o l : 02 ratio was found in the range of 1.5 to 1.2. During the transition from p h a s e I to p h a s e II the e n h a n c e m e n t of the cycle time as a function of the methanol-oxygen ratio does not change, i. e. the cycle time is i n d e p e n d e n t from the transition from an oscillating oxide sample state to a p e r m a n e n t l y 0xidised sample. T h u s it can be concluded t h a t the rate oscillations in p h a s e I a n d p h a s e II have their origin in the same m e c h a n i s m . The macroscopic visible periodic sample oxidation-reduction m u s t be a side effect. Obviously, the oxidation a n d reduction cycles of the copper during the oscillations at a lower oxygen content regenerate the catalyst, while with a p e r m a n e n t l y oxidised sample the catalytic activity decreases c o n t i n u o u s l y indicating t h a t copper oxide Cu20 is a catalyst poison. The same conclusion was also reached from earlier steady state in-situ m e a s u r e m e n t s u s i n g NEXAFS as probe for the oxidation state of the copper [3-6]. Both in p h a s e I a n d p h a s e II (Fig. 3) the m a g n i t u d e s of the a m p l i t u d e s of CO2 are s u b s t a n t i a l l y larger t h a n those of formaldehyde. Thus, u n d e r lean conditions the oscillations of the partial oxidation of m e t h a n o l are probably a side effect which is in s h a r p c o n t r a s t to the observations of Werner et al. [8] who operated in rich conditions. With a higher m e t h a n o l partial p r e s s u r e the selectivity for formaldehyde is higher t h a n t h a t for CO2 a n d also the oscillation a m p l i t u d e s for the formaldehyde volume fraction increase (e.g. Fig. 11). Moreover, the sign of the c h a n g e s in the gas p h a s e volume fractions of formaldehyde a n d of CO2 in a single oscillation period are opposite, s u c h as after re-oxidation (e.g. after the p r o n o u n c e d CO2-peak at 53 min in Fig. 3) the formaldehyde volume fraction rises a n d the CO2 volume fraction falls. This observation strongly point to the existence of different active sites for total a n d partial oxidation, an observation which was also found in the insire NEXAFS experiments [3].

3.6. Modification of the gas-phase heat conductivity In order to reduce the g a s - p h a s e h e a t conductivity, 45.7 % of the He feed gas was replaced by the same flow of Ar (~ - h e a t conductivity - at T = 600 K: E(He) = 252.4 m W / K m, MAr) = 30.6 m W / K [27]). The other p a r a m e t e r s remained u n c h a n g e d .

66 The gas exchange led to an explicit modification of the oscillatory behaviour. The m a s s spectra show oscillations with comparatively long cycle times (approx. 4 min) a n d large a m p l i t u d e s in the case of high g a s - p h a s e h e a t conductivity (Fig. 5a). On the other h a n d , periods are s u b s t a n t i a l l y shorter a n d the a m p l i t u d e s on average smaller at low g a s - p h a s e h e a t conductivity (Fig. 5b). Close inspection of the m a s s spectra reveals a folded p a t t e r n of a long period of a b o u t 5 with a series of short periods with cycle times in the second range, reminiscent to the observations of Werner et al. [8]. In both experiments, the sample was always visibly reduced for a s h o r t time, w h e n the activity j u m p of the long oscillation process occurred. FIG. 5. Change of the rate oscillations with variation of the gas p h a s e h e a t conductivity by partial s u b s t i t u t i o n of He feed gas by Ar: Methanol to oxygen ratio 2.4 : 1 (flow), sample: polycristalline copper foil, T = 688 K. Top: Feed gas 87 vol-% He, bottom: feed gas mixture of 47 vol-%He a n d 40 vol-% Ar.

The bulk sample t e m p e r a t u r e w a s controlled b u t the surface of catalyst O may have been at temperatures deviating significantly from the m e a n 2 bulk value. The s t e a d y state between a c o m b i n a t i o n of exothermic a n d e n d o t h e r m i c reactions (see above) a n d the 0 5 10 15 20 h e a t t r a n s p o r t t h r o u g h the gas p h a s e which is affected Time / min by the heat capacity of the gas mixture d e t e r m i n e s the surface t e m p e r a t u r e . Since the conditions in the sample b u l k are likely to be similar in both experiments the long period oscillations in both experiments are probably c a u s e d by a m e c h a n i s m i n d e p e n d e n t of the h e a t dissipation t h r o u g h the gas phase. The short-time processes, on the other h a n d are significantly more p r o n o u n c e d in their a m p l i t u d e s w h e n the h e a t dissipation t h r o u g h the gas p h a s e is reduced. It is likely t h a t the m o s t exothermic reaction CH30H + 3/2 02 > CO2 + 2 H20, A R H ~ = -674 k J / m o l Is gaining a significant control over the whole system u n d e r conditions of 4

-.,,..

,,...

. v J

67 poor h e a t dissipation. This conclusion is in good a g r e e m e n t with general catalytic conjecture t h a t selective oxidation requires efficient control of energy t r a n s p o r t .

4. Discussion of possible origins of the oscillatory behaviour The gas p h a s e oscillations m a y be controlled by a process at the catalyst surface (red-oxidation) or in co-operation between the surface a n d the bulk (re-structuring). It is proposed t h a t the latter is effective, b e c a u s e in case of a sole surface process at least the cycle time should be drastically influenced by the visible sample reduction a n d re-oxidation, since thereby the surface is completely re-arranged. The h y p o t h e s i s of an oscillation m e c h a n i s m controlled by the catalyst bulk state is s u p p o r t e d by the experiment with variable h e a t conductivity of the gas phase: the b r a n c h of the p h e n o m e n o n which is heat-dissipation i n d e p e n d e n t indicates t h a t a significantly larger sample volume t h a n the near-interface region which should always be in t h e r m a l equilibrium with the gas p h a s e is taking part. Therefore, it is a s s u m e d t h a t the oscillation characterised by long cycle times a n d large a m p l i t u d e s are controlled by a bulk process. Moreover, it is m o s t probable t h a t the oscillation of the macroscopic surface oxidation state n a m e l y metal or oxide as seen by visible inspection is only a secondary process t h a t does not control the oscillating mechanism. The observed rate oscillations reach their m a x i m a in the visibly reduced states. Their occurrence m a y be c a u s e d by the t e m p e r a t u r e rise from the oxide reduction by excessive m e t h a n o l which is c a u s e d by excessive c o n s u m p t i o n of oxygen in the total oxidation reaction.

4.1. Model for a copper-bulk-oxygen based oscillation mechanism XAS (EXAFS)-investigations of copper catalysts u n d e r real catalytic conditions indicated the presence of a metallic copper bulk a n d reversible small c h a n g e s of the C u - C u n e a r e s t n e i g h b o u r distances a n d coordination n u m b e r s correlated with the m e t h a n o l - c o n v e r s i o n a n d with the oxygen c o n t e n t in the gas phase. These c h a n g e s were interpreted as the formation of a nanocrystalline copper bulk s t r u c t u r e by reversible intercalation of atomic oxygen at the interface of the nanocrystallites a n d not in the regular Cu lattice [5]. In situ X A S - m e a s u r e m e n t s of Cu at the oxygen K-edge u n d e r conditions of the partial m e t h a n o l oxidation revealed t h a t a copper oxide different from Cu20 is responsible for the partial oxidation of m e t h a n o l [3, 4, 6]. A surface species of Cu20 is reacting to COx a n d is c o n s u m e d by the m e t h a n o l reaction. Variation of the a b u n d a n c e of this copper oxide m a y therefore cause the oscillations of the CO2 production. The surface analysis d a t a indicate t h a t oxygen from the copper bulk can segregate to the catalyst surface where it forms copper-oxygen c o m p o u n d s similar to Cu20 as well as a novel CuOx-transient state. In this way oxygen stored in the catalyst bulk can influence the surface catalytic properties.

68

C h a n g e s in the a b u n d a n c e of b u l k oxygen or in its t r a n s p o r t kinetics (e.g. by t e m p e r a t u r e fluctuations) can so effect variations of the catalytic activity. We s u p p o s e t h a t the i n d i s p e n s a b l e high t e m p e r a t u r e p r e - t r e a t m e n t of the copper sample generates a n a n o c r y s t a l l i n e copper b u l k s t r u c t u r e with dissolved oxygen. In the p r e s e n t oscillation p h e n o m e n a the CO2 volume fraction oscillates with s u b s t a n t i a l a m p l i t u d e s , w h e r e a s the formaldehyde shows smaller or no oscillations. Therefore, the overall oscillations are mainly due to a variation of the total oxidation activity (cf. E q u a t i o n 3 a n d 4) w h e r e a s the c h a n g e s of the partial oxidation activity are of m i n o r i m p o r t a n c e which is d u e to the choice of a lean reaction mixture. The catalytic m e c h a n i s m of the total m e t h a n o l oxidation m a y be a redox m e c h a n i s m in a m a n n e r first p r o p o s e d of Mars a n d van Krevelen [28]. According to this m e c h a n i s m the c o n s u m e d lattice-oxygen is c o n t i n u o u s l y restored by oxygen from gas phase. It r e m a i n s open w h a t the exact n a t u r e of "lattice oxygen" m e a n s . It m a y well be t h a t this term does not designate a physically discernible oxygen species b u t r a t h e r m e a n s a rate c o n s t a n t of a solid state reaction (the formation of the active oxide) m u c h different from the rate c o n s t a n t s of the catalytic processes. On the basis of this r e s u l t s a model for the c h a n g e s of catalytic activity is p r e s e n t e d which is b a s e d on c o n c e n t r a t i o n c h a n g e s of mobile oxygen species dissolved in the catalyst bulk. The m e t h a n o l - c o n v e r s i o n r e a c h e s its m a x i m u m if the surface is covered with a m a x i m u m of the active oxygen p h a s e which is according to the p r e s e n t surface analytical d a t a a n d to the XAS results a surface p h a s e of Cu20 a n d not a sub-oxide. The active oxygen species is c o n s u m e d by the m e t h a n o l - c o n v e r s i o n a n d is c o n t i n u o u s l y replaced by oxygen from the catalyst bulk. The b u l k reservoir is r e g e n e r a t e d by gas p h a s e oxygen. However, at high m e t h a n o l conversions a large p a r t of the gas p h a s e oxygen is c o n s u m e d by catalytic reactions a n d is so not available for replenishing the b u l k oxygen reservoir. Therefore, the catalyst b u l k oxygen a b u n d a n c e is lowered gradually a n d with it the c o n c e n t r a t i o n of the active surface oxygen species. If now the m e t h a n o l conversion is r e d u c e d a n d vice v e r s a the gas p h a s e oxygen c o n t e n t is increased the c a t a l y s t bulk oxygen reservoir c a n be r e g e n e r a t e d by gas p h a s e oxygen. This process seems to be greatly facilitated by short episodes of r e d u c e d metal surface p a t c h e s which originate from the m e t h a n o l excess due to preceding excessive activity for total oxidation. It is clear, however, t h a t this additional facilitation is not critically required to m a i n t a i n a rate oscillation. 5. Conclusion

From in-situ macroscopic o b s e r v a t i o n s coupled with on-line p r o d u c t analysis a n d ex-situ surface analysis of massive copper u s e d as c a t a l y s t s in m e t h a n o l oxidation the following c o n c l u s i o n s can be d r a w n for the d y n a m i c b e h a v i o u r of the system:

69 9 Two independent pathway for selective and total oxidation exist. Under conditions of lean gas phases the total oxidation pathway exceeds in a rapid reaction the selective oxidation. 9 A surface phase of Cu20, not too thick to inhibit the reaction but thicker as a surface oxide ad-layer is the active phase for total oxidation. Its formation requires a bulk process, be it either segregation of oxygen or diffusion of copper atoms. 9 The steady state of energy flow to the active surface controls critically the reaction scenario, possible surface restructuring effects could not be observed with the crude analytical techniques used here. 9 The origin of the dynamic behaviour is the incompatibility of rate constants between the formation of the active phase which is not pure copper nor a bulk oxide phase with the rate constants for the main gas phase processes. This incompatibility is traced back to the essential participation of the bulk of the catalyst in the generation of the active phase.

Acknowledgement This work was supported by the German Science Foundation (DFG) through its Schwerpunktprogramm "Brfickenschl~ige zwischen Oberfl&chenphysik und Heterogener Katalyse".

References [111. Wachs, R.J. Madix, J. CataL, 1978, 53, 208 [211. Wachs, R.J. Madix, Surf. Sci., 84 (1979) 375. [3]A. Knop-Gericke, M. H&vecker, Th. Schedel-Niedrig and R. Schl6gl, Topics in Catal. 10 (2000) 187-198 [4]Th. Schedel-Niedrig, M. H&vecker, A. Knop-Gericke and R. Schl6gl, PCCP 2(15) (2000) 3473-3481 [51I. B6ttger, Th. Schedel-Niedrig, O. Timpe, R. Gottschall, M. H&vecker, T. Ressler and R. Schl6gl, Chem. Eur. J. 6, No. 10 (2000) 1870 - 1876 [6]A. Knop-Gericke, M. H&vecker, Th. Schedel-Niedrig and R. Schl6gl, Catal. Lett. 66(4)(2000)215-220 [7]F. Schtith, B. E. Henry u n d L. D. Schmidt, Adv. Catal. 39 (1993) [8]H. Werner, D. Herein, G. Schulz, U. Wild und R. Schl6gl, Catal. Lett. 49 (1997) 109 [9]T. Inui u n d T. Iwana, Stud. Surf. Sci. Catal. 19 (1984) 205 [10]S. Arsalane, R. Brochu, C. R. Acad und M. Ziyad, C. R. Acad. Sci. Ser. 2, 311 (1990) 1303 [11 ]A. Amariglio, O. Benall und H. Amariglio, J. Catal. 118 (1989) 164 [12]L. Petrov, C. Vladov, A. Eliyas, N. Kirkov, K. Tenchev, C. Boney, D. Filkova und L. Prahov, J. Mol. Catal. 54 (1989) 237 [13]E. Eckert, V. Hlavacek u n d M. Marek, Chem. Eng Commun. 1 (1973) 95

70 [14]X. Bao, J. V. Bart, G. Lempfuhl, R. Schuster, Y. Uchida, R. Schl6gl, G. Ertl, Surf. Sci. 284 (1993) 14 [15]H. Schubert, U. Tegtmeyer R. Schl6gl, Catal. Lett. 28 (1994) 383 [16]D. Briggs und M. P. Seah Hrsg..: Practical Surface Analysis, Vol. 1: Auger and X-ray Photoelectron Spectroscopy, J o h n Wiley & Sons, Chichester, New York; Salle+Sauerl~inder Aarau (1990) 635, 210 [17]T. Ressler, WinXAS 97, version 1.2, software for spectra analysis [18]R. Courts, B. Cord, H. Wern, H. Saalfeld u n d S. Hfifner, Sol. State. Comm., 63(7) (1987) 620 [ 19]A. Spitzer and H. Lfith, Surf. Sci. 118 (1982) 679 [20]A. Compaan, Sol. State Comm., 16 (1975) 293 - 296 [21]P. Dawson, M. M. Hargreave u n d G. R. Wilkinson, J. Phys. Chem. 34 (1973) 2201-2208 [22]J. C. W. Taylor und F. L. Weichman, Can. J. Phys. 49 (9171) 601 - 605 [23]J. C. Irwin, J. Chrzanowski u n d T. Wei, Physica C 166 (1990) 456 - 464 [24]U. R. Evans u n d H. A. Miley, Nature 139 (1937) 283 [25]F. H. Constable, Proc. Roy. Soc., A, 117 (1927) 376 [26]B.A. Pepply, J.C. Amphlett, L.M. Kearns and R.F. Mann, Appl. Catal. A., 179 (1999)32 [27]CRC Handbook of Chemistry and Physics, 76th. Ed., Chemical Publishing Company, Boca Raton (1995) 6-251 [28]P. Mars and D. W. van Krevelen, Chem. Eng Sci. 3 (1954) 41


71

Novel Reactor Configurations and Modes of Operation Dan Luss Department of Chemical Engineering, University of Houston, Houston, TX 77204, USA

Introduction Reaction engineers are expected to transform laboratory discoveries of new synthesis routes or design concepts into economic, safe, and environmentally compatible processes. The highly competitive industrial environment has added the need to shorten the time interval in which this task has to be completed and to decrease the production price. This motivated several innovations. The first was development of novel catalysts, which increased the yield in existing processes, such as the novel Kellogg ammonia-synthesis process, which uses the much more active BP catalyst. Other catalysts were designed to provide either new synthesis routes, such as the production of synthesis gas by direct oxidation, or new products, such as production of novel polymers by metallocene catalysts. Attempts to use cheaper feed stocks led to the development of new processes. For example, the drive to replace olefins as reactants by paraffins led to the development of processes in which ethane rather ethylene is used to produce vinyl chloride, propane rather than propylene to produce acrylonitrile, and butane rather than benzene to produce maleic anhydride. The drive to produce more-economical synthesis gas from methane has motivated various novel process developments. Moreover, environmental regulations and needs caused modifications of many processes in order to minimize the production of pollutants. Most reactors are designed to handle a relatively narrow range of feed concentrations and space velocities. A different design approach has to be used if reactors are to destroy pollutants, as they have to operate at high conversion over a very wide range of feed compositions and feed rates. The development of the new processes led in some cases to the development of novel reactor configurations and/or mode of operation. I shall review here some of these developments and the corresponding design and operation problems and challenges. Specifically, I shall describe some of the challenges encountered in four processes: 1. Reactors for conducting very fast chemical reactions. 2. Periodic operation of chemical reactors. 3. Fuel-cell systems. 4. Gas-phase synthesis ofpolyolefins. I shall review here only a small fraction of the many related contributions. The goal is to point out the problems and challenges encountered in either the design of these reactors

72 or their mode of operation, rather than to present a comprehensive review. contributions of many engineers and scientists will not be described here.

The

Reactors for conducting very fast chemical reactions Almost all commercial heterogeneous catalytic reactions are conducted with a contact time of at least several seconds. Many require much longer contact time. There are, however, a few commercial catalytic processes in which the reaction is completed in a very short time (order of milliseconds). These very fast reactions are usually affected by heat- and mass-transport resistances. They are usually carried out either on catalytic gauzes or on a reticulated structure or monolith coated with a very thin washcoat on which a catalyst is impregnated. A very well-established fast reaction is the oxidation of ammonia to NO over a Pd/Rh gauze [ 1]: 4NH3 + 502 "-)' 4NO + 6H20

AH = -227 KJ/mol NH3,

commercialized by Ostwald in 1906. The NO yield is about 97%-98% in atmospheric commercial reactors and about 95% in reactors operating at 0.8 or 0.9 MPa. A related fast reaction, developed by Andrussov in the early 1930s, is the synthesis of hydrogefi cyanide [ 1]: CH4 +NH3 +1.502 --) HCN + 3H20

AH =-482 KJ/mol.

These fast commercial reactions have been considered to be rather uncommon exceptions. In recent years, the research group of L. Schmidt at the University of Minnesota [2] has discovered several very fast reactions (contact time of the order of 1 ms.) that have several advantages. They are more selective than existing ones, can reduce the amount of undesired by-products, feature lower capital costs, and can use lessexpensive reactants. The first reaction they studied was the production of syngas by direct oxidation of methane, CH4 + 0.5 02 ")' C O + 2H2

AH =-8.5 KJ/mol.

on a reticulated (foam) catalyst on which Rh was impregnated, having a metal surface area of about 100 cmZ/gr. The reactor lengthwas 0.5-10 mm and the inlet velocity about 1 m/s. at a residence time of several ms. About 90% of the methane and nearly 100% of the oxygen were converted at a residence tome of several ms., yielding more than 95% CO yield and over 90% H2 yield. This very fast reaction was insensitive to the form of the catalyst and the metal-loading, but rather sensitive to the specific metal catalyst. A 1ft.-diameter reactor with this very thin catalyst layer is predicted to produce 1 ton/d, of syngas. The same fast reaction was carried out in a laboratory fluidized-bed reactor using 100-~tm alpha-alumina spheres with 5-~tm-diameter holes within them. These beads did

73 not have a washcoat and were impregnated with Rh. The conversion and selectivity obtained at a residence time of 0.05-0.25 seconds were very high and close to those predicted from equilibrium calculations. Schmidt's group discovered another important class of very fast reactions - - the oxidative dehydrogenation of olefins over monolith catalysts. For ethane, the reaction is C2H6 + 0.5 02 -') C2H4 + H20

AH=-25.1 Kcal/mol.

Using a Pt catalyst, they obtained up to 70% selectivity to C2H4 at 80% conversion of C2H6 and nearly complete conversion of O2. The product distribution was very different from that attained under thermodynamic equilibrium, according to which CH4 should be the dominant product and only a small percentage of C2H4 should be obtained. A major advantage of these two processes found by Schmidt's group is that they lead to nearly complete conversion of the reactants with high yield of the desired product, avoiding the usual tradeoff between conversion and yield. Schmidt's experiments involved small hydrocarbons and noble metals. They suggest that by use of other feeds and catalysts, it should be possible to design other efficient processes for fast partial oxidation and maybe even other classes of exothermic reactions. The discovery of these reactions, catalysts, and operating conditions will be a major contribution. Many features of these very fast exothermic reactions need to be elucidated to enhance our understanding of the interaction between the (usually nonselective) homogeneous and heterogeneous reactions and their dependence on the catalyst structure, washcoat diffusion, and fluid velocity. It was suggested that in some cases the exothermic surface reactions generate free radicals which initiate gas-phase reactions that lead to some unstable species. These homogeneous reactions are quenched by mixing with colder gases passing through the gauze. This rapid quenching enables one to obtain valuable nonequilibrium-mixture products. Knowledge of the rates of both the homogeneous and heterogeneous reaction rates is essential for selecting the optimal catalyst configuration and operating conditions. This difficult task of separating the contributions of these two modes of fast reactions requires use of both detailed models of the kinetics of the many reaction steps and of the flow and transport within the complex reticulated structure. An example of such a model is that by Duetschmann and Schmidt [3]. The prediction of the transport coefficients under the extremely steep temperature and concentration gradients introduces additional uncertainties to the simulations. The theoretical model predictions need to be tested and modified by critical experimental data. An intriguing question is what the impact is of the nonuniform pore-size distribution on the performance of this very rapid reaction, especially in systems that involve several reactions, each having a different rate of heat release or consumption. For example, during the fast partial oxidation of methane, the oxygen is completely consumed in the upstream section of the reticulated structure while steam-reforming occurs further downstream.

74 The pressure-drop in shallow catalytic beds or gauzes is very small. Thus, in the design of large-diameter reactors, special attention must be devoted to maintain a uniform velocity and concentration across the reactor. Moreover, the rather short time constant of the fast chemical reactions relative to that of heat conduction in the radial direction raises the possibility of nonuniform temperature distribution on large-diameter industrial reactors. This nonuniform temperature may have a deleterious impact on the reactor performance and may even lead to safety problems. It is also important to prevent radiation-heating from initiating unselective homogeneous reactions from occurring before the hot catalyst bed is reached. The highly exothermic fast reactions may be handled safely only in a bounded region of temperatures and fuel/oxygen ratio. At high temperatures, the catalyst deactivates or melts. At low temperatures, extinction occurs. At a high fuel/O2 ratio, coke forms and deactivates the catalyst. At a low fuel-to-air ration, the reaction becomes homogeneous and may lead to an explosion. It is essential to be able to predict these operation regions to guarantee safe operation and to guide the search for novel fast reactions.

M

Figure 1: Short-Contact-Time Fluidized Reactor (SCTFR), after Bartholic [5]. Many commercial reactions are carried out in fluidized-bed reactors. Several very interesting developments of a short-contact-time fluidized reactor (SCTFR) have been reported in recent years. Professor Bergougnou and his associates at the University of Western Ontario [4] developed a "downer" SCTFR concept with the goal of rapid gasification of biomass. These novel downflow reactors facilitate operation at a short residence time of about 1 second. Ensyn Engineers of Ottawa scaled up the process to 100 kg/min and reported several commercial applications. Bartholic [5] patented an SCTFR for catalytic cracking which has been commercialized by Coastal Corporation. In one configuration, the atomized feed contacts a falling mountain of hot catalyst (above

75 540 ~ C) for a short time (less than 0.5 second). The higher catalyst temperature and shorter residence time of the SCTFR relative to that of a traditional riser reactor increase the fuel-oil conversion at the same coke yield and reduce hydrogen-transfer activity while increasing the yield of light olefins. This novel reactor configuration is expected to find important commercial applications. When the rate of the chemical reaction is very fast (short characteristic time) relative to that of mixing of the reactants, micromixing has an important impact on the conversion and/or selectivity or properties of the product. For example, the mixing in the high temperature conversion of TIC14 to TiO2 strongly affects the particle-size distribution of the product. The quality and hence price of TiO2 used in paints is strongly dependent on this particle-size distribution. Similarly, the mixing intensity in a polymerization reactor can affect the initiator efficiency and the final molecular-weight distribution [6]. Variation of the position of the feed in a crystallizer may lead to large variation in the average crystal size [7]. Unfortunately, a large uncertainty is still associated with the design of reactors involving very fast reactions and mixing. It is essential to improve our ability of a rational design of the configuration of these reactors and mode of operation and to develop a reliable scale-up methodology. It will be very beneficial to discover additional fast reactions and develop reactors in which they can be conducted safely and at a high rate. Their successful commercial operation would be of significant industrial importance and academic interest. Periodic Operation of Chemical Reactors Most chemical reactors operate at a steady state. Theoretical studies predict that the timeaveraged conversion or yield following a periodic variation of the feed concentration and/or temperature usually differs from that of a steady state for the same time-averaged feed conditions. Many experimental studies verified these theoretical predictions that a forced periodic operation can affect the time-averaged performance of a reactor. One of the few reported industrial applications is the cycling of a chain-transfer agent in order to affect the polydispersity of polypropylene [8]. Only two types of periodic processes have found wide technological applications so far. The first is a fluidized-bed reactor in which the catalyst undergoes periodic changes as it circulates through the reactor and regenerator. The first proposed application was by Lewis et al. [9] and was the selective oxidation of methane by a circulating copper catalyst, which was later oxidized in a regenerator. Callahan et al. [10] attempted to apply that concept to the oxidation and ammoxidation of propylene. DuPont has recently applied this concept to the oxidation of butane to maleic anhydride [11]. While this reactor configuration and mode of operation have many potential applications, its success is critically dependent on the ability to develop a catalyst that can reversibly bind a large amount of oxygen, as this determines the required rate of catalyst circulation.

76 The second type of reactor is the reverse-flow reactor (RFR), which has attracted significant interest in recent years. In that operation, the direction of the flow in a packed-bed reactor is periodically reversed to trap a hot zone within the reactor. The cold feed is regeneratively heated up as the high-temperature zone moves downstream. Before the hot zone exits the bed, the feed-flow direction is reversed. A detailed review of the various applications of the RFR was presented by Matros and Bunimovich [12]. Experimental and theoretical studies have shown that many (usually more than 100) flowreversals are needed before the reactor converges to the periodic state. There is still a need to determine the best control policy when the reactor is subjected to changes in the feed concentration and/or load.

Reverse flow reactor

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Figure 2" Schematic of reverse flow reactor and temperature profile at the beginning and end of a semi-cycle The RFR has been used so far in two major applications. One is to conduct an exothermic equilibrium-limited reaction so that its conversion increases as the reactants pass over the leading front of the moving hot zone. The second is to conduct reactions in which the adiabatic temperature is lower than that needed to carry out the r e a c t i o n - for example, the catalytic destruction of a dilute mixture of volatile organic compounds (VOCs). It was recently proposed to conduct coke-forming reactions in a RFR [13]. In this operation, the desired reaction is conducted in the upstream section of the reactor while the downstream section is regenerated at the same time from the coke deposited in the previous half-cycle. All the reported applications of the RFR were for systems involving a single reaction. There is a need to determine if it is useful and how to apply it in systems in which several reactions occur simultaneously. Levenspiel [ 14] suggested that one potential application is to conduct an exothermic reaction during the first semicycle and then an endothermic one in the second semicycle. An example of a potential application is conducting the

77 endothermic dehydrogenation of ethylbenzene in one semicycle and the exothermic combustion of hydrogen in the second [15]. Several analyses of this mode of operation [ 15,16] indicate that this periodic operation and its dynamics are rather intricate and that it is difficult to obtain high yields. A major difficulty is that it is not possible to efficiently store the heat of the exothermic reaction without reaching excessive temperatures in the reactor. The operation could be enhanced significantly by placing within the bed some inert material that undergoes a reversible phase-change at the desired operating temperature. The development of robust RFR operation for coupling of endoand exothermic reactions still awaits an imaginative novel approach. When moderately or highly exothermic reactions are carried out in a RFR, cooling is needed to prevent excessive temperatures. Recent theoretical studies have shown that this cooling may lead to rather complex dynamics, including chaotic behavior [17]. Various modes of cooling the RFR have still not established the most efficient and robust fashion of conducting this cooling. Several modifications to the operation of a cooled RFR operation have been suggested. A countercurrent reactor may be used instead of using an RFR with a high frequency of flow-reversal. A major advantage of this operation is that it obviates the need for expensive valves and the operational difficulties encountered during the frequent flow-reversals. Analysis indicates that the same performance and a more robust dynamic operation may be obtained in a cooled folded countercurrent reactor than in a regular one [18]. Gilles' group has developed a novel countercurrent reactor configuration, called the circulation reactor [ 19]. The original development of the RFR was for trapping of a hot-temperature zone within the reactor. It is also possible to use this operation for trapping of reactant zones. Agar and Ruppel [20] proposed to conduct a reverse-flow operation in which a zone of an adsorbed reactant, rather than a high temperature, is trapped in the reactor and used to destroy environmental pollutants. They proposed to use the RFR for isothermal, catalytic destruction of NO by the selective catalytic reaction (SCR), NO

+

NH 3 + 1/4 02 --)'N2 + 1.5 H20

AH = -407 KJ/mol.,

exploiting the catalyst's strong adsorption capacity of the ammonia. The periodic operation consists of four steps. In the first, ammonia in excess of the stoichiometric amount is added to the pollutant stream fed to the bed. In the second step, only the NO is fed to the reactor, and it reacts with the adsorbed excess ammonia. At the end of the semicycle, the flow direction is reversed and the first two steps are repeated. This periodic operation forms a trapped region of adsorbed ammonia in the bed. This periodic operation enables a tight control and prevention of ammonia slip, even under sudden changes in the load or NO concentration. The success of such a process depends on the availability of a catalyst with a high adsorption capacity of the added reactant under the desired operating conditions. In order to form sharp concentration fronts, the reactant adsorption rate should be comparable to that of the reaction, and its desorption rate should be much slower.

78 Agar and Ruppel [20] described the removal of about 1000 mg/m 3 of NO from a feed fed at a space velocity of 300 1/hr to such a RFR. During the first step, ammonia at a concentration of about 2000 ppm was added to the feed for 67 minutes. The ammonia feed was then stopped for the next 83 minutes, following which the flow direction was reversed and the two steps were repeated. The NO conversion exceeded 99% and the effluent contained less that 1 ppm of ammonia. Noskov et aL [21] pointed out that several safety problems are associated with the process. Specifically, ammonium salts that form below 80 ~ C can decompose at higher temperatures. They suggested circumventing this problem by introducing the ammonia to the hot central section of the reactor, rather than at the reactor ends. Snyder and Subramanian [22] have shown that the operation may be more efficiently conducted by injecting the ammonia as a side stream. Significant industrial interest exists in the application of an SCR process for the destruction of Nox. The RFR mode of operation is a promising SCR reactor configuration. Its application would depend on the familiarity and understanding of design engineers of the potential advantages of this periodic mode of operation. Several differences exist between that of trapping a reactant and a high-temperature zone. When an exothermic reaction is carried out in a RFR, increasing the feed temperature increases the maximum temperature of the trapped zone. However, increasing the feed concentration does not increase the maximum concentration of the trapped reactant zone it merely broadens it. Moreover, the width of the trapped concentration zone changes during each semicycle, which consists of two steps. At the 2000 Los Angeles AIChE annual meeting, Bjornborn proposed using the RFR operation for "homogeneous catalyst-trapping." He suggested trapping within a RFR an expensive organometallic catalyst that is reversibly attached to an inert support in order to decrease the rate of catalyst-loss from the loss that occurs in a unidirectional operation. This process may be used to produce fine organic chemicals. Since the catalyst is not consumed by the reaction, it is not possible in this case to obtain sharp fronts that bound the zone containing the catalyst. In fact, dispersion will broaden this trapped catalyst region. It will be interesting to find out if a successful application of this catalysttrapping can be accomplished. This suggested process is just another example of a potential application of a periodic operation. It is realistic to presume that several novel applications of the RFR operation will be commercialized in the future. Most previous theoretical studies of dynamic operating reactors used kinetic expressions obtained under steady-state operation. These models do not account for the detailed dynamics of the adsorption and desorption rate processes, and they may lead to erroneous predictions in periodic operation of the reactor. Thus, simulations of periodic processes may require use of kinetic models that are much more detailed than those used for predicting steady-state operation. These dynamic models also need to account for the rate of adsorption, desorption, and adsorption capacity of the catalyst. As mentioned above, the hot-temperature zone in a cooled RFR may exhibit complex dynamic features under

79 certain operating conditions [ 17]. It would be of interest to see whether complex dynamic behavior may be attained by a trapped concentration zone in an isothermal RFR, and under what conditions.

Fuel-Cell Systems William R. Grove, a British physicist, demonstrated in 1839 that a battery fed with oxygen and hydrogen can generate electricity. The fuel cell is a device that relies on electrochemistry to produce electrical power through noncombustive fuel oxidation. It consists of an anode and cathode separated by an electrolyte. Fuel (e.g., hydrogen) liberates electrons at the anode and forms ions that move towards the cathode. The ions react with oxygen at the cathode, forming water as a by-product. A catalyst such as platinum is usually added to the electrodes to increase the reaction rate. In 1950, Francis T. Bacon developed the first operational 5-kilowatt fuel cell. The first major application of fuel cells was by NASA in the Gemini and Apollo flights. The fuel-cell technology promises many advantages, such as high energy-conversion efficiency, very low chemical and noise pollution, fuel flexibility, cogeneration capacity, modular design, rapid load response, and distributed capacity. The technology is being touted for various applications, including cleaner power for buses and cars; stationary large fuel cells for electric utilities and small ones for homes; and miniature fuel cells for cellular phones, laptop computers, camcorders, and other consumer products. Many research groups around the world are aggressively attempting to lower the price of fuel cells to enhance the competitive position of this technology. Ballard has recently developed a fuel cell for automobiles that costs about $275 per kilowatt. It is estimated that to penetrate the automobile market, their price should be below $50 per kilowatt [23]. The outcome of the efforts to develop economical fuel cells may alter the world energy map. The configuration and mode of operation of the reactors associated with various fuel cells are rather different from those of conventional chemical reactors. Their design and mode of operation raise many challenging reactor-design problems, which require creative application of chemical reaction engineering methodology and tools. Currently, the major force behind fuel-cell development is the huge potential market in the transportation industry. Passenger cars represent the largest potential market, as annual worldwide sales are about 30 million cars. The worldwide urban bus market is about 17,000 vehicles sold annually, of which about 30% are sold in the USA. A major advantage of fuel cells is a more efficient and clean operation of automobiles, and virtually every major manufacturer is now working on the technology. Early demonstration vehicles operated by fuel cells are already on the road. Among others, several prototype Ballard fuel-cell buses have been operated and tested during the last two years in Chicago and Vancouver. Daimler-Benz and Ford expect to produce 100,000 fuel-cell-powered vehicles annually by 2005. The London based Zero Emissions Vehicle Co. (Zeveco) has been testing a prototype fuel-cell-driven taxi.

80 At present, it appears that fuel cells on automobiles will most probably use hydrogen as the fuel. While other fuels may, in principle, be used, their products deactivate the current cells. The leading technology for automobile fuel cells is the proton-exchange membrane (PEM), which uses synthetic polymers as the electrolyte. The membrane needs to be thin, and impermeable to both the fuel and oxygen. Currently, the membranes are made of a fluoropolymer (Nation) on which sulfuric acid groups are stung and supported on a water-repelling mesh (to prevent the membrane from becoming soggy and expanding). The cathode of one cell and the anode of an adjacent cell are separated by a plate that connects them electrically in series. Special gas channels or porous materials are used to distribute the hydrogen and oxygen to the electrodes on each side of the plate. The thickness of an assembly is of the order of 2.5 mm. A series of assemblies and plates is bolted together to a stack that supplies the power. The 45%-60% energy efficiency of the PEM cells is approximately 25% higher than that of the internal combustion engine. The high price of the platinum needed in the fuel cell is a major factor holding back commercial application. Considerable efforts are dedicated at present to reduce the large amount of expensive platinum catalyst needed in these cells to enhance their competitive position relative to that of internal combustion engines. Moreover, the worldwide production of platinum is not sufficient to provide each new car with a PEM fuel cell, unless the amount of platinum needed per cell can be reduced. Zeveco is pursuing the use of alkaline fuel cells, the platinum used in which is one-fifth of that of the PEM models. If fuel cells are to be widely used in vehicles, it is essential to develop on-board efficient hydrogen-storage or-production systems. About 3 kilograms of hydrogen are needed to drive a small car 500 kilometers. At a pressure of 200 atm, this would require a 180-liter tank. Another option is to use liquefied hydrogen. This will reduce the volume of the fuel tank but will lead to about 30% loss in the energy efficiency. Another solution is absorb the hydrogen by metal hydrides or nanotubes. Two other important impediments facing hydrogen-powered vehicles are the lack of infrastructure for delivering them, and safety considerations. Extensive research is currently directed at developing fuel cells that generate the fuel on board from a cartier fuel, such as methanol or even gasoline. The development of these compact reforming reactors is a major challenge for reaction engineers. Gasoline is an obvious fuel option, as the infrastructure for delivering it to cars exists. This option will require the development of a compact reformer that can handle the large mix of hydrocarbons in the gasoline and the impurities present in it. The steam required for the reforming will be generated by burning some of the fuel. This application will require an efficient removal of sulfur-containing compounds to avoid poisoning the catalyst in proton-exchange membrane cells. The technological problems associated with the development of such a reformer have led to the search for other alternative fuels. At present, major efforts are directed at the use of methanol as the fuel. While its energyrelease is only half of that of gasoline, it can be handled and transported more easily than hydrogen. A possible solution is to reform methanol with steam at 280 ~ C in the presence

81 of a catalyst, and burning some of the hydrogen to produce the steam. Phosphoric-acid cells can work well with a methanol-reformer because their relatively high operating temperature can be exploited to provide the steam; this enables them to resist poisoning by a small amount of carbon monoxide. It is much more difficult to combine methanolreforming with a proton-exchange membrane fuel cell, as an additional step is needed to avoid poisoning by the carbon-monoxide product. Major improvements in the reformer are needed. The design of an efficient reformer that produces hydrogen on board a car from a cartier fuel such as methanol or gasoline is a major technological challenge. Its success and wide application will have tremendous economic impact. The use of a reformer on cars will also enhance public understanding and interest in the operation of chemical reactors. Stationary fuel cells have been developed for use by utilities, industrial facilities, and homes. They can use a variety of fuels, and they operate relatively cleanly and silently. Unfortunately, the stationary fuel cell costs about $3000 per kilowatt, as compared to $500-$1000 for the combustion turbines commonly used by utilities [24]. As the concerns mount about the effects of greenhouse gases from power plants, a wider use of fuel cells is expected. The stationary fuel cells usually require use of an external or internal reformer to convert the fuel to a feed suitable for the reaction at the anode. While the reformer emits some pollutants, that level of release is much smaller (between a tenth to a thousandth) than that in a conventional combustion turbine. The energy efficiency of the fuel cell may be increased by recovering the waste heat. Five types of stationary fuel cells have been developed: phosphoric acid, molten carbonate, solid oxide, protonexchange membrane, and alkine [24]. The phosphoric-acid fuel cell is the most mature technology. However, its high price of about $4000/kilowatt and the even higher price of the alkaline fuel cells have motivated the search for other technologies. These include molten carbonate, proton-exchange membrane and solid-oxide fuel cells (SOFCs). The SOFC operates at high temperatures (around 1000 ~ C ) in order to enable the oxygen ions to migrate through a ceramic membrane to the anode. This high temperature requires use of very expensive materials. Recently reported advances [25] may enable operation at the much lower range of 6000-700 ~ C. The goal of the research in this area is to reduce their cost to about $400/kilowatt, which is in the ballpark of plants operating with natural gas. The development of more-efficient fuel cells will require development of new materials, reactor configurations and modes of operation, and creative utilization of the waste heat. This presents a great challenge and opportunity for the Reaction Engineering community to lead a major technological breakthrough.

Polyolefin synthesis The use of metallocene catalysts revolutionized the synthesis of polyolefins [26]. These catalysts consist of two ring molecules connected by a metal (Wilkinson catalyst). The two metallocene tings are also sometimes connected by a bridge molecule. The industrial applications of these catalysts was motivated by the discovery by Prof. Kaminsky, who

82 discovered that their activity can be increased by several orders of magnitude by the presence of a co-catalyst, methylalumoxane (MAO). A major advantage of these catalysts is that they enable a control chemistry (stero-regulation) and rational engineering of the polymer stero, which affects their properties. This control is achieved via changes in either the ring structure, the ring substitute, the presence and type of bridge, or the metal connecting the two tings. For example, polypropylene (PP) has the same backbone as polyethylene, with the exception that a methyl group is sticking out at each pair of carbons. The PP may be either syndiotactic, so that the methyl groups are added on alternating sides of the polymer; isotactic, so that all the methyl groups are always on the same side of the polymer chain; or hemi-isotactic. Ewen [27] provided several examples of this stero-regulation. Additionally, the metallocene catalysts have the ability to produce polymers and copolymers that cannot be synthesized by the Ziegler-Natta catalysts. These advantages and the high rate that they offer have led to the use of metallocene catalysts in many processes. This use has motivated extensive industrial research activity to generate novel products and processes, as well as academic research of their performance. Gaseous polymerization of polyolefins is carried out in fluidized-bed reactors. The conversion per pass is kept rather low (order of 2%) so that the heat of polymerization is removed via convection by the unconverted reactants, which are cooled externally and recycled to the reactor. In some cases, an inert liquid is added to the feed so that its evaporation helps remove some of the heat of the reaction. The Unipol process is one of the most common processes used to carry out this gaseous polymerization. Other processes for producing polyolefins include Hypol, Novolen, Catalloy and Spheriline.

R

R H

Me2

CI ..

CI

structure syndiotactic

r.-.9

I-

[

I

r, I

1

[~,. .u. .~

1-- 1

t-butyl

isotactic

LIA.... I I I 1.......t

methyl

hemiisotactic

[ J I ' I ~ I ' I

!

I

l

Figure 3: Stereo-regulation of polypropylene structure by a metallocene catalyst, after Ewen [27] These processes have been used successfully for many years to carry out gas-phase polymerization of p01yolefins using the Ziegler-Natta catalysts. It was expected that replacement of the Ziegler-Natta by metallocene catalysts in fluidized-bed reactors will be rather straightforward. However, the adaptation of the more active metallocenes to existing commercial fluidized-bed reactors led to many operational problems. For

83 example, it was reported [25] that technical problems led Exxon to declare force major on the production of polyethylene by metallocene at its Mont Belvieu plant in May 1997. The heat carried by the hot, recycled, unconverted ethylene is usually removed in watercooling towers. Thus, the feed temperature to the fluidized-bed reactor is usually in the range of 20~ ~ C, while the melting temperature of polyethylene (depending on the type) is about 450-50 ~ C. Thus, the exothermic polymerization in the fluidized-bed reactors has to be carried out at a rather narrow temperature range to avoid melting or softening of the polymer. One operational problem that hampered the operation of various fluidized-bed reactors using metallocenes was the formation of sheets of fused polymer (sheeting), which required termination of the operation. The sheeting occurs due to the temperature-increase of some particles, which in turn causes the polymer to melt or soften. This causes the sticking of other particles to it. The decreased surface-area-tovolume ratio of this lump of particles further increases its temperature and leads to the formation of a sheet of molten polymer. Several explanations have been proposed for the formation of these sheets of molten polymer, but it has not yet been established what exactly initiates and causes this undesirable event. One proposed explanation is that an electrostatic field attracts small particles to the reactor walls. This, in turn, decreases their velocity and heat-transfer coefficient, which leads to the excessive temperature-rise. Another explanation is that fine particles settle on the expended part of the reactor. The decreased heat-transfer rate of the settled particles and the high reaction rate of these fine particles cause their overheating. The resulting molten or soften polymer initiates sheet-formation. This sheet-formation clearly illustrates the potential problems that may be encountered during use of an existing reactor configuration to conduct the same reaction but with a more active catalyst. It is essential that we enhance our understanding of this deleterious sheeting in order to prevent its occurrence in novel processes utilizing this superior catalyst. The interaction of diffusion and reaction in heterogeneous catalysts is very well understood at present and can be reliably be predicted. Unfortunately, we still lack complete understanding and ability to predict all the interactions among the various rate processes occurring on growing polymer particles. The metallocene and MAO are usually supported on small silica particles (diameter of the order of 50g). As the polymer forms in the pores of the support, it fragments the support to smaller particles. The outward motion of these fragments enables the growing particle to maintain a high activity in spite of the continuous growth of the particle diameter. (The diameter of the polymer particle is usually at least ten times larger than that of the silica catalyst particle.) In the absence of this fragmentation and movement, the diffusional resistance would significantly decrease the reaction rate as the diameter size becomes larger. A recent description of the fragmentation can be found in [28]. It is essential to develop a better ability to predict the dependence of this fragmentation on the properties of the catalyst and the polymerization reaction rate. Too rapid a fragmentation may generate fine

84 particles, which can initiate sheeting. Moreover, it may distort the shape of the particle and hence of the product, which usually replicates that of the catalyst pellet. Another puzzling effect is the sharp increase in the polyethylene activity following the addition of a small amount of a comonomer (such as 1-hexene). The explanation may be due to polymer swelling which increases both the solubility and diffusivity of the ethylene. However, it may be due to a kinetic effect, such as an increase in the propagation rate constant or secondary site activation. The cause may of course be a combination of the physical and chemical effects The determination of the cause of this effect and ability to predict its impact is of both industrial importance and intrinsic academic interest. At present, the interaction of diffusion and reaction in a growing polymer particle is usually described by the multi-grain model [29,30], multi-layer model [31], or some modification of either. These models do not account for any of the rate processes that cause the fragmentation of the catalyst. Thus, they are unable to explain, let alone predict, the known impact of the mechanical properties and structure of the support on the observed reaction rate. It is important to enhance our understanding of the interaction between the reaction transport and catalyst fragmentation in order to enable an optimal choice of the support. Many commercial heterogeneous catalysts are not impregnated in a uniform fashion. For example, various precious-metal catalysts consist of an exterior active shell and an inert core in order to enhance the effectiveness factor. Several automobile-muffler catalysts have a carbon-monoxide-oxidation catalyst in one shell and an NOx-reducing catalyst in another shell. Our understanding of the reaction-diffusion interaction facilitated this rational design of the optimal profile of catalyst-activity distribution and shape. It would be of both practical importance and academic interest to develop a rational procedure for enhancing the performance of metallocenes by their nonuniform impregnation on the support.

Concluding remarks The examples presented here illustrate that application of novel reactor configurations and modes of operation present reaction engineers with important opportunities and challenges. The specific cases discussed here are of important applications, but not necessarily the most important or promising ones. Many future reactor-design and operational innovations are expected to involve nontraditional areas of application. The fuel-cell systems discussed here constitute one example, but many others exist, such as the synthesis of electronic materials and controlled drug-release. A recent novel reactor configuration not discussed here is miniature reactors. These may find novel applications in the production of hazardous chemicals at a site at which they are reacted; in military gas masks to neutralize toxic gases; as a tool in combinatorial catalysis and miniature fuel-cell systems for consumer products; etc. I am convinced that many important

85 contributions will be made in the future by imaginative engineers developing novel reactor configurations and creative, novel modes of operation.

Acknowledgments I would like to thank the National Science Foundation for continued support of my research in reaction engineering.

References 1. C.N. Satterfield, Heterogeneous Catalysis in Industrial Practice, 2 nd edition, McGrawHill, New York (1991 ). 2. L.D. Schmidt, M. Huff and S.S. Bharadwaj, Chem. Eng. Sci., 49 (1994), 3981. 3. O. Duetschmann and L.D. Schmidt, AIChE J., 44 (1998), 2465. 4. A.S. Bassi, C.L. Briens, A.I.M. Bergougnou, "Proc. 4 th Intl. Conf. of CFB" (A. Avidan, ed.), AIChE, New York (1993). 5. D.B. Bartholic, U.S. Patent #2,121,733 (1991). 6. P. Sahm, PhD thesis, Institute National Polytechnique de Lorraine, Nancy (1978). 7. G. Tosun, 6th Europ. Conf. on Mixing, Pavia, Italy, (1988), 161. 8. B.E. Claybaugh, J.R. Griffin and A.T. Watson, U.S. patent #3,472,829 (1969). 9. W.K. Lewis, E.R. Gilliland and W.A. Reed, Ind. Eng. Chem., 41 (1949), 1227. 10. J.L. Callahan, R.K. Grasselli, E.C. Milberger and H.A. Strecker, Ind. Eng. Chem. Prod. Res. Dev., 9 (1970), 134. 11. R.M. Contractor, U.S. patent #5,021,588 (1991). 12. Y.S. Matros and G.A. Bunimovich, Catal. Rev.-Sci Eng., 38 (1996). 13. A.M. de Groote, G.F. Froment and T. Koblinski, Can. J. Chem. Eng., 74 (1996), 735. 14. O. Levenspiel, Chem. Eng. Sci., 43 (1988), 1427. 15. G. Kolios and G. Eigenberger, Chem. Eng. Sci., 54 (1999), 2637. 16. M.S. Kulkarni and M.P. Dudukovig, Ind. Eng. Chem. Res., 37 (1998), 770. 17. J. Khinast, Y.O. Jeong and D. Luss, AIChE J., 45 (1999), 299. 18. R. Garg, D. Luss and J.G. Khinast, AIChE J.,46 (2000), 2030. 19. X. Hua, M. Mangold, A. Kienle and E.D. Gilles, Chem. Eng. Sci., 53 (1998), 47. 20. D. Agar and W. Ruppel, Chem. Eng. Sci. 43 (1988), 2073. 21. A.S. Noskov, L.N. Bobrova and Y.S. Matros, Cat. Today, 17 (1993), 293. 22. J.D. Snyder and B. Subramanian, Chem. Eng. Sci., 53 (1998), 727. 23. J. Appelby, Sci. Amer., 281 (July 1999), 74. 24. A.C. Lloyd, Sci. Amer., 281 (July 1999), 80. 25. R.F. Service, Science, 288 (2000), 1957. 26. Chem. Week (21 May 1997), 23. 27. Ewen, Sci. Amer., 276 (May 1997), 86. 28. J. Zechlin, B. Steimetz, B. Tesche, G. Fink, Macromol. Chem. Phys., 201 (2000), 515. 29. I. Yermakov and V. Zakharov, Adv. Cat., 24 (1975), 173. 30. S. Floyd, K.Y. Choi, T.W. Taylor and W. H. Ray, J. Appl. Polym. Sci., 31 (1986), 2231. 31. J.B.P. Soares and A.E. Hamielec, Polym. React. Eng., 3 (1995), 261.


Studiesin SurfaceScienceand Catalysis133 G.F. Fromentand K.C.Waugh(Editors) Publishedby ElsevierScienceB.V.,2001

87

New methodologies and reactors for catalytic process development M. P. Harold a, P. L. Mills b and J. F. Nicole b aDepartment of Chemical Engineering, University of Houston, 4800 Calhoun St., Houston, TX 77204-4792 USA

bDuPontCompany, Central Research and Development, Experimental Station, E304/A204 Wilmington, DE 19880-0304 USA Recent advances in the development and application of laboratory-scale reactor systems that are intended to accelerate the discovery and development of new catalytic processes are summarized and analyzed. Emphasis is placed upon multiphase gas-solid and gas-liquid systems that use either solid heterogeneous or soluble homogeneous catalysts since these provide the basis for most of the existing and anticipated future process technologies. The emerging role of various reactor configurations and supporting devices that are intended to accelerate discovery of new catalyst compositions is also reviewed. The impact of this approach on the next level of catalytic process development and future challenges is also discussed. 1. INTRODUCTION

1.1. Background The manufacture of a large variety of intermediate and end-products used in commercial and consumer applications is largely based upon process technology where either a soluble organometallic complex or solid heterogeneous catalyst is used to transform reactants into value-added materials. Examples of products and technologies where catalytic processes are widely used include petroleum processing, bulk and commodity chemicals, rubber and plastic products, specialty chemicals and pharmaceuticals, and processes for conversion of undesired by-products from stationary or mobile emission sources into environmentally friendly or recyclable products [ 1]. These products and technologies had shipments that were valued at $US 6.4 x 1011 with chemicals and allied products accounting for nearly one-half of this total [2]. Within the latter category of products, the greatest value was created by sales of drugs, industrial organic chemicals, and plastic materials and synthetics, which is illustrated in Figure 1 [3]. Most of the processes used in these and other related technologies are based on multiphase reaction systems with reactants and products existing as gas, liquid, and/or solid phases along with soluble or solid catalysts. Various reviews summarize catalyst technology commercialized in the United States, Europe and the Far East within the past decade [4-8]. These developments, along with information on the economic drivers associated with the products from these processes, leads to the conclusion that catalyst science and reaction engineering technology are essential for existing processes to maintain economic viability, and for creation of the next generation of efficient processes.

88

Fig. 1. Relative distribution of the shipment value of chemicals and allied products.

1.2. Catalyst development cycle It is generally accepted that discovery, development, and commercialization of catalyst technology for either an existing process or for a novel breakthrough route requires laborious and painstaking efforts by a dedicated team of technologists. Even if diligence and the latest tools for invention are utilized, a breakthrough is not guaranteed. An overview of the various activities that are typically involved in the evolution of a new process catalyst are shown in Figure 2. Each of the phases of the catalyst invention (i.e., discovery, development and commercialization) effort is analogous to a catalytic cycle with both feed-forward and feedback of information. Research

Commercial

Market

Concept "~r ~ Input Catalyst

Application

Modification~,I~ Evaluations~ , ~ \

~-

Catalyst , . .

Synthesis ~..~ Catalyst Characterization

Preliminary ~ Process Economics |

Lab Reactor

Detailed Process Economics

Ra~pldCatalyst ~\9~ Catalyst ~ _~~ 9 ~. BaslcJ~~n " Life _~More Detailed Screening

Discovery "GoodScience.

Kinetics

.

1"

Studies

--~

9Development . .

Kinetics

Pilot-Plant Evaluation

-'~ Catalyst Scale-Up

~ Commercialization

Good Engineering"

Fig. 2. Catalyst invention activities from discovery, development through commercialization.

89 As shown in the upper left of Figure 2, the impetus for discovery of a new catalyst occurs from a new concept that emanates from either research or business-driven market needs. In the discovery stage, new catalyst compositions must be postulated using either knowledge of the given chemistry or by use of previous experience with a related chemical system. The resulting catalyst compositions are synthesized, characterized using various analytical techniques, and subjected to a first level of performance evaluation using a lab-scale catalyst testing system. The information derived from these experiments, such as reactant conversion, product selectivity or yield, activity and global reaction rate can be used to provide basic information needed to obtain a preliminary estimate of the reaction kinetics. The latter information is sometimes used as a key input for assessment of preliminary process economics from which realistic targets for catalyst performance during the discovery stage of research can be surmised. Once a catalyst that meets the minimum required performance standards is identified, the research effort is then shit~ed from discovery toward a development type of activity as shown in the middle cycle. Hence, more detailed evaluation of the catalyst candidates is conducted using a more sophisticated reactor system. This system should be designed so it can provide experimental data that can be used as the basis for discrimination between various proposed kinetic mechanisms and the associated kinetic rate parameters. It should also be capable of providing information on catalyst activity versus time-on-stream for quantification of catalyst deactivation. Since the cost of periodic catalyst replacement or regeneration to maintain plant productivity can have a significant impact on process economics, information on catalyst activity and the catalyst performance parameters over a range of activities is critical for identifying more precise catalyst research milestones. If the minimum required level of catalyst performance versus time-on-stream is not attained, it may be necessary for additional discovery work to be undertaken. Assuming that a catalyst having the required performance is identified that meets the criteria defined by process economics, the development effort can now be focused on the efforts shown in the commercialization part of the development cycle shown in Figure 2. The key activities here include scale-up of the catalyst recipe, performing additional catalyst performance evaluations in a pilot-scale reactor system, interpretation of reaction kinetics and catalyst deactivation, reactor modeling, and translation of these results into a form that can be used to create a more accurate process economic model. It can be surmised that the catalyst would most likely be subjected to a more harsh reaction environment when compared to previous discovery and development cycles. Hence, the potential for unforeseen factors to have a negative impact on catalyst performance increases at this final and critical stage of development. This may require ongoing research effort to solve any issues around the catalyst itself, such as recipe modification, preparation variables, and activation protocol. Once the catalyst meets the specifications that have been defined for commercialization, the technical team responsible for incorporation of the catalyst into the full-scale process would take the lead.

1.3. Catalyst development methodologies The above discussion suggests that utilization of a sequential approach for catalyst discovery and development reactor selection, and catalytic process design would not be preferred since it would be too time consuming and hence difficult to justify from a project funding perspective. A parallel approach in which various key activities in the catalyst discovery, development and commercialization cycles are simultaneously initiated with

90 closely coupled interaction between the cycles is clearly needed to be economically justified. This paradigm of conducting simultaneous catalyst development, reactor selection, and process synthesis is embedded in the concepts first set forth by Villermaux [9] and later expanded by Lerou and Ng [10]. However, before any serous efforts can be devoted to reaction engineering and postulation of various process concepts, an experimental lead that demonstrates feasibility for the catalytic chemistry of interest is preferred. In the last few years, the development and implementation of parallel or high throughput screening methods for catalyst discovery have gained increasing attention. This is partly supported by Figure 3, which shows that a notable increase in the number of US patents that are generally concerned with the development and application of combinatorial methods for catalyst discovery or development. The application of these methods in catalysis is an extension of similar concepts that have been used to accelerate drug discovery in pharmaceuticals research [ 11]. Most of the open literature on high-throughput combinatorial catalysis is concerned with development of the experimental hardware and demonstration of a proof-of-concept by using a particular gas-solid model reaction. While significant progress is being made toward developing and using high-throughput combinatorial catalysis methods, there have been no company announcements to commercialize a breakthrough process where the key catalyst invention is based on this approach. 250 200 150 100 50 =

~

9 ~--~=-~~,

1988 1990 1992 1994 1996 1998 Year

Fig. 3. Number of US patents yearly issued containing the words "combinatorial" and "catalyst". The main objective of this paper is to provide an overview of recent developments in experimental techniques, supporting devices, and data handling methods as applied to highthroughput combinatorial methods in the context of catalytic process development. Emphasis is placed upon developments that apply to gas-solid and gas-liquid multiphase catalyzed chemistry owing to their technological importance both now and in the future. Another objective is to present a brief overview of various advanced experimental reactors that can provide unique data for a particular catalytic system, but are not yet designed as highthroughput devices. Future needs and areas where advances in experimental devices, analytical methods, and informatics would be useful for accelerating the invention of the next generation of commercial catalytic processes are also discussed.

91 2. HIGH-THROUGHPUT EXPERIMENTAL DEVICES Implementation of high-throughput methods for catalyst discovery and development has prompted the creation of new experimental devices for catalyst synthesis and catalytic performance evaluation. These systems have provided the impetus for development of various devices and methods that are used as sensors, control elements, and analytical characterization, to name a few. ~In this section, an overview is given on the various research tools that have been reported in the literature that address the above aspects of highthroughput methods.

2.1. Catalyst library synthesis One of the claimed advantages of using a combinatorial method for catalyst library synthesis is that it allows, in principle, all possible combinations of the experimental variables to be generated and evaluated. This is a significant departure from the conventional serial approach for generating catalyst candidates in which the composition, preparation, activation, and other variables are studied one at a time. The possible number of combinations is given by Nr = n!/[r!.(n-r)!] where n is the total number of variables and r is the total number of independent variables. Hence a catalyst library containing five elements with eleven discrete compositions would require generation of 1,001 discrete samples. The effort that would be required to accomplish this task using manual synthetic techniques by a single researcher would most likely require at least six months. Development and application of specialpurpose devices for generation of catalyst samples that use parallel high-throughput methods are essential for development of a database that can be used to guide a discovery. The requirements for generation of heterogeneous catalyst libraries are more demanding when compared to soluble organometallic complexes for several reasons [12]: (1) the final heterogeneous catalyst structure exists far from an equilibrium state; (2) heterogeneous catalysts are strongly affected by the synthetic procedures, activation protocol, and other reaction environment procedures; and (3) the elemental or molecular building blocks used to generate the desired composition are often ill-defined from a solid-state structural characterization perspective. Other differentiating factors are that homogeneous catalysis draws upon a large technology base from organic and organometallic chemistry whereas heterogeneous catalysis relies upon a relatively small knowledge base from synthetic inorganic and solid-state chemistry [12]. Hence, different combinatorial approaches must be applied to these two catalyst classes. Some illustrations of the differences between homogeneous organometallic and heterogeneous combinatorial approaches can be seen in recent review by Jandeleit et al. [ 13]. Preparation of catalyst libraries for the discovery phase can be broadly divided into two categories: (1) solution-based methods, and (2) thin-film deposition/based methods. The solution-based methods include those that utilize: (1) multi-tasking robotic workstations [ 1416]; (2) inkjet printhead technology [17], and (3) microjet technology [18]. The thin-film deposition-based methods employ electron beam and thermal evaporation, sputtering, pulsed laser ablation, and chemical vapor deposition (see [60] in [13] for a list of references). A detailed discussion of how these have been applied is provided elsewhere [13]. Web sites provide links to various aspects of combinatorial chemistry (www.combichem.net and www.5z.com). Here, the hardware aspects are lightly sketched. A typical commercially-available multi-tasking robotic workstation for high-throughput solution-based synthesis is shown in Figures 4a and 4b. The key hardware features include:

92 (1) two independent piercing probes for pipetting liquid solutions; (2) reaction blocks containing 6-2m vials where m = 0, ..., 4; (3) a built-in orbital shaker; (4) independently controlled filtration stations for solid-phase unit processes; (5) six on-line solvent sources; (6) inert gas blanketing; and (7) two independent reagent racks that can accommodate up to 60 glass vials. The system software allows a synthesis strategy to be designed and downloaded for implementation through a convenient user interface. Details are available elsewhere [ 14].

A typical array of heterogeneous catalyst powders for a three-component metal oxide that was generated using a robotic workstation is shown in Figure 5. In this example, the reaction block containing the solution was manually placed in a heated vacuum oven and evaporated to dryness. Additional unit operations are needed before the powders can be tested in a catalytic reactor. Typical ones include: (1) removal of the powder from each array holder; (2) additional drying and calcination; (3) granulation and sieving to the desired particle size; and (4) transfer and loading a predetermined amount of catalyst granules into the catalytic reactor. These steps assume that catalyst activation is either performed in situ or not included as part of the testing protocol. Future advances in robotic systems will most likely be able to perform all of the above steps, since these consume the greatest of the total time required to prepare the catalysts for testing. Figure 6 shows a commercially available robotic system that can automatically dispense precise quantities of a powder into a rack of vials, which is one of the time-consuming steps involved in reactor loading. However, this particular system is designed to dispense a single powder into multiple vials as typically encountered in pharmaceutical discovery research. In catalyst discovery, the robotic design would require modification to handle multiple powders that would be created in a synthesis robot. A key requirement would be to minimize cross sample contamination. The utility of using one of the thin-film deposition methods is that handling of liquid and solids-laden is avoided. This may seem as reducing the complexity, but other concerns are introduced. Some of these concerns include obtaining a source of organometallic species that contain the elements of interest, design and safe operation of a vacuum system for containment and delivery of the vapors to the desired surface location, and whether or not the resulting mixed-metal layer structure exhibits catalytic activity that mimics one synthesized using solution-phase methods. These and other related issues are partly discussed in [13] from the context of materials applications versus catalysis.

93

2.2. Combinatorial screening techniques Various techniques have been proposed in the literature for high-throughput screening of catalyst candidates. These techniques can be broadly classified according to a particular type of multiphase catalytic reaction system. This system includes gas-solid reactions, gas-liquid catalyzed reactions, and multiphase polymerization reactions. In this section, recent progress in development of novel hardware configurations will be reviewed.

Gas-solid catalytic techniques. Development of experimental techniques for highthroughput screening of gas-solid catalyzed reactions have received the most attention when compared to either gas-liquid catalyzed or polymerization reaction systems. The two key :lifferences between the techniques for gas-solid systems are in the analytical methods used to extract one or more standard measures of catalyst performance, such as activity, conversion, or selectivity, and the catalyst form that is used. Table 1 gives a summary of the various techniques along with their distinguishing features. The use of a scanning mass spectrometer to detect reaction products from a catalyst array was pioneered by researchers at Symyx Technologies [20-27] who used various model reactions to demonstrate feasibility. Figure 7 shows two of the key system features of the Symyx device that is designed for catalyst discovery [24]. A sectional view of the reaction chamber is shown in Figure 7a which shows the scanning mass spectrometer and a movable surface that contains the catalysts library. Figure 7b shows the method used to expose a particular library member to reaction gas and the subsequent transfer of the product gas to the scanning mass spectrometer. Heating is accomplished by using a selective IR source so that adjacent library members are not in the source beam. This technique was recently applied to screen ternary noble metal catalysts for the oxidative dehydrogenation of ethane [26, 27]. The same concept of using a precision mass spec probe positioning system to sample the reaction product gas from a high-density monolith array has also been described [28], which is shown in Figures 8a and 8b. A system that uses reactors manufactured using microfabrication technology coupled to a mass spec capillary sniffing device has also been demonstrated [29]. An automated high throughput design that can accommodate 10 to 100 micro fixed-bed reactors where each uses up to 5 ml of catalyst powder with sequential sampling of the reaction product gas has been recently commercialized [36]. It can also be

94 customized so that a user-specified analytical device, such as a gas chromatograph or mass spectrometer, can be used. A custom-designed rapid-screening system that used 10 fixed-bed microreactors to identify preferred metal oxide catalysts for the oxidative dehydrogenation of propane was described by Baems and co-workers [28, 37]. Table 1. High throughput screening methods for gas-solid catalyzed systems Method

Substrate

Scanning MS

Si Wafers, Quartz Ceramic monolith Mixed oxide Metals, Ceramics, Glass,... Catalyst support

Scanning MS Stationary MS Scanning MS

REMPI Resonance Enhanced Multiphoton Ionization IR Thermography Alumina disk

Catalyst Form Film

Sampling Mode Sequential

Method of Detection Quadrupole MS

Ref.

[20-

27] Film

i Sequential Quadrupole MS

[28]

Particles Films, particles,

Sequential i Quadrupole MS Sequential i Quadrupole MS

. [28] [29]

Particles

Particles on ),-alumina AMM AMM

IR Thermography Slate Slate Emissivity Corrected IR Thermography (ECIRT) Particles Mixed Parallel oxides, micro reactor supports, ... Particles Mixed Parallel oxides, micro reactor supports,... Film Fluorescence Film FT-IR imaging AMM: Amorphous microporous mixed oxides

Parallel

Microelectrodes with REMPI,

[3032]

Parallel

TOF MS for product verification IR camera

[33]

Parallel Parallel

IR camera IR camera

[34] [35]

Sequential

GC, FT-IR, MS

[36]

Sequential

Multiport valve with MS

[28, 37]

Parallel Parallel

Fluorescence FT-IR

[38] [39]

Another approach that allows parallel screening to be conducted using IR thermography has been investigated as an alternate to the use of mass spectroscopy [33-35]. This method monitors the heat changes that occur as a result of the heat of reaction on the catalyst library surfaces. One of the claimed advantages of this method is that the activity of the catalyst library is monitored as a function of time-on-stream in parallel. However, no information on reactant conversion or product selectivity is obtained so that ranking is based upon

95 temperature versus time data. Using this basis could be misleading depending upon the type of catalytic chemistry that is under investigation.

Fig. 7a. Partial sectional view of the reaction chamber showing the scanning mass spectrometer and library [24].

Fig. 7b. Detailed view of the method used for exposing a library member to the reaction gas and reactor heating [24].

An illustration of the above point is shown in Figure 9a where the selectivities of C02 and maleic anhydride relative to n-butane are compared versus time-on-stream using a fresh vanadium-phosphorus oxide catalyst prepared according to an early patent [40]. The experiment was conducted in a laboratory-scale fixed-bed microreactor at 380~ using a feed containing 9% n-butane, 10% oxygen, 9% nitrogen and 72% helium. The catalyst initially exhibits high selectivity for n-butane to CO2, which is a highly exothermic reaction with (-AHr = 664 kcal/mole). As the catalyst undergoes a solid-state transformation with increasing time-on-stream to the active phase [41 ], the selectivity to maleic anhydride relative to CO2 increases. Figure 9b shows the total heat flow, based upon the addition of the heat flows from the butane oxidation to maleic anhydride reaction and the combustion of butane to CO2. The sudden increase of the heat flow that occurs during the initial part of the catalyst time-on-stream is a consequence of the reaction light-off. Thereafter, the behavior of the heat flow versus time on stream undergoes a slight increase, which is due to the increase in the catalyst activity and change in selectivity. Use of IR thermography for this particular reaction would lead to an incorrect conclusion on catalyst performance since the contribution of selective and non-selective reactions to the heat flow cannot be determined. Hence, the ability to measure product gas composition is essential to properly interpret catalyst performance data along with heat flow or temperature measurements. In summary, high-throughput methods for performing a first-level screening of gas-solid catalyzed reactions are in an early stage of development. Case studies where complex multicomponent catalysts have been identified for a reaction having commercial significance, such as a selective oxidation reaction, have not been reported in the open literature. Methods that allow conventional catalyst synthesis and any pre-treatment steps (e.g., drying, calcination, activation, size reduction, pelletization, etc.) to be conducted at high throughput have not been demonstrated. In addition, the direct coupling of a high-throughput powdered

96 catalyst synthesis system to a micro-fabricated catalyst screening unit that automates all of the unit sub-processes is an unsolved challenge.

Fig. 8a. Microreactor catalyst screening system.A: Array microreactor; B: Capillary sampling probe; C: Mass spectrometer; D: Catalyst pellets; E: Aluminum heating block; F: Insulation [28].

Fig. 8b. Monolith array for high-throughput catalyst discovery [28].

Gas-liquid catalytic techniques. Experimental techniques for high-throughput screening of gas-liquid reactions have been the subject of fewer studies when compared to those for gassolid systems. This mode of operation introduces another level of complexity for several reasons: (1) the magnitude of gas-to-liquid and liquid-to-solid external mass transfer resistances (in the case of heterogeneous catalysts) relative to the intrinsic kinetic resistance should be minimized; (2) in the case of heterogeneous catalysts, solids suspension should be uniform and internal catalyst diffusional resistances should be negligible relative to the interparticle and intrinsic particle resistances: (3) liquid volumes should be small enough to satisfy safety requirements and to minimize the demand for liquid reactants, but large enough so that the design and operation of an individual reactor does not introduce other difficulties; (4) analytical characterization of the reaction products is complicated by the presence of multiple phases, sample volume requirements, and reaction quenching after sampling; (5) operation at elevated temperature and pressure imposes limits on the selection of materials and reaction vessel design requirements; and (6) the choice of an appropriate catalyst loading (i.e., catalyst weight/liquid volume) that provides reasonable global reaction rates without creating mass transfer limitations is difficult without previous data. Other typical concerns that occur when dealing with microscale multiphase systems include the generation of adequate gas-liquid interfacial area through creation of micro bubbles, and creating ideal flow patterns for the gas, liquid and solid phases and maintaining an isothermal reaction environment. Table 2 gives a summary of the various techniques that have been recently proposed for high-throughput screening of gas-liquid catalyzed systems along with their distinguishing features. Figure 10 shows a multiple reactor station that can be used for either liquid phase or gas-liquid catalyzed, which is typical of commercially available systems and some of the systems described in Table 2. Most of these systems are intended for reactions that operate at relatively mild temperatures and atmospheric pressure, which are below the ranges of most commercial gas-liquid catalyzed systems [1, 4-7]. The only exception is the oxidation micro-

97 reactor described in [43], which can operate at 140~ and 10 atm of oxygen. This also points out one of the real advantages of using microreactors to study oxidation reactions, since the small volumes are safer and can be more easily operated in a normal laboratory versus a barricade. However, another issue is that on-line sampling and analysis of microscale gasliquid catalyzed reactors is more complicated than for gas-solid catalyzed microreactors. Capture of a consistent small sample volume, such as 1 to 10 ~tL, and transfer to an on-line analytical device, such as a GC or HPLC, requires the development of special methods that can handle multiple reactors. A special purpose on-line sampling valve that can capture these small volumes with flash vaporization into a GC transfer line was recently demonstrated [50]. This points to one of the future challenges for high-throughput screening methods of gasliquid catalyzed systems. 45

12

30

"~ 15

-

10

o

8

t~ fJ

._Z2

--

Maleic Anhydride

0

.I-

--CO2 r

T

-

.

-

~

5 10 15 Time-on-Stream, hr

6

~

20

I

0

i

l

5 10 15 Time-on-Stream, hr

20

Fig. 9b. Total heat flow versus time-onFig. 9a. Selectivity of maleic anhydride and stream. C02 versus time-on-stream. Conditions: Fresh vanadium-phosphorus oxide catalyst; feed = 9% n-butane, 10% 02, 9% Nz, 78% He, T = 380~ wc = 0.56 g, Q = 37.5 mL/min (NTP), x = 1 s. In summary, experimental systems for high-throughput screening of gas-liquid catalyzed reactions will require a significant development before they can be used on a routine bases under industrially relevant conditions. The work that has been done this far has been mainly focused upon demonstrations using relatively simple chemistries under mild reactions conditions. While it appears feasible to design and fabricate multiphase reactors that have volumes on the order of 1 mL or less, development of systems for catalyst handling, reactant addition, automation, control and on-line sampling with integrated and rapid product analysis have not been demonstrated. Polymerization techniques. High-throughput methods for discovery of new polymerization reactions and catalysts are also in an early stage of development, which is similar to gas-liquid catalyzed systems. Most of the efforts in this area in the open literature have been focused upon combinatorial approaches for olefin polymerization catalysts, which are reviewed in detail in [13]. In one case, a custom high pressure parallel polymerization reactor with a series of 48 reactions chambers was used for screening of ethylene polymerization catalyst performance (see figure 14 in [ 13]), analysis of the polyethylene was performed by high-speed gel permeation chromatography. A combinatorial approach for the synthesis of biocatalytic polymers was recently described [51]. It is expected that techniques that can be applied to a broader range of polymerization systems will soon emerge.

98 Table 2. High throughput screening methods for gas-liquid catalyzed systems.

Catalyst type

Method Description

Application

Analytical Method

Operating Range

Ref.

Immobilized lipase

10 mL stirred batch

Offline GC with FID

65-75~ wc = 0.03 g/ml X ~ 90%

[42]

Ag, Ni, Platinum metals

4 ~tL reactor channel

Diglycerol monolaurate ester synthesis Oxidation of benzyl alcohol

Thin film doped TiOa

96 wells with thin film catalyst

Hydrogen production from water

Chiral catalyst

96 well glass plate 25 stirred reactors

Chiral screening

WO3/Pd Chemooptical sensor Imaging polarimetry Off line GC

..

Catalyst powder 50 ~tm catalyst particles or structured array Homogeneous or heterogeneous

Micro machined packed-bed reactor 12 stirred glass tubes

3 -phase hydrogenations and oxidations 3-phase hydrogenation of cyclohexane Liquid or multiphase reactions

140~

[43]

10 atm 02 wc - 0.25 g/m1 X > 90% Not given

[44]

Not given

[45]

Not given

[46]

Off line GC

Not given

[47, 48]

Off line GC, HPLC, or others

T < 160~ P < 1 arm

[49]

Off line GC with FID

Fig. 10. Carousel reactor station for parallel evaluation liquid phase or multiphase reactions. Key system features include 12 reactors with magnetic stirring, hot plate type heating system, central gas inlet and radial gas distribution system with gas tight PTFE caps and water cooled reflux head [49].

3. CONCLUSIONS Invention of the next generation of catalytic processes will be accelerated by development and application of high-throughput testing. Until now, these methods have been developed without considering reaction engineering issues, such as transport-kinetic interactions. For these methods to have practical applications, they must be benchmarked using more realistic reactions having complex product distributions and catalysts that exhibit a range of activities. It is expected that future methods will be developed for gas-liquid catalyzed and polymerization since these provide the bases for new emerging materials.

99 REFERENCES

1. K. Weissermel and H.-J. Arpe, Industrial Organic Chemistry (2 nd Edition), VCH Press, Weinheim (1993). 2. Anonymous, Stanford Research Institute, Statistical Abstracts, January 1996. 3. Anonymous, Facts and figures for the chemical industry, Chemical and Engineering News, 76 (26), June 29, 1998. 4. J.N. Armour, New catalytic technology commercialized in the USA during the 1980's, Appl. Catal., 78, 141-173 (1991). 5. M. Misono and N. Nojiri, Recent progress in catalytic technology in Japan, Appl. Catal. 64, 1-130 (1990). 6. N. Nojiri and M. Misono, Recent progress in catalytic technology in J a p a n - A supplement, Appl. Catal. A: General, 93, 103-122 (1993). 7. M. Enze and Z. Peiling, Progress in catalytic technology in the People's Republic of China during the 1980's, Appl. Catal. A: General, 95, 1-20 (1993). 8. S.M. Csicsery, Catalysis research in India, The Catalyst Review, 5-12 (19xx). 9. J. Villermaux, Future challenges for basic research in chemical engineering, Chem. Engng. Sci., 50, 2525 (1993). 10. J. J. Lerou and K. M. Ng, Chemical reaction engineering: A multiscale approach to a multiobjective task, Chem. Engng. Sci., 51, 1595 (1996). 11. S. R. Wilson and A. W. Czamik, "Combinatorial Chemistry- Synthesis and Application", John Wiley & Sons, New York, 1997. 12. M Dieterle, G. Mestl, Th. Ressler and R. Schlrgl, Combinatorial heterogeneous catalysis; Partial oxidation with MMO materials, Paper presented at the 2nd Annual Conference on Combinatorial Approaches for New Materials Discovery, San Diego, CA, January 23-25, 2000. 13. B. Jandeleit, D. J. Schaefer, T. S. Powers, H. W. Turner and W. H. Weinberg, Combinatorial Materials Science and Catalysis, Angew. Chem. Int. Ed., 38, 2494-2532 (1999). 14. Charybdis Technologies Inc., Carlsbad, California, USA, (www.charybdis.com). 15. Argonaut Technologies, San Carlos, California, USA, (www.argotech.com). 16. Mettler-Toledo Bohdan, Vernon Hills, Illinois, USA, (www.bohdan.com). 17. E. Reddington, A. Sapienza, B. Gurau, R. Viswanathan, S. Sarangapani, E. S. Smotkin, and T. E. Mallouk, Science 1998 June 12; 280, 1735-1737. 18. S. M. Senkan, and S. Ozturk, Discovery and Optimization of Heterogeneous Catalysts by Using Combinatorial, Angew. Chem. Int. Ed., 38, 791-795 (1999). 19. TAP Inc., Wilmington, DE, USA, www.autoprt.co.uk 20. P. C. Cong, R. D. Dooloen, Q. Fan, D. M. Gianquinta, S. Guan, E. W. McFarland, D. M. Poojary, K. Self, H. W. Turner and W. H. Weinberg, Angew. Chem. Int. Ed., 38, 484 (1999). 21. W. E. Weinberg, E. W. McFarland, P. C. Cong, S. Guan, Mass spectrometers and methods for rapid screening of libraries of different materials, US Patent 5,959,297, assigned to Symyx Technologies, September 28, 1999. 22. B. Jandeleit, H. Weinberg, Putting catalysis on the fast track, Chemistry and Industry, 19, 795 (1998). 23. R. F. Service, High-Speed Materials Design, Science, 1997 July 25,277, 474-475.

100 24. E. McFarland, P. Cong, S Guan and W. H. Weinberg, World Intellectual Property Organization, Publication Number WO 98/15969 (1998). 25. B. Jandeleit, H. W. Turner, T. Uno, J. A. M. van Beek and W. H. Weinberg, CatTech, 4, 101-123(1998). 26. J. R. Engstrom and W. H. Weinberg, Combinatorial Materials Science: Paradigm Shift in Materials Discovery and Optimization, AIChE Journal, 46 (1), 2-5 (2000). 27. Y. Liu, P. Cong, R. D. Doolen, S. Guan, D. M. Poojary, H. W. Turner and W. H. Weinberg, High-throughput synthesis and screening of mixed metal oxides for ethane oxidative dehydrogenation to ethylene, Paper ISO 46 presented at the 6th European Workshop on Selective Oxidation, September 9-10, 1999, Rimini, Italy. See also: Y. Liu, P. Cong, R. D. Doolen, H. W. Turner, W. H. Weinberg, High-throughput synthesis and screening of V-A1-Nb and Cr-A1-Nb oxide libraries for ethane oxidative dehydrogenation to ethylene, Catalysis Today, 61, 87-92 (2000). 28. M. Baems and M. Buyevskaya, Strategies in the Development of Heterogeneous Catalyst for the Partial Oxidation of Propane by Combinatorial and Evolutionary Methods, Paper presented at the 2nd Annual Conference on Combinatorial Approaches for New Materials Discovery, San Diego, CA, January 23-25, 2000. 29. T. Zech and D. Hoenicke, Efficient and reliable screening of catalysts for microchannel reactors by combinational methods, Proceedings of the 4th International Conference on Microreaction Technology, AIChE Spring Meeting, Atlanta, March 5-9, 2000. 30. S. M. Senkan, High-throughput screening of solid-state catalyst libraries, Nature, 394, 350-353(1998). 31. S. Senkan, K. Krantz, S. Ozturk, V. Zengin, I. Onal, High-throughput Testing of Heterogeneous Catalyst Libraries Using Array Microreactors and Mass Spectrometry, Angew. Chem. Int. Ed., 38, 2794-2799 (1999). 32. S. M. Senkan, High-throughput, Parallel Testing of Heterogeneous Catalyst Libraries, Paper presented at the 2nd Annual Conference on Combinatorial Approaches for New Materials Discovery, San Diego, CA, January 23-25, 2000. 33. F. C. Moates, M. Somani, J. Annamalai, J. T. Richardson, D. Luss and R. C. Willson, Infrared Thermographic Screening of Combinatorial Libraries of Heterogeneous Catalysts, lnd. Eng. Chem. Res., 35, 4801-4803 (1996). 34. A. Holzwarth, H.-W. Schmidt and W. F. Maier, Detection of Catalytic Activity in Combinatorial Libraries of Heterogeneous Catalysts by IR Thermography, Angew. Chem. Int. Ed., 37, 2644-2647 (1998). 35. A. Holzwarth and W. F. Maier, Catalytic Phenomena in Combinatorial Libraries of Heterogeneous Catalysts, Platinum Metals Review, 44, 16-21 (2000). 36. Zeton Altamira, Pittsburgh, Pennsylvania, USA, Automated High Throughput Screening Micro-Reactor System for Combinatorial Catalysis, Product Bulletin Rev 1., www.zetonaltamira.com/hts.html. 37. D. Wolf, O. V. Buyevskaya and M. Baems, An Evolutionary Approach in the combinatorial Selection and Optimization of Catalytic Materials, Catalysis Today, submitted (2000). 38. P. Gibbs, R. C. Willson, Imaging Polarimetry for High Throughput Chiral Screening, Paper 336c presented at the 2000 AIChE Annual Meeting, Los Angeles, November 1217, 2000.

101 39. J. A. Lauterbach, G. Oskarsdottir, C. M. Snively, Novel Imaging System for Screening of Combinatorial Catalyst Libraries, Paper 312f presented at the 1999 AIChE Annual Meeting, Dallas, November 1-5, 1999. 40. R. A. Schneider, Process for producing a mixed oxide of vanadium and phosphorus having an improved intrinsic surface area, US Patent 4,043,943, assigned to Chevron Research Company, August 23, 1977. 41. G. J. Hutchings, Effect of promoters and reacting concentration on the selective oxidation of n-butane to maleic anhydride using vanadium phosphorus oxide catalysts, Applied Catalysis, 72, 1-32 (1991). 42. E. Garcia, F. Ferrari, T. Garcia, M Martinez and J. Aracil, Use of microreactors in biotransformation process. Study of the synthesis of diglycerol monolaurate ester, Proceedings of the 4 th International Conference on Microreaction Technology, AIChE Spring Meeting, Atlanta, March 5-9, 2000. 43. S. Isogai, M. W. Losey, M. A. Schmidt and K. F. Jensen, The Application of Microfabricated Reactors to Catalyst Testing--Liquid Phase Oxidations at High Temperatures and Pressure, Paper 336b presented at the 2000 AIChE Annual Meeting, Los Angeles, November 12-17, 2000. 44. C. Br'~xidli, E. W. McFarland and T. F. Jaramillo, Combinatorial Photocatalysis for Hydrogen Production, Paper 336e presented at the 2000 AIChE Annual Meeting, Los Angeles, November 12-17, 2000. 45. P. R. Gibbs and R. C. Willson, Imaging Polarimetry for High Throughput Chiral Screening, Paper 336c presented at the 2000 AIChE Annual Meeting, Los Angeles, November 12-17, 2000. 46. F. SchtRh, C. Hoffmann, A. Wolf, S. Thomson, D. Farrusseng and T. Johann, Parallel Reactor Technology for the Discovery of Novel Catalysts. Paper 330d presented at the 2000 AIChE Annual Meeting, Los Angeles, November 12-17, 2000. 47. M. W. Losey, R. J. Jackman, S. Ajmera, M. A. Schmidt and K. F. Jensen, Micromachined Devices for the Study of Heterogeneous Chemical Processes: Transport and Catalysis in microfabricated Fixed-Beds, Paper 350f presented at the 2000 AIChE Annual Meeting, Los Angeles, November 12-17, 2000. 48. S. Ajmera, M. W. Losey, M. A. Schmidt and K. F. Jensen, Microfabricated Reactors for Catalyst Testing: Novel Design for Kinetics Extraction in Heterogeneous Gas Phase Processes, Paper 336f presented at the 2000 AIChE Annual Meeting, Los Angeles, November 12-17, 2000. 49. Zinsser Analytic GmbH, Frankfurt am Main, Germany, (www.zinsser-analytic.com). 50. P. L. Mills, J.S. McCracken and T. M. Delaney, Novel Micro-Scale Reactor for Liquid Phase and Gas-Liquid Catalyzed Reactions with On-Line GC Analytical System Sampling, Paper 288d presented at the 1999 AIChE Annual Meeting, Dallas, November 1-5, 1999. 51. D.-Y. Kim and J. S. Dordick, A New Methodology to Novel Polymer Materials Discovery: The Generation of Polymer Libraries via Combinatorial Biocatalysis, Paper 336d presented at the 2000 AIChE Annual Meeting, Los Angeles, November 12-17, 2000.



Studies in Surface Science and Catalysis 133 G.F. Froment and K.C. Waugh (Editors) ~9 2001 Elsevier Science B.V. All rights reserved.

105

Hysteresis Kinetics of Propene Oxidation Characterized by a Ag-Re Supported Membrane Reactor Masayoshi Kobayashi, Tohru Kanno, Jun-ichi Horiuchi, Shigeyuki Hoshi, Noritaka Hattori and Junya Togawa Department of Chemical System Engineering, Kitami Institute of Technology, 165 Koencho, 090 Kitami, Hokkaido, Japan The partial oxidation of propene (PP) to propylene oxide over a Ag-Re catalyst immobilized in a porous ceramic tube membrane reactor was studied by a continuous forced flow system under a differential reactor condition. The steady state rate equations for the production of propylene oxide (PO) and carbon dioxide were separately proposed by two different reaction pathways as follows. rpo =kKppKoPppP02/{ (KppPpp+ 1)(KoPo2+ 1) } rc02 = kKppKoPppPo 2/(1+KppPpp+KoP02) 2 In the CO2 -pathway, a characteristic hysteresis was observed according to the propene concentration increase (C-up) or decrease (C-down) operation each of which was separately characterized by own kinetic parameters. An identical stable reaction intermediate (In) which was commonly formed on the surface in the courses of C-up and C-down operations at a given reaction steady state was quantitatively evaluated by the transient response method using two different reactions, In+O2 --~CO2 and In + H2 -~C3H8 ,monitoring the amounts of CO2 (qco2) and propane (qH) produced. The value of qco2/qH gave about three without depending on the concentration of PP and C-up or C-down operation indicating the intermediate to be a structure of C-C-C bond. The surface coverage of In for the C-down was about six times higher than the C-up operation at most. The kinetic modification due to surface silver rearrangement induced by the strong adsorption of In was possibly proposed for the explanation of the hysteresis behavior. The selectivity to PO (S) was drastically varied as S=13 - 55 and S=5 - 12 % depending on the C-up and C-down operations respectively. The total gas flow rate effectively contributed to the enhancement of S as S=I 8 - 4 1 % with increasing the flow rate, indicating an advantage of the membrane reactor (which was characterized by a convection flow in membrane pores) rather than conventional packed bed reactors (which were characterized by molecular diffusion in the pores of catalyst particles).

106

1. I N T R O D U C T I O N Although the oxidation of propene to propylene oxide in heterogeneous catalysis has been challenged by a large number of researchers, no solution is developed still now. To break this difficulty, a drastic breakthrough of reactor system including catalyst modification might be requested. As one possible break though, a membrane reactor system should be considered even though the application to industrial systems has not effectively been realized yet. In laboratory scale application, a large number of researchers have studied to develop new membrane systems with high efficiency of objective products [ 1-6]. Almost of them used the membrane for a selective penetration of reactant gases through it. An idea to use the pores of membrane as a path of reactant convection flow could be proposed instead of the use as a penetration path for gas diffusion. In our previous paper [7], it was mathematically demonstrated by using a computer simulation method that the selectivity and yield of exothermic consecutive reactions in catalytically active porous membrane reactor were sensitively influenced by the convection flow of reactants in the pores. In the conventional packed bed reactors of continuous flow system, it has been demonstrated that the convection flow in the large pore catalyst effectively contributes to enhancement of reactant conversion [8-11 ]. In the present study, our interest is focused on a membrane reactor in which the membrane can be employed as a large pore catalyst for the convection flow of reactants. Propylene epoxidation to produce propylene oxide (PO) is taken as a model reaction of parallel-consecutive reactions by using a Re-Ag immobilized in the pores of a micro-porous-glass membrane (MPG). The objectives of this paper are to clarify (1) the reaction kinetics in partial and total oxidation differing from the packed bed reactors, (2) the advantage of the forced flow membrane reactor for the PO-selectivity (S), and (3) quantitative evaluation of the selectivity efficiency derived from the convection flow in the membrane macro pores.

2. E X P E R I M E N T A L The membrane (Micro Porous Glass=MPG, Ise Industry Production; SiO2 70-80, A1203 4-6, BO3 10-12 wt.%) used in this study was 10 mm OD, 8 mm ID and 50 mm length with the pore diameter of 0.34 micron meters. For the immobilization of the catalyst in the membrane pores, Ag and Re were separately supported. A silver ketenide was formed in the membrane pores by the reaction of the two solutions, a mixed solution of acetic anhydride (3.0 g) and pyridine (50 cm3), at room temperature for 12 hr

107 in a dark room. The silver ketenide supported membrane was dried at room temperature for 36 hr.

Re was supported by a successive procedure of the impregnation of an

ammonium perrhenate solution (0.05g/cm 3) at room temperature for 12 hr and the membrane was heat-treated in a N2 stream at 553K for 24 hr.

The catalyst supported

membrane reactor obtained thus consisted of 0.161 g-Ag, 0.0178 g-Re/1.45g-MPG, and was characterized by the pore diameter of 0.22 micron meters, the pore volume of 0.37 cm3/g and the BET surface area of 3.35 m2/g. For the catalyst stabilization, the membrane was pretreated by a reaction gas mixture of PP(11)-O2( 18)- N2( 71 vol %) at 553K for 24 hr. The reaction was operated by a continuous flow reactor under atmospheric pressure at a total gas flow rate of 50-150 cm3/min and 453-493 K. The gas composition was regulated at Ppp= 0.035-0.4 and Po2=0.03-0.15 atm by using N2 as a balance gas. The total reaction conversion was kept less than 5% through all the reaction conditions satisfying a differential reactor. 3. RESULTS AND DISCUSSION 3.1. Steady state rate expressions for PO and CO2 production

The steady state rates for propylene oxide (PO) and carbon dioxide (CO2) production were separately characterized as a function of PP as shown in Fig. 1. The rate of CO2 production clearly indicated a hysteresis depending on the concentration rise

2.0[ '= t / .~t~ 1.5

/ON

!.0

Q:up

N,~

C):down ~ ' ~

0.8 .~ o

~

0.6 "~

1.o 0.4

x (~ 0.5

x 0.2

0.0 0

0.1

0.2 PC3H6 ] atm

0.3

0.0 ).4

Fig. 1. CO2 and PO production rates as a function of PP at 473 ~ and

PO2 =

0.1 atm.

108 (C-up) or decrease (C-down) of PP, indicating a gradual increase for the C-up and a volcano type for the C-down. The rate of PO production gave, however, an identical PPdependency without depending on the C-up and -down. From this extra ordinal difference, the different reaction mechanisms for the two products were proposed based on a Langmuir-Hinshelwood model assuming a surface reaction controlling. The CO2 production rate equations for the C-up and C-down were commonly derived by a single equation of (1) but with different kinetic parameters. rc = kcKppKoPppPo#(1 +KppPpp+KoPo:) 2

(1)

The linear analysis of the rate data obtained evaluates two different kinetic parameter sets for the C-up and C-down operations as kc=7.23x10 -5 and 3.93x10 -5 mol/g-cat.min, Kpp=4.35 and 147 atm l, and Ko=4.35 and 114 atm -1 at 453K respectively. On the C-down, the surface reaction rate constant is about 0.54 times of the C-up whereas the adsorption equilibrium constant of PP is about 34 times than the C-up, suggesting 2.0

A

C OC((up)

1.5

~)(~

1.0

~ 9

_

1)

0.5

~

0.0

I~ 0

O~'X,~

/,,,~O !

5

~

10

~ 0

~

C -

p~

o..o o.,o

085

,.o

5

1.4 1.2

A B P(m,,+ 0.05 -

I)" -

1.0 e=

0.8% 0.6 • 0.4 r

"-~,.,~----qco2(down)

0.2

qco~(up),

0.0

10

15

20

L

m

"~ 6.0

-

25

30

35

40

5.0

c

E

r

4.0

j 4.0

x eo

2.0

-

3.0~

_

2.0

- 1.o~

(b) 0.0

0.0

0

5

10

0

5

10

15

20

25

30

35

40

T i m e / rain

Fig. 2.

xo o

The transient responses of CO2 (a) and

C3H8

(b) formed by the

reaction of In with 02 and H2 at 493 ~ respectively. The (up) and (down)

correspond

to

the

intermediate

formed

by

the

PP-concentration-increase and-decrease operations respectively.

109 higher surface coverage of PP during the reaction. The PO production rate, on the other hand, was proposed by a single equation of (2) without depending on the C-up and C-down. (2)

rpo=kpoKppKoPppPo2/{ (KppPpp+ 1)(KoPo2+ 1) }

The values of kpo, Kpp and Ko evaluated were 5.64 x10 "6 mol/g-cat.min, 120 and 120 atm -~ at 453K respectively. The PO selectivity (S) was evaluated as 53-12% for the C-up and 12-5% for the C-down with increasing PP-concentration.

3.2. Dynamic behavior of In during the reaction For the quantitative evaluation of the reaction intermediates, the transient response method was applied for the intermediates formed by the two different ways of the C-up and C-down. After the reaction steady states were separately prepared by the C-up and C-down operations at the gas composition of PP(5)-O2 (10)-N2(85vol%), the catalyst membrane was flushed by a N2-stream for 10 min to remove the reaction gas mixture in the

reactor.

Two

different

gas

B-gas:H2(10)-N2(90vol%), were then

mixtures,

A-gas:O2(10)-N2(90vol%)

and

separately introduced into the reactor. Fig.2

illustrates a comparison between the two evolution responses of CO2 and propane formed by the reaction of In with two different reactants 02 and H2 at 493K. According to the C-up and C-down operations, the A-gas produced CO2 of 0.53 x 10.4 (=qco2up)and 2.3x10-4 mol/g (=qco2aow~) respectively, and the B-gas produced propane of 1.71x10 4 (=qHup) and 7.0x10 -4 mol/g (=qHdown) respectively, where qHup and qHdown were commonly evaluated by reducing the blank amount (1.65x10 3 mol/g) which is inactive for the main reaction. Fig.3 illustrates the amounts of In formed in the course of both the C-up and C-down operations as a function of PP-concentration at Po2=0.10 atm and 493K. Comparing the amounts of CO2 and propane formed at any reaction gas composition, one can recognize a rough ratio of about 3 for the two components indicating an identical intermediate species (In) that would keep the C-C -C bond of PP molecule. This intermediate is similar to ones observed on a modified silver oxide catalyst as demonstrated in our previous paper [12]. Noting the evidence that the amount of the intermediate formed by the C-down was higher than the C-up, the characteristic hysteresis behavior would be attributed to the surface coverage difference which rearranged the catalytic active sites of silver surface structure. Our interest is focused on a real intermediate in the CO2 production path. To respond this request, the amount of In (qln) can be related to the CO2 formation rate.

110

0,0-qco~

/0~ / ~~ -o

l l , 5 " qH

6

[

5

\\

/

~ : , l up Q),_[f]down

~

4

3

qH.

__ 0

0.05

O. 1

O. 15

0.2

0.25

0.3

0.35

PC3H6 / atm

Fig. 3. The amounts of In evaluated as CO2 and C3H8 formed by the reaction with 02 and H2 at 493 ~ respectively. Fig.4 illustrates the steady state rate of CO2 formation as a function of qco2 which was formed in both the C-up and-down operations. As can be seen from the figure, a good linear relation was obtained through the two operations suggesting a strong contribution of In to the CO2 production path. 3.4. C h a r a c t e r i s t i c s

of the forced flow membrane

reactor

0" up

~

Q)" down o

8 6

X r

L

4

2 1

3

qco2

5 •

7

10 4 ] mol/g

Fig. 4. The steady state rate of

CO2

formation as a function of

qco2 at Ppp - 0.05 atm, Po2 = 0.1 atm and T = 493 ~

111 Pem (-) 50

5

6

7

8

9

!

!

I

I

!

v.L

o.~

40

~ o

30

Vm= F (-)

!0

////J /

/

-~ 9 2o .. "~

10

0

60

I

I

I

I

I

I

I

70

80

90

100

110

120

130

140

F / ml/min

Fig. 5. The selectivity to PO as a function of the total gas flow rate at at Ppp - 0.1 atm, Po2 = 0.1 atm and T = 483 ~ Fig.5 illustrates the PO selectivity (S) as a function of mass Peclet number. The selectivity was increased with increasing the total reaction gas flow rate (F) passing through the membrane pores, such as S=18-41% according to F - 70-130 cm3/min at 483K, indicating an effective role of the convection flow in the membrane pores for the enhancement of S compared to intraparticle diffusion in spherical catalyst solid supports used for conventional packed bed reactors. This result experimentally proves the validity of our previous work based on the mathematical analysis [7]. 4. C O N C L U S I O N S The efficiency of the forced flow membrane reactor was interestingly examined by using the partial oxidation of propene to propylene oxide over a Re-Ag immobilized in the pores of membrane. The differential continuous flow reactor was designed as to be a parallel

reaction

producing

CO2

and

PO

mainly

and

the

two

different

Langmuir-Hinshelwood models were separately applied based on the two different reaction pathways. The CO2 production pathway has a stable intermediate, which strongly

regulate

the

catalyst

surface

states

inducing

two

different

Lngmuir-Hinshelwood mechanisms according to the PP-concentration increase (C-up) or decrease (C-down). The surface coverage of the intermediates on the active sites in the C-down is about six times larger than the C-up at PP-0.05, Po2=0.10 atm and 493K. The surface state change induced by the strong adsorption of In might contribute to both

112 the surface reaction rate constant reduction for the C-down operation as about 0.54 times of the C-up and the PO-selectivity reduction from 53 to 12 %. The convection flow in the membrane pores effectively contributes to the PO-selectivity enhancement indicating an advantage rather than the conventional packed bed reactor systems.

REFERENCES [1] H. P. Hsieh, Catal.Rev.-Sci.Eng., 33(1991)1. [2] H. P. Hsieh, Inorganic Membranes for Separation and Reaction, Elsevier Science, Amsterdam, 1996. [3] J. N. Armor, J. Membrane Sci., 147(1998)217. [4] G.Saracco and V.Secchia, Catal.Rev.-Sci.Eng., 36(1994)305. [5] J. Zaman and A. Chakma, J.Membrane Sci., 92(1994)1. [6] K. K. Sirkar, P. V. Shanbhag and A. S. Kovvali, Ind. Eng. Chem. Res., 38(1999)3715. [6] M. Harold, V. T. Zaspalis, K. Keizer and A. J. Burggraaf, Chem. Eng. Sci., 48(1993)2705. [7] B. Golman, K. Shinohara and M. Kobayashi, J.Chem.Eng.Japan, 30,3(1997)507. [8] A.Nir, Chem. Eng. Sci., 32(1977)925. [9] A.Nir ands L. M. Pismen, 32(1977)35. [ 10] A. E. Rodrigues, B. J. Ahn and A. Zoulalian, AIChEJ., 28(1982)541. [11 ] J. C. Lopes, M. M. Dias, V. G. Mata and A. E. Rodrigues, Ind. Eng. Chem. Res., 35(1995)148. [12] M. Kobayashi, Canadian J. Chem. Eng., 58(1980).

Studies in SurfaceScienceand Catalysis 133 G.F. Fromentand K.C. Waugh(Editors) Published by ElsevierScienceB.V., 2001

KINETICS OF OXIDATIVE DEHYDROGENATION OF LPG TO OLEFINS ON Dy-Li-CI-Zr-O CATALYST M.L.Kaliya, O. V. Malinovskaya, M.V. Landau and M. Herskowitz Blechner Center for Industrial Catalysis and Process Development Department of Chemical Engineering Ben-Gurion University of the Negev, Beer-Sheva, Israel P.F. van den O0sterkamp Technip Benelux B.V., Zoetermeer, The Netherlands

Abstract The kinetics of oxidative dehydrogenation of LPG, a mixture of 50mo1% propane, 25mo1% butane and 25mo1% isobutene, on a high-performance Dy-Li-C1-Zr-O catalyst was studied. Experimental dath measured over a range of concentrations and temperatures was the basis for the development of a kinetic model. Sixteen reactions that include oxidative dehydrogenation and cracking to olefins and oxidation to carbon monoxide and carbon dioxide were included in the model. The kinetic parameters were estimated using SIMULSOLV. Comparison of predicted and .experimental data measured in two different modes of operation yielded a good agreement.

Introduction Oxidative dehydrogenation of light alkanes to olefins may offer a viable altemative to commercial processes of catalytic dehydrogenation and steam cracking. It displays no chemical equilibrium limitations like the dehydrogenation process and does not require high temperature and high heat fluxes needed in steam cracking. An oxidative route to light olefins requires a selective oxidation of hydrogen without substantial oxidation of hydrocarbons to carbon oxides, as this will reduce the yield of desired products. In addition, the oxygen conversion should be as high as possible, as oxygen is not desired in downstream (cryogenic) separation of the products. Recent reviews of experimental results [1, 2] indicate that olefins selectivity higher than 70wt% was measured at alkanes conversion below 50%. Attempts to modify catalyst composition [3], application ofzeolites (ZSM-5) and boralite did not improve the catalyst selectivity [4]. Replacing oxygen with CO 2 [5, 6] or N20 [7] reduced the contact time to 5 ms [8] also were not enough effective. Application of membrane reactors aimed at controlling oxygen supply to the catalyst [9,10] displayed significant improvements of olefin yield. Supported Dy-Li-C1-O catalysts displayed high selectivity and yield in oxidative dehydrogenation of light alkanes [11, 12]. Zirconia support displayed specifically high performance level [13]. This catalyst was selected for detailed kinetic study of oxidative dehydrogenation.

113

114 The main scope of this study was to develop an empirical kinetic model to examine various reactor configurations designed to improve performance.

Experimental

The experimental rig consisted of a quartz tubular preheater and reactor in series, 17 mm ID and 250 mm long with a 5.5 mm thermowell. Electric tapes controlled by Eurotherm heated the reactor and preheater. 2g of catalyst (granules, diameter of 2mm, 1,5 mm length) diluted with 8g of inert SiC pellets to keep the reactor isothermal, was located between two layers of SiC particles of 3-4 mm in diameter. Oxygen staged experiments (SOF - staged oxygen feeding) were performed in a reactor 17 mm ID, 550 mm length with six side inlets for oxygen and a central thermowell, 5.5 mm OD. A schematic setup of the reactor and catalyst loading for experiments with oxygen staging is shown in Fig.1. Hydrocarbons and nitrogen or steam with a fraction of total oxygen were fed at the main inlet of reactor while rest of the oxygen was split between 5 side inlets. 32 g of catalyst was loaded into 6 layers separated by the layers of SiC Butane, isobutane, propane, oxygen and nitrogen were fed from cylinders. Brooks mass-flow meters controlled the flow rate. Water was fed by a metering pump and mixed with other gases in the preheater. Two condensers were staged in series after the reactor, both cooled to 15~ Condensable organic compounds and water were condensed while other gases flowed to GC for analysis. The analysis of volatile reaction products was performed on line with GC HP-5890 that contained four columns - 45/60 Molecular Sieve 13X, 10 ftx 1/8"; 50 m x 0.53 mm Plot A1203; 80/100 Hysep Q 4 ft x 1/8" and 1 ft x 1/8" with internal switching valves and two detectors TCD and FID controlled by ChemStation analytical software. Analysis of water collected in the condensers displayed only traces of oxygenates.

Fig.1 Schematic Diagram of SOF reactor

115 The catalyst was prepared by impregnation of zirconium oxide pellets MEL-706 (supplied by MEL Chemicals), 3-4 mm long and 2 mm in diameter, calcined for 2h at 550~ The pellets were evacuated for lh at 25~ and impregnated with aqueous solution containing 19 wt% Dy(NO3)3, 34% LiNO3 and 42 wt% NH4C1 where the weight ratio of solution to pellets of three. The pellets were filtered, dried at 110~ for 16h and calcined at 550~ for 2h. This procedure was performed twice, with a final calcination at 750~ for 16h. V-Mg-O and Mg-Dy-Li-C1-O catalysts were prepared as reported in [ 14]. Catalyst characteristics are listed in the Table 1. The catalysts chemical composition was analyzed by EDAX method (JEM-35CF SEM, Jeol Co. Japan) excluding Li, which was detected by atomic adsorption spectrometry (Varian Spectrometer AA 375) after dissolution of sample. The catalysts surface area was measured (BET method, ASTM D3663-84). Table 1

Characteristics of catalysts .

.

.

.

SA, m2/g

Catalyst

Chemical composition, %wt

Zr-Dy-Li-C1-O

Dy203-11,9; Li20-4.2; CI-6.2; ZRO2-64.6, A1203-7.6

20

V-Mg-O

V205-7; MgO-93

60

Mg-Dy-Li-CI-O

MgO-80; Li20-9; Dy203-0.2; CI-10

20

Results and discussion The olefins selectivity, defined as weight of formed olefins divided by the weight of converted LPG, measured on Mg-Dy-Li-CI-O[14] and Zr-Dy-Li-CI-O catalysts was similar as indicated by data in Table 2. However, superior selectivity was registered in comparison with V-Mg-O catalyst due to higher rates of hydrocarbon cracking that yielded ethene, propene and methane. All the results reported in this paper were measured with the Zr-Dy-Li-C1-O catalyst. The catalyst performance depends, as expected on the oxygen to hydrocarbon feed ratio (Fig. 2) and the temperature (Fig. 3). The olefins selectivity increased with decreasing oxygen to LPG feed ratio and increasing temperature. The yield of olefins was essentially constant over a range of oxygen to LPG feed ratio and increased mildly with temperature. Table 2

Performance of catalysts V-Mg-O and Zr-Dy-Li-CI-O and Mg-Dy-Li-CI-O T=630~

Catalyst

Conv.

Selec

wt%

wt%

V-Mg-O

31.2

Zr-Dy-LiC1-O

Mg-DyLi-CI-O

O2/LPG=0.8, WHSV=4.1h"1

C2H4

C3H6

n-C4H8

i-C4Hs

C4H6

CH 4

CO 2

44.9

0.2

8

0.3

0.3

6

2

55

28-

33.1

54.7

17.6

18

0.8

0.6

0.3

10.7

38

13.7

12.3

76.4

18

26

3

2

0.4

6

28

15

CO

116

Fig.2 Effect of oxygen on the process parameters. Conditions as at Fig3

70

Fig.3 E f f e c t o f temperature on the catalyst performance Initial Po=, PUG, P m o - 0 . 0 8 . 0.2, 0 . 6 5 a t m , W H S V - 4 . 1 h "1

60 50. 40.

/ /

30.

f

20-'*- Conver. LPG -4- Yield LPG - * - Selectivity

.

o.1

.

.

0.2

.

0.3

.

9

0.4

0.5

0.6

.

0.7

0

.

0.9

8

--.- Cony. LPG - - - Selectivity --*-- Yield LPG

100 590

600

610

620

630

640

650

660

Tempe~re, eC

Table 3 Oxidation of pure paraffins at low oxygen T=650~ Alkane

WHSV=4.5h -l

Ethane

Propane

Butane

Isobutane

LPG conversion,%

10.1

17.2

19.2

23.8

Oiefin selectivity, wt%

50.3

70.5

70.0

71.6

Olefin yield, wt%

5.1

12.1

17.5

17.0

43.4

40.2

69.1

56.6

57.4

3.1

wt% of C3H6 in olefins wt% of C2H4 in olefins

100

27.5

wt% of i-C4H $ in olefins 2.4

wt% of C4H $ in olefins 0.8

0.9

15.3

CH4/C3H 6

1.0

1.1

0.7

C2H6/C2H 4

0.05

0.14

0.8

CH4/C2H 4

0.1

Oxidative dehydrogenation of pure hydrocarbons at low oxygen partial pressure was tested to determine the performance at low conversion. The partial pressure of alkane, oxygen and steam was 0.2, 0.02 and 0.6, respectively. Nitrogen was used as diluent. The results listed in Table 3 indicate that lack of oxygen in the feed yielded negligible conversion (below 0.5%). Increasing the oxygen partial pressure to 0.02 atm increased

670

680

117 conversion from 10% for ethane to 24% for isobutene. Cracking was an important route to production of olefms for propane, isobutene and especially butane. A detailed kinetic study of oxidative dehydrogenation of propane, isobutane, n-butane (23 runs) and LPG (27 runs) was conducted over a wide range of partial pressures of pure and mixed hydrocarbons (0-0.3 atm), oxygen (0-0.2 atrn) and steam (0.2-0.7) atm and temperature 600-670~ Oxidation of H 2, C3H6, C2H4, CH 4 and CO was also tested at 600-650~ A set of reactions was selected based on the distribution of products: 1. Oxidative dehydrogenation and oxidative cracking of alkanes to alkenes C3H 8 + 0.502"-C3H 6 q- n 2 0

(K-l)

n-C4Hlo + 0.502 --n_C4H8 + H20

(K-2)

n-C4Hl0 + 0.502 = 2-c,t-C4H 8 + H20

(K-3)

n-Chill0 + 0.502 "-2C2H 4 + H20

(K-4)

i-C4Hlo + 0.502--i-C4H 8 + H20

(K-5)

2. Cracking of alkanes to alkenes

C3H8 = C2H4 + c n 4 n-C4Hlo = C2H4 + C2H6

n-C4Hlo = C3H6 + CH4

(K-6) (K-7) (K-8)

3. Partial oxidation of alkanes to carbon monoxide CH4 + 0.502 "- CO -k 2H2

(K-9)

4. Partial oxidation of alkanes to carbon monoxide C2H4 q" 02 -- 2CO + H2

(K-10)

C3H6 + 1.502 = 3CO + H2

(K-11)

5. Partial oxidation of alkanes to carbon dioxide

CH4 + 02 = CO2 + 2H2

(K-12)

6. Partial oxidation of alkanes to carbon dioxide

C2H4 + 202 = 2CO2 + 2H2

(1(-13)

C3H6 + 302 = 3CO2 + 3H2

(1(-14)

7. Oxidation of carbon monoxide and hydrogen

CO + 0.502 - CO2

(K-15)

HE + 0.502 "- H20

(K-16)

Analysis of data using SIMULSOLV software proved that this set of reactions yields a good fit. Furthermore, specific experimental observations, like the ratio of methane to olefins and the negligible dehydrogenation rates, were considered.

118

Considering the complexity of the kinetic system, the simplest empirical expressions that fitted the data were selected. Reactions (1)-(4) were modeled by a LangrnuirHinshelwood expression: ri= kiyjY4/(l+K4jY4), i=1,4

j=l,2

(1)

Reaction (5) was found to be a first-order reaction with respect to isobutane rs=ksY3

(2)

where Yl, Y2, Y3 and y4 are the partial pressures of propane, butane, isobutane and oxygen, respectively. Power-law type expressions were employed for all other reactions (6-16): O~i ~i

ri = kiYj Y4

(3)

where yj is the partial pressure of the reactant. The kinetic constants derived from the analysis of data using SIMULSOLV are listed in Table 4. Table 4

Kinetic Parameters

ki at 650~

Reaction

Units

czi

l~j

E/R*I 0 "3, K

kg/kg cat. sec. atm 2

4•

2

2100• i i 205•

kgJkg cat.

4+0.1

3

1120•


5•

4

i 1250-a=85

kg,/kg cat. sec. atm 2

4:t:0.1

5

11+1.8

kg/kg cat. sec. atm

4•

6

130•

k~kg cat. sec. atm 1"5

1

0.5

7

7•

kgJkg cat. sec. atm

1

0

10•

8

170•

kg/kg cat. sec. atml5

1

0.5

10•

9

4060•

kg/kg cat. see. atm 2

1

1

2•

10

3+0.17

kg/kg cat. sec. arm

0.5

0.5

5•

11

176•

kgJkg cat.

0.5

1

8•

12

2140•


1

1

2•

13

45•

k$/kg cat. see. atm

0.5

0.5

4•

14

1.5•

kg/kg cat. sec. atm

0.5

0.5

7•

15

1580•

kg/kg cat. sec. atm 1"5

1

0.5

4•

16

1000•

kg/kg cat. see. atml5

1

0.5

7•

K.41

73•

atm -1

K42

31•

arm -1

1

I

sec. atm 2

s e c . a t m 1"5

11•

119

A parity plot of experimental and predicted of ethene and propene mol% in the product shown in Fig. 4 represents the good agreement between model predictions and data. Fig.4 Comparison of predicted and experimental values of ethene and propene content 0

E =~ 3.s

Y

u ~

2.5

~

2

~ r

1.5

g

1

8

t

9

*

predicted values of propene 9Predicted values of ethenq

~- O.5 D_ 0_ I

,

0

0.5

1

1.5

2

2.5

3

3.5

Pmpene, ethene c~ntent expedmental, %mol.

Fig.5 SOF affects on the oxygen conversion

and yield of

CO and C02. 100

98

0
-

4

88 0

0 4

6

Number of oxygen inlets

8

12

120

The kinetic model facilitates a detailed analysis of the optimal operating conditions that could be demonstrated experimentally. Keeping the oxygen concentration low improves selectivity to olefins. Therefore, the staged feeding option was modeled and tested experimentally.

Stagedfeeding reactor An isothermal plug flow model that employed the kinetic model demonstrated that staging oxygen feed enhanced yield of olefins. Model predictions at 650~ shown in Fig. 5, indicate that staged oxygen feeding (SOF) decreased oxygen conversion and carbon dioxide production compared with single point feeding (SPF). Oxygen affects not only on the rate of oxidative route but also on the cracking of LPG. Therefore improvement in oxygen distribution along catalyst layer increased LPG conversion. The improved performance of SOF operation is illustrated in Table 7. Data measured with the reactor shown in Fig. 1 demonstrated the capability of the kinetic model coupled with the proper mass balances to predict LPG conversion and olefin selectivity.

Table 7

Comparison of experimental and predicted results in SPF and SOF operation WHSV = 1.1h"1, 650~

Molar ratio 02 inlet split (ratio) O2/LPG

_

PLvc = 0.2 atm, Pmo = 0.6 atm

LPG conv

LPG conv

Olefin selec.

pred., wt%

exp., wt%

pred., wt%

Olefin selec. exp., wt%

0.40

1:0:0:0:0:0

41.9

39.7

66.7

66.0

0.40

1:1:1:1:1:1

49.0

50.3

69.2

69.5

0.45

1:1:1:1:1:1

48.9

50.4

69.8

69.5

0.45

2:1:1:1:1:1

51.3

49.8

70.4

70.3

0.66

4:2:1:1:1:1

56.0

55.3

66.9

66.8

0.72

4:2:1"1:2:1

57.8

56.6

66.2

65.3

Conclusions The complex system of oxidative dehydrogenation that includes different types reactions was modeled by a set of sixteen reactions. The kinetic model was tested simulations of various operating conditions. Simulations indicated that staging oxygen feed improve performance. Experimental data measured over a range conditions was in good agreement with the model simulations.

of in of of

121

References 1. F.Cavani, F.Trifiro, Catal. Today, 24, (1995), 307. 2. F.Cavani, F.Trifiro, Catal. Today, 51, (1999), 561. 3. A.Khodakov, O. Olthof, A.BeI1, E.Iglesia, J. Catal. 18 I, 2, (1999), 205. 4. G.Centi, F.Trifiro, Appl. Catal., A, 143 (1), (1996), 3 5. Ge Hin, Zou Hu, Li Kang, Zhang Weiguang, Shen Jianyi, Cuihua Xuebao 20,5 (1999), 515. 6. M. Kralik, V.Macho, E.Jurecekova, L.Jurecek, Chem. Pap, 52, 5, (1998), 682. 7. A.Kenishi, I.Makoto, K.Kazuhiro, J. Chem Soc., Faraday Trans, 1 (1987), 83 (10), 3139.. 8. M.Huff, L.Schmidt, J. Catal. 149, 1, (1994), 127. 9. A. Pantazidis, J.A. Dalmon, C. Mirodatos, Catal. Today 25, (1995), 403 10. G. Capannelli, E. Carosini, F.Cavani, O. Monticelli, F.Trifiro, Chem. Eng. Sci., 51, 10, (1996), 1817. 11.S.J. Conwey, J.H.Lunsford, J.Catal. 131, (1991), 513. 12. Landau M.V, Kaliya M.L, Herskowitz M, van den Oosterkamp P.F, Bocque P.S, CHEMTECH, 24-29, February, 1996. 13.1.Landau M.V, Kaliya M.L, Herskowitz M., EP 804287 (1997). 14. Landau M.V, Gutman A. Kaliya M.L, Kogan L.O., Herskowitz M, Studies in Surface Science and Catalysis, 110, 315-326, 3rd World congress on oxidation catalysis San Diego 1997. 15. F.A. A1-Sherehy, A.M. Adris, M.A. Soliman, R.Hughes, Chem. Eng. Sci., 53,23, (1998)


Studies in Surface Science and Catalysis 133 G.F. Froment and K.C. Waugh (Editors) 9 2001 Elsevier Science B.V. All rights reserved.

123

De Donder relations and the theory of reaction routes Ilie Fishtik and Ravindra Datta Department of Chemical Engineering, Worcester Polytechnic Institute, Worcester, MA 01609-2280, USA There is now a great deal of interest in utilizing the microkinetic approach in modeling rates of catalytic reactions despite the lack so far of reliable rate constants of elementary reactions on different catalytic materials. However, the alternative approaches that provide a simple means of understanding, explaining and predicting the kinetic behavior of complex heterogeneous catalytic reactions continue to be invaluable. The main approximations that are conventionally used to simplify the detailed kinetics are [ 1]: a) Quasi-steady state (QSS) approximation b) Quasi-equilibrium (QE) and rate determining step (RDS) approximations c) Irreversible step approximation d) Most abundant reactive intermediate (MAR1) approximation These approximations also provide elegant physicochemical insights into the mechanism as summarized in two recent papers by Stoltze [2] and Dumesic [3]. Thus, as shown by Stoltze [2], under the RDS approximation many characteristics of the microkinetic mechanisms, e.g., rate equations (De Donder relations), apparent activation energies, apparent reaction orders, etc., may be naturally partitioned into a sum of contributions associated with a special class of reactions involving only one surface species. Dumesic [3] extended the idea to general systems. In this work, we address the problem of partitioning De Donder relations into contributions associated with different types of reactions from the point of view of the theory of reaction routes, or mechanisms. Thus, derivation of reduced routes allows reaction rate analysis in terms of the QSS, RDS and QE formalisms.

1. Notation and Definitions Consider the general case of a heterogeneous catalytic chemical reaction system. The species comprising the elementary reactions that describe the detailed chemistry of the catalytic process are explicitly divided into active sites on the surface of the catalyst S, intermediate (surface) species I1, 12, ..., Iq and terminal (gas phase reactants and products) species T1, T2,...,T,. Thus, the chemistry is assumed to be described by the following set of p elementary reactions q

n

rj = ajoS+ ~,ajkI k + ~_,flj, T~ = 0 ; k=l

j = 1,2,...,p

(1)

i=l

An elementary reaction is one that is amenable to theoretical predictions of its energetics. It is assumed, as usual, that the stoichiometric coefficients ajo,ajk and flj~ take positive values for products and negative for reactions. The rank of the stoichiometric matrix is m < p , i.e., the elementary reactions are not necessarily linearly independent. Each of the elementary reaction is characterized by its affinity that is defined as usual [ 1]

124 ~,q 1 R T Aj = - l n K j + aj0 ln00 +~..~ajk ln0 k + ~.~flj, lnP~ k=l

(2)

i=1

Here Kj. is the equilibrium constant of thej-th elementary reaction, 00is the fraction of the free (uncovered) surface of the catalyst, 0 k is the fraction of the surface covered by the intermediate Ik and, Pi is the partial pressure (for gaseous reaction systems) of the terminal species Ti. The surface coverages are subject to the site balance q

00 + ~ 0 k = 1

(3)

k=l

For simplicity, it is assumed here that all intermediates occupy a single active site each. The rates of the elementary reactions are given in the form of De Donder relations [ 1] Vj = V+j - V j = II+2 1-exp - - ~

,

j = 1,2,...,p

(4)

where V+yand V_j-are rates of the forward and backward reactions. 2. Reaction Routes

A linear combination of the elementary reactions rl, r2,...,rp that eliminate a specified number of species (either intermediates or terminal species) is called a reaction route (RR), or, a mechanism. In the most general case, a RR may be expressed as [4,5] P

(5)

R = ~ ~rjrj j=l

where a~,a 2..... ap is a set of real numbers, called stoichiometric numbers. Each RR produces an overall reaction (OR) if it involves only terminal species or an intermediate reaction (IR) if it involves some intermediates that may be obtained substituting the elementary reactions, Eq. (1), into Eq. (5)

R=s

s q=l

k=l

q=l

O'q~qiWi--O

(6)

i=1 ~, q=l

The above definition of RRs is a generalization of that used in the literature. Conventionally, a RR is defined as a linear combination of elementary reactions that eliminates all of the intermediates. The reason for this generalization will become clear later on. Because, the elementary reactions normally are linearly dependent, the RRs and, consequently, the ORs may be defined and derived in an infinite number of ways. In general, it is useful to limit the number of such mechanisms to a finite and a unique set. This may be done invoking the concept of directness (uniqueness) of both RRs [6] and ORs [7,8]. Here, for brevity, however, we employ the concept of directness only as applied to RRs. We will define two types of direct RRs that avoid arbitrariness by limiting the number of involved elementary reactions. 3. Intermediate Reaction Routes

The first type of RRs involves only q elementary reactions from the set of p reactions while eliminating ( q - 1) intermediates, i.e., all of the intermediates but one

125 (plus vacant sites). Such a RR is called an intermediate reaction route (IRR) while the reaction produced by an IRR is called an intermediate reaction (IR). Let the q elementary reactions involved in an IRR be rj,,rj,,...,rjq (1 < jl < J2 221 ~ gasoline (C5-221 ~ LPG (C3-C4), dry gas (C1-C2) and coke. The gasoline fraction was further divided into paraffins, olefins, naphthenes and aromatics. The conversion was defined as the weight fraction of all products boiling below 221 ~ plus coke. The amount of coke was measured by the TPO-method (Temperature Programmed Oxidation). The catalyst was regenerated after each cracking experiment with 0.5 wt% oxygen in helium by increasing the reaction temperature from ambient temperature up to 830 ~ at the rate of 5 ~ The formed product stream containing CO and CO2 was split into two streams. One stream was passed through a CO2-trap, and the other was bypassed. Both streams were converted into methane over a Ni/),-A1203 catalyst (350 ~ and further analysed by FID. From the amounts of methane in these two streams the amounts of CO and CO2 could be calculated and thus the amount of carbon in the formed coke could be determined. The

167 experimental set-up and kinetic modelling of coke combustion is described by Kanervo et al. [8]. 3. MATHEMATICAL MODELLING 3.1. Mass and heat transfer model

The mass transfer outside catalyst particles was modelled with rigorous Maxwell-Stefan equations, along with simultaneous heat transfer [9]. The multicomponent mass transfer coefficient matrix was calculated by assuming a film model for mass transfer and estimating thejo factor for the packed bed [ 10]. The mass transfer coefficient correlations were extended into multicomponent systems by approximate matrix function calculations [ 11 ]. Temperature and compositions on the catalyst surface (to be used in the kinetic modelling) were iterated with a non-linear algebraic equation solver. The total mass transfer flux was also used as an iterated variable. This was not necessary since the total flux can be calculated from reaction rates with estimated compositions, but this additional variable usually linearizes the system and improves convergence with only a minor increase in computational work. The residuals for the equation solver were obtained by comparing the calculated mass transfer fluxes and the fluxes obtained from the reaction rates. The surface temperature was iterated to satisfy the energy balance. This general approach is necessary for complex reaction schemes where overall stoichiometric coefficients relating single component mass transfer fluxes cannot be found. Characteristic of the mass transfer modelling of this reactor was a relatively high total (convective) flux that results from the non-equimolarity of the reactions. The effect of the high total flux was also included in the mass transfer flux algorithm by a linearized high flux correction described by [ 12]. The following equations were thus used for the mass and heat transfer calculations (N) = ct[k'](XB-Xs) + (XB)Nt

(1)

h(Ts - T.) = ~"~HiN i

(2)

i=l

The effect of intraparticle gradients was assumed to have been included in the estimated parameters, i.e. the reaction rates at the whole catalyst particle was calculated with the surface conditions. The intraparticle gradients were not calculated, because a commercial FCC catalyst was used in the experiments. The bulk gas temperature was assumed to remain constant along the reactor length due to the surrounding heating oven. Binary diffusion coefficients were calculated by the Wilke and Lee equation and the required parameters, characteristic Lennard-Jones energies and lengths for each lump, were calculated by using the average critical temperatures and volumes of the lumps [13, p. 396,587]. The critical properties were calculated from the average product compositions of the lumps (dry gas, LPG and gasoline) and from the distillation curve (gas oil). The heat transfer coefficient was obtained by using the analogy for mass and heat transfer (jD=jn). The overall heat of cracking reactions was assumed to be 380 kJ/kg gas oil converted. In addition, the heat flux by radiation from the reactor tube to the surface of the catalyst particles was accounted for by the Stefan-Boltzmann law [ 10].

168 3.2. Kinetic model and parameter estimation Eight cracking reactions were included in the model: second-order gas oil cracking to gasoline fraction paraffins, olefins, naphthenes and aromatics and to LPG, first-order gas oil cracking to dry gas and to coke and first-order cracking of gasoline olefins to LPG. Product yields for the reactor were calculated by integrating molar flows over the bed volume with the initial values of the kinetic parameters. Rate coefficients at the selected reference temperature, ki(Tref), and activation energies, Eai, for every reaction were determined by the non-linear regression program KINFIT using a combination of the Simplex [14] and the Levenberg-Marquardt [15] methods. The concentrations and temperatures used in the kinetic rate equations were the calculated values at the catalyst surface. The objective function minimised was the square of the differences between the estimated and the observed weight fractions (wt%) of the products at the reactor outlet. The estimated weight fractions were functions of the independent variables (feed molar flow, temperature, pressure and the mass of the catalyst) and the kinetic parameters (rate coefficients and activation energies). 3.3. Reactor model The steady-state, plug flow reactor model was used. Although the reactor was operated by a pulse mode, the reactor model was continuous. This assumption was considered to be valid, because the volume of the vaporised oil is three times as large as the volume of the catalyst bed and the mixing of the oil with the cartier gas was assumed to happen only at the edges of the pulse. These edges were supposed to represent only a minor fraction of the total pulse volume, which means that the mixing of the hydrocarbons with the carrier gas has a negligible effect. 4. RESULTS AND DISCUSSION Measured and estimated yields of gas oil, gasoline and LPG and dry gas and coke as functions of the reactor temperature are shown in Figs l a and lb.

25-

i

,

_~ .......

=

Ii "i

I1 40 .~

...., 30-

15.

-15

._~ >..

10 2O 5

10

0

550

600

Temperature, 9 e -

....

-dr . . . . .

'5O0

_

= 9

Oil meas, wt% LPG est, wt%

650

est, wt% 9 Gasoline meas, wt%

- o -Oil

700

500

600

550

Temperature,

~

9 LPG meas, wt% Gasoline est. wt%

- A .

Fig. l a. Measured and estimated yields of gas oil, gasoline and LPG as functions of the reactor temperature. C/O-ratio 4.0 and gas residence time 55 ms.

9 9

Dry Gas meas, wt% Coke meas, wt%

650

700

oC

- Q - Dry - A -Coke

Gas est, wt% est, wt%

Fig. lb. Measured and estimated yields of dry gas and coke as functions of the reactor temperature. C/O-ratio 4.0 and gas residence time 55 ms.

169 The measured and estimated yields of paraffins, olefins, naphthenes and aromatics in the gasoline fraction as functions of C/O-ratio are shown in Fig. 2.

30 25

A

9 A -

=

9

"&-.~ 9

* 20

i

.-.x

, .O

~15 --

9

~.x~ 9

--

Paraffins meas, wt%

- -o- - Paraffins est, wt%

-'~'-" -o. 9 "-:--.-.- |

9 Olefins meas, wP~

9

- -z~ - Olefins est, wt%

e 99 -

|

9 Naphthenes meas, wt% - -o- - Naphthenes est, wt%

>'10

f" ~;

9o n

~

H H

-'='

--

9 Aromatics meas, wt% - -o- - Aromatics est, wt%

5

~r

0

*

i

0

10

i

20

!

30

i

40

50

C/O-ratio

Fig. 2. Measured and estimated yields of gasoline fraction paraffins, olefins, naphthenes and aromatics as functions of the C/O-ratio. Reactor temperature 550 ~ and gas residence time 55 ms. As can be seen, the estimated yields correspond well with the experimental results. The model is able to predict the yields accurately as functions of temperature and at low C/Oratios. However, at high C/O-ratios the gasoline fractions are not well predicted suggesting that there exist secondary reactions, such as hydrogen transfer and further cracking. The estimated rate coefficients and activation energies with 95 % confidence limits are shown in Table 1. The confidence limits show that the parameters are well identified. Table 1 Estimated rate coefficients and activation energies with 95 % confidence limits. Cracking reaction Gas oil to gasoline paraffins Gas oil to gasoline olefins Gas oil to gasoline naphthenes Gas oil to gasoline aromatics Gasoline olefins to LPG Gas oil to coke Gas oil to dry gas Gas oil to LPG

Rate coefficient (at 850 K), m 3n mol 1-n kgcat-1 s -1 0.015 0.064 0.013 0.035 0.15 0.042 0.045 0.019

+ 0.002 + 0.004 + 0.002 + 0.002 + 0.01 + 0.008 + 0.008 + 0.002

Activation energy, kJ mo1-1 55 + 13 99 + 6 75 + 14 95 + 8 132 + 12 100 + 16 186 + 14 97 + 18

170

The high activation energy for gas oil cracking to dry gas is an indication that this reaction is of a thermal nature. In Figures 3a and 3b the calculated mole fractions in the bulk gas phase and at the catalyst surface for gas oil, gasoline and LPG are shown as functions of the reactor length at reactor temperatures of 550 and 650 ~ In Figure 4 the calculated temperature differences between the bulk and the catalyst surface are shown as functions of the reactor length at reactor temperatures of 550 and 650 ~ The profiles are calculated with estimated kinetic parameters.

09/ 1,0

tJas

Ull

Gasoline

[

~....~

'

I~ ~ I

(

LPG

~

Gas Oil,bulk Gas Oil,surf. Gasoline,bulk Gasoline,surf. LPG,bulk LPG,surf.

'

~-

o,a Gas Oil

0,7

!~

0,6 |

0,5 ~ ; ; ; = = , . . . . . , . ~

0.40,3[~(~

~nsolina

Lp,'~

Gas Oil,bulk Gas Oil,surf. Gasoline ,bulk Gasoline,surf. LPG,bulk LPG,surf.

0,2 ~ ~ - - - - - ' - ' - ~ - 0,1

0,0 ! 0

10 20

30

40

50 60

70

0

80 90 100

. . . . . . . 10 20 30 40 50 60

70 80

90 100

Reactor length, %

Reactor length, %

Fig. 3a. Mole fractions of gas oil, gasoline and LPG in the bulk and at the catalyst surface at the reactor temperature of 550 ~

Fig. 3b. Mole fractions of gas oil, gasoline and LPG in the bulk and at the catalyst surface at the reactor temperature of 650 ~

80 70

.: 60

50

oC) dT(550 oC)

- - - - dT(650

40

30 20

!

i

i

i

i

i

10

20

30

40

50

60

70

80

90 1

Reactor length,%

Fig. 4. Temperature differences between the bulk and the catalyst surface at reactor temperatures of 550 ~ and 650 ~

171 These results show that in the initial part of the reactor large differences in concentrations and temperatures exist between the bulk and the surface. This means that the rate of the mass transfer is rate determining. As the temperature and therefore the rate of the reaction increases these gradients become steeper. As an example, the estimated rate coefficient of gas oil cracking to gasoline olefins at 850 K is 0.045 m6/(mol s kgcat) when the concentration and temperature gradients are neglected and 0.064 m6/(mol s kgcat) when these effects are included in the estimation. Therefore, neglecting the gradients would lead to the low rate coefficients being too low. The effect on the activation energies is very small. From these results it is obvious that in estimating the kinetic parameters from the experimental results ignoring mass and heat transfer gradients would lead to parameter values which are affected by these gradients. Therefore the parameters would not be directly applicable to other reactor types were these gradients may be totally different due to the different hydrodynamic conditions. NOTATION

(N) [k'] XB, XS Nt, Ni H ct TB, Ts HI

vector of mass transfer fluxes high flux corrected mass transfer coefficient matrix mole fractions in the bulk and at the surface total flux, mass transfer flux of component i heat transfer coefficient total concentration temperature in the bulk and at the surface partial molar enthalpy of component i

(mol/m2s) (m/s) () (mol/m2s) (W/m2K) (mol/m 3) (K) (J/mol)

REFERENCES

1. Perego, C. and Peratello, S., Catal. Today 52 (1999) 133-145 2. Jakkula, J., European Refining Technology Conference, 17-19 November, London (1997) 3. Lee, L.-S., Chen, Y.-W., Huang, T.-N. and Pan, W.-Y., Can. J. Chem. Eng. 67 (1989) 615-619 4. Ancheyta-Jufirez, J., Lrpez-Isunza, F., Aguilar-Rodriguez, E. and Moreno-Mayorga, J.C., Ind. Eng. Chem. Res. 36 (1997) 5170-5174 5. Pitault, I., Nevicato, D., Forissier, M. and Bernard, J.-R., Chem. Eng. Sci. 49 (1994) 4249-4262 6. Jacob, S. M., Gross, B., Voltz, S.E. and Weekman Jr., V.W., AIChE. J. 22 (1976) 701-713 7. Lipi~iinen, K., Hagelberg, P., Aittamaa, J., Eilos, I., Hiltunen, J., Niemi, V.M. and Krause, A.O.I., App. Catal. 183 (1999) 411-421 8. Kanervo, J.M., Krause, A.O.I., Aittamaa, J.R., Hagelberg, P.H., Lipi~iinen, K.J.T, Eilos, I.H, Hiltunen, J.S and Niemi, V.M., accepted for publication in Chem. Eng. Sci. 9. Taylor, R. and Krishna, R., Multicomponent Mass Transfer, Wiley, New York (1993) 10. Brodkey, R.S. and Hershey, H.C., Transport Phenomena, McGraw-Hill, New York (1988) 11. Alopaeus, V. and Nordrn, H. V., Computers & Chem. Eng. 23 (1999) 1177-1182 12. Alopaeus, V., Aittamaa, J. and Nordrn, H. V., Chem. Eng. Sci. 54 (1999) 4267-4271 13. Reid, R.C., Prausnitz, J.M. and Poling, B.E., The Properties of Gases and Liquids, McGrawHill, New York (1988) 14. Nelder, J.A. and Mead, R., The Computer Journal 7 (1965) 308-313 15. Marquardt, D.W., J. Soc. Indust. Appl. Maths 11 (1963) 431-441



173

KINETICS OF ETHYLENE POLYMERISATION OVER C R Y ZEOLITES y. Zhanga, b, X. Wang a' b, M.A.N.D.A. Lemos a, F. Lemos a, R.T. Henriques a, M.M. Marques a aDepartamento de Engenharia Quimica, Instituto Superior T6cnico, Av. Rovisco Pais, 1049-001 Lisboa, Portugal. bInstitute of Industrial Catalyst, Dalian University of Technology, 158 Zhongshan Road, 116012 Dalian, China. In the present study, the use of a chromium-exchanged Y zeolite for the polymerization of ethylene was studied, with a particular emphasis on the kinetic analysis. Chromium/zeolite catalysts have been prepared both by ion-exchange and by impregnation with chromium salts and the activity of the various samples was measured, as a function of time, at various temperatures and pressures. The rate of monomer consumption presented a typical shape, with a sharp increase at the beginning of the run, followed by a decrease that tended to an almost stationary rate. These results are discussed in terms of the possible mechanism for the reaction.

1. INTRODUCTION The polymerisation of light olefins is a very important industrial process because polyethylene and polypropylene have a large demand for a wide range of products. The catalysts have been evolving rapidly and Ziegler-Natta polymerization has been boosted by the discovery of very active and selective transition-metal catalysts. In recent years the heterogenization of these very efficient homogeneous catalysts has been attracting a lot of attention [ 1-8]. The use of supports like silica has the distinct disadvantage that the catalyst often leaches out of the support, depositing, for instances, on the walls of the reactor, and causing operational problems. Zeolites, with their particular porous system, seem to be likely candidates for the role of support since it is possible to entrap catalytic species inside their micropores.

2. EXPERIMENTAL 2.1 Materials

NaY zeolite (LZ-Y52) was supplied by Union Carbide. Cr(NO3)3-9HzO, trimethylaluminum (TMA), tri-i-butylaluminum (TIBA), diethylaluminum chloride (DEAC) were all analytical grade chemicals from Aldrich.

and

174 Ethylene (N25) was supplied by Ar Liquido and passed over molecular sieves before used in polymerisation.

2.2 Catalysts preparation and characterisation Cr/Zeolite catalysts have been prepared both by ion exchange and by impregnation. Ionexchange of zeolite was performed by using aqueous solutions of Cr(NO3)3 in the concentration range of 0.8M to 0.01M. The suspension of the zeolite in the Cr aqueous solution was stirred at room temperature for 2 hours. Then the solution was filtered and the solid calcined at 480 ~ for 8 hours. in the case of zeolite samples prepared by incipient wetness impregnation the aqueous solution of Cr(NO3)3, with a metal concentration of 4% w, was added drop by drop to the zeolite with stirring. The suspension was then evaporated to dryness, keeping the stirring to assure uniformity and the solid was calcined at 480 ~ The concentration of Cr incorporated in zeolite was determined by IPC and by a spectrophotometfic method based on the measurement, at 540 nm, of the intensity of the color of Cr (VI) obtained after oxidation of the sample with HC104 [9]. Cyclic voltammetry and ESR spectroscopy techniques have been used to characterise the catalyst, namely to identify the oxidation state of the metal. The ESR instrument used was a Bruker SR 300 X band. Cyclic voltamograms were obtained using a Radiometer (model DEA 101) digital electrochemical analyser controlled by a computer, which was also used to acquire the data. The electrode was preparared by mixing the sample to be inspected with graphite, pressing the mixture into a pellet and mounting the resulting pellet in a special container in an electrolyte solution. Details on the experimental procedure can be found elsewhere [ 10]. 2.3 Polymerisation Procedure and Kinetic Measurements The polymerisation apparatus, polymerisafion procedures and polymer work up have been described in previous papers [ 1I, 12]. Typical kinetic measurements involving the heterogeneous catalysts in a slurry process are made by using an apparatus described in more detail in a previous communication [ 13]. in this apparatus, the inlet of ethylene is made by way of a mass-flow controller (Hastings instruments, ModelHFC-202). This mass-flow controller can be used both to measure the intake of monomer to the systems with reasonable precision, thus allowing us to make precise mass balances, and to control the flow, so that a constant pressure is achieved. This pressure is measured by a pressure transducer (Lucas, P416X-000X-050PV). Both the signals from the flow meter, and from the pressure transducer are acquired by a computer using a 12 bit A/D converter. The control of the mass-flow control valve is also made by the computer using a D/A converter; Labtech Notebook is used to acquire the data and control the valve, in most of the experiments, the control valve was always kept fully open, and was exclusively used for measuring the monomer intake flow. 2.4 Polymer Characterization FTIR spectra were recorded either on a Perkin Elmer spectrophotometer1600 or Bio-RAD FTS-3000MX Excalibur Spectrometer.

175 3. RESULTS AND DISCUSSION

3.1

Catalyst

Characterization

In table 1 we present the composition of the various catalysts that were covered in this study. Table 1. Characteristics of the catalysts used in this study. For each catalyst the method of preparation is indicated as well as the final chromium content of the zeolite, corresponding exchange ratio, and steady state activity*.

Sample ref. Cr/Y-6

% Cr in Zeolite 4.3

Ion exchange degree (%) 72.0

Preparation method Ion Exchange

Activityx 10-5 (~/molcat. M.h) 0.110

Cr/Y-1

3.2

53 9

"

0.146

Cr/Y-5

1.4

23.4

"

0.606

Cr/Y-9

0.6

10.0

"

1.070

Cr/Y- 11

0.3

4.9

"

-

Cr/Y-12

0.2

3.3

"

1.219

Cr/Y-7

4.0

-

Incipient Wetness

9

=

09 o

p

=

*Experimental conditions [Cr] 400 ~tM; A1/Cr= 300; T=25 C; Ethylene 30psi; activity obtained diving the amount of polymer (in grams) by the concentration, in molar, of the catalyst and monomer, and time. The chromium/zeolite catalyst sample prepared by ion exchange with a Cr concentration of 0.6% was studied by ciclic votametry and ESR spectroscopy 9 The ESR spectrum shows a main peak centered at g~ = 1.97 and two other small peaks at g2 = 4.27 and g3 = 5.57 (Figure 1). The main peak corresponds to Cr (111) and as expected (s=3/2) it is splitted into four bands (Figure 1 insert). The peak at 4.27 and 5.57 is likely to correspond to an iron impurity [ 14]. These results show that Cr (nI) species are present in the zeolite in spite of having been calcined at 480 ~ under air, which was supposed to oxidise all the chromium to Cr(VI). When inspected by cyclic voltammetry, the catalyst show a reduction wave (see figure 2), attributable to the reduction of chromium VI to chromium 111, and which is present even for low chromium loads. This reduction wave presents the usual shape for the reduction of a species contained in the electrode itself and can not be attributed to species in solution. The amount of chromium VI that is detected in the voltamogram increases sharply with the increase of the chromium content.

176

4-

4

,

_

!

0

2-2 3300

,

.

3350

3400

.

.

.

3450 Field (G)

3500

3550

3600

.

,~ 0

-

~

r~_l_ -2

-

500

I

I

I

I

I

1

1500

2500

3500

4500

5500

6500

7500

Field (G)

Figure 1. ESR spectrum of the sample Cr/Y-9 displaying the peak assigned to chromium(m).

1,0 /

........

i

__i i i -1500

-1000

-500

0

1

...... 500

1000

1500

E (mV)

Figure 2. Cyclic voltammogram of a CrY sample, at a 20 mV/s scan rate, showing a clear reduction wave which can be attributed to the reduction o chromium(VI) to chromium(m). The combination of the results indicated that the fresh calcined catalyst contains a distribution of chromium in at least two different oxidation states (+6 and +3).

177 3.2 Kinetic Results

As stated above, the polymerisation reactions were carried out at different pressures and temperatures in a stirred semi-batch reactor with continuous in-flow of ethylene; the catalyst was used in suspension in toluene.

3.2.1 Influence of the Chromium Content The behaviour seems quite different for catalysts prepared by ion-exchange or by incipient wetness, the latter showing a much lesser activity than the former. This indicates that, while the chromium in the cages, coordinated to the zeolite structure itself, is relatively active, the chromium present in the zeolite prepared by incipient wetness, and which should contain oxide forms of chromium, is much less active. In the case of the catalysts prepared by ion-exchange, the results, presented in table 1, indicate that the specific activity, in the steady-state, of the chromium cations decreases steadily with increase in load of cation. This indicates that only the first few cations that are introduced are active, and the addition of further cations, beyond a certain threshold, will not improve the activity of the sample. The most promising samples, in terms of total activity, were Cr/Y-9 and Cr/Y- 12.

3.2.2 Influence of the Ethylene Pressure In figure 3 we can see typical runs obtained at different partial pressures of ethylene.

Figure 3. Rate of consumption of ethylene as a function of time. Experimental conditions: P(C2H4) = 10.8 PSI (O), 20.8 PSI (O), 30.5 PSI (ll), T = 25 ~ Symbols are experimental data and the line corresponds to the kinetic model (see text below). The consumption of ethylene clearly shows an induction period, which may be associated with the formation of the first reactant-chromium complexes, followed by a deactivation process, which leads to an almost stationary rate of comsumption of ethylene. It is clear from

178 the results that are presented that the deactivation rate seems to decrease with the increase of the ethylene pressure and that for the lowest pressure, the deactivation rate, which is very fast at the beginning, decreases for intermediate times. Thus, the increase in partial pressure not only increases the rate of the reaction but also decreases the rate of deactivation.

3.2.3 Influence of the Temperature Figure 4 represents the rate of consumption of ethylene for different temperatures. As it can be seen, the global activity increases with temperature. The rate of deactivation also increases with temperature.

Figure 4. Rate of consumption of ethylene as a function of time for different temperatures. Experimental conditions: P(C2H4) = 10.65 PSI, T = 25 ~ (O), = 55 ~ ( I ) . Symbols are experimental data and the line corresponds to the kinetic model.

3.2.4 Kinetic Model The decrease of the deactivation rate as the pressure increases seems to indicate that the deactivation is not due solely to internal mass transfer limitations, which could be caused by the obstruction of the zeolites pores by the growing polymer chains themselves. Besides, the observed behaviour is also very similar to those observed for homogeneous catalysts, and, thus, we are led to believe that the deactivation is due to the deactivation of the active sites themselves. A simple coordination-reaction mechanism that includes a deactivation step was fitted to the experimental results with success (see figure 5). This mechanism is an extension of the mechanism proposed to explain the behaviour observed for homogeneous Kaminsky-type catalyst [ 15].

179

kid

( Insertion

Cr* k,

~ M

kc,~~Cr*M Insertion ko~~,~

k2

k,dCrd

Cr* - active species

_ _ ~ CrdM M

Crd- active species k~

Cr*M2.,

~ CrdM2

k-3d Figure 5. Proposed scheme for the polymerization of ethylene over CrNaY zeolites. The mechanism assumes that each site is capable of binding more than one ethylene molecule and that the rate of insertion of ethylene into the growing polymer chain is dependent on the number of ethylene molecules. This mechanism is supported by the fact that second order dependencies on ethylene partial pressure are often found in homogeneous catalysts, thus indicating that two monomer molecules are involved in the mecvhanism until the rate-determining step. Application of this kinetic scheme to the experimental results gives very good fittings (see figure 3 and 4), both for the partial pressure variation and for the temperature variation, although activation energies could not be computed due to the limited amount of data still available as a function of temperature.

Table 2. Kinetic rate constants obtained by fitting the model described to experimental data.

kc2

T (~ 25

kl 0.77

k-1 0.03

kcl 0.08

k2 28

k-2 0.17

6.7

55

1.85

0.05

0.16

33

0.17

15.0

T (~ 25

kld 2.1 x 10.3

k-ld 1.1 x 10-4

k2d 4.0x 103

k-zd 8.9x 10.3

k3d

k-3d

7.4x 10-3

5.2x 10-3

55

5.7x10 "3

1.1xl0 -4

7.9x10 -3

1.5x10 -z

1.5x10 -2

4.0x10 -3

In the case of the catalysts considered in this paper, the assumption that more than one molecule of ethylene can be coordinated to the chromium active site is also relevant in relation to the deactivation. As stated above, when the partial pressure is increased, not only does the global activity increase, but also the deactivation rate decreases. This increase in stability can be interpreted, in terms of the scheme shown in figure 5, as a shift towards the species with more ethylene coordinated, which may be less sensitive to deactivation.

4. CONCLUSIONS Chromium-exchanged Y zeolite seems to be a reasonablly active catalyst for the polymerization of ethylene, although its performance is highly influenced by the load level.

180 The scheme proposed seems to be able to describe the general behaviour of the catalyst, although there is still no indication on the mechanism for deactivation. It is likely that it will be linked with changes in the redox state of the cations. This is consistent with the fact that both Cr(ffl) and Cr(VI) were observed in the catalyst. The reducibility of the cation under operating conditions is likely to be different for the species with no ethylene coordinated and for mono and di-coordinated ones. The increase in stability with the increase in partial pressure indicates that the species with more ethylene coordinated is probably the most resistent to oxidation. Further studies will be required to clarify the nature of the deactivation process.

REFERENCES

1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13.

14. 15.

M.R. Ribeiro, A. Deffieux, M.F. Portela, Ind. Eng. Chem., 36 (1997) 1224. M.P. McDaniel, Ind. Eng. Chem. Res., 27 (1988) 1559. M.P. McDaniel, J. Catal., 67 (1981) 71. M.P. McDaniel, J. Catal., 76 (1982) 17. M.P. McDaniel, M.M. Johnson, J. Catal., 101 (1986) 446. B. Rebenstorf, Acta Chem. Scand., 43 (1989) 413. J.H. Lunsford, S-L. Fu, D.L. Myers, J. Catal., 111 (1988) 231. C.S. Kim, S. I. Woo, J. Mol. Catal., (1992) 249. Vogel's Textbook of Quantitative Chemical Analysis., fifth edition, chap. 17-24, 1989. M.A.N.D.A. Lemos, P. Sousa, F. Lemos, A.J.L. Pombeiro, F. Ram6a Ribeiro, Stud. Surf. Sci. Catal., 122 (1999)443. J. C. W. Chien, B. P. Wang, J. Poym. Sci. Part A: Polym. Chem. 20 (1988) 3089. Z. Yu, M. M. Marques, M. D. Rausch, J. C. W. Chien, J. Poym. Sci. Part A: Polym. Chem., 33 (1995) 979. S. Correia, F. Lemos, M.M. Marques, F. Ram6a Ribeiro, A.R. Dias, "Measuring Polymerization Kinetics Using a PD Pressure Control Loop" 7th Intemational Chemical Engineering Conference CHEMPRO'98, September 26-28, Lisbon, Portugal, 1998. J. Costa; R. Delgado; M. G.B. Drew; V. Felix; R.T. Henriques; J.C. Warenborgh, J. Chem., Dalton Trans., (1999) 3253. M.M.Marques, C. Costa, F. Lemos, F.R. Ribeiro, A.R. Dias, Polymer Int. 43 (1997) 86.


181

Investigations of the selective partial oxidation of methanol and the oxidative coupling of methane over copper catalysts H.-J. WSIk, G. Weinberg, G. Mestl and R. Schl(~gl aFritz-Haber-lnstitut der Max-Planck-Gesellschaft, Abt. Anorganische Chemie, Faradayweg 4-6, D-14195 Berlin

ABSTRACT Copper-based catalysts are of considerable importance for industrial reactions, e. g. partial oxidation reactions. This contribution reports on a broad study of the catalytic activity of copper in model redox reactions, e. g. methanol oxidation and oxidative coupling of methane. In addition the interaction of Cu with these reactive gases was investigated by thermoanalytic techniques (TG/DTA, DSC), temperature programmed oxidation and reduction (TPO/TPR) and thermal desorption spectroscopy (TDS). Scanning electron microscopy (SEM) and electron backscattering diffraction (EBSD) was additionally used to characterise the copper catalyst before and after catalytic action.

1. INTRODUCTION The oxidation of glycol and the sythesis of methanol are the most important industrial reactions which uses copper metal or copper based catalysts. In case of the methanol synthesis catalysts, small copper particles supported on a zinc oxidealumina system are used [1]. In general the catalyst for the glycol oxidation are copper metal ring. For structure sensitive reactions, generally the catalytic activity depends on the particle size and shape of the catalyst particles. Most of the investigations on copper for partial oxidation reactions described in the literature were performed on foils, single crystals or powders. We decided to use spherical copper particles with diameters between 20 pm to 250 IJm to possibly derive a correlation between copper nanoparticles and polycrystalline foils. Important morphological effects like nucleation [2] forced by the catalytic reaction can also be better observed and compared to macroscopic changes known from copper metallurgy [3]. Nucleation is for example one possible explanation for the catalytic activity of the ammonia catalyst [4]. Not very much is known about the system copper-oxygen in the temperature range between room temperature and 900~ especially, in relation to catalytic redox reactions. The phase diagram as reported in the literature is based on measurements in 1929 [5]. This gap in fundamental knowledge is also a motivation for this work. The methanol synthesis and the glycoloxidation [6] are reactions which make a more expendable equipment necessary. It was decided to use the methanol oxidation and the oxidative coupling of methane as test reactions for the low and high temperature regime, respectively Ion-molecule-reaction mass spectrometric gas analysis (IMR-MS) and in-situ XRD measurements have shown that during the partial oxidation of methanol over copper oscillatory reactions occur with different frequencies in the temperature

182 interval between 3 5 0 ~ 450~ [8]. The interest was in the explanation of the relation of the gas-surface interaction, the occurance of facetting and the simultaneous changes of the crystal structure during the catalytic reaction. Temperature, partial pressure, space velocity, diffusion, oxygen species and surface reconstruction and defects are the main parameters which have to be investigated by in sffu and ex situ techniques. The influence of many of these parameters for partial oxidation reactions over silver catalysts could be unravelled [7]. In this work it is attempted to derive a comparison between the morphological and surface changes of copper and silver induced by partial oxidation reactions. 1. RESULTS AND DISCUSSION

In 9 the reported study, Cu particles were used with diameters between 20pm and 250pm to investigate the relation between particle size and these oscillations. Cu particles of different diameter render possible the investigation of the influence of bulk structure on the catalytic performance, as compared to industrially used Cu rings. TG/DTA-experiments in the temperature interval between 25~ and 600~ were carried out to compare the redox behaviour of Cu with CuO and Cu20 in air and under reductive conditions (5% H2 in N2), respectively, in an attempt to interpret the catalytic results of methanol oxidation. At the moment, SEM is an ideally suited technique to determine this alteration of the particles in addition to conventional methods like TPR and TG/DTA. It is possible to conserve the surface of the catalyst by quenching in helium to room temperature. Unfortunately, it is not possible to determine the degree of oxidation of the copper samples by EDX due to its low oxygen sensitivity. The heating rates in the TG/DTA experiments were varied from 2 to 20K/min in order to determine the kinetics of the redox reactions. Maximum temperature was set to 600~ to avoid sintering.

100

c o ~

60

> r

-

o o I T

40

O

20

O

CH30~02

100

Fig." Methanoloxidation

o

200

300

temperature [~

400

183

In addition, spherical copper particles allow the detailed determination of the kinetics of the reduction process with respect to the nucleation and contracting sphere models of the reduction of transition metals as described in literature [9]. It should be possible to distinguish between both models by the interpretation of the signal profile of TPR experiments in combination with TPO experiments. It is possible to adapt several model dependent or model independent kinetic analyses to the TPR/TPO profiles to obtain informations about the activation energies and the mechanisms of reduction or oxidation. The kinetic analysis that is necessary to get a suggestive result should include the three diminsional aspect of the process together with diffusion controlled effects [10]. The alteration of the particle surface resulting from the copper/oxygen redox reaction is the main process for the observed increasing activity of the catalyst with time on-stream. This change depends on the oxygen partial pressure, the reaction temperature and the particle size. For instance, small particles exhibit catalytic activity at lower temperatures than larger particles under same conditions. In addition to the reduction of the particles, sinter effects are caused by the exothermicity of the catalytic reaction. The shrinkage of the particles possibly follows the same rules as defined for the behaviour of metal particles under high pressure conditions [11]. This sintering perturbes the nucleation and hence the resolution according to the contracting sphere model, and necessitates a more detailed definition of the reconstruction process of sperical particles in general. SEM-analysis prior and subsequent to the experiments showed the preservation of grain boundaries within the copper particles. In situ XRDmeasurements proved the co-existence of Cu/Cu20-phases between 280~ and 450~ The oxidation of Cu20 to CuO is an exothermal process between 250 to 445~ DSC measurements showed three exothermal signals at 254~ 428~ and 444~ The last one correlated with the main weight gain as detected by TG/DTA. The oxidation of the copper surface already starts at 25~ and is terminated in case of small Cu particles (1001Jm 80

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Fig. 4. OCM-reaction

600

700

temperature [~

800

186 The selectivity to ethane is controlled by the oxygen partial pressure and thus the copper oxygen interaction. Experiments under steady state conditions at different temperatures shows that the deactivation of the catalyst corresponds to the formation of carbon and the destruction of the facettes on the catalyst surface. After a reaction time of about 6h the defined character or the facetted surface is exchanged by an undefined spongelike appearance. Both, the formation of carbon deposits and the change of the surface character of the facettes seems to be resposible for the deactivation of t.he copper catalyst during the reaction. H2 + RHn. 1

CO2 + H20

RH n

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Fig.5: Ag-model For silver catalysts it was possible to determine by photoelectron spectroscopy (XPS/UPS) and thermal desorption spectroscopy (TDS) that different oxygen species are situated at the silver surface or are dissolved in the bulk, respectively, under partial oxidation reaction conditions. O~ is dissociatively-adsorbed, atomic surface oxygen. The second species,Oi}, is formed when O~ diffuses into the bulk. The third species, Oy, is formed at higher temperatures via the subsequent segregation of O~ to the (111) surface. It is suggested that O~ induces the catalytic activity. This mechanism is indicated by the formation of different surface structures. The formation of Ag(111) surfaces during the reaction leeds to an increase of the catalytic activity In the attempt to extend the Ag model to Cu EBSD was also used subsequently, to obtain informations about the main orientations of the copper surfaces [13]. As mentioned above, all samples showed a well defined facetted surface after the reaction. The copper (111) surface orientation was preferred. Moreover O2-TDS experiments showed the evolution of two high temperature Ospecies between 600 and 900~ depending on the heating rate and treatment. SEM of the copper particles after this treatment under UHV conditions showed comparable morphological changes, like sintering and facetting, as obtained by catalytic

187

reactions.. This result leeds to the assumption that the Ag-model can be transfered to copper in selective partial oxidation reactions. 3. CONCLUSION

In the present study, the influence of the catalyst particle size on the methanol oxidation and the oxidative coupling of methane over copper has been investigated between 350~ and 900~ TPRfTPO and TG/DTA was used for the determination of the redox behaviour and SEM for the morphological characterization at the different stages of the reaction. During the methanol oxidation, oscillations due to the exothermicity of the reaction were observed in DTA signals which allowed a relation to earlier IMR-MS results. A low oxygen partial pressure during the OCM reaction leads to the conversion of methane to ethane and ethene. Both reactions result in facetted surfaces. The combined results are a first evidence for a high structural sensitivity of these reactions over copper surfaces comparable to that observed for silver catalysts. Three different oxygen species intercalated in Cu(111) seem to be responsible for the catalytic activity as suggested by O2-TDS. The main goal of further investigations is the extension of the validity of this reaction model to copper catalysts. REFERENCES

[1]M. Bowker, R.A. Hadden, H. Houghton, J.K.L. Hayland, K.C. Waugh, J. Catal., 109 (1988) 263. [2] A. Pattek-Janczyk, Appl. Catal. A, 124 (1995) 267 [3] E. Arzt, Acta mater., 46 (1998) 5611. [4] R. Schl0gl, in J.R. Jennings (Ed.), Catalytic Ammonia Synthesis: Fundamentals and practice, Plenum, New York, 1991, Ch. 2, p. 38. [5] Gmelin Institut. Kupfer. Gmelins Handbuch der Anorganischen Chemie. VCH, 1958. [6] L.A. Arkatova, O.V. Vodyankina, Russ. J. Appl. Catal., 72 (1999) 1081. [7] A. Nagy, G. Mestl, Th. R0hle, G. Weinberg, R. SchlOgl, J. Catal., 179 (1998) 548. [8] H. Werner, D. Herein, G. Schulz, U. Wild, R. Schl5gl, Catal. Lett., 49 (1997) 109; M.H~vecker, A. Knop-Gericke, Th. SchedeI-Niedrig, R. SchlOgl, Angew. Chem. 110 (1998) 2049. [9] N.W. Hurst, S.J. Gentry and A. Jones, Catal. Rev.-Sci. Eng., 24 (1982) 233. [10] S. Vyazovkin, C.A. Wight, Annu. Rev. Phys. Chem., 48 (1997) 125. [11] T.E.M. Staab, R. Krause-Rehberg, B. Vetter and B. Vieback, J. Phys.: Condens. Matter 11 (1999) 1757. [12] A.K. Datye, Topics Catal., 13 (2000) 131. [13] V. Randle, J. Microscopy, 195 (1999) 226.



189

A new kinetic model for hydrodesurfurization of oil products. T. Mejdell a, R. Myrstad a, J.Morud a, J.S. Rosvoll c, P. Steinerb and E. A. Blekkan b aSINTEF Applied Chemistry, N-7465 Trondheim, Norway bNorwegian University of Science and Technology, Department of Chemical Engineering, N-7491, Trondheim, Norway CSTATOIL Research Center, N-7005 Trondheim, Norway

A new kinetic model for hydrodesulfurization (HDS) of oil products is proposed. The model is based on a discretization of the entire sulfur GC-AED spectrum into small pseudocomponents of only 1~ boiling point range. In addition, a few components with particularly low reactivity, like 4-Me-DBT and 4,6-Me-DBT, are identified and modeled separately. Experiments on a lab reactor are used to identify the kinetic parameters for an LGO feed. A good fit was obtained, and the model is now used to predict the conversion in an industrial trickle bed reactor.

1. INTRODUCTION The hydrotreating of middle distillates is a class of processes with ever increasing importance. Recently there has been a renewed interest in these processes due to the more stringent regulations on diesel fuel quality introduced both in Europe and the USA. The limit on sulfur contents is one of several quality factors being addressed in the new regulations. A sulfur limit of 50 wppm S implies that a typical gas-oil containing 0.5 wt% S must be hydrotreated with a sulfur conversion of 99% to fulfil the requirement, a very difficult task requiring very good control of reaction kinetics and reactor design and operation. Numerical modelling is a very important tool in this respect, and these new requirements require better models to be able to predict reactor behaviour with better precision at these very high conversions. It has been common to treat the sulfur components in the feed as one or at most a few lumps with similar kinetic behaviour. This approach has been adequate in the past, but to be able to understand and predict the behaviour at conversions close to complete a different approach is needed. Hydrotreating of middle distillates is a complex system, numerous parallel and consecutive reactions involving sulfur removal, saturation of double bonds and aromatics, as well as nitrogen removal and some hydrocracking. The sulfur is present as a range of organic compounds, with reactivities ranging from quite high to extremely low [ 1]. It has been shown [ 1,2] that the substituted dibenzothiophenes like 4-methyl dibenzothiophene (4-Me-DBT) and 4,6-dimethyl-dibenzothiophene (4,6-Me-DBT) constitute the remaining sulfur compounds in deeply desulfurized samples. The modelling of complex systems like this can be approached using several methods. Froment and co-workers [3,4] use a molecular approach, relating the activity of substituted compounds to the activity of the non-substituted ring-structure. This gives good results, but the model is complex and involves extended analytical work to identify the components. Alternatively, one could use continuum theory of lumping as discussed by e.g. Chou and Ho [5], and applied by Sau et al. [6] for HDS kinetics. In this approach it is assumed that the

190

reactivity is a function of some other parameter, e.g. true boiling point (TBP). Sau et al- [6] assumed that the reactivity decreases exponentially with TBP for a certain class of sulfur compounds like thiophenes, benzothiophenes and dibenzothiophenes. This exponential function will typically have 3 parameters that must be fitted to the experimental data. The activation energy is modeled with a similar function of TBP or assumed constant. In this work, we propose a new kinetic model for HDS. The model is based on using GCAED (Gas Chromatography - Atomic Emission Detection) to characterise the sulfur components in the oil fraction. The spectrum is divided into many small parts corresponding to approx. 1 ~ TBP, each one considered as a pseudo-component. In addition, a few components with particularly low reactivity, like 4-Me-DBT and 4,6-Me-DBT, are identified and modeled separately. These components are important when the model must be accurate at high conversions. By combining a continuous-like approach with treating a small number of refractive components separately, we get a model that is both precise and requires limited analytical work to provide data input.

2. E X P E R I M E N T A L

2.1. Experimental procedure The up-flow unit used in collecting the experimental data is shown schematically in Figure 1. This unit has earlier been demonstrated to be well suited for catalyst testing, not only in catalyst ranking, but also in terms of production planning and product quality [7]. The gas feed rates were controlled by mass flow controllers and the liquid feed was fed into the reactor by a high-pressure pump from a reservoir connected to a balance. Downstream of the reactor the products were quenched and depressurized and the liquids and gases were separated in a low-pressure separator. Total gas flow was analyzed on a wet gas meter and the gas-composition was analyzed online by a GC.

ENCH

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The reactor was a stainless steel tube (1 m x 19.3 mm i.d.) with a total volume of 260 cm 3. The first part of the reactor inlet zone was filled with SiC in order to preheat the feed and

191

minimize temperature gradients in the catalyst zone. The reactor was placed vertically in a bronze block and heated in an electric furnace with two separate PID-controlled heating zones. The reactor was equipped with an axial thermowell containing 4 thermocouples, 3 of which were in the catalyst bed. The catalyst used in this study was a commercial CoMo/AI203 used as received from the suppliers. The liquid feed used was light gas oils (LGO) drawn from a commercial unit. The liquid products are sampled and stripped with N2 to remove dissolved H2S and NH3. For further details about the experimental procedure, see [7].

2.2. Experimental conditions 28 experimental runs were accomplished with the following range of experimental conditions: - Temperatures 250- 400 ~ - Pressures 2 0 - 80 barg, - L H S V = I - 8 h -1 - H2:oil-ratio 200 Nm3/m 3 (the same for all experimental runs) 2.3. Analysis The sulfur components in the feed and the products were identified using GC-AED analysis. The six special components were first identified using standard integration routines. Then the signal was divided into 132 parts corresponding to a TBP of approx. 1 ~ ensuring that all peaks in the spectrum corresponded to only one pseudo-component. The area of each pseudo-component was calculated, and the area of the six real components was subtracted from the corresponding pseudo components. The total sulfur content was also analyzed using an ANTEK 7000 analyzer, and this value was used to normalize the individual sulfur component concentrations. 3. M O D E L I N G

3.1. Kinetic model Each of the 138 components, 6 real components, and 132 pseudo-components, is modeled by the following reaction kinetic model:

r,

(l+Knzs.Cn2s)2

,

k,=k,.s73.e

(1)

This expression takes into account both hydrogen pressure variation and the inhibiting effect of H2S. For each component i there are 2 parameters to fit, ki,573, the rate constant at 300 ~ (573 K), and the activation energy, Ei. In addition the parameter KHZS must be found. The total kinetic model then consists of 2* 138+ 1=277 parameters.

3.2 Lab reactor model The up-flow lab reactor is modeled as a plugflow (integral) reactor. It is assumed an equilibrium between the gas and the liquid in the reactor. The temperature gradient in the reactor is small, never larger than 4-5 ~ so isothermal conditions are assumed. A weighted average of the temperature profile is used as the isothermal temperature.

192 For all other components than H2 and H2S the Henry laws constant Hi is assumed to be zero. The constants for these two components have been calculated as a function of temperature from HYSYS (a commercial process simulation package) simulations. The simulated oil had the same boiling point distribution as the actual gas oil. From these simulations we also got the oil vapor pressure. The reactor equation for all compounds is then: dC~

riPcat

dx

vL + vc zRr

(2)

Hi

Here vL and vc are the space velocity of the liquid and the gas respectively. Estimates of the compressibility factor z gave values very close to 1.0, i.e. ideal gas law could be used. 3.3. Parameter fitting procedure The kinetic expression (1) shows that each of the 138 sulfur components may be treated separately except for the H2S concentration and the parameter KH2s. We therefore used a recursive procedure to fit the data: 1. Fit the data to a one-lump version of equation (1), and use this model to estimate the H2S concentration profiles in the lab reactor. 2. For i= 1 to 138 calculate the ki and Ei by using the estimated sulfur profiles and Kms. 3. Refit the parameter Kri2s by using the rigorous model with all the parameters found in 2. Estimate new H2S concentration profiles. 4. Go to 2 and re-estimate the parameters until convergence.

4. RESULTS In Figure 1 and 2 the selected individual components are the first 6, the remaining lumped components are put in increasing TBP order. Figure 1 shows that the reactivity generally decreases with increased TBP, and that the reactivity of the individual 6 components is much lower than the rest. The activation energy also decreases with TBP as shown in Figure 2. x 10"s t

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Figure 2. The activation energy constant for the 138 components.

9

193 The reason why the parameters are quite scattered in both ends of the TBP range is that the concentrations here are very small. On the other hand, the contributions of these components are also small for the predictions of the total sulfur. For illustration, Figure 3 shows the fit for one single component. The 28 experiments are sorted in increased sulfur content order. The actual component is component no 30, and has a TBP between 269-270~ In Figure 4 the prediction of the total sulfur is shown. It is based on the fitted parameters of all the components and show that the total prediction is very good.

2.5

Sulfur comp no 30

x !o ~

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/

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- --

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Figure 4. The fit of the total sulfur content based on all the 138 components

5. DISCUSSION Figure 1 and 2 pinpoint many of the advantages and disadvantages with the continuum theory approach. The data in Figure 1 confirm that most of the sulfur components have a decreasing reaction rate with TBP, and an exponential function to describe this decrease seams not unreasonable. However, at the heavy end of the spectrum, the activity seems to increase slightly again, and all in all, it is not obvious that an exponential function will represent the reactivity-TBP relationship best. Also, the activation energy decreases with increased reaction rate, but here it looks more like a linear decrease than an exponential one. In any case, the data shows that one cannot assume that the activation energy is constant. Most important, however, is that some of the components have large deviations from the general TBP-reactivity tendency. The six first components in Figure 1 show this clearly. For high conversion kinetics these components are very important, and consequently, the continuum theory approach is not very applicable. Consequently, it is better to treat these six components separately. For the rest of the spectrum we could have used the continuous function to filter out randomness in the parameter values. However, we think that most of the variation in reactivity and activation energy is not random, but reflects real differences in chemical reactivity. One exception is at

194 the end of the spectrum where the signal to noise ratio is very small. Here it might be advantageous to do some smoothing of the values. However the contribution to the total sulfur content is minimal so in most cases the difference would be negligible. With this new model we think we have a good kinetic model for desulfurization reactivity for other LGO feeds as well. A GC-AED analysis of the new feed is all that is needed. An assumption for this claim is of course that the reactivity for the 132 lumps is the same for different oils. The kinetic model has recently been implemented in a model for an industrial trickle bed reactor with the same type of catalyst pellet. So far the model seems to describe the conversion very well.

6. CONCLUSIONS -

A new approach for desulfurization kinetics is proposed which is based on a discretization of the GC-AED spectrum. For LGO it was discretisized into 132 pseudo components 6 components with very low activity are treated separately. The kinetics of these components are important at high conversions The kinetic parameters were identified and the whole model was tested against data from a commecial reactor unit.

Acknowledgment This work has been supported by Borealis, Dyno, Statoil, Hydro, Reichhold, and the Research Council of Norway, and is a part of the research project "Reactor Technology in Petrochemistry and Polymer Industry" (REPP).

REFERENCES 1. H.TopsCe, B.S. Clausen and F.E. Massoth, "Hydrotreating Catalysis", Springer-Verlag, Berlin 1996, p. 114-115. 2. P. Steiner and E.A. Blekkan, "Catalytic hydrodesulfurisation of light gas oil: Detailed product analysis and kinetics", CHISA 2000, Prague, 27-31 August 2000. 3. G.F. Froment, G.A. Depauw, and V. Vanrysselberghe, Ind. Eng. Chem. Res 33 (1994) 2975. 4. V. Vanrysselberghe and G.F. Froment, Ind. Eng. Chem. Res 37 (1998) 4231. 5. M.Y. Chou and T.C. Ho, AIChE Journal 34 (1988) 1519. 6. M. Sau, C. S. L. Narasimhan, and R. P. Verma, "A Kinetic Model for Hydrodesulfurization", Stud. Surf. Sci. Catal. (1997) 421. 7. R. Myrstad, J.S. Rosvoll, K. Grande, E.A. Blekkan, Stud. Surf. Sci. Catal., 106 (1997), 437


195

Low-temperature, carbon-catalyzed, solvent-washed, trickle-bed sulfuric acid process Peter L. Silveston b, Mahiyar A. Panthaky a, Kirk Duval b, Radu V. Vladea b, Ali Lohi c, and Robert R. Hudgins b a Present Address: Brantford Chemicals Inc., Station Main, P. O. Box 1976, Brantford, Ontario, Canada, N3T 5W5 bDepartment of Chemical Engineering, University of Waterloo, Waterloo, Ontario, Canada N2L 3G1 c Dept. of Chemical Engineering, Ryerson Polytechnic Institute Toronto, Ontario, Canada

Abstract Production of sulfuric acid as a concentrated acid or at any desired dilution can be carried out at ambient temperature and pressure for a humidified SO2 and 02 stream using a trickle-bed reactor flushed with a solvent such as acetone or ethyl ether. Activity of the carbon catalyst in a methyl ethyl ketone washed trickle bed is about 100 times the activity of conventional vanadia catalysts operating at 350 to 400~ Measurements of removal. conversion to acid and productivity of the catalyst in g of H2SO4/g carbonoh were done using moderately polar organic solvents acetone, methyl ethyl ketone (MEK), and methyl isobutyl ketone (MIBK) m as well as water as the flushing liquids for the oxidation of SO2 to sulfuric acid in a trickle-bed reactor containing a structured packing coated with Centaur T M activated carbon. SO2 removal from the gas sent to the trickle bed reached 95% and conversion to acid based on SO2 fed attained 94% under some operating conditions. Productivity was measured at 2.9 g of H2SOa/g carbon~ with 6 vol% SO2 in the gas feed. Pressure was ambient and temperatures ranged from 18 to 21~ in all experiments. MEK and MIBK undergo reaction as batch distillation proceeds, but otherwise 98% solvent recovery was achieved. Key words: SO2 oxidation, sulfuric acid production, trickle-bed reactor, activated carbon catalyst b To whom correspondence should be addressed.

196 1. INTRODUCTION The well established activity of activated carbon as a low temperature catalyst for the oxidation of SO2 to SO3 coupled with our recent discovery of the solubility of sulfuric acid in various organic solvents appears to open up prospects for a novel sulfuric acid process suited to small installations and tunable to produce acid of any concentration for local use. Oxidation activity of activated carbon for SO2 at remarkably low temperatures has been known for almost fifty years (Davtyan and Ovchinnikova, 1955). Carbon, however, is rapidly poisoned.by the SO3 product so its high activity is difficult to exploit. Hartman and Coughlin (1972) produced sulfuric acid by carrying out the reaction in a trickle bed packed with carbon under continuous co-current washing of the bed with water. Continuous flushing, however, yields a dilute acid. Researchers at the University of Waterloo have been investigating the application of trickle beds to stack-gas clean-up since 1989. Acid concentrations as high as about 2N can be reached by intermittently flushing the trickle bed with dilute sulfuric acid while achieving better than 90% removal of SO2 (Haure et al., 1989; Metzinger et al., 1994). In subsequent studies, Lee et al. (1995) demonstrated that a recently developed activated carbon, Centaur TM, had an activity up to tenfold greater than those of carbons previously studied. The Centaur carbon has been chemically bonded to Sulzer wire mesh static mixers and to "jelly-roll" wire screens (Suppiah et al., 1987) to form a low density trickle-bed packing. Vladea et al. (1997) observed that these packings reduce the pressure drop in a trickle bed by a factor of about 5 while maintaining the same degree of SO2 removal. The objective of our program was to study the use of moderately polar organic solvents as flushing agents for SO2 oxidation in a trickle bed as the first step towards the development of a new, low-temperature acid process. We were particularly interested in the effect of trickle-bed operating variables on reactor performance and the productivity of the catalyst. Some preliminary experiments on separating sulfuric acid from the flushing agent are reported. 2. SOLVENT-BASED SULFURIC ACID PROCESS Figure 1 shows a schematic of our proposed solvent-based process, assuming sulfur dioxide (SO2) comes from the burning of sulfur in either air or oxygen. After cooling the SO2 containing stream is humidified and chilled to 20 to 30~ with cooling water. The temperature and the relative humidity set the concentration of the acid formed in the trickle bed. Acid is formed in the trickle bed operated in down-flow. A chilled solvent is introduced to the top of this bed and flows co-currently with the SO2 and Oz gas stream. Two options are possible: continuous co-current flow or alternating flow of gas and liquid. The attraction of alternating flows is lower solvent emissions and greater catalyst productivity. The off-gas from the reactor passes into a condenser cooled by chilled brine and then into an adsorbent bed for solvent capture and recovery. The solvent phase descending through the trickle bed absorbs the sulfuric acid formed on the activated carbon surface, leaves the bed, and then flows to a heater and on to a packed bed distillation column. Overhead from the distillation is condensed, a portion returned to the column as recycle, and the remainder sent to solvent hold-up tankage. Further process details are contained in Canadian, European and American patent applications (Vladea et al., 1998).

197

CARBON TRAP

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Product

SulfudcAcid REBOILER PRODUCT COOLER Figure 1 - Flow diagram of a trickle-bed process for sulfuric acid production using activated carbon as catalyst and flushing with an organic solvent.

3. SOLVENT CONSIDERATIONS Acetone (AC) and methyl ethyl ketone (MEK) were the main organic flushing agents employed in this study. A few runs were made also with methyl isobutyl ketone (MIBK). Ethyl ether was not used because of its high vapor pressure at 25~ Since experiments were undertaken near room temperature, the other solvents could be used without large vaporization losses. Another important consideration was the solubility of SO2 and 02 in the solvents. The solubility of 02 in AC at 20~ is almost tenfold greater than in water. SO2 is about six times more soluble in AC than in water at 20~ Unfortunately, solubility data for SO2 and 02 in ethyl ether and MEK were not available. The solubility of SO3 in AC and MEK is very low (Won et al., 2000). 4. EXPERIMENTAL A schematic of our trickle-bed system has been published previously (Lee et al., 1995). The trickle-bed was a 500-mm glass tube of 45-mm I.D. A glass cap attached to the top of the tube had ports for the inlet gas mixture and the flushing liquid.

198 Flushing liquid

~~

Fluegases

O-Ring

__~~~__

IIF

"~'Activatedcarb. . . .

analyzer

H2SO4 solution

Figure 2 - Schematic of the co-current, downflow trickle bed packed with carbon coated Sulzer Static Mixer Packing: T = thermocouple location, P = pressure tap Instead of the packed-bed used by Lee et al. (1995), a high-efficiency structured packing was employed. Static mixers (type CV) supplied by Sulzer Chemtech were coated with Centaur carbon by Atomic Energy of Canada Limited, using their proprietary binder technology. A Teflon TM binder used as a coating rendered the surface hydrophobic. The coating method is described by Suppiah et al. (1987) who observed that a Teflon-to-carbon ratio of 0.5:1 reduced the original surface area by about one-half. Our measurements on the coated Sulzer packing showed that the surface area was 140 m2/g. Mesopores contributed 68% of the surface area while the micropores made up the remaining 32% (Vladea et al., 1997). In most experiments, three stacks of the catalyst were used, equivalent to a bed depth of 45 cm. The carbon catalyst showed no signs of deactivation over a period of two years during which numerous runs were performed. Feed gases containing various amounts of SO2 and O2 were used in the experiments. Air and nitrogen were saturated with water vapour before mixing with SO2. For an operating temperature of around 20 to 23~ the saturated concentration of water vapour in the feed was calculated to be 2.9 vol%. Feed gas mixture was introduced into the trickle-bed reactor co-currently with the liquid flow. An overhead tank stored the flushing agent, which was distributed over the bed packing by a spray nozzle. Liquid flow rate was regulated by a gear-type micropump. Temperatures in the trickle-bed at two different points near the outer wall of the catalyst stack were measured by thermocouples. Pressures were measured by manometers just above and below the bed. Gases and liquid leaving the TBR were separated just below the packing. The liquid was collected continuously to facilitate sampling. The gaseous effluent went through a series of

199 ice-traps to remove any moisture and acid carryover. 802 concentration in the gaseous effluent stream was continuously monitored in ppm. Steady state was quickly achieved. Nevertheless, most of the experiments were run for about two hours. SO2 concentration in the outlet gas stream was continuously monitored and used to determine the steady-state condition. Once this was reached, 25-mL liquid samples were taken to determine the amount of SO2 converted into sulfuric acid and the SO2 dissolved in the effluent solution using conventional titration methods. 5. RESULTS 5.1. Trickle

Bed Studies

Performance of the trickle-bed reactor was gauged by SO2 removal and its conversion to sulfuric acid. SO2 removal is the sum of the SO2 converted to H2SO4 and SO2 dissolved in the wash liquor; hence, it is always greater than the conversion. The productivity of the catalyst, defined as the mass in g H2SO4 produced/g carbon.h, measures the size of the reactor and thus indirectly the pressure drop. It is used to compare

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Figure 3 - Trickle-bed performance with acetone as the flushing fluid for 0.3 vol.% 502 at SLV = 0.953 mm/s, O2 = 5 vol%, H20 = 2.9 vol% and T = 18~ measurements made under different experimental conditions. It also provides a means of comparing the catalyst used in this study with other catalysts used for SO2 oxidation. Experimental measurements are summarized in Figures 3 to 5. With the ca. 45-cm long carbon bed, recovery of SO2 from the gas reached 95%. Conversion to sulfuric acid attained 90% at high flushing rates, but generally fell in between 60 and 80%. Activity of the carbon catalyst with acetone (AC) flushing, as measured by productivity, is about twice as high as flushing with water (Table 1). The effect of superficial gas velocity (SGV) on the performance at an inlet SO2 concentrations of 0.3 is shown in Figure 3 using AC as the flushing solvent. The average superficial liquid velocity (SLV) was 0.953 mm/s and the experiments were performed at a bed temperature of 18~ Oxygen concentration was 5 vol% and water was 2.9 vol%. As expected from contact-time considerations, recovery of SO2 from the gas and its conversion to acid decreases with increasing SGV. The removal curve extrapolates smoothly to approximately 100% as the gas velocity goes to zero. Since conversion is measured on the

200 basis of S O 2 fed to the trickle bed, identical slopes of the removal and conversion curves indicates that the conversion of SO2 recovered from the gas is independent of SGV. Productivity increases with SGV, again as expected. Increasing SGV means more SO2 passes across the carbon catalyst per unit time. The SVG experiments were repeated at 1% SO2, but at an SLV of 0.723 mm/s. Results are similar to those given in Figure 3. Removal of SO2 was slightly higher and conversion to acid slightly lower. Productivity was significantly increased as would be expected from a threefold increase of SO2 in the feed. Despite these differences, the 1% SO2 measurements indicate good reproducibility of our measurements. Earlier work in our laboratory (Haure et al., 1989; Metzinger et al., 1994; Lee et al., 1995; Hinrichs, 1996), employing water for flushing the carbon bed, found that the volume of liquid used for flushing affected reactor performance. Consequently, this operating variable was examined. Figure 4 shows the effect using AC as the liquid phase for a feed concentration of 0.75 SO2 vol%, 5 vol% O2, 2.9 vol% H 2 0 , S G V = 0.1 m/s and T = 21~ Removal from the gas exhibits a slight minimum for the liquid flow rate (SLV) that could be an experimental artifact; however, there is a significant minimum in the conversion to acid at an SLV of around 0.77 mm/s. Productivity increases with SLV above a velocity of 0.8 mm/s. Minimums in conversion or in reaction rates have been observed previously for SO2 oxidation in water-flushed trickle beds (Mata and Smith, 1981; Haure et al., 1989). They have been attributed to opposing effects of wetting of the carbon surface and increasing mass transfer resistance as the liquid velocity

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Figure 4 - The effect of liquid flow rate on the performance of a trickle bed with acetone as the flushing liquid: SO2 = 0.75 vol%, O2 = 5 vol%, H 2 0 = 2.9 vol%, SGV = 0.1 m/s and T = 21~ decreases. Another notable feature of Figure 3 is the 95% S O 2 removal and 93% obtained at the highest liquid flow rate of 1.16 mm/s. Increased transport of SO2 to the carbon surface at higher liquid flow rates probably accounts for this improved performance. Experiments reported heretofore are for feeds containing low levels of SO2; the lowest concentration used, 0.3 vol%, is about that found in stack gases from power plants burning

201 high sulfur coals. To evaluate trickle-bed performance with solvent flushing at higher concentration, several runs were made on feed streams containing up to 6 vol% SO2. This level is within the range encountered for stack gases from smelters; however, it is below levels obtained by burning liquid sulfur. Limitations imposed by our UV spectrometer prevented use of higher concentrations. Oxygen and water vapour concentrations were 5 vol% and 2.9 vol%, respectively; for all the runs, SGV was 0.1 m/s, while SLV was about 0.888. Acetone was the solvent. Figure 5 shows that the productivity of the Centaur carbon increases strongly with SO2 concentration in the feed gas. Productivity reaches about 2.9 g H2SOa/g carbonoh at 6% SO2 and seems to approach a plateau. The high plateau productivity shown in Figure 5 is remarkable considering that the SLV used is in the range where removal, conversion, and productivity depend on SLV indicating that mass transfer,

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probably of 802 to the surface of the carbon particles, is influencing the oxidation rate. Even with this possible transport limitation, the productivity of the Centaurcarbon

Figure 5 - Effect of concentration on activated carbon productivity with an acetone solvent and 02 = 5 vol%, H20 = 2.9 vol%, SGV = 0.1 m/s, SLV = 0.888 mm/s and T = 20~ using solvent flushing is about two orders of magnitude higher than the productivity on the basis of weight of industrial vanadium pentoxide catalysts operating between 350 and 400~ Experiments on the effect of SGV on trickle-bed performance using MEK as the flushing agent in place of acetone were undertaken at 1 vol% SO2. Although MEK has a higher molecular weight than acetone, this solvent exceeds acetone in SO2 removal, conversion to acid and productivity. A possible explanation of the difference lies in the solubilities of SO2 and O2 in the solvent. It appears that at the SLV used in the experiment, mass transport is important. Methyl isobutyl ketone and water were also tested as flushing liquids. Measurement conditions and results are summarized in Table 1. Both these liquids exhibit SO2 removals, conversions to acid and productivities of the carbon well below those for acetone. We suspect that the explanation for the poorer performance

202 with MIBK and water results from the lower solubilities of S02 and 02 in those liquids. Both SO2 conversion and removal are approximately 49% with MIBK. Thus, all the SO2 absorbed in the solvent is converted to acid. It is possible that SO2 is less soluble in MIBK than in acetone or in MEK. With the large excess of O2 used in the experiment, there would be adequate O2 dissolved in the liquid to oxidize all the SO2 dissolved. Iodometric titration revealed no trace of dissolved S02. Table 1 Comparison of flushing agents under similar conditions Flushing Agent Water Acetone Acetone Methyl (Medium (High Liquid Isobutyl Liquid Flow) Flow) Ketone (MIBK) SO2 % (feed) 0.75% 0.75% 0.75% 0.75% SGV (m/s) 0.1 0.1 0.1 0.1 Flushing Velocity (mL/s) 1.54 1.22 1.85 0.333 SO2 Conversion 38% 59% 94% 49% SO2 removal 81% 87% 95% 49% Productivity (g H2SO4/g carbon) 0.21 0.34 0.53 0.28

5.2. Reactivity/Distillation Experiments Economic feasibility of our low-temperature, solvent-based sulfuric acid process depends on solvent recovery. Some preliminary experimental results obtained by mixing 5 mL of 100% or 50% acid with 45 mL of solvent and heating the mixture showed strong discoloration of MEK and AC at 40~ and 55~ (with 50% acid) indicating a reaction. Indeed MEK showed a yellow tinge in the trickle bed experiments. Ethyl ether exhibited no discoloration up to its 35~ boiling point. These results suggest sulfuric acid promotes a slow decomposition reaction for the ketones. Dehydration in the presence of acid probably occurs, although a reaction with impurities in the solvents remains possible as only technical grade solvents were used. Because the ketones reacted slowly, batch distillations using a mixture of 100 mL of solvent and 1 mL of 100% HzSO4 were undertaken in a closed flask fitted with a condenser but without reflux. Using acetone, the solvent-acid solution turned yellow at 55~ after 75% recovery of solvent. Distillation was discontinued at 57~ with 96% recovery. The distillate was clear but the solution remaining was yellow with tiny amounts of a black residue. Repeating the experiment with MEK: 75% recovery of a clear distillate was obtained after the temperature reached 78~ Recovery was 98% at 81~ but the residue was deeply colored and contained a black residue. Employing 10 mL of 100% acid with MEK gave similar results, but only 91% recovery was obtained at 103~ and the residue was a dense black sludge. These experiments demonstrate that ethyl ether is a potential solvent, but the TBR would have to operate under pressure to minimize solvent evaporation. Acetone can also be used, but the separation step would have to be carried out under vacuum to keep temperatures below 40~

203 6. DISCUSSION Our experiments demonstrate that sulfuric acid can be produced over activated carbon in a trickle bed flushed continuously with a solvent. Catalyst activity in this system is very high so that trickle-bed heights and thus pressure drops will be reasonable. As expected, gas and loading liquid are important parameters governing SO2 recovery from the gas phase and productivity of the catalyst. Of the solvents investigated, acetone and ethyl ether appear to be feasible, although vacuum distillation will be needed to separate acid and acetone and a pressurized process will be required if ethyl acid is the solvent. There are three areas of uncertainty that must be addressed in further development. The first is the continuous separation of solvent and acid. Lack of data on solubility, activity coefficients, and heats of mixing suggests that the separation would be easiest to study experimentally in a plate column. The second question is the solubility of the solvent in the acid product and whether or not ppm-levels of solvent affect the use or value of the acid. The third question is the loss or consumption of solvent in the process, particularly solvent loss in the gas phase leaving the trickle bed. There is a complex relationship between the liquid loading, requirements of acid-solvent separation and loss of solvent to the gas. Optimization is called for but this cannot be undertaken until more experimental data are available.

7. A C K N O W L E D G M E N T S Financial support of The Natural Sciences and Engineering Research Council of Canada in the form of a Strategic and Operating Research Grants (to PLS and RRH) is gratefully acknowledged. So too is the help of Atomic Energy Canada Limited, Chalk River, in providing the carbon coating of our trickle-bed packing. REFERENCES 1. O.K. Davtyan and E. N. Ovchinnikova, Doklady Akad. Nauk S. S. S. R. 104 (1955). 857-860. 2. M. Hartman and R. W. Coughlin, Chem. Eng. Sci. 27(1972) 867-880. 3. P. Haure, R. R. Hudgins, and P. L. Silveston., AIChE J. 35 (1989) 1437-1444. 4. J.K Lee, R. R. Hudgins, and P. L. Silveston., Chem. Eng. Sci. 50 (1995) 2523-2530. 5. A.R. Mata and J. M. Smith, Chem. Eng. J. 22 (1981) 229-235. 6. J.V. Metzinger, A. Kuhter, P.L. Silveston, and S.K. Gangwal, Chem. Eng. Sci. 49 (1994) 4533-4546. 7. S. Suppiah, K. T. Chuang, and J. H. Rolston, Can. J. Chem. Eng. 65 (1987) 256-263. 8. R.V. Vladea, N. Hinrichs, R. R. Hudgins, S. Suppiah, and P. L. Silveston, Energy & Fuels 11 (1997) 277-283. 9 R.V. Vladea, R. R. Hudgins, and P. L. Silveston, Canadian Informal Patent Application No. 2 237 744 (August 10, 1998). 10. W.Y. Won, A. Lohi, R. R. Hudgins, and P. L. Silveston, Note submitted to Canadian Journal of Chemical Engineering, Feb. 2000.


Studies in Surface Science and Catalysis 133 G.F. Fromeni and K.C. Waugh (Editors) 9 2001 Elsevier Science B.V. All rights reserved.

205

Butane oxidation to maleic anhydride over a VPO catalyst following the riser regenerator approach St. Hess, H. Freund, M. A. Liauw and G. Emig Lehrstuhl fiir Technische Chemie I, Universit~it Erlangen-Ntirnberg, Egerlandstr. 3, D-91058 Erlangen, Germany* With a given pretreatment, furan selectivities of up to 10 % are achieved during the butane oxidation over vanadium pyrophosphate (VPO). A reaction model based on a Mars-van Krevelen mechanism including the formation of furan is presented. The agreement with the experimental results is satisfactory. While it is derived from steady-state experiments, the model still allows for simulating unsteady-state experiments, too. The course of the selectivities to maleic anhydride and furan is consistent with the notion of VPO as an oxygen storage. Diffuse reflectance vis spectroscopy seems to be a convenient in-situ method for validating this model. 1

INTRODUCTION Vanadium pyrophosphate (VPO) is unique in its capacity to selectively oxidize n-butane to maleic anhydride (MA). Gas phase oxygen is made responsible for the total oxidation of butane and MA to carbon oxides. It is known that oxygen stored in the crystal structure of the VPO catalyst is required for the desired reaction. Thus it is possible to run the reaction without gas phase oxygen in order to reduce the undesired total oxidation and to increase selectivity and yield, and to reoxidize the catalyst after oxygen depletion. Du Pont has suggested to separate the two steps in space by the so called "RiserRegenerator Concept". The (fast) reaction of butane to MA is performed in a transport reactor where a narrow residence time distribution of the particles is achieved. The reoxidation of the catalyst is performed in a bubbling fluidized bed to ensure a maximum oxygen uptake. Temperatures of reaction and reoxidation can be optimized separately and no attention has to be paid to explosion limits because gas phase oxygen and butane are never fed simultaneously. The kinetic studies presented here have been carried out in a laboratory fixed bed at steady state and under periodic variation of the feed composition, respectively. One major focus is the formation of furan in the low conversion region. 2

2.1

EXPERIMENTAL

Catalyst preparation and characterisation

In this study, VPO catalysts were prepared via an "organic route" (as opposed to the "aqueous route"). The synthesis after a recipe by Uihlein [1 ] has been systematically monitored and improved. The reproducibility of the synthesis could be significantly enhanced by *St. Weiss and H. Hansemann have contributed to this paper. Financial support by the DFG (FOR 262 / 2-1) is gratefully acknowledged.

206 using a nearly alkaline free V205 (Purity > 99.6 %) with a definite size < 0.15 mm (GfE, Niimberg), ensuring an efficient reduction of the V205 and complete crystallization of the VPO during the synthesis. The precursor (P/V = 1/1.1, BET surface = 7.9 mVg) is characterized by XRD, FTIR, BET, REM, ICP and titrimetric measurements. The XRD spectrum of the resulting precursor is in good accordance with literature data [2 - 4]. The catalyst is then made abrasion resistant by encapsulation in a hard porous silica eggshell, resulting in encapsulated particles of 60 - 120 pm diameter. The achieved abrasion resistance is in the. same range as commercially used FCC-catalysts. EDX-mapping of ground particles revealed an enrichment of SiO2 near the surface of the particles. The encapsulated precursor is calcined at 450~ for one hour in nitrogen (heating rate = 5 K/min), then one hour in air and afterwards activated under reaction conditions (2 vol.-% n-butane in air at 430~ for 14 days to obtain an equilibrated catalyst. This catalyst showed yields that were about 10 % higher than those of an equally tested industrial reference catalyst when activated under the same conditions.

2.2

Experimental Setup

Kinetic measurements were carried out in a lab-scale setup. The concentration can be varied from 1 to 15 vol.-% n-butane in "synthetic air" (20 vol.-% oxygen in neon), the temperature from 300~ to 650~ The setup is designed to allow steady state experiments as well as periodic operation. Analysis of the products is performed by quantitative mass spectroscopy (HIDEN RGA HAL/3F RC 301). Calibration of the MS was performed with a GC. Ne (m/e 20) was used as inert gas instead of nitrogen (m/e 28) and as internal standard. Overlapping peaks were deconvoluted by subtraction. In order to simulate the conditions in the riser-regenerator, a special minireactor was designed that allows fast switching between different feeds without considerable broadening of the pulse fronts (Fig. 1). In combination with fast-switching VALCO-valves (switching time about 10 ms) and a specially optimized coupling to the mass spectrometer (redesigned MS sampling system), pulses could be kept as low as 150 ms with the reactor installed. The Erlangen Micro Reactor Optimized for Pulse operation, EMROP, consists of a 60 mm long stainless steel tube with 4.6 mm inner diameter. Supply and off line are done in 1/16" silicosteel. The widening from 1/16" to 1/4" and back is realized by modified VALCO reducers. The electrical heating consists of separate elements for preheating and reaction zone. Care was taken that no transport limitations impaired the kinetic measurements.

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207

Fig. 2: Selectivity to maleic anhydride (left) and furan (right) as a function of reactor temperature and butane inlet concentration at a flow rate of 60 Nml/min. 2.3

Kinetic measurements

The applied butane concentration for both steady state and periodic operation varied from 2 to 8 vol.-% in "synthetic air", the temperatures from 400~ to 440~ to cover common reaction conditions. The volumetric flow rates were 40, 60 and 80 Nml/min, respectively. Steady state measurements were performed with increasing butane inlet concentrations, two hours for each concentration, where the last 30 minutes were taken to collect the data. 3

EXPERIMENTAL RESULTS The degree of conversion was about 10 - 40%. The selectivity to maleic anhydride increased with decreasing reactor temperature and decreasing butane inlet concentration, respectively. The maximum selectivity to furan of 10% was found at low temperatures and high butane inlet concentrations (Fig. 2). This is in agreement with Sananes et al. [5] who found appreciable amounts of furan at low conversion. 4

MODELING In various competing reaction networks proposed for the butane oxidation [6, 7], furan is an intermediate. The extent to which it is formed has been linked to the oxidation state of the VPO. Rodemerck et al. prepared VPO catalysts exhibiting six different oxidation states ranging from 3.18 to 4.90 [8]. These are the average valence states as determined by potentiometric titration. The catalytic behavior of the respective samples as determined with a TAP reactor setup is summarized in Table 1. These observations were explained by the special role of V 5+ centers that are assumed to be indispensable for the oxidation of furan to MA. All other reaction steps may occur on V 4+ centers. This interpretation may also apply to the results depicted in Fig. 2. An increase in butane inlet concentration, i.e. an increase in reducing conditions, leads to a decrease of V 5+ centers, and hence to a decreased oxidation of the intermediate furan to MA. The observed temperature dependence of the selectivities, too, is in agreement with the given explanation: with increasing temperature, the oxygen uptake of VPO is enhanced, leading to a rise of the oxidation state of the catalyst and hence to a decrease of the amount of detected furan.

208 Oxidation state

Color

Catalytic behavior

4.90

yellow

Initially CO2 formation only. MA or furan are only detected after several minutes of continuous butane exposition.

4.10 4.02 3.96 3.73 3.18 Table 1: Catalytic behavior in oxidation states after [8].

green gray gray black

Observation ofMA, furan and CO2.

Only butenes, butadiene, furan and CO2, no MA a TAP experiment of VPO catalysts with different average

The following mathematical model is based on a network (Fig. 3) where butane (Bu) may be oxidized to maleic anhydride (MA) via furan (Fu). All three compounds may undergo total oxidation to the carbon oxides. The oxidant in all cases is VPO in its oxidized form which is then reoxidized by molecular oxygen. This is in contrast to the approach that total oxidation is mainly due to molecular oxygen in a direct manner. The rate equations for this Mars-van Krevelen mechanism are:

(~) (~) (~)

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(4)

(3)

(~) r5 = k5 "PMA "| (5) (6) Oxidation rox = kox "Po~ "l~)Cat,Red Parameters for this model were fitted using the software SiaMod. The results were checked with independent data in a satisfactory manner. 5

MODELING RESULTS The proposed model yielded good results. The parity plots reveal that the reactant exit concentrations are reflected properly within 10% error by the model (Fig. 4). There is a slight systematic deviation in the oxygen exit concentrations for higher flow rates. The product exit concentrations scatter within 20-25% error, respectively, reflecting the low absolute concentrations. It is noteworthy that within this range of conversion there is virtually no difference if the total oxidation of maleic anhydride and furan is neglected. This suggests that these compounds are sufficient stable at the given conditions.

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,.COMA

'b

CO / CO 2 Fig. 3: Reaction network. All steps require oxidized sites on the VPO.

209 200

80 +

.~

60-

150-

40-

~9 100 t-,q

O

20-

I

0

~

I

50-

I

20 40 60 p(Bu), exp., mbar

I

80

0

8.0

1.2

6.0

,~ 0.9

4.0

0.6

I

I

50 100 150 p(O2), exp., mbar

200

+25%

-25% ~2.0

"~ 0.3

0.0

0.0 0.0

I

!

I

2.0 4.0 6.0 8.0 0.0 0.3 0.6 0.9 1.2 p(MA), exp., mbar p(Fu), exp., mbar Fig. 4: Parity plots of the outlet partial pressures of butane, MA, furan and oxygen. 0.0

MA _

=i

"fi

LA.aA~

0

I

I

I

I

5

10

15

20

time, s

25

~

Bu

"

.=

0

5

10

15

20

25

time, s

Fig. 5: Time series of the outlet partial pressures (arbitrary units) when a butane pulse is led over the preoxidized VPO catalyst at 5 s < t < 20 s. Experiment (left) and simulation (right).

210 Unsteady state measurements show that a cyclic steady state may be attained where the response of the catalyst to concentration cycling does not change over hundreds of cycles. Upon butane admission after oxidation and a purge, maleic anhydride is formed over the oxidized catalyst (Fig. 5). With increasing time, the butane conversion as well as the MA selectivity decreases. According to the model, this is due to a decrease in oxidized sites. Consequently, the furan selectivity displays a certain increase with time. This behavior is also found in numerical simulations using the parameters taken from steady-state experiments. 6

OUTLOOK A validation of this approach should be possible by using in situ DR (Diffuse Reflectance) Vis spectroscopy to monitor the color of the VPO under process conditions. Following Rodemerck et al., this should reflect the oxidation state of the catalyst. It is noteworthy that Golbig and Werther [9] report different colors of the VPO catalyst behind the riser (brown) and behind the regenerator (dark green) in their lab-scale riser-regenerator experiments. Preliminary experiments with high temperature (< 600~ fiber optics show satisfactory spectral and temporal resolution. REFERENCES 1. Uihlein, K., Butanoxidation an VPO-Wirbelschichtkatalysatoren, Karlsruhe, 1993, Ph.D. thesis. 2. Cavani, F., Trifiro, F., Vanadium/phosphorus mixed oxide from the precursor to the active phase: Catalyst for the oxidation of n-butane to maleic anhydride, 1995, Preparation of catalysts IV, 1. 3. Hutchings, G. J., Sananes, M. T., Sajip, S., Kiely, C. J., Burrows, A., Ellison, I. J., Volta, J. C., 1997, Improved Method of Preparation of Vanadium Phosphate Catalysts, Catal. Today 33, 161. 4. Sananes, M. T., Hutchings G. J., Volta J. C., 1995, On the Role of the VO(H2PO4)2 Precursor for the n-Butane Oxidation into Maleic Anhydride, J. Catal. 154, 253. 5. Sananes, M. T., Hutchings, G. J., Volta, J. C., 1995, n-Butane Oxidation to Maleic Anhydride and Furan with no Carbon Oxide Formation Using a Catalyst Derived from VO(H2PO4)2, J. Chem. Soc., Chemical Commun., 243. 6. Kubias, B., Rodemerck, U., Zanthoff, H.-W., Meisel, M., 1996, The Reaction Network of the Selective Oxidation of n-Butane on (VO)2P207 Catalysts: Nature of Oxygen Containing Intermediates, Catal. Today 32, 243. 7. Xue, Z.-Y., Schrader, G. L., 1999, Transient FTIR Studies of the Reaction Pathway for nButane Selective Oxidation over Vanadyl Pyrophosphate, J. Catal. 184, 87. 8. Rodemerck, U., Kubias, B., Zanthoff, H.-W., Wolf, G.-U., Baerns, M., 1997, The Reaction Mechanism of the Selective Oxidation of Butane on (VO)2P207 Catalysts: The Influence of the Valence State of Vanadium, Appl. Catal. A 153,217. 9. Golbig, K. G., Werther, J., 1997, Selective Synthesis of Maleic Anhydride by Spatial Separation of n-Butane Oxidation and Catalyst Reoxidation, Chem. Engng Sci. 52, 583.

Studies in Surface Science and Catalysis 133 G.F. Fromentand K.C. Waugh (Editors) Published by Elsevier Science B.V., 2001

Effects of Particle Size and Modified SAPO-34 on Conversion of Methanol to Light Olefins and Dimethyl Ether Michael G. Abraha, Xianchun Wu, and Rayford G. Anthony

Kinetics, Catalysis and Reaction Engineering Laboratory Department of Chemical Engineering Texas A &M University College Station, Texas 77843-3122 Abstract Selectivity of SAPO-34 and modified SAPO-34 to ethylene and propylene in the MTO process was studied using a fixed bed reactor at 1 atm pressure and a temperature range of 400-450~ Particle size of 1.1 mm and powder of 1.1 gm SAPO-34 catalyst were compared. Powder catalyst gave higher selectivity to hydrocarbons compared to bigger particle sized SAPO-34. Also conversion of methanol to light olefins was studied by using modified SAPO-34(O) with larger pore size prepared using a surfactant. TPD-NH3 and FTIR-NH3 showed that SAPO-34(O) had stronger acid sites and also more coke was deposited. SAPO-34 (O) had greater average pore size and less surface area compared to SAPO-34. Increased stronger acid sites and increased average pore size may have led to increased selectivity to ethylene and propylene. Keywords: Methanol, Ethylene, Propylene, SAPO-34, Conversion

Introduction Light olefins especially ethylene (C2 =) and propylene (C3 =) can be formed from methanol in the MTO process (Chang et al., 1979) using catalyst SAPO-34. Several other catalysts like ZSM-5 (Marchi and Froment, 1991), and Chabazite (Liu et al., 1984) have been tested. Physical and chemical properties of the catalyst influence its selectivity to hydrocarbons. The physical factors that affect the selectivity of the catalyst are temperature, pressure of the fixed bed reactor, and space velocity of the feed. Other physical characteristics that influence selectivity are crystal size, crystal size distribution, pore size and pore size arrangement. The chemical characteristics that influence the selectivity are acid site density, strength of acid sites, and type of surface acid groups. The selectivity to ethylene and propylene is defined as

Selectivity to ethylene(C~)=

2 *(f'l~% H..... t-1;IC2H4,in ) f'I('H~OH,in -- f'l(.,H3OH,oul

Selectivity to propylene(C;)= Where h~,m and n,,ou, product.

3*(hc3H.....,-/;/C~H.... )

~'ICH3OH,in-- ~'l(,H~()H,oul is the mole flowrate of component i in the feed and

211

212

Experimental Equipment A fixed bed reactor at atmospheric pressure and reaction temperature of 400450~ was used. SAPO-34 pelletized, crushed and sieved into 1.1ram catalyst particles was packed into a stainless fixed bed reactor with the top bed diluted with 25 wt % alpha alumina, middle with 50 wt % alpha alumina, bottom 75 wt % alpha alumina. The total amount of SAPO-34 was 5g diluted with 5g total alpha alumina. Also another experiment was done with 1.1 ~tm SAPO-34 powder packed into the reactor with oralumina. The stainless steel fixed bed reactor was surrounded by a three-zone heater controlled independently to keep the reactor bed isothermal at 400~ 425 ~ or 450 ~ The feed contained 80 mol % water and 20 mol % methanol, which was pumped by a Bodine Electric Company pump to the top of the catalyst bed. The reaction products are cooled with a copper coil immersed in an ice bath. The liquid product was collected in a bottle and the gaseous hydrocarbons flowed through a bubble flow meter. Regeneration of the coked catalyst takes place with a mixture of 2 mol % 02 and 98 mol % nitrogen where the oxygen percentage is slowly increased to a maximum of 21%. During this time the temperature slowly increased from 400 to 500~ The gas product was analyzed by using a Carle Chromatograph and Shimadzu GC-17A with a TCD and Propack Q and the liquid part was analyzed using Shimadzu with Propack Q. SAPO-34 was prepared using procedures in a patent (Lok et al., 1984). SAPO-34 (O) was prepared by adding a surfactant. The catalysts were prepared by initially mixing aluminum isopropoxide, phosphoric acid, tetraethyl orthosilicate and templating chemical tetraethyl ammonium hydroxide (TEAOH) in a stirred polyethylene plastic bottle using a magnetic stirrer. Silicoaluminophosphate catalyst of the form (TEA)zO:0.3SiOz:AlzO3:PzOs:50.0H20 (Lok et al., 1984) was produced using TEAOH. Surfactant was used as part of the template to form catalysts with different secondary pore size. An autoclave equipped with Reliance TR motor driven stirrer was used to stir the reaction mixture at 300 rpm and a temperature of 220 ~ The autoclave was equipped with an Autoclave Engineers heater and K type thermocouple to control temperature accurately.

Catalyst Characterization Surface area, micropore volume, and pores size measurement was done by using N2 with Micromeritics ASAP 2000 equipment and ASAP 2010 software. Surface acidity was determined from temperature programmed desorption of ammonia (TPD-NH3) and equipment used was Micromeritics 2000. Prior to starting tests, catalyst was dried under 30 ml/min Helium continuous flow at 300~ The temperature at a rate of 10~ was increased from 27~ to 650~ under a continuous flow of 30 ml/min Helium. Surface groups were determined using diffuse reflectance fourier transform infra red (FTIR) spectroscopy with Nicolet Magna-IR 560 Spectrometer. The catalyst was dewatered at

213

300~ for 1 hour under 20 ml/min He flow before tests were performed. Enviro II ICAP Spectrometer and ThermoSPEC Version 6.20 software was used to analyze composition of the catalyst. Catalyst was dissolved in HF and standard samples purchased from SIGMA. X-ray diffraction (XRD) of all catalysts was made by a XDS 2000, Scintag Inc equipment with Cu Kc~ radiation. The angle ranged from 2 to 85 ~ at a scan rate of 4 ~ Crystal size measurements were done using scanning electron microscopy (SEM) at magnification of X 3,500. Results and Discussion

XRD results for SAPO-34 and SAPO-34(O) are shown in Fig. 1. The XRD results of SAPO-34(O) show different peaks that shows it is a different crystal structure compared to SAPO-34. The FTIR at 300~ shows four absorption bands on the surface of SAPO-34 and SAPO34(0). There are two absorption bands at 3601 and 3627 cm -I which are associated with strong acid sites (Fig. 2). These two bands also disappeared for the coked catalyst unlike the two weak acid sites at 3674 and 3734 cm l (See Fig. 3). The same conclusion can be reached using ammonia FTIR where ammonia was injected at 25~ Prior to injecting ammonia, the sample was dewatered for 1 hour under the continuous flow of helium at 300~ The weak acid sites 3674 and 3734 cm l do not absorb ammonia as strongly as the acid sites at 3601 cm -~ and 3627 cm l. The sample was heated at a rate of 10~ in steps of 50~ The weak acid sites appear at low temperature versus the strong acid sites, which appeared at 400-450~ The strong acid site at 3601cm -1 was attributed to A1-OHSi and the weak acid site at 3674 cm l corresponds to extra framework A1-OH and 3734 cm -1 corresponds to Si-OH. FTIR using pyridine to determine the ratio of Bronsted and Lewis acid sites was conducted but due to stearic effects pyridine was not able to diffuse into the pores of SAPO-34 with 0.43nm channel size except on the outside of the crystal. The temperature program desorption using ammonia (NH3-TPD) at 10~ illustrated in Fig. 4 shows two peaks, one for weak acid site and another for strong acid site. The results for measured and calculated acid site density, and temperature where peaks appeared are shown in Table 1. Also from composition measurements using ICP listed in Table 2 and assuming that there is one acid site for each Si. We can determine that there is approximately one Si or one acid site for each 12 tetrahedra atoms in SAPO-34 and one acid site for each 14 tetrahedra atoms for SAPO-34(O). This is in agreement with results shown that increased A1/Si ratio, decreased acid sites (Marchi and Froment, 1991). Also Kang et al. (1998) showed that increased A1/Si ratio decreased the acid sites. Results in Table 1 show Si/A1 for SAPO-34(O) > SAPO-34. This could have resulted in increased acid site density. Calculated values shown in the table assuming one acid site for each Si is less than the measured acid site density. This indicates that there are more weak acid sites formed that do not involve Si. This was found to be true as shown in FTIR analysis in Fig. 2 where the existence of extraframework A1-OH at 3674 c m 1 w a s confirmed. The peak at 3627 cm -~ is yet to be determined. The measured values of 3028 gmol/g total acid site density for SAPO-34 (O) is greater than 2637 ~tmol/g total acid site density for SAPO-34 as shown in Table 1. The increased

214 acid site density leads to faster deactivation of SAPO-34(O). SAPO-34 and chabazite showed that reduced acid site density reduced deactivation at low space velocities (Dahl et al., 1999). Conversion of methanol was high initially for the high acid site density of both SAPO and Chabazite and decreased fast. Inui and Kang (1997) showed that decreased acid site density on the internal surface of crystals correlated with increased ethylene selectivity at GHSV (Gas Hourly Space Velocity) of 1000 h L (equivalent to WHSV=5.4 h l ) and 1 h TOS. Internal acid site density of 4 gmol/m 2 showed the best result combined with other factors like crystal size. SAPO-34 has super cages of 0.75 nm X 0.82 nm, which can be reached through 0.45 nm X 0.41 nm. As illustrated in Table 2, BET surface area measurement show 512 m2/g and average pore size of 0.52 nm using nitrogen Micromeritics 2000 equipment. Activity tests for a 1.1ram particle size with total 5g of SAPO-34 which was divided into three beds where the top bed has 25 wt % SAPO-34, middle has 50 wt % and bottom has 75 wt % SAPO-34 was prepared. Another fixed bed reactor with total of 0.61g SAPO-34 size of 1.1 ~tm powder catalyst and (z-alumina was also prepared. The results are shown in Figures 5 and 6. A comparison between SAPO-34, 1.1 mm catalyst and powder SAPO-34 (-~l.lgm) was made. The results show that 1.1 ~tm powder catalyst produces more ethylene and propylene than dimethyl ether (DME) compared to the 1.lmm SAPO34 bed. The most probable cause is that there is diffusion limitation to DME. The DME was not able to diffuse into the 1. l mm particle acid sites and form ethylene and propylene. MeOH .~

~ DME ~---~Propylene IL Ethylene

~ Higher hydrocarbons (Park and Froment. 2000)

It was shown that uniform crystal size gives a uniform reactant residence time, therefore better control. Crystals with size of 0.85 gm showed high selectivity towards ethylene (Inui and Kang, 1997) using Ni-SAPO-34. However, Chen at al. (1999) using SAPO-34 state that crystals in the range 0.4-0.5 ~tm give the largest olefin production whereas 0.25 ktm size crystals form a large amount of dimethyl ether that leave the crystals before converting to olefins. Larger crystal sizes produced less olefins. Coke content was greater for larger crystals than smaller crystals. The results in Fig. 5 and Fig. 6 suggest that reduced size give more ethylene and propylene than DME. It was also shown with crystal sizes 0.25, 0.49, and 2.50 ~tm that particles > 2 ~tm to be diffusion limited using ethanol and 2-propanol over SAPO-34 (Dahl et al., 1999). The conversion of MeOH is usually higher initially for higher temperature, example 450~ compared to 400~ and decreases faster as a result of more coke accumulating in the catalyst at higher temperature. See Fig. 7. The equivalent conversion tbllows the same trend.

215 The selectivity to ethylene and propylene is higher using powder SAPO-34 (O) compared to using powder SAPO-34 (Fig. 8). This was attributed to stronger acid sites present in SAPO-34(O) as obtained by using TPD-NH3 and FTIR-NH3. Comparing TPD-NH3 and FTIR-NH3, strong acid site peak from TPD-NH3 must belong to the FTIR-NH3 acid site of 3601 and 3627 cm~ which appeared between 400~ and 450~ The strong peak for SAPO-34(O) comes at 404~ greater than peak for SAPO-34 which appears at 384~ This shows that SAPO-34(O) has stronger acid sites. The stronger acid sites of SAPO-34(O) tend to protonate DME and convert it to higher hydrocarbons like ethylene and propylene. It is known that DME can form on weak as well as strong acid sites but olefins form only on strong acid sites. This was also shown by Compelo et al. (2000) where SAPO-34 and SAPO-5 prepared with stronger acid site density produced more hydrocarbons than DME compared to SAPO-11. The other reason is that since part of the pores of SAPO-34(O) are made bigger using a surfactant template. DME with larger kinetic diameter and which has larger molecular weight will diffuse though the pores. Therefore, DME gets access to acid sites. This must have led to higher conversion of methanol to ethylene and propylene when SAPO-34(O) is used. Note the same strong acid sites disappeared for the coked catalyst shown in FTIR absorbance Figure 3. It can be concluded that coke forms on strong acid sites, Si-OH-A1. The FTIR shows that coked SAPO-34(O) aliphatic intensity is higher which could mean higher coke. The peaks at 2957, 2929 and 2868 cm -1 are associated with aliphatic stretching. Table 1. Ammonia TPD using 10~ and 30ml/min He flow. Average temperature belongs to peaks of strong and weak acid sites. Catalyst Type Weak T(~ Strong (~ Acid density Calculated Acid density (~tmol/g) Acid density (j.tmol/m2) (pmol/g) SAPO-34 117 384 2637 1214 3.48 SAPO-34(O) 105 404 3028 1215 5.65 Table 2. Composition in moles using ICP and Surface area measurements Composition BET Surface Micropore Pore size Crystal Catalyst Area (m2/g) Volume (cm3/g) (nm) size (pm) SAPO-34 All.oP0.7oSio.14 512 0.19 0.52 0.6-1.5 SAPO-34(O) All.0P0.89Sio.16 391 0.15 0.53 0.9-3.2 References

J.M. Campelo, F. Lafont, J.M. Marinas, and M. Ojeda, Appl. Catal., 192,(2000), p. 94. C.D. Chang, W.H. Lang, and R.L. Smith, J. Catal., Vol. 56, (1979), p. 169. D. Chen, K. Moljord, T. Fuglerud, and A. Holmen, Microporous and Mesoporous Mater., vol. 29, (1999), pp. 191-193. I.V. Dahl, R. Wendelbo, A. Anderson, D. Akpriaye, H. Mostad, and T. Fuglerud, Microporous and Mesoporous Mater., 29, (1999), pp. 159-171.

216

PS 948. t

8.838

4.t36

2. ~176 2.~52

1.723

1.~41

1.743

1.19e

10~

as3.

9o I

758.

eo I

663.6-

70

i 568 8j L

,

60 i

""?t.i Fig. 1. Continuos X-Ray powder diffraction using Cu Ka radiation source at a scan rate of 4 ~

it,,t

,2j

1o:

f

Jt

~ --

i

,~ .J

//

/

i

,~ i ) t....~,,.,. AS I + w

(1)

0~= CwK~.C, +k-~CA

A + S2

< K2 " AS 2

(2)

02 = 1+

W + S2

< K~, > WS 2

(3)

02w =

K2 "CA + Kzw" Cw

(la) (2a)

K 2 9CA

K2w "Cw I + K 2-C A+K2w.C w

(3a)

AS I + s 2

~

> E+WS I+s 2

(4)

R 3 = k3.01-(1-02-02w )

(4a)

AS 1 + AS 2

k, > D + WS 1 + S 2

(5)

R 4 = k 4.0t.0 z

(Ja)

A, E, W, D are isopropanol, propene, water and diisopropylether, Ol and 02 are the fractions of sites 1 and 2 occupied by isopropanol, 02w is the fraction of sites 2 occupied by water. In theoretical studies Thullie and Renken [7] showed that operating under periodic variation of the inlet concentration leads to a considerable increase of the reactor performance exceeding the maximal performance obtainable under steady state conditions. The attainable enhancement under periodic operation depends on the sorption equilibrium constants and the sorption rates. Assuming instantaneous equilibrium on $2 and irreversible adsorption on S1 the model calculations lead to a maximal performance under periodic operation for cycle periods approaching zero (tp--->0) Experimental studies of the dehydration of ethanol confirmed the predicted rate enhancement, but due to limitations of the used conventional fixed bed reactor cycle periods shorter than 600 s could not be realised [8].

1.2 Periodic operation in microreactors It is known that microchannel reactors, due to their small dimension and well defined structure have many advantages compared to conventional fixed bed reactors. The main ones are an efficient temperature control and well defined flow patterns. As the channel diameters are in the order of micrometers, microreactors operate under laminar flow conditions resulting in a parabolic velocity profile. But, due to the short radial diffusion times the radial concentration profile is flat, resulting in a narrow residence time distribution of the reactant. The latter characteristic is of crucial importance in the actual study. Only reactors with an uniform residence time can be used to get meaningful kinetics information under periodic operation at short cycle periods [9].

241 2. MATERIALS AND METHODS

2.1 Experimental set-up The experimental set-up consists of two separate feed sections converging to a 4-way valve, a multichannel microreactor and a quadrupole mass spectrometer. The variations of the inlet concentration are generated switching the 4-ways valve. The specifically designed microchannel reactor consists of microstructured stacked plates (Fig. 1). The geometry of the plates and of the stack itself is optimised to avoid mixing in the entrance and outlet area and to distribute evenly the flow between the different channels [ 10]. Since the flow in the structure is always laminar, the pressure drop can be easily estimated, and the optimisation of the reactor design is facilitated. The reactor, plates as well as the housing, is constructed of stainless steel. Each plate contains 34 rectangular channels of 300 ~tm width, 240 ~tm depth and 20 mm length. It contains 5 reaction plates (coated with catalyst) and 5 thermostatisation plates glued together. The deposited catalyst mass is 75 mg, the measured reactor volume, VR, is about 0.5 cm 3. For the experiments described here, the reactor block is heated from the outside by an electrical tape. As the enthalpy of the model reaction is small and low reactant concentrations are applied, a homogeneous temperature distribution is obtained. The temperature inside the reactor is measured using a K-type thermocouple (Philips AG, Dietikon, Switzerland) introduced into the outlet tube. The catalyst, y-alumina is deposited in the microchannels according to a special experimental procedure, allowing a uniform catalytic layer in the channels [6]. The efficient use of the microchannel reactor for periodic operation at cycle periods as short as l s was evidenced in an experimental study [6].

l~lg. 1: Microreactor (lnstitut tiir Mikrotechnik, Mainz (IMM)): left: sketch of fluid flows. Right: picture of the whole reactor block

2.2 Reactants Nitrogen (>99.95%, Carbagas, Lausanne, Switzerland) was used as carrier gas. Isopropanol (>99.5, Fluka Chemie AG, Buchs, Switzerland) was provided by a temperature controlled bubble-column fed by N2.

242

2.3 Procedure Prior to each experiment, the reactor was heated to 400~ for 1 hour to activate the catalyst. After cooling to 200~ the reaction was run under stationary conditions for 14 hours in order to stabilise the catalyst activity. The concentrations of diisopropylether, isopropanol and propene were determined by measuring the intensities of the masses 87, 43 and 42 respectively every 2 s with the mass spectrometer. All experiments were carried out at 200~ and 130 kPa. The average inlet concentration of isopropanol was fixed to 0.45 mol/m 3 with variations of +2% between the experiments. The inlet concentration was switched between 0 and 0.9 mol/m 3. Each experiment was repeated during 30 min, in order to obtain cycle invariance.

3. RESULTS

3.1 Transients experiments The evolution of the propene concentration after a switch of isopropanol to inert is shown on Fig 2. There is a two- to threefold increase of the propene formation rate in the seconds following the switch due to the evacuation of the inhibiting isopropanol and water adsorbed on $2. The propene formation rate then decreases, due to the consumption of the isopropanol adsorbed on the sites S1. A successful strategy of periodic operation has to take advantage of this increase in productivity, generating it repeatedly. To have a high productivity, the feed of isopropanol has to be resumed as soon as the formation rate of propene begins to decrease and the cycle phase in which the reactor operate under stationary conditions has to be as short as possible. Q=1.33cmals

I

Q=3.33 cma/s e9 oc~'

soy ~" -1to t.)

2 s O

t.)

t

ov (1=0.67 crn3/s OV

r t.)

t.)

f

,

-so

,

o

,

,

,

,

.

so

.

.

.

|

~oo

,

,

|

15o

Time Is]

Fig. 2: Transient behaviour of the C 3 H 6 outlet concentration (relative to the steady state concentration) after a switch isopropanol/Inert. T=200~ P=l.3 bar, CiPrOH switches from 0.86 mol/m 3 (STP) to 0 at t=0 s

3.2 Periodic operation After a switch from isopropanol ---> inert the outlet concentration of the alcohol falls quickly and desorption from the S2-sites occurs resulting in a drastic increase of propene production whereas the ether formation slows down. At long cycle times the propene concentration

243 passes through a maximum within the half-period under inert (Fig. 3B). This can be explained with the consumption of ASI species. When the isopropanol feed is restarted the alcohol is at first completely adsorbed on the surface displacing the water from the catalyst. As a consequence $2 sites are blocked and the propene production drops drastically. After some time the rate of propene production is accelerated as the initially produced water is expelled from the catalyst and the coverage of Sl with isopropanol increases simultaneously. At short periods (Fig. 3A), the desorption time is insufficient and an important fraction of S2sites are still blocked when the isopropanol feed is restored. Therefore, the propene production rate does not reach is maximum possible value and a considerable amount of ether is produced throughout the whole period. The oscillations are more symmetric and the production rate of ether is roughly proportional to the isopropanol concentration.

0.7 ,4--'

~

0.4

o .E. 0.3 C .o ,.. .m 0.2

x Propene/~ iPrOH ~ \_

~_ . . . . . .

:

itp=30s i !~=0.175 =

C O O

.

:

if

0.4

:

._ . . . . .

:

..,

:B

tp=180 s

=0. 0s i

_8

r 0.1 0 0.01 ~

: 0.00

O

0

20

40 Time [$]

60

~ 1o.7 6

0

0

100

200

300

T i m e Is]

Fig. 3: Evolution of the outlet concentrations during a period T=200~ P= 1.3 bar, split 0.5, C~PROH,~.=0.43mol/m 3 (STP); CiPrOH switches from 0.86 mol/m 3 (STP) to 0 at t=0 s. Experimental datacompared to model simulations (lines) 4. DISCUSSION 4.1 Reactor model Residence time distribution experiments have shown that the reactor behaves almost like a plug flow tubular reactor with a small dispersion [6]. The RTD can be described using a tanks in series model with 35 ideal mixers. As the simulated reactor behaviour based on the kinetic model is only slightly influenced by the number of ideal mixers for more than 8 tanks, this value was used for all simulations in order to reduce the calculation time needed for parameter optimisation. 4.2 Parameter estimation

The parameters to be estimated are the equilibrium constants K1, K2 and K2w, the reaction constants k3 and k4 and the densities of sites S1 and $2 defined as Z1 and Z2 respectively. Zl can be estimated separately. At steady state, the equilibrium (eq. 1) is almost completely displaced to the right. After a switch isopropanol---~inert isopropanol adsorbed on $1 reacts

244 and leaves the surface as propene, while the alcohol adsorbed on S2 is desorbed rapidly. Therefore, ZI can be estimated by integrating the amount of propene formed atter a stop of the reactant feed. The other parameters of the kinetic model must be estimated simultaneously by fitting the concentration time course under periodic operation. The optimal set of parameters was obtained using the build-in optimisation routine of the Madonna software (Macey and Oster, Berkeley, 1998).

4.3 Factors influencing the time average productivity and selectivity Cycle period For constant space time the average propene and ether productivities are dependent of the cycle period as shown in Fig. 4. For extremely long periods (tp-~Oo), the reactor operates at quasi steady-state, and the productivities correspond to the half of the stationary productivity with Cm,i~OH--2" C~.i~OH. At the contrary when the period is extremely short (tp--}0), the relaxed steady-state is reached. The productivities and selectivities correspond to those obtained at steady state with Cm~P~OH-- C=~o H . The domain of interest lies between these two extremes. The behaviour at optimal cycle length is shown on Fig. 3B. In this case, the half cycle under inert is sufficiently long to reach the maximal transient propene production rate, but, as soon as this rate begins to decrease, the reactant feed is restarted in order to reload the catalyst surface. For shorter periods, the maximal propene production rate is not attained, causing a decrease of the average productivity. Similarly, if the period is too long, the propene concentration falls to low values at the end of the inert phase and the phase with isopropanol feed exceeds the time needed for reloading the catalyst. The optimal length of period is depending of the rate of the the different reaction steps. With increasing reaction rates optimal periods will decrease. The selectivity is also influenced by the length of period. The formation of ether requires the simultaneous presence of isopropanol adsorbed on S1 and on $2 sites, on the contrary to the formation of propene, which is accelerated by vacant $2 sites. It is interesting to note that at maximal reactor productivity the selectivity for propene reaches its maximal value as well. 0.35 ~0.30

•

m

O

Isopropanol Propene 9 Ether ~ . _ _ . ~ _ ~

0.90: 0.85~

E. 0.25 r3 0.20 = o 0.15 o ~" 0.10

[] x

1.000.95 ~

Selectivity

Conversion

~~~---[]

_ o.8o~ N

Q -,=

o

0.4

0.01

0.3tx

Q

IE 0.00

o

160

260 360 460 " Cycle Period [s]

co

,

o

,

,

,

,

loo 200 360460

Cycle period [s]

560 660

Fig 4 Average outlet concentrations (lett), Selectivity and conversion (right) as function of the cycle period. T 200~ P 1.3 bar, split 0.5, Ci~oH~e~ 0.45 mol/m 3 STP, ~ 0.175 s Experimental: points. Model: Lines

245 P,,,i

"ca"200 =

150

O

100 50

_o 9 o

== ~ ~ ~ ~ % =

A

m

E

~ o

=~

0

0.0

012

014

Space time,

016 9[sl

1.80 ='1.75 ~:~ 1.70

9 9 "

"

~

It~~ 1.60

00

012

o14

016

Space time, ~ [s]

Fig. 5 Optimal cycle length and productivity increase under periodic operation (Rta ot~/Rtp-~) Experimental" points. Model" Lines Space time The optimal length of period leading to maximal average productivity is influenced by the space time: With increasing space time, the optimal cycle length becomes longer (Fig. 5A). This is due to the time needed to wipe out the inhibiting species aiier the switch isopropanol/inert. At the same time the obtainable maximal productivity under periodic operation decreases with increasing space time as indicated in Fig. 5B. This result is in line with the observation that the maximal transient propene production increases with decreasing space time and that the maximum is shitted to shorter times (Fig. 2).

5. CONCLUSIONS The catalytic dehydration of isopropanol was studied under transient conditions in a catalytic microreactor. The reaction is characterised by educt inhibition and shows a pronounced "stopeffect". Therefore, the average productivity under forced periodic operation can be considerably higher compared to the maximal productivity obtainable at steady state. For high rates of the sorption processes and surface reactions involved, the optimal cycle time for the forced concentration variations lies in the order of seconds. As microreactors are characterized by low mass storage capacity and narrow residence time distribution, they are particularly suitable for periodic operation at relatively high frequencies. Tis could be demonstrated in the present study. The experimental results could be described satisfactorily on the base of an adequate kinetic scheme in combination with a reactor model. The complete model can now be used to predict the performance of different reactor types and the influence of backmixing or maldistribution in multichannel reactors. ACKNOWLEDGEMENTS The authors are grateful to the Swiss National Science Foundation for financial support.

246 SYMBOLS A Ci

D E ki Ki m~t MS N P

Isopropanol concentration of compound i, mol/m 3 (STP) diisopropylether olefin (propene) rate constant, kg/mols equilibrium constant, various units catalyst amount in the reactor, kg mass spectrometer total number of tanks in the model propene

P Q R RTD Si STP t ta T W Zi

pressure, bar gas flow, cm3/s reaction rate, mol/(skgcat) residence time distribution active site of type i standard conditions (25~ 1 bar) time, s cycle period, s temperature, ~ water concentration of Si, mol/kg

x

space time, s

Greek letters

01, 02, 02w

fraction of occupied sites

LITERATURE Koubek, J., J. Pasek, and V. Ruzicka, Exploitation ofa Nonstationary Kinetic Phenomenon for the Elucidation of Surface Processes in a Catalytic Reaction, in New Horizons in Catalysis. 1980, Elsevier-Kodansha: Amsterdam-Tokyo 853-862. 2. Thullie, J. and A. Renken, Model discrimination for reactions with a stop-effect. Chem. Eng. Sci., 1993.48: 3921-3925. 3. Thullie, J. and A. Renken, Transient behaviour of catalytic reactions with stop-effect. Chem. Eng. Commun., 1990. 96: 193 - 204. 4. Golay, S., O. Wolfrath, R. Doepper, and A. Renken, Model Discrimination for Reactions with Stop-effect, in Dynamics of surface and reaction kinetics in heterogeneous catalysis, G.F. Froment and K.C. Waugh, Editors. 1997, Elsevier: Amsterdam. Studies in Surface Science and Catalysis, 109, 295-304. 5. Golay, S., R. Doepper, and A. Renken, In-situ characterisation of the surface intermediates for the ethanol dehydration reaction over y-alumina under dynamic conditions. Appl. Catalysis A: General, 1998. 172: 97-106. 6. Rouge, A., B. Spoetzl, S. Schenk, K. Gebauer, and A. Renken, Microchannel reactors for fast periodic operation: The catalytic dehydration ofisopropanol. Chem. Eng. Sci., 2000. in press: . 7. Thullie, J. and A. Renken, Forced concentration oscillations for catalytic reactions with stopeffect. Chem. Eng. Sci., 1991.46: 1083 - 1088. 8. Golay, S., R. Doepper, and A. Renken, Reactor performance enhancement under periodic operation for the ethanol dehydration over y-alumina, a reaction with stop-effect. Chem. Eng. Sci., 1999. 54: 4469-4474. 9. Liauw, M.A., M. Baems, R. Broucek, V.B. O, J.M. Commenge, J.P. Corriou, K. Gebauer, H.J. Hefier, O.U. Langer, M. Matlosz, A. Renken, A. Rouge, R. Schenk, N. Steinfeldt, and S. Walter. Periodic Operation in Microchannel Reactors. in IMRET 3. 2000. Frankfurt am Main 224-234. 10. Commenge, J.M., L. Falk, J.P. Corriou, and M. Matlosz. Optimal design for flow uniformity in microchannel Reactors. in IMRET 4. 2000. Atlanta 23-30. 1.


A Dynamic Study of Steam-Methane Reforming by M.A.E1-Bousiffi, Petroleum Research Centre, PO Box 6431 Tripoli-Libya and D.J.Gunn, Process Engineering Systems, PO Box 349, Swansea SA2 7YS U.K. Abstract A microreactor has been operated over cycles of temperature, feed composition, and pressure for the steam-reforming of methane over a nickel-based catalyst. The temperature range was 600~ to 840~ and the pressure range was 1-10 bar. The influence of cyclical temperature changes upon the reactor gas exit compositions was usually reversible, but the introduction of experiments at 800~ increased the catalyst activity for subsequent operation at 700~ for many hours of operation. Introduction The formulations of catalysts developed for the steam reforming of naptha fractions included high nickel content, but catalysts subsequently developed for the commercial exploitation of natural gas were reduced in nickel content and operated at higher temperatures. Commercial reaction conditions now vary in the ranges of 600-900~ and 5-40 bar. A high temperature culminating in a reactor exit temperature of 800-900~ is necessary to overcome the adverse effect of lower temperatures on equilibrium yields, but these conditions are not reflected in all experimental programmes. Akers and Camp(1955) studied the reaction between steam and natural gas within the temperature range of 340-640~ at 1 atmosphere using 3mm diameter pellets of nickel produced by the reduction of nickel oxide supported on kieselguhr. There was little reaction at temperatures less than 600~ and they found first order dependence of the rate of disappearance of methane upon the partial pressure of methane at 640~ They reported that there was no significant variation of catalyst activity with time. Following studies of steam reforming of methane over nickel foil (Bodrov et al, 1965), Bodrov et al, (1967), measured the kinetics of the reaction of methane with water vapour over a commercial catalyst Giap-3, and a catalyst containing ot alumina impregnated with 4% nickel by weight at temperatures between 700~ and 900~ The rate of reaction could be represented as first order with respect to methane, inhibited by adsorption of water and carbon monoxide in a similar manner to their findings on nickel foil. There was evidence of diffusional inhibition due to the size of the catalyst grains that ranged from 1.2-5.3mm. The reaction was started at 700~ after preliminary reduction with hydrogen and the activity gradually dropped for 30-40 hours when a stationary value was reached. Ross and Steel (1973) studied reforming of methane over a coprecipitated nickel-alumina catalyst at temperatures between 500 and 680~ in the pressure range from 1-10 Torr. Their results supported first order dependence of the rate of disappearance of methane upon methane partial pressure together with evidence of water inhibition. Agnelli et al (1987) measured rates of reaction in a 10mm diameter reactor between 642 and 738~

247

248 They reported similar findings to those ofBodrov et al (1967) including the attainment of a steady state after 40 hours and they represented the kinetic data by similar expressions. Xu and Froment (1989) studied the steam reforming of methane over a crushed nickelmagnesium-alumina catalyst containing 15% nickel in a reactor of 10.7mm diameter at temperatures between 500 and 575~ and pressures between 5 and 15 bar. Soliman et al (1992) studied steam reforming over a nickel-calcium aluminate-spinel catalyst in a commercial microreactor 6mm in diameter at temperatures from 475 to 550~ and pressures from 2 to 4 atmospheres. It is a feature of all studies that there was a common difficulty in establishing a steady state for kinetic measurements. Catalyst activity declined sometimes sharply particularly at the outset, with measurements made at a time when the rate of decline was slower and often corrected for the decline in activity. In this paper we present some results from a dynamic study of a commercial nickel-based catalyst using a newly-developed microreactor operating under computer control. The physical size of the reactor was very small and as the flow capacity of the reactor system was high, equilibrium limitations were not significant. The range of operating temperature for which kinetic data were collected was 600-850~ at pressures from 1.5-10 bar.

Experimental equipment and methods The catalyst contained approximately 15% nickel oxide, with calcium oxide and alumina. The reactor consisted of a length of Inconel tube of 3mm bore packed with catalyst crushed to 0.2-0.3mm in diameter. Reactants were delivered to the reactor from pressurised gas cylinders with individual flow controllers for each gas set by output signals from the digital computer. The temperatures in the several parts of the equipment were measured by chromel-alumel thermocouples inserted in the reactor and preheater shells, boiler, feed gases, cooler and other strategic points. The temperatures of the reactor, feed preheater and boiler required for continuous temperature control were connected to set channels of a sixteen channel amplifier and analogue to digital converter. A multiplexer addressed by the computer connected other selected thermocouple outputs to a single channel of the amplifier. Temperature control of the preheater, reactor and boiler was arranged by calculating control actions from thermocouple measurements and passing the control voltages via an eight channel digital to analogue converter to the analogue controllers of high capacity thyristors connected to resistive heaters. Feed gases from cylinders metered by individual controllers, passed through a preheater and were delivered at controlled temperature to the reactor. Steam generation was arranged by feeding demineralised water at a controlled rate to a temperature-controlled boiler and the reactants were mixed at the preheater outlet and fed to the reactor. Outlet gases from the reactor passed into a high capacity cooler and thence into a pressure control module in which the pressure registered by a transducer was used to control the position of a variable-opening solenoid valve in the fluid effluent line. From the outlet control valve, condensate was separated into a receiver for subsequent measurement, and the gases were passed to a bypass and sample valve that allowed sampling and elution of the sample into a chromatograph. In operation methane, carbon monoxide and carbon dioxide were determined from the chromatograms by means of calibrations, and hydrogen was found from the total sample and the methane, carbon

249

monoxide and carbon dioxide samples. The instrument and flow arrangements for the assembly are shown in the diagram.

Inl

PREHEATER

~

ov

~ ~

REACTOR

[~--ER

~ RESERVOIR

_

CH4

PRESSURE CONTROL MODULE

H2 C0/C02

~

valve isolaUonvalve non-returnvalve

solenoid

~

condensate

controlvalve

Diagram of microreactor Aspects of the Operating Dynamics Control settings and values of flowrates, pressures and temperatures were monitored by the computer at chosen intervals. A new set point for temperature was achieved within about 25 minutes to a precision of about +_2~ as shown in Fig.2.

800

/

780 =

760

o

740

~

720

700

mmmmm mm--m-m-m

795~

k

/

7~176162

70o~

mllmlm~m_m_m_n~m_ m

0

I

100

l--I~i--_l_l--I

'

I

200

'

I

300

Time minutes

Fig.2 Temperatureresponse ofmicroreactor

250 The catalyst was activated at 700~ for 18 hours at a steam to hydrogen molar ratio of 7 and pressure of 5 bars. Subsequent steam-methane reforming experiments were carried out for a total of 590 hours using the same catalyst bed. The operating time was determined as the period during which methane was fed to the reactor and excluded preparation when the reactor was brought to operating conditions, and shut down when the methane feed had been stopped and the reactor was cooled in a flow of hydrogen. The first three experiments were carried out at 700~ at a methane flowrate of 202 nml/min, a steam to methane molar ratio of 2.94 and hydrogen to methane molar ratio of 0.5 at a pressure of 5 bars. The experiments produced exit gas compositions that could not be distinguished from equilibrium. However subsequent experiments within a few hours at the same conditions showed that equilibrium had not been attained evidently because there had been a fairly sharp decline from the initial level of catalyst activity. Additional periods of activation were intermingled with the reforming experiments using activation by a mixture of steam and hydrogen as described in accounts of industrial practice. The hydrogen to methane feed ratio was set at 0.5 for the first sixty hours of operation that included eleven separate runs each including about six hours of reaction. After forty hours when the activity of the catalyst continued to decline the catalyst was reconditioned for periods from one to fifteen hours using steam and hydrogen, and hydrogen alone during the period from forty to sixty hours. Equilibrium in the reforming process may be attained by the set of reactions, CH4 + H 2 0 CO + 3H2 (1) C H 4 + 2H20 CO2 + 4H: (2) CO2 + H2 CO + H20 (3) From gas analyses and the feed composition, the approach to equilibrium conversion of methane was calculated as the ratio of methane converted to that converted at equilibrium. The change in activity in the early stages is shown in Fig3.

100

80

9 IE. \ -\.

\\\

\,

0 |

\ .1\ /

"1

20

'

I

20

'

I

40

'

I

60

'

I

80

'

I

100

'

9

I"

I

120

'

I

140

Operating time hours Fig.3 Initial changes in catalytic activity

251 By this time it had become clear that although the period and procedures for shutdown were standard, the condition of the catalyst at startup for subsequent runs varied, being affected possibly by the variable period of shutdown between runs. The uncertainty in steady-state composition could be reduced and a clearer pattern of experimental behaviour could be observed by introducing dynamic experimentation. After startup in a dynamic experiment when a first steady state had been reached, changes in temperature, flowrate or feed composition were introduced by computer program until a second steady state was reached. The experiment was then stopped or continued either by reverting to the first set of conditions, or introducing further changes in operating conditions.

Fig.4 shows the composition response to a change in temperature from the initial setting of 840~ to 797~ followed by a return to 840~ Other operating conditions of feed composition, flowrate and pressure were kept constant. The time allowed for the development of each steady state was significantly greater than the time required to reach a new steady state in temperature only and in this particular example, the activity of the catalyst showed little variation from the initial condition over the first period until the temperature set points were changed. The fall in temperature caused a reduction in the rate of methane oxidation, associated with a significant fall in carbon monoxide content from 18.5% to 15%. The change in carbon dioxide concentration was slight, with an increase from 4% to 5%. This experimental procedure was adopted for subsequent experiments. In this particular example it appears that the step change from 840~ to 797~ and back to 840~ was reversible and steady state compositions have been obtained. The gas compositions leaving the reactor are effectively the same at the same temperature levels and therefore the compositions were attained reversibly within the time required for temperature stabilisation.

252 The reversibility over a temperature cycle was generally obtained for other temperature levels. However for cyclical changes of temperature between 800~ and 700~ a significant difference was observed that was induced at the lower temperature following a period of operation at 800 ~ as shown in Fig.5 illustrating the effect of a temperature change from 706 ~ to 795 ~ and then to 701 ~

Fig5 Compositiondynamics between700~ and 800~

It is clear that the composition changes were not reversible since the effect of the intermediate change in temperature from 706~ to 795~ has increased the catalyst activity for subsequent operation at 701 ~ Following the temperature change from 705~ to 795~ the rate of oxidation of methane gradually increased with the methane content in the exhaust gas falling from 17% to 13%. In subsequent operation at 700~ the catalyst showed significantly enhanced activity compared to the first period at 705~ An inverse change when the temperature was changed from 795~ to 701 ~ and then back to 795~ was reversible in composition. In addition to the catalyst activation at 700~ caused by immediate previous operation at 800~ the experimental record showed that the insertion of experiments at 800~ had a subsequent effect upon experiments at 700~ over many hours of operation. Some change was induced in a matter of an hour or so, but some influence of this change was retained during subsequent operation at 700~ diminishing over a period of some twenty hours. Furthermore this was a clear effect of activation, and at this stage of the experiments there was no evidence of the catalyst deterioration reported in the literature by workers who experimented only at lower temperatures. The experimental record of standard experiments at 600~ 700~ and 800~ expressed as steady state equilibrium conversion of methane is shown to 500 hours in Fig.6. Other experimental conditions that were examined during this period are not shown in this figure. The different temperatures that were examined were chosen to cover the range of industrial interest.

253

There is a clear increase of catalytic activity at 700~ once experiments at 800~ were included. The insertion of four runs of about 30 hours in duration at 800~ spread over the period from 160 hours to 260 hours increased conversion at 700~ from 30% of equilibrium conversion of methane to approximately 40%. An increase in the frequency of runs at 800~ over the period 260 to 340 hours was associated with a further increase in conversion of methane at 700~ to 65% of equilibrium. Operation at 800~ was discontinued at 450 hours and over the next 40 hours conversion fell to 30% of equilibrium at 700~ the same conversion as was found over the period 70-130 hours. Evidently periods of reactor operation at 800~ and 750~ increased catalyst activity for subsequent operation at 700~ with the level of enhancement dependent upon the frequency of operation at 750-800~ The increase in activity induced in this manner decayed with time falling to the unactivated condition some 30-50 hours after activation. Increasing the frequency of operation at higher temperature reduced the rate of decline of activity at 700~ leading to increased activity at the lower temperature as shown in Fig.5. However the temperature activation was not sufficient to restore the activity found in the earliest operations and it appears that the initial decline was irreversible. Although not as marked as the increase in catalytic activity at 700~ shown over the period to 500 hours, catalytic activity at 800~ approached an apparent asymptote attained at about 500 hours. There was no indication in the dynamic records that operation at 700~ increased catalytic activity at 800~ Rather it appears that catalytic activity at 800~ was inhibited by earlier operation at 700~ and the asymptotic activity at 800~ was only obtained by increased frequency of operation at that temperature. Particularly for experimental measurements at 700~ it is apparent that the application of this kinetic study to industrial reactors depends upon the extent to which the temperature history of the industrial catalyst can be found in the kinetic studies described here. If catalyst in the industrial reactor at 700~ has not been exposed to earlier operation at a

254 higher temperature, then kinetic information will have to be taken from operating periods free from activation due to previous operation at a higher temperature. The temperature distribution along the length of a conventional primary reformer increases from about 600~ at the entrance to 800-900 ~ at the reactor exit and enhancement of kinetics in the industrial reactor at 700 ~ by earlier operation at 800 ~ does not occur; neither are the kinetics at 800-900 ~ diminished by preceding operation at 700 ~ Thus kinetic measurements made in the experimental reactor have to be interpreted to remove the effect of temperature enhancement and diminution before possible application to the industrial reactor. If the operation of the industrial reactor is cyclical, the kinetic studies presented here may provide kinetic information by matching the cycles of operation between industrial and laboratory scale. It is possible that some improvement in reactor productivity may be achieved by deliberately employing the principle of temperature activation but the reactor operation will be much more complex.

References

Agnelli, M.E.,DeMicheli, M.C., and Ponzi,E.N., (1987),"Catalytic deactivation of methane reforming catalysts",Ind.Eng.Chem.Research, 26,1704-1713. Akers, W.W., and Camp, D.P.,(1955), "Kinetics of the methane-steam reaction", A.I. Ch.E. Jnl., 1,471-475. Bodrov, I.M.,Apel'baum, L.O., and Temkin, M.I., (1965), "Kinetics of the reaction of methane with steam on the surface of nickel", Kinet. Catal., 5,696-705. Bodrov, I.M.,Apel'baum, L.O., and Temkin, M.I., (1967), "Kinetics of the reaction of methane with water vapour catalysed by nickel on a porous cartier", Kinet. Catal., 8, 821-828. Ross, J.R.H.,and Steel,M. C.F., (1973), "Mechanism of the steam reforming of methane over a coprecipitated nickel-alumina catalyst", J.Chem.Soc.Faraday Trans., 1,10. Rostrup-Nielsen, J., (1984), "Activity of nickel catalysts for steam reforming of hydrocarbons", J. Catalysis, 31,173-199. Soliman, M.A.,Adris, A.M.,A1-Ubaid, A.S.,and E1-Nashaie,S.S,(1992), "Intrinsic kinetics of nickel/calcium aluminate catalyst for methane steam reforming",J.Chem.Tech. Biotechnol., 55,131-138. Twigg, V.M.,(1989), "Catalyst Handbook", 2nd ed. ,Wolf Publishing Co.,London. Xu,J., and Froment, G.F.,(1989), "Methane steam reforming, methanation and water-gas shift: 1.Intrinsic kinetics",A.I.Ch.E.Jnl.,35,88-96.

Studies in Surface Scienceand Catalysis 133 G.F. Fromentand K.C. Waugh (Editors) cc32001 ElsevierScienceB.V. All rights reserved.

255

Model-based optimization of the periodic operation of the Fischer-Tropsch synthesis* R.M. de Deugd a+, S.M. Ypma a'b#, F. Kapteijn a, F.M. Verheijen b

Meeuseb,J.A. Moulijn a, P.J.T.

aSection Industrial Catalysis, Department of Chemical Technology, Delft University of Technology, Julianalaan 136, 2628 BL Delft, The Netherlands bprocess Systems Engineering Section, Department of Chemical Technology, Delft University of Technology, Julianalaan 136, 2628 BL Delft, The Netherlands The feasibility of increasing the selectivity of the Fischer-Tropsch synthesis by periodic operation is investigated. The process is modeled in a dynamic form using a CSTR reactor. The dynamic behavior of the model corresponds with in literature reported experimental results. The simulation results show a 20% increase in selectivity to Diesel range products compared to the best steady-state solution using a blockprofile. The profile has a cycle time of 1.104 seconds and consists of a base composition of almost pure carbon monoxide and a pulse of 95% hydrogen and 5% carbon monoxide during 10.8% of the cycle time. 1

INTRODUCTION

The Fischer-Tropsch synthesis is a process to convert synthesis gas (a mixture of carbon monoxide and hydrogen) to hydrocarbons that can be used as for instance transportation fuels. In the process all (straight chain) hydrocarbons from methane to heavy waxes are produced. In general this product distribution can be described by an Anderson-Schulz-Flory distribution based on a constant chain growth probability. As a consequence the selectivity towards diesel production is limited. When the diesel fraction is defined as C10 till C20, the maximum fraction of diesel that can be obtained is 39.4%, reached at a chain growth probability of 0.87. Raising this selectivity can be done in several in ways. One option is producing heavier products by using a catalyst with a high chain growth probability. The long products can be cracked subsequently by a hydrocracking unit to the desired product range. This method is being used on commercial scale in the SMDS process [ 1]. Periodic operation of the Fischer-Tropsch synthesis may be another option. Periodically changing one or more variables such as feed composition or feed rates could lead to

* This work is part of the Delft InterdisciplinaryResearch Centre "Mastering the Molecules in Manufacturing" + Correspondingauthor: [email protected] Present address: Twister Supersonic Gas Solutions, Volmerlaan 8, 2280 AB Rijswijk, The Netherlands

256

enhanced selectivity towards certain products or product ranges. In literature some indications can be found that cyclic operation of the Fischer-Tropsch synthesis may lead to better yields. Adesina et al. [2] show increased selectivities to certain, low molecular products, while Dautzenberg et al. [3] report a narrower product distribution in a periodically operated system. Most studies, however, only present experimental results [2] or use the data to derive kinetic constants [3]. No examples are known of studies focussing on maximizing the yield to certain products or product ranges. In this contribution it is explored what theoretically is the maximum Diesel yield defined as the C 10-C20 product range by simulation and optimization of the feed composition profile, i.e. by imposing a cyclic concentration variation. 2

APPROACH

In this simulation study a blockprofile is imposed on the feed composition of a FischerTropsch reactor. Subsequently, the blockprofile is optimized to reach maximum selectivity to Diesel. To simulate the dynamic behavior of the Fischer-Tropsch synthesis a reactor description and a set of detailed kinetic equations and constants are needed. In literature much is known about Fischer-Tropsch reactors (e.g. [1]), but the detailed kinetics is lacking. For calculation of conversions or selectivities towards certain (light) products or fractions rather simple reaction kinetics is enough, but the description of the reaction rates of both reactants and products requires more detailed information about the reaction mechanism and the constants in the rate equations. 2.1 Reactor model The effect of periodic operation on the performance of the catalyst can be examined most clearly if the applied concentration change is immediately and fully felt by the catalyst. This situation is reached by defining a small amount of catalyst in a well-mixed reactor (CSTR) with large flow rates resulting in high space velocities. The system is assumed to be operated in a regime where mass transfer limitations can be neglected. Furthermore, the system is assumed to be isotherm. Although the Fischer-Tropsch synthesis is very exothermic the CSTR character and the low conversions due to the high space velocity ensure the validity of this assumption. Moreover, the low conversion combined with temperature and pressure validate the assumption of single gas phase conditions in the reactor. The pressure in the reactor is kept constant at 10 bar. Modeling with constant pressure in the reactor has as a consequence that the mass flow rates into the reactor will vary during the cycle period as the changing composition causes changing consumption rates. 2.2 Kinetic model The kinetic model is build upon generally accepted concepts. The mechanism of the FischerTropsch synthesis has been subject of many investigations, but is never fully agreed upon.

257 General consensus is reached on the carbide mechanism [4,5], which is used here to describe the formation of the monomer and the paraffinic and olefinic products. The main problem is that the constants in the reaction scheme are not available. Literature provides kinetic constants for simplified reaction equations, but these equations are not detailed enough. It was decided to calculate the necessary constants from steady-state experimental data as available in several literature sources (e.g. [6]). A steady-state model has been set up based on the same kinetics and reactor model as used in the dynamic model. Some assumptions were made regarding the activity of the catalyst and the product distribution. The chain growth probability is supposed to be 0.92 and the olefin content is presumed to be 10%. The activity of the catalyst was taken from Post et al. [6] and is characterized by the rate constant for CO consumption; 58"10 -4 m3/(m 3cat*s). Stating the feed composition (H2/CO =2) and the conversion level (60%) fixes the feed rates of hydrogen and carbon monoxide. Subsequently, the reaction scheme is fitted between the feed and consumption rates of the reactants and the formation rates of products. Some additional presumptions were added to reduce the number of unknowns in the system. Most reactions are considered as reversible reactions. Only the reaction of surface oxygen to water (in two steps), the propagation reaction and the termination reaction towards paraffins are modeled as irreversible. All except one of the reversible reactions are considered to be near equilibrium. Only the rate determining step, the dissociative adsorption of hydrogen [7], is modeled to be far from equilibrium. Finally, some estimates have been made for the surface coverage of the catalyst. As the basis for these estimates several literature sources have been used. Van der Laan [8] concludes based on statistical analysis that a site balance consisting of the surface coverage of carbon, water and vacant sites fits better experimental data than a balance without vacant sites. Mims et al. [9] state that the surface coverage by intermediates is very low. Adesina et al. [10] mention low coverage by hydrogen and high coverage by carbidic species. Van Neer [11 ] concludes that a high surface coverage is needed to find resonance phenomena in catalytic reactions anyway. As effects of cyclic operation are experimentally reported (e.g. [2,3,10]) the fraction vacant sites on the catalyst must be limited. In this work a high coverage by carbide and water is supposed (0.4 and 0.2 respectively). Vacant sites are supposed to be present with a concentration of 0.1. The rest of the surface is assumed to be covered with not and partially hydrogenated oxygen (total coverage of 0.08), partially hydrogenated carbide (totally 0.18), hydrogen (0.02) and products (0.02).

2.3 Model validity As no complete data sets with both reactor and the kinetic data are known in open literature validating the model is difficult. Therefore, the steady state behavior of the model used has been determined and compared to the behavior as reported in literature [ 12]. Raising the hydrogen to carbon monoxide feed ratio reduces the chain length, raises the methane selectivity and lowers the olefin selectivity. All these three characteristics are in agreement with the findings of Schulz [12]. Higher pressure in the system leads to a higher

258

chain growth probability both in our calculations as in literature [12]. Decreasing the space velocity results both in the calculations as in the literature [12] to lower olefin selectivity. Introduction of water in the feed lowers the conversion as may be expected for a reaction product. The steady-state behavior of the catalyst in this system is in agreement with literature reports. Regarding the dynamic behavior is less certainty, as no validation to experimental data is possible.

2.4 Optimization strategy The basic idea of periodic operated Fischer-Tropsch synthesis is periodically removing the growing hydrocarbon chains from the catalyst surface by pulsing a hydrogen-rich synthesis gas mixture over the catalyst. In the time span between the pulses a more CO-rich syngas mixture should facilitate chain growth. By optimizing the concentration profile the fraction products with the desired chain length should be increased. In this study a blockprofile is used to change the feed composition. The block profile is defined by four parameters: cycle time, split ratio, base ratio and pulse ratio. The split ratio is defined as the fraction of the cycle time in which the catalyst is exposed to a hydrogen-rich synthesis gas mixture. The base ratio is the H2/CO ratio of COrich synthesis gas mixture and the pulse ratio that of the H2-rich pulse mixture. The products are divided into three categories; lights (C1-C9), Diesel (C10-C20) and heavies (C20+). It can be calculated from the ASF distribution that the steady state selectivity towards Diesel never can exceed 39.4 % regardless the chain growth probability. Cyclic operation should result in higher selectivity than this steady-state value. As the goalfunction for the Diesel selectivity optimization the following relation has been defined: ~

goalfunction=

~ (n~=lo(Fout...... paraffins-t-Fout........lefins cycle

~(Fo,, HC)dt

(1)

......

cycle

The optimizations have been performed in two ways. First, gOPT, the optimization tool in the software package gPROMS, has been used to find the maximal Diesel selectivity. Furthermore, gridsearches have been executed to gain more insight in the influence of the parameters of the block profile on the Diesel selectivity. 3

SIMULATION RESULTS

3.1 Optimized blockprofiles Using the gOPT software in the gPROMS package we obtained a blockprofile leading to a Diesel selectivity of 47.1%. This Diesel fraction is about 20% higher than the best steadystate solution. Table 1 and Figure 1 summarize the characteristics of the profile used.

259

~

15

lO

0 230000

. 232000

.

. 234000

.

236000

238000

Time (s)

Figure 1 Optimal block profile (cycle time = 1.104 seconds, base ratio = 0.001, pulse ratio = 19, split ratio = 0.108

Table 1 Characteristics blockprofile Cycletime Split ratio 1"104 s 0.108

Base ratio 0.001

pulse ratio 19

Goalfunction 0.471

The average H2/CO feed ratio, 0.19, is far away from the stoichiometric ratio of slightly above two. Consequently, the conversion of carbon monoxide is low, 5.1%. The hydrogen conversion is at a higher level, 53%. The response of the product fractions to the applied feed cycle is shown in Figure 2a. The Diesel fraction never exceeds 36 wt%, but this selectivity maximum coincides with a maximum in the hydrocarbon production rate as shown in Figure 2b. Upon switching to the H2-rich mixture the feed mass flow rate decreases strongly due to the low density of hydrogen. After switching back to the CO-rich synthesis gas the mass flow shows a sharp peak because of the uptake of carbon monoxide.

0.07 I 0.06 t

~l, .... 1 .... 1-~

~ o.3

~oo2}~~ A

1~ o.2 0.1

o 230000

o.ool

....

232000

234000

236000

238000

240000

. . . . . . Time (s) I --lights - - Diesel - - hea . . . . I

Figure 2a Response of the product fractions to the optimal block profile

~176

230000

,

232000

~ ,

234000

lirne (s)

O.0005 ~

,

,

236000

238000

0

240000

r - mass flow reactants ,n ~ mass flow hydrocarbons Out I

Figure 2b mass flow rates of the reactants and the product fractions

260

Figure 3 Product distribution of the optimal simulation result compared to the optimal ASF distribution

Figure 4 Gridsearch with varying split ratio and pulse ratio. Base ratio - 0.001, cycle t i m e - 1.104 seconds

The mass flowrates show large changes over the cycle period. However, the changes of the volumetric flows are much smaller. The volumetric output flow and reactor pressure are maintained constant, leading to volumetric inlet rates varying up to 20%. Figure 3 shows a comparison between the product distribution found in this simulation study and the optimal ASF distribution. The maximum in the distribution shifts from C7 to C12 within the desired product range. 3.2 Gridsearches

To gain more understanding in the factors that influence the selectivity some gridsearches have been performed. In these searches one or more parameters of the blockprofile are kept constant while the others are varied. This approach yields, depending on the amount of varied parameters lines, surfaces or spaces with for each point a value for the goalfunction. In this study searches with one (linesearches) and two (gridsearches) parameters are done in the which the linesearches can be regarded as part of a gridsearch. Figure 4 shows an example of a gridsearch in which the split ratio and the pulse ratio are varied. The optimal solution presented earlier lies outside this picture on the dark ridge of high selectivities. The whole central dark area which continues up to pulse ratios of about 47 shows higher selectivity than the steady state solution. The decrease in the goalfunction at lower split ratios can be explained by short duration of the hydrogen pulse.

3.3 Alternative cases

In addition to this base case the sensitivity of the dynamic behavior to the assumption of the relative speed of the reversible hydrogen adsorption reaction is investigated. The base assumption is that the dissociative adsorption of hydrogen is 500 times faster than the backward reaction of adsorbed hydrogen to gaseous hydrogen. In the alternative cases the ratio between the forward rate and the backward rate is 100 and 10. In Table 2 the results of these two cases are compared to the basecase.

261 Table 2 Influence on block profile and the Cycletime in 104 s Basecase 1 rx/r.x= 100 1.25 rx/r_x= 10 1.25

goalfunction Split ratio 0.108 0.100 0.100

of speed of hydrogen equilibrium Base ratio Pulse ratio Goalfunction 0.001 19 0.471 0.001 11 0.467 0.01 10 0.482

The position of the equilibrium has an effect on the simulation results, but the goalfunction still shows a higher selectivity towards the Diesel range components. The longer cycle time can be attributed to the slower forward reaction. 3.4 Discussion

Literature (e.g. [3, 13]) shows increases in the selectivity to certain products and fractions in the Fischer-Tropsch synthesis. These studies are limited to products below C10. In this contribution the effect of periodic operation on a higher product fraction is explored. This difference makes comparing of the presented results difficult, but the trends are similar. The optimization results show a significant raise (20%) in the selectivity towards the desired Diesel range products. Still a large fraction of the products consists of low and high molecular products, but the increase underlines the potential of the approach. It is striking that the optimal composition during the pulse is almost pure hydrogen. However, the gridsearches show that influence of the pulse ratio on the goalfunction is rather flat around the optimal split ratio. Using synthesis gas compositions closer to the composition in industrial operation leads to loss of selectivity, but the selectivity towards Diesel stays above the best steady-state value The time scale of the cycle time is different from other results, as this time scale is dependent on the product range. Dautzenberg et al. [3] report time scales between 102 and 103 seconds for products shorter than C10. Adesina et al. discusses experiments focussing on products between C1 and C6 with cycle times up to 80 minutes (5000 seconds). Increased selectivities are observed up to 40 minutes (2500 seconds). The presented results show an optimal cycle time of (1-1.25).104 seconds. Regarding the longer cycle times for higher products [3] the calculated optimal cycle time is in agreement with the earlier reports. The alternative cases show that the influence of the position of the hydrogen adsorption equilibrium is not crucial. The goalfunction shows some variation for the various cases, but still the Diesel selectivity is far above the best steady-state value. A drawback of the present approach is the nature of the model. The kinetic model shows a steady-state behavior in agreement with literature reports, but the dynamic behavior can hardly be related to other reports. The time scale of the optimized cycle period can be related to literature reports, but the improved selectivities towards C10 products is, as far as we know, not reported earlier. The reactor model is very simplified to be able to focus on the behavior of catalyst. The productivity of the reactor system suffered from this approach (5.1% CO conversion). Furthermore, the reactor behavior can also have a serious impact on the overall result. For instance, using a plug flow reactor may even yield better results as the

262 concentration changes move through the reactor as a front rather than mixed with the contents of the reactor. In this study using a small reactor, yielding a low conversion, prevents this mixing problem. The implications of periodic operation on the level of the total process will be discussed in an upcoming publication [ 14]. 4

CONCLUSIONS

This study explores the potential of periodic operation for the Fischer-Tropsch synthesis aiming at Diesel range products. The approach followed is modeling the process in a dynamic form using a simple CSTR reactor configuration. The kinetic scheme is based on steady-state data reported in literature. The steady-state behavior is in agreement with experimental observations reported earlier by various research groups. The results show a significant increase (20%) of the selectivity. The optimal cycle time is around 104 seconds. The cycle consists of a blockprofile with a base composition of almost pure carbon monoxide and a pulse of 95% hydrogen and 5% carbon monoxide. The pulse time is 10.8% of the total cycle time. The carbon monoxide is, due to the assumptions made in the model, low (5.1%). The hydrogen conversion is 53%. Although the model contains many assumptions the results can be regarded promising. The steady-state behavior of the model is in agreement with literature reports. The reported cycle time is also in line with earlier reported work in literature. The assumption made about the equilibrium constant of the rate determining step does not influence the results dramatically. Experimental verification of the presented approach will be the future challenge. REFERENCES [ 1] [2] [3] [4] [5] [6]

S.T. Sie, Rev. Chem. Eng., 14 (1998) 109. A.A. Adesina, R.R. Hudgins and P.L. Silveston, Catal. Today, 25 (1995) 127. F.M. Dautzenberg, J.N. Helle, R.A. van Santen and H. Verbeek, J. Catal., 50 (1977) 8. M.E. Dry, Appl. Catal. A, 138 (1996) 319. G.P. van der Laan, and A.A.C.M. Beenackers, Catal. Rev. Sci. Eng., 41 (1999) 255. M.F.M. Post, A.C. van 't Hoog, J.K. Minderhoud and S.T. Sie, AIChE J., 35 (1989) 1107. [7] H. Chen and A.A. Adesina, J. Chem. Tech. Biotechnol., 60 (1994) 103. [8] G.P. van der Laan, Ph.D. thesis, University of Groningen, (1999). [9] C.A. Mims, J.J. Krajewski, K.D. Rose and M.T. Melchior, Catal. Lett., 7 (1990) 119. [ 10] A.A. Adesina, R.R. Hudgins and P.L. Silveston, J. Chem. Tech. Biotechnol. 50 (1991) 535. [11] F.J.R. van Neer, Ph.D. thesis, University of Amsterdam, (1999). [12]H. Schulz, Pure & Appl. Chem. 51(1979) 2225. [13]A.A. Adesina, R.R. Hudgins, P.L. Silveston, Catal. Today, 25 (1995) 127. [14] F.M. Meeuse, R.M. de Deugd, F. Kapteijn, P.J.T. Verheijen and S.M. Ypma, to be presented at ESCAPE 11, Kolding, 27-31 May 2001.

Studies in Surface Scienceand Catalysis 133 G.F. Fromentand K.C. Waugh(Editors) Publishedby ElsevierScienceB.V., 2001

263

D i r e c t d e t e r m i n a t i o n of p e r i o d i c s t a t e s of cyclically o p e r a t e d p a c k e d b e d reactors T.L. van Noorden ab, S.M. Verduyn LuneP and A. Bliekb ~Department of mathematics, Vrije Universiteit, Amsterdam, De Boelelaan 1081A, 1081 HV Amsterdam, The Netherlands bDepartment of chemical engineering, University of Amsterdam, Nieuwe Achtergracht 166, 1018 WV Amsterdam, The Netherlands The convergence behaviour of cyclically operated packed bed reactors is studied. Rather than converging to a steady state, a cyclic process approaches a periodic state. The determination of this periodic state by dynamic simulation techniques is computationally inefficient. Several convergence accelerating techniques are studied and applied to four pressure swing adsorption systems. A comparison is made with respect to CPU time. 1. I N T R O D U C T I O N Many processes in the chemical industry are operated under periodically varying, dynamic conditions. Examples of such processes are temperature swing and pressure swing separation, the simulated moving bed reactor-separator, the reverse flow reactor, and the recently developed pressure swing reactor. The basis for both design and optimization of chemical processes are their limiting states: they are the state a process will be operated in for most of the time, after an initial transient start-up phase. The typical limiting state of a cyclic process is a so-called cycle invariant or periodic state. A natural and obvious way to calculate a periodic state of a cyclic process is to simply simulate the dynamic behaviour of the process. The simulation will approach, as does the physical system, a periodic state. The disadvantage of this strategy is that the initial transient phase may be very long, in particular in the presence of large capacity terms and slow kinetic terms. All example reactor types given above are based on packed beds, where a column is filled with a solid sorbent and/or catalyst. Typically the solids inventory is very high and so is the buffer capacity in terms of the adsorption capacity or heat capacity. For such systems the dynamic simulation of the process may need tens of thousands of cycles in order to converge to a periodic state. This long transient phase enforces the use of over-simplified models to be able to do any computations within a reasonable CPU time. When more computationally efficient methods are available, this will also allow for the development of more realistic simulation models for cyclic processes. In the last decade a number of convergence acceleration methods have been discussed in literature (see, for example, [1] and [2]). Most of these methods are designed for specific systems and it is not a priori clear which method to use

264 for a given system. What is needed is a generic approach for general cyclic processes. The philosophy of the present work is to continue the approach of [3], which is the first systematic comparison between different methods, and of which we believe it is the first step towards a more general understanding of periodic states and methods to compute them. Our goal is to develop a generic strategy for the analysis and optimization of cyclic processes. The efficient determination of periodic states of these processes is a crucial part in such a strategy. We use four different methods to calculate periodic states of four pressure swing adsorption test systems. Two of these systems exhibit linear behaviour. Three of these four methods, i.e. the dynamic simulation, Newton's method and Broyden's method, have previously been applied to compute periodic states of cyclic chemical processes. The fourth method, which is the Newton-Picard method [4], is more recent and new in chemical engineering literature. The focus of the present work is on convergence acceleration methods; the pressure swing adsorption systems, which are all taken from literature, are mainly used to illustrate these methods. The example systems are all well documented and thoroughly studied. Here they are used to benchmark and, more importantly, to compare the acceleration methods.

2. T H E P R O C E S S E S In this section the four systems that are used in the comparison of the numerical methods are described. The systems are called A,B, C and D. The first two are modeled with linear equations and the last two with nonlinear equations. S y s t e m A: System A is used to adsorb water from an air-water mixture, using alumina as the sorbent material. The adsorption is assumed to be 100% selective towards H20. The used modeling assumptions remove all the nonlinear terms from the model. The model and the parameters for this systems are taken from [3]. S y s t e m B: Systems B is used to separate CO2 from He gas using silica gel. The CO2 is the sole component to adsorb. The modeling assumptions for this system are exactly the same as for System A, resulting in the same linear model. The parameters are taken from [3]. The adsorption equilibrium constant for System B is much smaller than for System A. This means that the buffer capacity for System A is larger than for System B. In addition, the adsorption rate constant for System A is two orders smaller than for System B. It is thus expected that a dynamic simulation of System B will approach a periodic state much faster than for System A. S y s t e m C: System C is used to separate CO2 from N2 using 5A zeolite. The CO2 is the only component to adsorb. The composition of the feed gas is 10% CO2 and 90% N2. The parameters and model are taken from [5]. In contrast to System A and B, this system exhibits nonlinear behaviour. S y s t e m D: System D is used to remove H2S from natural gas using 5A zeolite. The model natural gas consists of 1000 ppm H2S, 5% CO2 and 94.9% CH4. This system has also been studied in [6] under a slightly different operating scheme. The physical and chemical parameters are all taken from [6].

265

3. D E T E R M I N A T I O N

OF PERIODIC

STATES

A common approach is to model a packed bed reactor as a system of partial differential equations and in order to investigate the behaviour of such a reactor numerically, these model equations have to be discretized. For all four methods discussed in this section, first the space variable of the dimensionless equations is discretized. This leads to a large system of N (where N depends on the space discretization) time periodic ordinary differential equations. This system can be written as x'(t) -- f(x, t ) , where f(-, t + T) = f(-, t ) ,

(1)

where x ( t ) is a N dimensional vector, and f(-, t) a function from the N dimensional vectors (denoted by R N) to the N dimensional vectors. A periodic state of the chemical process is now equivalent to a T-periodic solution x (i.e. a solution with x(O) = x ( T ) ) of this system. The map F that assigns to the initial data at time zero, x(0), the solution after one cycle, at time t -- T, x(T), is called the Poincar4 or period map: one may write F ( x o ) -- x(T),

(2)

Where x ( t ) is the solution of equation (1) with initial condition x(0) - x0. So, evaluating the map F is equivalent to simulating one cycle of the process starting with a given initial condition Xo. A T-periodic solution is a solution of the fixed point equation x - F ( x ) . Methods to solve this equation are called shooting methods. The four methods for obtaining a periodic state of a cyclic chemical system differ in the way the fixed point equation is solved.

3.1. D y n a m i c Simulation The dynamic simulation "method" consists of directly integrating the model equations for a number of cycles starting from an initial condition. In this way the dynamics of the process is simulated. However, the process may demonstrate a very slow convergence to the periodic state. In this case a large number of "redundant" cycles of the process have to be simulated before a periodic state is reached. The dynamic simulation can also be viewed as the iteration of the period map F: Xi+l - F ( x O , where F is defined by (2) and xi - x ( i T ) is the solution of the systems after i cycles. 3.2. N e w t o n ' s M e t h o d The dynamic simulation approximates fixed points of the map F by iterating the map itself. If we define the map G = F - I , then fixed points of F become zeros of G. Newton's method generates approximations of a zero of G using the iteration scheme X i + l -~ X i -

av(xi)-lG(x,)

,

where J c ( x i ) -1 is the inverse of the Jacobian of G at xi. Note that each new iteration requires the computation of Je at the previous iteration. The evaluation of Jv requires the computation of the dependencies of G on each of the N variables introduced by the discretization of the space dimension. Using first order finite differences to compute these dependencies results in simulating N + 1 cycles with different initial conditions.

266

3.3. B r o y d e n ' s M e t h o d Rather that computing the full Jacob,an of G, Broyden's method uses an approximation of Jc which is updated each iteration. To be more precise, Broyden's method produces approximations to a zero of G using the following iteration scheme Xi+l :

Xi +

HiG(xi) ,

with Hi iterative approximations to - J a ( x i ) -1 defined by

(p, + U,g,)pTU, Hi+ l - H i -

pT Hig i

,

where g, = G(x,+l) - G(x~) and p, = X,+l - x,. The only information that is used in updating the Jacob,an is the function value of G at the new iteration. The method thus uses only one cycle simulation in each iteration. This is a large advantage in comparison to Newton's method, which uses N + 1 cycle simulations each iteration. As the initial approximation of the inverse of the Jc we use Ho = I. It can be shown that for linear problems, Broyden's method converges in at most 2N iterations. 3.4. T h e N e w t o n - P i c a r d M e t h o d The philosophy underlying the Newton-Picard method [4] is to combine the good properties of both Newton's method and the dynamic simulation. The procedure followed in this respect is described below. The Newton-Picard method, like Broyden's method, is a quasi-Newton method. This means that it produces approximations to a fixed point of F using the iteration scheme Xi+l "-- X i "nt- A X i where Ax, is some approximation of Axi ~ - ( J ( x , ) - I ) - I ( F ( x , ) - x,),

where J(xi) denotes the Jacob,an of F at xi. Let us assume the eigenvalues of J(xi) are known. They can be ordered by decreasing modulus

I~1 >_ I#=1 _> ... ___ I,~1 > P > I#p+ll > ---> I~NI, where p is a free parameter of the method. Let U be the subspace of the (discretized) state space R N spanned by the (generalized) eigenvectors of J(xi) corresponding to the eigenvalues #,, i = 1,...,p , let Vp E R Nxp be a basis of U (the columns of Vp are orthogonal and span U) and let Vq E RNx(N-p) be a basis of the orthogonal complement of U (this subspace is denoted by U J-). Now the correction Axi can be decomposed into a component in U and a component in U j-. For this decomposition we construct the orthogonal projectors e := VpVf (where VpT denotes the transpose of the matrix Vp) and Q := VqVqT = I - P. The correction can now be written as Ax, = P A x , + Q A x i = vpWTAx, + v q V T A x , -- VpAp + VqAq,

where the notation AI~ := V f A x , approximate solution of (J(x,) - I ) A x , = - ( F ( x , ) - xi).

(3)

and Aq := V T A x i is used. We want A x i to be an

(4)

267

Now one may substitute (3) in (4) and multiply the left and right hand side of (4) by (Vp Vq) T. The result is

0

V q T ( j ( x , ) - I)Vq

Aq

= --

"v~T(F(xi) -- xi)

"

The last row of this equation does not contain an expression in Ap. Therefore this row (that contains N - p equations) can be solved with the Picard iteration scheme Aq,+l - v T j(xi)VqAr=l, + VT (F(x,) - x,).

(5)

The number of performed Picard iterations is denoted by I. Because of the way the two subspaces U and U • are constructed, this scheme will converge much faster than the direct substitution method applied to the original system of equations. We refer to [4] for more details about the convergence analysis. If the approximation of the solution of (5) is again denoted by Aq then the equation for Ap becomes

vpT j(xi) VqArt 4- vpT ( j ( x , ) - I) VpAp = --vpT (F(x,) - x,). Solving this equation only requires the evaluation of F (i.e. a cycle simulation) p times. In our implementation for each iteration of the Newton-Picard method we need 2p 4- 4 cycle simulations (and if [G(x~)[ is small enough only 4). For more details about our implementation of the Newton-Picard method, we refer to [7]. If p is small compared to N, one iteration of the Newton-Picard method will be much more efficient than an iteration of Newton's method. 4. R E S U L T S 4.1. S y s t e m A and B After the spatial discretization for System A and B a system of 512 ordinary differential equations is obtained. Hence N - 512 for the first two systems. In Fig. 1, left graph, the error ~ (:= IF(x~) - x~)l) versus the CPU time in seconds is given for the dynamic simulation of System A and B. We see that System A converges very slowly (769 cycle simulations, or iterations of F) to a periodic state, whereas System B converges very fast (31 iterations of F). In Fig. 1, middle graph, the error r versus the CPU times for Newton's method is shown. On comparing this graph with the left graph of Fig. 1, it is seen that Newton's method accelerates the convergence to a fixed point for System A, but not for System B. For System A, Broyden's method converges in 99 iterations, and for System B in 14 iterations. Thus, for both System A and B, Broyden's method converges much faster than the dynamic simulation and also in far less than N + 1 steps, which is the minimum number needed for Newton's method. Recall that Broyden's method has a theoretical bound of 2N iterations for linear systems (see Section 3.3). For the two present systems this clearly is an overly conservative bound. Since in the first iteration of the Newton-Picard method the Jacobian is computed, this method will need at least as much time as Newton's method, which, because of the linearity of the systems, needs to compute the Jacobian only once. Therefore the NewtonPicard method is not optimal for the linear Systems A and B and it was not applied to these systems.

268

le+10 |

1

le-10 le-201 0

Dynamic Simulation

o i

e-ol

Newton's Method

....

+

I

. . . . . I 10000 20000 30000 le-20 "1"7~)00 ' 18000 ' 18i00 ' CPU time in seconds CPU time in seconds

,_ 1

Br~

Method

le-10

le-20L

0

2-~

4-~

-.J

6oo0

CPU time in seconds

Figure 1. The error versus the CPU time in seconds for the different methods for System A (O, or thick black line) and System B (+).

4.2. System C After the spatial discretization a system of 120 ordinary differential equations is obtained. Hence N -- 120. In the left graph of Fig. 2, the deviation from the periodic state is plotted against the CPU time for the four different methods. One may observe that the convergence rate of the iterations of the dynamic simulation is constant. This resembles very much the convergence of a linear system. The fact that Newton's method needs more than a few iterations to converge, illustrates that System C is nonlinear. Only the first two Newton iterations are shown in the left graph of Fig. 2, for it takes 7 iterations to converge to a periodic state. Newton's method is thus very inefficient for this system. The optimal choice of the parameter p in the Newton-Picard method depends on the problem under consideration. There are no good rules for making the choice, except a rule of thumb gathered from experience. In [4] it is reported that values of p around 0.5 result in good performance of the method. Therefore we used p - 0.5 for the Newton-Picard computations shown in the figure. Note that the first iteration of the Newton-Picard method needs much more time than the next iterations. Figure 2 shows that Broyden's methods does not have this lengthy first iteration. From this one may identify the major drawback of the Newton-Picard method: it needs to calculate the slowly converging subspace before it makes its first iteration. In our implementation the determination of the slow subspace needs N + 1 cycle simulations. The computations shown in the left graph in Fig. 2 all started from a saturated bed. For nonlinear systems it can be expected that the performance of a method depends greatly on the the initial conditions. For System C, however, it turned out that when the computations were started from an almost empty bed, results similar to Fig. 2, left graph, were obtained. Thus the convergence of System C does not depend much on the initial conditions. For System D this will be no longer the case. 4.3. System D The discretization of the model equations results for System D in a system of 120 ordinary differential equations. Hence N - 120 for System D. System D is unique among the test systems in that the convergence behaviour of the different methods depends on the initial condition. All the methods are started from two different initial states. The

269

System c

le+10

1 systemD; empty bed

1 System D; saturated, bed

s le-10 le-1

le-20

0

20000 40000 60000 CPU time in seconds

le-10

0

40000 80000 CPU time in seconds

0

40000 80000 CPU time in seconds

Figure 2. The error versus the CPU time in seconds for different methods for System C and D. The computations for System D are initiated from two different starting values: an empty bed and a saturated bed; --: dynamic simulation, O: Broyden's method, +: Newton-Picard method, El: Newton's method.

first initial state is an empty bed and the second is a saturated bed. When comparing the middle and right graph of Fig. 2, it is seen that the dynamic simulation needs slightly more time to converge to a periodic state starting from a saturated bed than starting from an empty bed. The asymptotic convergence rate, however, is the same. Newton's method yields no improvement for these initial conditions as compared to the dynamic simulation. Broyden's method does not converge from every starting point. Starting from an empty bed, Broyden's method performs very well. However, when Broyden's method is started from a saturated bed, after a few iterations the method produces badly posed initial conditions for the cycle simulation, so that the step length of the time integration becomes to small and time integration stops. In order to eliminate this problem the computation can be started with a number of dynamic simulation iterations with the aim to obtain initial conditions closer to the periodic state and in the region of attraction for Broyden's method. It is found that 200 cycles have to be simulated before a state is obtained close enough to the periodic state for Broyden's method to converge. The computation that starts with the 200 dynamic simulation cycles is shown in Fig. 2. This computation is still faster than the dynamic simulation. The Newton-Picard method suffers from the same problem when the computation is started from a saturated bed. Here the problem can be eliminated by choosing a smaller value of p so that the dimension of the slow subspace becomes larger. For the computation shown in the figure, the value p - 0.1 is used. 5. C O N C L U S I O N S Presently we have tested four methods for obtaining periodic states of cyclic processes, namely the dynamic simulation, Newton's method, Broyden's method and the NewtonPicard method. These are compared with respect to computational efficiency. To this end the four methods were applied to four PSA test systems. Two of these test systems (System A and B) exhibit linear behaviour. An important

270 difference between these two systems is that System A has large capacity terms and System B comparatively small capacity terms. This difference explains why the dynamic simulation of System B converges much faster to a periodic state than that of System A. For System A Broyden!s method reduces the CPU time with a factor five, whereas for System B it is only slightly faster than the dynamic simulation. For nonlinear systems we have introduced the Newton-Picard method as an alternative to methods more commonly used in chemical engineering literature. It is found that for the weakly nonlinear System C, Broyden's method is fastest and that the NewtonPicard method needs a CPU time equal to the dynamic simulation. For the more strongly nonlinear System D, Broyden's method is again the fastest but here the Newton-Picard also reduces the CPU time as compared to the dynamic simulation by a factor two. The main drawback of the Newton-Picard method is the lengthy first iteration, which arises from the construction of the Jacobian in order to determine a basis for the slowly converging subspace. In practice, often the periodic states need to be determined for optimization ends. This requires the computation of periodic states over a region in a multidimensional parameter space, with the aim to find an optimal parameter set. Especially in such a case the efficiency of the various methods to identify periodic states is of considerable importance. Thus, the ultimate test for these methods is to compare their performance in combination with optimization or continuation procedures. In future work the combination of the various methods for obtaining periodic states with both continuation and optimization procedures will be addressed. The final goal is a generic strategy for the analysis, design and optimization of cyclically operated processes. REFERENCES 1. O.J. Smith IV, A. W. Westerberg, Acceleration of cyclic steady state convergence for pressure swing adsorption models, Industrial and Engineering Chemistry Research. 31 (1992) 1569-1573. 2. D.T. Croft, M. G. Levan, Periodic states of adsorption cycles-I. Direct determination and stability, Chemical Engineering Science. 49 (1994) 1821-1829. 3. H. M. Kvamsdal, T. Hertzberger, Optimization of PSA systems - studies on cyclic steady state convergence, Computers and Chemical Engineering. 21 (1997) 819-832. 4. K. Lust, D. Roose, A. Spence, A. R. Champneys, An adaptive Newton-Picard algorithm with subspace iteration for computing periodic solutions, SIAM Journal on Scientific Computing. 19 (1998) 1188-1209. 5. H.M. Kvamsdal, Studies on modeling, simulation and optimization of PSA systems, Ph.D. thesis, University of Trondheim (1995). 6. E.S. Kikkinides, V. I. Sikavitsas, R. T. Yang, Natural gas desulfurization by adsorption: Feasibility and multiplicity of cyclic steady states, Industrial and Engineering Chemistry Research. 34 (1995) 255-262. 7. T.L. van Noorden, S. M. Verduyn Lunel, A. Bliek, Acceleration of the determination of periodic states of cyclically operated reactors, submitted to Chemical Engineering Science.

Studies in Surface Science and Catalysis 133 G.F. Froment and K.C. Waugh (Editors) ~_c-~2001 Elsevier Science B.V. All rights reserved.

271

Dynamics and control of a radial-flow ammonia synthesis reactor N. S. Schbib, M. N. Pedernera, D. O. Borio Planta Piloto de Ingenieria Quimica (UNS-CONICET) Camino La Carrindanga, Km. 7, 8000 Bahia Blanca, ARGENTINA e-mail: [email protected]

Abstract The steady- and non steady-state simulation of an industrial ammonia converter is presented. The reactor includes two adiabatic radial-flow catalyst beds in series. An interbed (gas-gas) heat exchanger is used to preheat the feed stream. The steady-state results showed good agreement with plant data. The influence of different disturbances (feed composition and temperature, reactor pressure) on the dynamic evolution of the main variables is analysed. The open-loop and closed-loop operation is compared from the standpoint of the reactor stability. 1. INTRODUCTION In large-scale plants, ammonia synthesis is frequently carried out in adiabatic multi-bed reactors with either quench gas or interstage cooling. Two basic flow configurations exist: axial and radial-flow reactors. Most of the large-scale ammonia plants use radial-flow reactors because this arrangement enables to handle large catalyst volumes and small particle diameters (high catalyst efficiency) without considerable pressure drops. In industrial converters the feed stream is usually preheated before entering the catalyst beds using the heat of reaction. Recently, the internal heat exchange has been explored to improve the outlet conversion in an adiabatic ammonia reactor (Abashar, 2000). The feedback of heat, which is intrinsic to autothermal processes, is a source of reactor instability (van Heerden, 1953; Pedernera et al., 1997; Morud and Skogestad, 1998; Mancusi et al, 2000). The existence of steady-state multiplicity and its connection with the optimal operating points has been analysed in detail in a previous contribution (Pedernera et al., 1999). In the present paper, the non-steady state simulation of a large-scale radial-flow ammonia converter is carried out, aiming to study the influence of the operating variables on the dynamic and stability of the reactor. By means of a detailed mathematical model, the time responses of the main variables are obtained and compared for two different conditions: open loop (without feedback control) and closed loop.

2. PROCESS DESCRIPTION A large-scale ammonia converter consisting of two adiabatic radial-flow catalyst beds with interstage cooling is considered in the present paper (Figure 1). The sense of motion in both catalyst beds is centripetal. The feed stream is preheated in an interbed countercurrent heat

272 exchanger, using the hot gases leaving the first bed. In the closed loop operation, the inlet temperature of the first bed (T01) is regulated by feedback control, manipulating the valve located on the cold by-pass stream (X~). The inlet temperature of the second bed (T02) can also be controlled by adjusting the feed temperature (Tf) in a heat exchanger outside the converter (not considered in the present paper).

(1-X)r

T f, ~)

[

i~.

R1

To1

Tout2

v Ttout Figure 1" Scheme of the ammonia synthesis reactor 3. MATHEMATICAL MODEL The catalyst beds are represented by a pseudo-homogeneous one-dimensional model. The converter is assumed to be isobaric. The kinetic expression reported by Temkin (1950) is adopted. The balance equations for the catalyst beds and the heat exchanger are the following:

Catalyst beds Mass balance: aC N2 aFN2 s . . . . at av Energy balance: OT

otR NH3 r l ~ 2

(1)

OT

~--~- = -,Cog - ~ + rlCZRN. 3 (- AH r )

(2)

Heat Exchanger Energy balance (tube side) ~ft -OT ~ t = -d~t Cpg L OTt + UA(Tsh - "It ) az where: ~bt = ~b(1- L)

Energy balance (Shell side)

(3)

pg aTsh + UA(Tsh - T t) 7sh aTsh ",,. d[ - -d~c L az

Boundary conditions First bed

Second bed 0

ION2 (V,O)= ON2 (Vlss2

V =0

ICN2 (O,t)= CN2 (VR1, t) [T--,t-(O) To2(t)

t

=

'

(4)

273

Ic,,,~(v.o)=c.~ (Vl.s.. t=0

LT(V'O)=T(V~,s,1

IcN~(o.t)=c.~., v=0

[T(O,t)=Tol(t)

where : To1

=

XTf + (1 - X)Ttout

Heat exchanger

Irt (z,O) = rt (Z~ss t= o

Lr,.h

(z.O)=r~, (z~

z = 0

ITch

(o,t) = T02 (0

Ir,(o,t)=r,

z=L

T t (L,t) = Ttout (t) Tsh (L, t) = Toutl (t~

The efectiveness factor (11) was assumed to be equal to one, based on the small particle diameters used in radial-flow ammonia reactors. This assumption was verified by calculating 1"1through a heterogeneous reactor model (Pedernera et al., 1999). The radial coordinate in the catalyst beds (Eqs. 1, 2) and the axial coordinate in the heat-exchanger (Eqs. 3, 4) are discretized using finite differences, leading to a set of ODE's which are simultaneously integrated along the time using a Gear algorithm. The same set of equations is used to solve the initial steady-state, by setting the time derivative equal to zero in Eqs. 1-4 and solving the resulting algebraic system with a Quasi-Newton method. 4. RESULTS AND DISCUSSION 4.1 Steady state results The influence of the cold by-pass fraction (X) on the outlet molar fraction of ammonia is shown in Fig. 2, for different values of the feed temperature (Pedernera et al., 1999). A maximum outlet content of ammonia is obtained at Z, _=_0.417 for the curve corresponding to Tf = 230 ~ This optimal operating point, which is located near the extinction point, has been selected as the reference steady state to obtain the dynamic results shown in item 4.2.

274

0.20

0.16 O4

~o.12 0.08

l/Y/ 7/~245 2~5 _930

0.04 0.00

,

225

|

,

!

0.20

,

i

0.40

,

0.60

0.80

Figure 2: Inlet temperature of the first bed as affected by the cold by-pass, for different Tf. 9 Optimal operating point Figure 3 shows the steady-state radial temperature profiles for the two adiabatic catalytic beds operating at conditions of the optimal point. The corresponding axial temperature profiles in the interbed heat exchanger are also included in Fig. 3, for the tube side (Tt) and shell side (Tsh). The simulation results have been compared with industrial data corresponding to a large scale ammonia converter. The deviations at the reactor outlet were less than 0.2% (relative error) in composition and 14 ~ in temperature (Tout2). x[c]

550

First bed (R1)

Toutl ...

Heat exchanger

~

First bed (R2)

]

500 450

S

400 350 300 250 200

,

I

,

i

,

f

,

J

,

,

i

,

i

,

i

,

i

,

!

,

i

,

i

,

i

,

i

,

J

0.0 0.2 0.4 0.6 0.8 1.0 0.0 0.2 0.4 0.6 0.8 1.0 r* 0.0 0.2 0.4 0.6 0.8 1.0 r* Z*

Figure 3" Temperature profiles in the catalyst beds and heat exchanger (Tf = 230 ~ ; ~,=0.417). r*=(r0 - r)/(r0 - ri), here r0, ri = outer and inner radius of each bed; z*=z/L, where z= axial coordinate and L = tube length in the heat exchanger. Although a large amount of heat is being transferred in the interbed heat exchanger, the autothermal reactor, when it is considered as a whole, is an adiabatic reactor where a single exothermic reaction is being carried out. Therefore, an almost linear relationship exists between the outlet conversion (XN2) and the overall temperature rise (ATovl= Tout2 - Tf). As it can be seen in Fig. 4, the variable ATovl is an indirect index of the conversion of N2 per pass.

275

30

~20 Z X 10 40

' 8'0 ' 189 ' l~i0 ' 200 ' 240 Tout2 - T f

Figure 4: Outlet conversion XN2(~ VS. overall temperature rise (Tout2 - Tf). 4.2 ]Dynamic results The time responses of the temperature (deviation variables) at different positions in the reactor are included in Fig. 5, for a positive step change of 10 ~ in the feed temperature, Tf (see Fig. 1). An instantaneous change of around 4 ~ appears at the inlet of the first bed (T01), due to the sudden increase in the temperature of the cold by-pass (L~). The time evolution of T01 is a consequence of the heat feedback in the interbed heat exchanger (see the similar shape of the Ttout and T01 curves). A slight inverse response is observed at the outlet of the first bed (Toutl), which is a typical behaviour of fixed bed catalytic reactors. This "wrong way behaviour" influences the temperature at the inlet of the second bed (T02), which oscillates at t~_150 sec. 30.80

_

f

f

/

~ _ ~ ,

/

'''S

~

ToutTt~

30.40

,"

e

Z

I

30.00

/Tout2 /

0 !

0

f

400

s

i

!

|

i

800 1200 time [see]

|

|

1600

|

2000

Figure 5: Dynamic evolution of the temperature at different points of the reactor, for a step change ATf=10~ Deviations from the initial (steady-state) values.

29.60

z

0

i

400

|

|

!

!

800 1200 time [see]

i

|

1600

|

2000

Figure 6" Dynamic evolution of the conversion of N2 at the reactor outlet, for a step change ATf= 10~

276 The differences between the shape of the Toutl and T02 curves arise from the dynamics associated to the interbed heat exchanger. An inverse response also appears at the outlet of the second bed (Tout2), with a considerable delay with respect to that of the first bed. Fig. 6 shows the transient behaviour of the outlet conversion for the same operating conditions of Fig. 5. As expected, a final drop in conversion appears, because the disturbance in Tf shitts the reactor from its optimal operating conditions. The final conversion drop is in concordance with the results shown in Fig. 5. In fact, for a positive change of 10 ~ in Tf, a temperature rise of around 8 ~ appears in the variable To,t2 (final steady state), i.e, ATovldrops around 2 ~ when the outlet conversion also decreases from 30.2 to 29.7 %. The transient behaviour of the variable XN2 (Fig. 6) is determined by the slow response of the temperature (high heat capacity of the catalyst), which is the source of the overshoot observed in the XN2 curve. The influence of the cold by pass fraction (~,) on the temperature and conversion responses is shown in Figs. 7 and 8. As it was indicated in Fig. 1, ~, is one of the manipulated variables when the reactor is operated under feedback control. When a negative step change in ~, occurs, the temperature T01 shows an instantaneous increase of around 10~ (Fig. 7). This step change in temperature at the inlet of the first bed leads to an initial inverse response in Toutl (wrong way behaviour). This inverse response affects the variable Ttout in the heat exchanger, and To1 is also affected as a consequence of the heat feedback (see Fig. 1). The inlet temperature of the second bed (To2) follows the same initial trend than the Toutl curve. However, in contrast to the behaviour observed for Toutl, the final value of To2 (t=2000 sec.) is lower than the initial one, because of the higher heat transfer rate in the heat exchanger (lower ~, implies higher gas flowrate in the tube side). As before, the variable Tout2 shows an inverse response. Finally, the outlet conversion exhibits a more oscillating response than that observed for the case of a step change in the Tf (compare Figs. 6 and 8).

30.40 10 .,. Toutl

30.20

5 .,..~

'

30.00

Ttout

0

x

"

-5 't I,"'" V/ -10 0

. . . . . . . . . 400 800 1200 1600 2000 time [sec]

Figure 7" Dynamic evolution of the temperature at different points of the reactor, for a step change in ~, from 0.417 to 0.375. Deviations from the initial (steady-state) values.

29.80 29.60 0

.

. . . . 3000 1000 2000 time [sec] Figure 8: Dynamic evolution of the conversion of N2 at the reactor outlet, for a step change in ~, from 0.417 to 0.375.

277 The closed loop and open loop reactor dynamics are compared in Figs. 9 and 10, for a step decrease in the reactor pressure. A SISO controller was tuned to control the inlet temperature of the first bed, T01, through the by-pass valve (see Fig. 1). When the reactor is operated without feedback control, the sudden reduction in the pressure leads to a decrease in the heat generation rate, the temperature Toutl (and consequently T01) decreases and the reactor moves to a lower (extinguished) steady state. This effect can be seen in Fig. 9, where the overall temperature rise drops from 230 ~ to 12 ~ Conversely, when the feedback control loop is included, the overall temperature rise shows a significantly lower decrease (from 230 to 199~ as a consequence of the gradual reduction of the cold by-pass (Fig. 10). A similar behaviour is found for the case of an increase in the ammonia contents at the inlet of the converter. In fact, when the molar fraction of ammonia in the feed stream increases, the equilibrium shifts to the reactants and the heat generation rate decreases. If the reactor is being operated under open loop, an extinction phenomenon appears due to the autothermal operation of the converter (Fig. 11). Under closed loop operation, the control action leads to a decrease in the cold by-pass fraction (Fig. 12). As a result, the reactor remains at the upper branch of the curve shown in Fig. 2 (ignited steady-state) and the outlet conversion drops slightly. 300v

0.42 open loop losed loop

200

0.40

! C'q

0.38

0

10o

closed loop

loop 0.36

10'00 20'00 time [sec]

'

3000

Figure 9: Time responses of the global temperature rise in the converter, after a negative step change in the reactor pressure (AP =- 10 atm.).

o

ooo 20'00 time [sec]

3000

Figure 10: Evolution of the cold by-pass fraction, after a negative step change in the reactor pressure (AP = - 10 atm.).

278 300

0.42 openloop closed loop

200

0.40

!

0.38

0

t- 100 closed loop 0.36 '

10'00

'

20'00

time [see]

'

o

3000

Figure 11: Time responses of the overall temperature rise in the converter, aider a positive step change in the amonnia content (YNH3)f = 0 . 0 3 9 ~ 0.05

100020003000 time [sec]

Figure 12: Evolution of the cold by-pass fraction, after a positive step change in the amonnia content, (YNH3)f = 0 . 0 3 9 ~

0.05

5. CONCLUSIONS By means of a detailed mathematical model, the steady- and non-steady state operation of a large scale ammonia converter have been analysed. The model takes into account the dynamics associated to each catalyst bed and the interbed heat exchanger. The steady-state results showed good agreement with plant data. The heat feedback associated to the autohermal operation is a source of reactor instability, leading to oscillating time responses and possible extinction of the reaction. This negative effect can be partially compensated by means of a feedback control action, aiming to keep the inlet temperature of the first catalyst bed in the set-point value. In case of changes in the operating conditions of the ammonia synthesis loop, this set-point value (T01) should be adapted to minimise production losses. This procedure could be carried out using an optimisation algorithm that contains the steady-state model of the plant. NOMENCLATURE A

heat transfer area, m 2

T = Cpcat1313

Cpg cp Cm

mean specific heat of gas, kJ/(kg K) specificheat, kJ/(kg K) molar concentration o f N 2 , k m o l / m 3 molar flow of N2, kmol/s length of the heat exchanger, m mass, kg reactor pressure, atm outer and inter radius of each bed, m intrinsicreaction rate, kmO1NH3/(m3 s)

)', "- Cpg mg, t+Cpstmst, t ~/-n Cpgmg, sh+Cpstmst, sh e bed porosity, m3/m 3 dHr heat of reaction, kJ/kmO1NH3

FN2 L m P ro,ri R~H3

"

~, r/ p~

-

cold by pass (fraction of ~) effectivenessfactor bulk density, kgcat/m3.

Subscripts

279 T t U V y z

temperature, K time,s overall heat transfer coefficient, kW/(m2K) reactor volume, m 3 molar fraction axial coordinate (heat exchanger), m

cat f g sh ss st t

catalyst feed gas shell side (heat exchanger) steady state steal tube side (heat exchanger).

Greek letters a

catalyst activity coefficient total mass flow rate, kg/s

REFERENCES

Abashar M. E. E., Chem. Eng. dr., 78 (2000), 69-79. Mancusi E., Merola G., Crescitelli S. and Maffetone P.L., AIChE J. 46 (2000), 824-828. Morud J. C., Skogestad S., AIChE Journal, 44 (1998), 888-895. Pedemera M. N., Borio D.O., Porras J. A. AIChE Journal 43 (1997), 127-134 Pedemera M. N., Borio D. O., Schbib N. S., Comp. & Chem. Eng. (1999) $783-$786. Temkin M., J. Phys. Chem. (USSR) 24 (1950), 1312. Van Heerden C.,Ind. Eng. Chem., 45 (1953), 1242-1247.



Studies in SurfaceScienceand Catalysis133 G.F. Fromentand K.C.Waugh(Editors) 9 2001 ElsevierScienceB.V.All rightsreserved.

283

Mass transfer through the gas-liquid interface at the presence of adsorbed active phospholipid monolayer L. Gradofi and T.R. Sosnowski Department of Chemical and Process Engineering, Warsaw University of Technology, Waryrlskiego 1, PL-00-645 Warsaw, Poland. E-mail: [email protected] The transfer phenomena in the thin liquid layer covered by the phospholipid monolayer are investigated. The case of dynamic oscillatory changes of the air/liquid area is considered. Such situation occurs in the pulmonary surfactant system being the natural biological structure in the breathing zone of the human respiratory system. It was shown that in those conditions the surface tension gradient is instantly generated and varied along the interface what results in intensified flow at the interfacial region. Experimental analysis conducted with the Langmuir film balance confirmed the theoretically predicted induction of the ordered superficial flow. Measurements of the interfacial mass transfer rate for the oxygen being absorbed by the liquid film showed the increase of the absorption rate in comparison with the static or phospholipid-free interface. It was postulated that for certain system parameters the hydrodynamic structure of the flow can be qualitatively changed what should result in a strong increase of interfacial mass transfer rate. 1. INTRODUCTION Chemical molecules, which posses both strong hydrophilic and strong hydrophobic parts, spontaneously migrate to the interface, where such asymmetric molecular structures can find the most favorable energetic state. This process - called adsorption - is common in aqueous solutions of many organic compounds, which tend to concentrate at the gas-liquid interface. The energy minimization of the system is manifested by the reduction of the surface tension as a result of the appearance of a monomolecular layer of the asymmetric molecules at the gas-liquid interface. The substances, which are spontaneously adsorbed and which decrease the surface tension, are recognized as the surface-active compounds, i.e. surfactants. They are widely used in the technology and in household applications. Also in the natural world there are many examples of such substances. One of the most important - due to its vital functions is the so-called pulmonary surfactant existing in the lungs and providing the mechanical conditions of breathing. It is mainly composed of phospholipids and some specific proteins, which are naturally produced and secreted by the epithelial cells. The existence of the phospholipid-rich film at the air-liquid interface of distal airways may have further physiological consequences.

Acknowledgment: This work was supported by the Polish Committee for Scientific Research (KBN) under the grant No. 3T09C 037 16.

284 The surfactant in the respiratory system undergoes dynamic excitations during oscillatory area changes of the breathing cycle.Such conditions can promote a specific activity of the phospholipid film, which may result in modified mass transfer conditions in the pulmonary region. This can be important for the efficiency of the pulmonary gas exchange or of the removal of aerosol deposits from the respiratory zone. Investigations of the underlying phenomena are presented in this paper. 2. THEORETICAL AND EXPERIMENTAL ANALYSIS 2.1. Physical model The experimental analysis of the discussed system was done using the Langrnuir film balance - LFB (Lauda, Germany). The device is schematically depicted in Figure 1. The LFB consists of a shallow rectangular trough filled with water (hypophase). The air-liquid interface is confined between two barriers. The left one is a part of the surface tension measuring system and it is immobile. The other is movable thanks to the rotation of the driving screw, what allows for the well-defined preprogrammed variation of the interfacial area in the vessel. If the phopsholipid (in our case - dipalmitoyl phosphatidylcholine, DPPC) monolayer is introduced at the air-water interface, the experimental system becomes a physical analogy of the pulmonary surfactant and allows for studying the specific phenomena, which may be related to the function of the surfactant in the lungs. The changes of the surface tension associated to the oscillatory variation of the interfacial area are automatically measured and recorded by the computer connected to the device. To be able to investigate the hydrodynamics of the system, the device was additionally equipped with a video camera (SONY, Japan) for observations of the displacement of tracer particles located at the gas-liquid interface. The experimental system could be also adapted for direct measuring of the mass transfer rate across the interface in the presence of the active phospholipid monolayer. For that purpose, the electrochemical system was developed [1], where the oxygen flux across the interface could be determined by the measurement of the electric current intensity. The results of experimental investigations will be presented in the further part of the paper. 2.2. Basics of the mathematical modeling The hydrodynamics of the experimental system can be described theoretically. Such approach is very important for correct interpretation of the experimental results, and for their extrapolation for the conditions not attainable in the existing experimental system. With the mathematical model the parametric study of the system is also possible, what can reveal the most important factors responsible for the occurrence of the specific transport phenomena. The model was presented in details elsewhere [2]. It was based on the equations of the momentum and mass transfer in the simplified two-dimensional geometry of the air-watersurfactant system. Those basic equations were supplemented with the equation of state for the phopsholipid monolayer. The resultant set of equations with the appropriate initial and boundary conditions was solved numerically and led to temporal profiles of the surface density of the surfactant, F [mol m2], surface tension, cr [N m-l], and velocity of the interface, Vs [m sl]. The surface tension variation and velocity field obtained from the computations can be compared with the results of experiments conducted with the LFB.

285

Fig.1. Schematic of the Langrnuir film balance (LFB)connected to the personal computer (PC): a- top view, b - side view.

3. RESULTS Before the results of the computations and the measurements are presented, let us discuss some general features of the analyzed system. The presence of the monomolecular layer of the phospholipid at the air-water interface introduces new qualitative features to the system. The interfacial tension is related in some proportion to the surface density of the surfactant. If the surface coverage is not uniform, it gives a rise to the surface tension variation along the interface. The non-uniformity of the surface tension means in fact the non-uniformity of the surface stress, which induces the interfacial flow to recover the surface equilibrium. Under dynamic conditions of oscillatory area changes (similar to those being typical for breathing cycle), the surface density of the phospholipid is instantaneously varied, and the superficial flow towards regions of elevated surface tension (i.e. reduced F) is propagated. This effect provides the mechanism for intensified hydrodynamic conditions, which may result in higher rates of the mass transfer in the interfacial region. Figure 2 presents the results of computed and measured superficial velocity propagated as the result of the perturbation of the surface density distribution by the barrier sweeping the interface in the LFB experiment. In the graph the velocity and location are presented in the dimensionless forms: the velocity is reduced in respect to the velocity of the moving barrier, and the distance - in respect to the temporal location of this barrier. It is seen that during linear surface deformation caused by the shift of the fight barrier (Fig.l), a quasi-stationary state is achieved, where the superficial velocity is a linear function of the reduced distance between the surface element and the immobile left barrier.

286 A

0.9 0.8 O

~

t~ O tD Q.. z~ o'J

-13 (D "13

0.7

/

I

0.6 0.5 0.4 0.3 0.2

experiment

9

0.1

-

I

0

i

i

i

i

- - computation

-

!

i

i

i

0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 reduced distance from the stationary barrier

Fig.2. The profile of the surface velocity in the air-water-phospholipid system in the LFB.

This result demonstrates a dramatic difference in comparison to the situation when no surface-active monolayer is present in the system. In the latter case no flow organization will be induced and, as observed in the laboratory, only surface elements remaining in close vicinity to the moving barrier will be displaced during its movement. Viscous dissipation consumes the huge part of the energy introduced to the system by the moving barrier, so the distant regions of the interface are not disturbed by this local stimulus. On the contrary, the phospholipid film provides the mechanism of the propagation the local disturbance along the whole interface. This mechanism is known as the Marangoni effect - the generation of the surface tension gradients, which result in the flow of the surface layer. Such flow can be of a great importance for the pulmonary surfactant system since it can provide the way of removal of deposited aerosol particles from the lungs towards ciliated airways [3]. When the interface is moving, the continuity conditions require that underlying liquid be also dragged with the surface. Such hydrodynamic effect should be related to the modification of the mass transfer condition across the active interface. Our measurements showed [ 1] that oxygen absorption rate is indeed higher when the monolayer is present in the system, which undergoes cyclic area oscillations, Figure 3. This experimental result is consistent with the theoretical prediction [4]. An interesting conclusion from these last measurements is the relation between the absorption rate and the depth of the hypophase. It seems that the mechanism responsible for a noticeably increased mass transfer rate may be more pronounced at the specific system conditions.

287

2.8 2.7

% ~ .It

,v-i

2.t5 S

2.5

%

/

~ 0

2.4

~

2.3

N

/

N

"0

/ O/ / /

~

2.:2

~

2.1

/

2.0 1

'

I

2

'

I

3

'

I

4

'

I

5

'

I

6

'

7

hypophase depth [ram] Fig.3. The mass transfer coefficient for oxygen absorbed from air to water across the moving interface with the phospholipid monolayer.

Accordingly, a hypothesis was formulated, which seeks the theoretical rationale of appearance of the specific flow structures inside the liquid layer. It is known, that under certain circumstances, the surface tension variations may lead to the flow instability and to the induction of convection cells [e.g., 5, 6]. Our preliminary theoretical analysis of the hydrodynamic stability of the system [7] indicated that it is possible, that for the Reynolds numbers exceeding the critical value, convection cells inside the hypophase can be formed. This should lead to the significant increase of the mass transfer rate. A novel technique of visualization of such "jump" of the system from the stable to the unstable flow conditions was tested. By observing the spreading of the pink spot of the phenolphthalein on the liquid surface in the LFB it was noticed, that instead of a continuous increase of the spot diameter due to diffusion, the spot immediately disappeared, Figure 4. Such effect suggests that a qualitatively different mixing mechanism can appear in the liquid layer. At the current stage of the project, also other flow visualization experiments are conducted to demonstrate the loss of flow stability at the specific conditions, which are responsible for generation of the surface tension gradients. Additional technique of measuring the mass transfer rate was recently proposed [8], based on the use of the pH-microelectrode and CO2 as the compound being absorbed by the liquid layer.

288

Fig. 4. The shape of a phenolphthalein spot at the expanding surface of water covered with the phospholipid monolayer.

4. CONCLUSIONS The specific momentum and mass transfer conditions were observed in the thin liquid layer covered by a phospholipid monolayer and subjected to forced variations of the interfacial area. Due to the Marangoni effects, the absorption rate for oxygen was increased when compared either to static condition or to area variations without the monolayer in the system. It was shown that for certain system parameters given by a critical Reynolds number, there is a possibility of propagation of the specific flow structures inside the liquid layer. The convection cells, which can appear, are similar to those observed in the Ryleigh-Benard experiment [9]. Such phenomenon can be very important for some air-water-phospholipid systems such as the pulmonary surfactant present in the natural mass exchanger - the lungs. The hydrodynamic system described in the paper can be a very useful tool for explanation of convective diffusion-reaction transport process in case of interaction of allergens with pulmonary epithelium, causing atopy. REFERENCES

1. T.R. Sosnowski, L. Gradofi, M. Skoczek, H. Dro~dziel, Int. J. Occup. Safety Ergon., 4 (1998) 391. 2. L. Gradofi, A. Podgrrski, T.R. Sosnowski, J. Aerosol Med., 9 (1996) 357. 3. A. Podg6rski, L. Gradofi, Ann. Occup. Hyg., 37 (1993) 347. 4. A. Podg6rski, Pr. Wydz. In~. Chem. Proc., 24 (1997) 15 - in Polish. 5. J.R.A. Pearson, J. Fluid Mech., 4 (1958) 489. 6. K. Warmuzifiski, Chem. Eng. Sci., 45 (1990) 243. 7. L. Gradofi, J. Biatecki, J. Hotyst, T.R. Sosnowski, In~. Chem. Proc., 21 (2000) 163 - in Polish 8. T.R. Sosnowski, L. Gradofi, M. Skoczek, J. Hotyst, Summaries of CHISA 2000, vol.3, 137. 9. Lord Raleigh, Phil. Mag., 32 (1916) 529.


Oxygenates and Olefins from Catalytic Partial Oxidation Cyclohexane and n-Hexane in Single-Gauze Chemical Reactors

289

of

Ryan P. O'Connor a and Lanny D. Schmidtb ~bDepartment of Chemical Engineering and Materials Science, University of Minnesota, 421 Washington Ave. SE, Minneapolis, MN 55455-0132 aE-mail: roconnor @cems.umn.edu bCorresponding author. E-mail: schmi001 @tc.umn.edu

Abstract The single-gauze millisecond reactor has shown promise for the production of oxygenates and olefins from n-butane [1-3]. In this work, experiments are performed by reacting cyclohexane or n-hexane with oxygen over Pt-10%Rh single-gauze catalysts with ~ 100-l.tm wire diameter and 20, 40, or 80 mesh (wires per inch). The 40-mesh catalyst provides the optimal amount of heterogeneous chemistry to initiate the necessary gas-phase reactions. Using the 40-mesh single gauze, we vary alkane/oxygen molar ratios (C6/O2) from 0.5 to 5.0, N 2 dilution from 10% to 75%, flow rates from 1.0 to 3.5 standard liters per minute (superficial catalyst contact times from ~1.0 to -0.3 ms, respectively), preheat temperatures from 100~ to 300~ and pressures from 1.2 to 2 atm. The alkane/oxygen ratio is the most important reactor operating variable because temperatures, reactant conversions, and product selectivities all change significantly as C6/O2 is varied. Parent oxygenates and olefins are maximized at C6/O2 = 2. Low dilution favors olefin formation while an optimum in N 2 dilution is observed for oxygenate selectivity in our experiments. Lower flow rates (longer contact times) promote higher COx selectivities. Lower feed temperatures favor the production of oxygenates over olefins. Higher reactor pressures (up to 2 atm) increase the yield of key products and allow complete oxygen conversion. Cyclohexane partial oxidation in the single-gauze reactor can produce ~85% selectivity to olefins and oxygenates at 25% cyclohexane conversion and 100% oxygen conversion, with cyclohexene and 5-hexenal as the dominant products [4]. Cyclohexene is more valuable than cyclohexane, and 5-hexenal is a 1,6-difunctional olefinic aldehyde which is potentially useful for the manufacture of polymers. Partial oxidation of n-hexane in the single-gauze reactor can produce 70% selectivity to oxygenated hydrocarbons, including >50% selectivity to C 6 oxygenates with--35% selectivity to 2,5-dimethyltetrahydrofuran at ~20% n-hexane conversion and 100% oxygen conversion [5]. Synthesis of polymers with special properties from the monomer 2,5-dimethyltetrahydrofuran is possible. Density-Functional Theory (DFT)with the B3LYP/6-31+G(d) method for geometry optimizations has been employed for the prediction of reaction enthalpies and rate-constant parameters for the partial oxidation of cyclohexane in single-gauze chemical reactors. The model includes 29 species and 46 irreversible (23 reversible) gas-phase reactions. The energetics of major gas-phase reaction channels are probed by locating stable reactants, products, and transition-state intermediates. One-dimensional reactor simulations are carded out using the DFT mechanism, and qualitative agreement with experimental data is obtained. Understanding the favored reaction pathways suggests ways to adjust reactor operation for desired product distributions. Detailed numerical simulations of the surface-assisted gas-phase process, including surface chemistry and fluid dynamics, will allow the investigation of experiments which are costly or potentially dangerous to carry out.

290

1. Introduction Oxygen-containing hydrocarbons (oxygenates)and olefins are important intermediates in the manufacture of specialty chemicals. Oxygenates are typically produced in the liquid phase through large multistage processes involving expensive separations and careful temperature control. Conversions of hydrocarbon reactants are often 10% or less. A single-stage method to selectively manufacture oxygenates as well as olefins could be significant for the manufacture of specialty chemicals. The millisecond single-gauze reactor, which successfully couples catalytic and gas-phase chemistry, can directly produce oxygenates and olefins from alkanes and oxygen at time scales much shorter than those in comparable industrial processes. A unique feature of the single-gauze chemical reactor is rapid preheat (on the order of 106 ~ followed by fast thermal quenching to prevent decomposition of reactive intermediates. Species residence times over the platinum-rhodium catalyst are very short (100-500 kts) and mass-transfer rates to and from the catalytic surface are high. It is believed that the heat from the exothermic surface reactions along with desorbed free radicals initiate a gas-phase reaction sequence which leads to the formation of unstable species. The large transparency of the single gauze facilitates rapid mixing of the colder gases passing between the catalyst wires with the hot gases and radicals leaving the gauze surface, resulting in fast quenching of the homogeneous reactions so that unstable oxygenates and olefins cannot decompose. The quenching action in the wake region of the single gauze kinetically freezes the product mixture so that valuable non-equilibrium species can be obtained.

2. Experimental The catalyst in all experiments was a single 40-mesh (77% transparency) woven gauze with 90% Pt (by mass), 10% Rh, and traces ( 3, oxygen conversion decreased drastically and then leveled off at approximately 40%. The transition at C6H12/O 2 = 1 was accompanied by a change in the cyclohexane-conversion trend. Dropping sharply until a cyclohexane/oxygen ratio of about unity, the cyclohexane conversion roughly leveled out at ~20% before eventually decreasing to 4% at C6H12/O2 = 5. Operation beyond C6H12/O 2 = 5 was not possible at a 2.5-SLPM flow rate due to pump limitations. The reaction was nearly extinguished at these conditions, as indicated by the low temperature and conversions. Figure l b shows the overall selectivities (on a carbon-atom basis) to alkanes, CO x, olefins, and oxygenates. Alkanes were only significant (5-10% selectivity) for C6H12/O 2 < 1. Methane and ethane were the dominant alkanes, although propane, n-butane, and n-pentane were present in small amounts (60% selectivity) for C6H12/O 2 < 1 and were produced with 35-40% selectivity for C6H12/O 2 = 1-5. Oxygenates, on the other hand, were ~10% selective for C6Hl2/O 2 < 1 but increased drastically to nearly 45% maximum selectivity at C 6 H J O 2 = 2.5. Oxygenates overtook olefins as the dominant product at a cyclohexane/oxygen ratio of about 1.75, and the yield of oxygenates was optimized (at --10%) for C 6 H J O 2 = 2.0. Panel c in Figure 1 indicates the major individual olefins. For low cyclohexane/oxygen ratios, ethylene and 1,3-butadiene were abundant products, giving evidence of cracking reactions. C2H 4 and 1,3-C4H 6 followed the same trends and were both minimal for C6H12/O 2 > 1.5. Propylene (not shown), another cracked product, was >5% selective at low ratios but was not produced for C6H12/O 2 > 1.5. Cyclohexene is an important olefin at all ratios, and it was maximum (>30% selectivity) at C6H12/O 2 = 2.0. Benzene was about 5% selective for C6H12]O2< 3 and then increased to > 10% selectivity for cyclohexane/oxygen ratios exceeding 3. Cyclohexadiene, the intermediate between cyclohexene and benzene, was just a few percent selective at low ratios ( C 6 H J O . < 1) and was not present at richer conditions. Figure ld shows the major oxygenates. At low cyclohexane/oxygen ratios (C6H12/O 2 < 0.8), the predominant oxygen-containing species were C2-C 3 compounds and small amounts of methanol (not shown). C2HxO includes acetaldehyde (CH3CHO) and ethanol (CH3CH2OH), and C3HxO includes propionaldehyde (CH3CH2CHO), propanol (C3H7OH), acetone (CH3COCH3), and 2-propenal (CH2CHCHO). For C6H12/O 2 > 1, pentanal (CsHI00) and especially 5-hexenal (5-c6nloO) became the significant oxygenated products. Selectivities of --25% to 5-hexenal and--15% to pentanal were achieved at cyclohexane/oxygen ratios of 2 and 3, respectively. 5-Hexenal selectivity decreased for cyclohexane/oxygen ratios from 2 to 3 but

292 100

(a) 800 % 9 oke olt~

60 9

TotalOlefins

700

80

\

~ ~0

\ xX

6o

X.

Total O x y g e n9a t e s e

30 i d~

~

0= 20

y............iiiiiiiiiiiiii

10

r,.)

0

20

0

1

2

3

4

5

0

300

0

i

i

1

2

/=~

i

"

3

i

i

4

5

C6H12/O 2

C6H12/O 2

9

-

c-CH,o

o

.,. ~;~ 25

9~,

25

"~

20

o

o

o r~

20

o "~

~r~

+

,,

lO

"\

88 8 9

5 0

c2 0

~

. ..."'*

~

1,3 C,H *

Z- ....... '*...... *

-.,.,_.= 2 ....

1

........ 9......

.7-_~,-_______. 3 4 5

C6H12/O 2

/

o

/

" ~

.....-"../--,, *." t

/. __ i."O C2 7' ' *I-Ix .

5 O 0

"o

. . . . ....................... 9

c-C H O c ! . '~ I0~~ 6o

9,, o----,. . . . . ~ . . . . ; I

I

1

2

N

?..,..---"~,,,0o-0 ". Qmli

"*""""'"e.."e....... 2

C_H O - ' ' ' ~ . . . . i 3

3

X

I

4

~ . . . . "~ I

5

C6H12/O 2

Figure 1. Effect of cyclohexane/oxygen ratio for 2.5 SLPM total flow (z'-- 400 ITS), 30% N 2 dilution, TO= 200~ and P = 1.2 atm. then increased significantly to >30% for C6H12/O2 - 5. Cyclohexanone (c-C6H100) was produced most selectively (--5%) for C6nl2/O 2 -- 4. The corresponding alcohol, cyclohexanol, was not detected (

~

20

20

.~ "/

Combustionratio I 1

t 2

I 3

L

Alkanes

l 1

I 4

t J200 5

0 0

.

,

--.

I

I

I

I

I

1

2

3

4

5

n-Era 1./O2. in Feed

n - C H J O 2 in Feed

x ~

"ca, .,.

C Oxygenates

30

30 x

2,

,

N
200 ~ and > 175 ~ for Ru) where isomerisation from the cis-form occurs [6,7,9]. However, no isomerisation or dehydrogenation was observed under the studied conditions where pure cis- or trans-decalin was used as a reactant in decane. This indicates that the hydrogenation occurs mainly through A~'9-octalin. Long experiments, resulting in low catalyst activity at the end of the experiment, led to increased cis-product. The change in isomer ratio was nevertheless less pronounced than observed with tetralin hydrogenation [ 10]. The cis-decalin content was also enhanced at the start-up and when tetralin conversion approached zero, as seen in Figure 3. Thus, A~'9-octalin is formed only through the isomerisation of A9'l~ and not by isomerisation and parallel hydrogenation of hexahydronaphthalene as we earlier suggested for tetralin hydrogenation [ 10].

The isomer ratio was only very weakly dependent on the process conditions of temperature, pressure and initial concentration of naphthalene. The stereochemistry indicates that the isomer ratio is governed by kinetic constraints, not by thermodynamic equilibrium.

1.00 0.75 OOOo0000

0.50 0.25 0.00

~ 000000~.~,., 0

1

2

3

4

61

62

63

~

64

,qw, v "

65

TOS, h

Figure 3. Cis-to-trans-ratio (filled squares) and tetralin conversion (open circles).

313 naphthalene

1l

tetralin

dihydronaphthalene

]l

A9,1~

hexahydronaphthalene

]l

IL

cis-decalin

trans-decalin

Al,%octalin Figure 4. Reaction scheme for the hydrogenation of naphthalene. 3.6. Reaction Scheme

The proposed reaction scheme (Figure 4) is based on the observations reported above. It includes stepwise cis-addition/cleavage of two hydrogen atoms. Thus, the reaction mechanism includes the formation of both di- and hexahydronaphthalene even though these were not detected. However, Weitkamp [7] observed different isomers of these intermediates in naphthalene hydrogenation on noble metal catalysts. Reversible hydrogenation steps were not detected in the present study but were observed as a minor reaction in our previous work with tetralin [ 10]. The irreversibility of the reaction from octalins to decalins was deduced from the absence of any traces of octalins or tetralin when pure cis- or trans-decalin was used as reagent in a decane. In turn, direct hydrogenation of tetralin to cis-decalin was omitted because of the high trans-decalin content, which is typically observed only on far more transselective palladium catalyst [7]. The isomer ratio was virtually independent of pressure indicating that there was no hydrogen addition or cleavage between decalin precursors (i.e. octalins). This is also noted in our reaction scheme. In the comparison of the hydrogenation rates of A 9'1~ and A~'9-octalin on PtO catalyst Smith and Burwell [13] found the rate of A~'9-octalin to be 25-fold that of A 9'10octalin. Weitkamp [7] reported an even larger difference at higher pressure on noble metal catalysts. At the same time, no isomerisation of Al'9-octalin to A9'~~ was observed [7,13], probably because of the high hydrogenation rate of Al'9-0ctalin; i.e. Al'9-octalin is hydrogenated to cis- and trans-decalin rather than isomerised to A9'l~ For these reasons, we assumed an irreversible isomerisation of A9'l~ to A l'9-octalin. The main hydrogenation route occurs, then, via Al'9-octalin. This route is also supported by the high concentration of trans-decalin. High cis-decalin content at the start-up and when the tetralin conversion ceased (Figure 3) indicates that the hexahydronaphthalene is hydrogenated only to A 9'1~ octalin, and not to A 9'10- and Al'9-octalin as proposed in our earlier studies [10] and by Weitkamp [7]. Al'9Octalin reacts to decalins with constant isomer ratio and it is not plausible that this ratio would change with time. Therefore, higher cis-decalin content is observed at the start-up

314 when hydrogenation occurs predominantly through A9'~~ than in steady state, when the main hydrogenation route is through Al'9-0ctalin. Similarly, excess of cis-decalin is observed when the tetralin conversion ceases because the precursor for the formation of transdecalin (Al'9-0ctalin) is used up before that for cis-decalin (A9'l~ 3.7. Overall Reaction Rate Deactivation of the catalyst forced us to integrate the rate over whole experiment instead of using rates in steady state. Transient mole balances for the gas and liquid bulk were

dt

an/ dt

"- FinG,i - VRNiaGL -- FoGut,i

(1)

= Fi~,i + VRNiaGL + mcat~',app --~Lt, i

(2)

Gas and liquid outlet flows were obtained by simulating a P-controller (equation 3) for which the liquid volume was evaluated from a step response experiment.

(3)

Fou t -- Xp(Vcalc - Vset)

The gas-liquid mass transfer was modelled with the two-film theory. Vapor-liquid equilibrium constants, Kgat, were calculated by the Soave-Redlich-Kwong equation of state. The gas and liquid side mass transfer coefficients were assumed to be high ( ~ a a t = 1.0xl0 2 and tcaaat = 1.0xl0 4 s -l) since our earlier results indicated no limitations at the gas-liquid interface [ 11 ]. The mole balance inside the spherical catalyst particle was described by equation 4, in which the catalyst activity, a, was based on equations 5 and 6: de i

Di,eff ( 0 2Ci

20C i ~

p

,

dt = eRep ~~z e + ---z Oz ) + --~ ri

(4)

(5)

ri* = a i r i

dai = k D iaai ' dt

(6)

Liquid-solid mass transfer resistance could be excluded because of the high agitation [ 11 ]. As a result, concentration at the catalyst surface was assumed to be the same as in the liquid bulk. The second boundary condition is defined by the symmetry (spherical catalyst particles) and is thus dc/dz = O. Overall reaction rates for the naphthalene and tetralin were based on generalised Langmuir-Hinshelwood kinetics [14] according to equations 7 and 8. Temperature dependency of rate constants kD, kN and kT was described by Arrhenius' law and adsorption equilibrium constants by the van't Hoff equation.

C ,N --rN

(1 + KNC N + Krc T + KHC H )1

(7)

315 k

- rr

tl T m T TCT C H

(8)

(1 + KNC N + K r c r + KHC" )'

Equation 4 was discretised by a 5-point central difference formula and thereafter first-order differential equations 1, 2, 4 and 6 were solved by a backward difference method. Apparent reaction rate was solved by summing the average rates of each discretisation piece of equation 4. The reactor model was integrated in a FLOWBAT flowsheet simulator [12], which included a databank of thermodynamic properties as well as VLE calculation procedures and mathematical solvers. The parameter estimation was performed by minimising the sum of squares for errors in the mole fractions of naphthalene, tetralin and the sum of decalins. Octalins were excluded from the estimation because their content was low ( 2 0 "Cintr a .

The diffusion coefficient of xenon in HZSM-5 being several orders of magnitude higher than that of benzene [7] we can admit that adsorbed xenon is always at equilibrium. The simulation of the experimental spectra, assuming Lorentzian lines, was obtained from the relationships of the 129Xe NMR parameters (chemical shift, linewidth, intensity) with the hydrocarbon concentration determined at equilibrium from preliminary experiments [8].

Time(h)

/t

8.0~

~

t/xi,~,

ff ~ ~

7.0

Time (h)

A

44

3.0

t/'l:intra

J

k

_ _~J~

28

1.95

W~.~

0.90

2.3 2.1

~.~J"

1.4

~

0 I

180

~ '

I

150

| | Mx....__

2.0

! I~

1.0

kX,~ '

I

120 8 (ppm)

L___

0 '

I

90

i

180

!

'

160

|

i

i

140

i

120

i

i

100

!

i

80

8 (oom)

Fig. 3. Experimental and simulated 129Xe NMR spectra during benzene (a) and paraxylene (b) adsorption in HZSM-5 ofbedlength 5 and 13 mm respectively.

129Xe NMR spectra are simulated from the adsorbate concentration profiles. The fit of experimental with calculated spectra (Fig. 3a) using the adjustable parameter, 1;intra, allows the determination of the intracrystalline diffusion coefficient of benzene in HZSM-5 zeolite during its adsorption under constant (saturation) pressure, Dintra = 7x10 -15 m2s -1. This value agrees with that obtained with other techniques and reported in literature [4]. The equilibrium time, too, of 8 and 13 h, depends on the bed length, 5 and 15 mm, respectively. The calculation

379 of "lTinter (according to eq. 1) gives 2.0 and 12.3 h. The comparison of too and "l;inter shows that diffusion in macropores is the limiting step for the 15 mm sample while both the diffusion in micro- and macropores controls the whole process for the 5 mm sample. For a given length, when the powder is compressed, the equilibrium time is larger due to the decrease in the size of the intercrystallite spaces (the tortuosity increasing) but a similar value of Dintra is obtained. We also investigated the paraxylene/HZSM-5 system with the same technique. Due to the lower intracrystalline diffusion rate the "non uniform" model was not able to simulate the NMR spectra correctly. A simpler "core shrinking" model was used [9]. The crystallites are divided into two zones: a core free of diffusing molecules and a shell with a uniform hydrocarbon concentration. During the adsorption there is a diffusion front in the crystallites, the shell region increases at the expense of the core. Then, the NMR spectra is simply the sum of two lines whose intensity inversely varies with time (Fig. 3b). The simulation gives a Dintra value of 3x10 q6 m2s "l which is in good agreement with the literature [4]. 3.2.1H NMR imaging 3.2.1. Single reactant The 1D images of benzene adsorbing in the HZSM-5 bed have rapidly a rectangular shape for the loose as well as for the compressed powder (Fig. 4a and 4b) proving that the benzene concentration is rapidly the same in any part of the bed and that the diffusion of hydrocarbon is controlled by the micropores as discussed by Heink et al. [10]. The compaction of the powder increasing the density of the sample, the NMR signal is better defined for the compressed sample. The signal intensity increases with adsorption time. Assuming that the benzene concentration in the gas phase is negligible, the total amount, M(t), of benzene adsorbed in the sample is directly proportional to the integral of the NMR profiles. The experimental kinetic curves M(t) are simulated using the solution of the diffusion equation, obtained by Cranck [ 11 ], applied to this system and its particular boundary conditions [ 12]:

3,

EJ 2 / 21 62expDin,an2,2tJ,2,

Mt =1 .... ~d~Taexp(-kt) 1"M oo

D-~n~aCOt Dintra + ~ (5) ~2Dintra ? =1 n 2 (n2n 2 - kR2 / D.mtra )

where k is the equilibrium constant of adsorption on the external surface of the crystallites. The simulation gives a Dintra value of about l x l 0 14 m2sl which agrees well with that obtained by 129Xe NNIR. The small (several orders of magnitude are often found in the literature) difference may be due to the absence of xenon in the NMR imaging experiment (no restriction to hydrocarbon diffusion) or simply to an imprecision of the simulation. When the zeolite is mixed with silica-alumina (60% weight of zeolite), the benzene concentration profiles present an adsorption front which lasts more than 30 min. On one hand the intercrystalline diffusion rate is decreased by the presence of the silica-alumina which has a mesoporosity centered around 50/~ (pore diameter); on the other hand benzene also adsorbs on the silica-alumina. This result shows the important influence on the mass transport of the binder in industrial catalysts. In the case of n-hexane adsorption on powder sample, the concentration profiles (not shown) become perfectly rectangular after only 2 min (almost 7 min for benzene). The value

380 obtained for Dintra (about 10 -13 m2s -1) has the same order of magnitude as that reported in literature using other techniques (zero length column, frequency response...) [4].

b

Temps (h)

C ~ (h)

J/-

--"-

/~\~--~--

Jr

4.25

~ \

12.oo

s.oo

~k...---- 5.40

J

,600 90o

2.25

Temps(h)

4oo o6o

J

~._j ~ - - ~ ~ - -

1.67

0.50

j

]

................................ .~....................................................... u............................

z:---e

z--':0

0.00

0.83 0.58 0.25

---~j r ~

.

;9 z-e

~ .

.

.

~ . ~ -

0.15 0.08

i z=0

Fig. 4. Concentration profiles in a loose (a), compressed (b) powder or in a mixture (zeolite / silica-alumina binder) (c) during benzene adsorption 3.2.2. Mixture of two reactants Different types of experiments were performed with n-hexane and benzene. Here we report the results conceming, first, their competitive adsorption, then the desorption of benzene when deuterated n-hexane was adsorbing. When the adsorption of n-hexane proceeds from a gas phase mixture of n-hexane and deuterated benzene (both under the same partial saturation pressure), the intensity of the profiles first decreases from the top to the bottom of the bed; but at intermediate time there is a slight excess of n-hexane in the bottom of the bed before a homogeneous distribution all over the bed is reached at equilibrium (Fig. 5a). As the intercrystalline diffusion of n-hexane is slower than in the pure gas phase experiment (due to the presence of benzene and a higher total pressure), the molecules adsorb in the first layers before the n-hexane pressure becomes equal all over the bed. The benzene molecules whose intracrystalline diffusion is much slower than that of'n-hexane, progress along the bed and adsorb then on the subsequent layers whose crystallites are free of any adsorbate. As time increases, benzene also adsorbs on the first layers, displacing n-hexane towards the bottom where n-hexane displaces in turn the benzene molecules to finally obtain a distribution governed by thermodynamics. The intermediate states of the system are the result of both competiting kinetic and thermodynamic effects, the diffusion of n-hexane being faster than that of benzene while this latter is more strongly adsorbed.

381

a

b

~

time (min)

~

522 305 260 61 28

time (min)

~N~'/]/,~~

~-,-._-.,.-..-~ 353 275

----~~/~ ~ /

I~'------.---.-.----~/ ' ~ k . ~

34 19

8

7 ,

top

,

4

........ top

bottom

i

bottom

Fig. 5. Concentration profiles of n-hexane (a) and benzene (b) during their competitive adsorption in HZSM-5 zeolite. This scenario is confirmed by the evolution of the benzene concentration along the bed recorded during the same experiment in which we used protonated benzene and deuterated nhexane (Figure 5b). These profiles clearly show that benzene first adsorbs preferably in the bottom layers before displacing the n-hexane adsorbed in the top layers and finally the equilibrium is reached all over the sample.

me (h~ 0 0.02 0.09 0.17 0.31 1.03

1.51 4.77 16.06 too

bottom

Fig. 6. Benzene concentration profiles when n-hexane is adsorbing.

In case of benzene being preadsorbed at equilibrium under 6.4 mbars and deuterated n-hexane adsorbing under its saturation pressure (Fig. 6), we can initially observe a rectangular profile corresponding to a uniform distribution of benzene (notice that time decreases from top to bottom of the Fig. 6). The variation of the signal (intensity decreases at the top and strongly increases at the bottom of the bed) shows that the nhexane adsorbs first in the upper layers; the benzene is pushed out and its concentration decreases in the top region. These benzene molecules then move preferably towards the bottom of the tube since the total pressure is lower than at the top. The local partial pressure of benzene and, in parallel its concentration in the bottom of the bed increase.

382 As the "wave" of n- hexane reaches the bottom this latter gas adsorbs on the lower layers; the two partial pressure becomes uniform along the sample, the benzene molecules can adsorb again in the upper layers until the thermodynamic equilibrium is obtained. The distribution of adsorbed gases is first determined by kinetics and then the system is governed by thermodynamics. The total amount of adsorbed benzene has decreased with respect to the initial state (about 30%). Quantitative determination of the diffusion coefficients are in progress. 4. CONCLUSION

129XeNMR spectroscopy of adsorbed xenon, largely used to investigate static properties of porous solids, appears very useful to study the diffusion of coadsorbed molecules when the local concentration of these molecules changes as for example during the adsorption process. Coefficient of intracrystalline molecular transport can be obtained from the simulation of the NMR spectra using the solutions (adsorbate concentration profiles) of the diffusion equations. The ~H NMR imaging directly gives instantaneous concentration profiles of hydrogencontaining molecules such as hydrocarbons along one or more directions. The changes of these profiles during the adsorption or desorption allow to determine transport coefficients by simulation of the kinetic curves (adsorbed amount in function of time) obtained from the variation of the total signal intensity. The shape of the signal reflects the variation of the local adsorbate concentration which may be drastically affected by kinetic effects. REFERENCES 1.

J.-L. Bonardet, J. Fraissard, A. Grdron and M.-A. Springuel-Huet, Catal. Rev.-Sci. Eng., 41(2) (1999) 115 2. M.-A. Springuel-Huet, A. Nosov, P. Ngokoli-Kekele, J. K~irgerl J. M. Dereppe and J. Fraissard, in "Fluid Cracking Catalysts", M. Occelli and P. O'Connor eds., Marcel Dekker Inc. Publi., page 191 (1997) 3. J.-L.Bonardet, T. Domeniconi, P. N'Gokoli-Kekele, M.-A. Springuel-Huet and J. Fraissard, Langmuir, 15 (1999) 5836 4. J. K~irger and D. Ruthven, Diffusion in Zeolites and Other Microporous Solids, J. Wiley & Sons, New York, 1992 5. E. Ruckenstein, A. S. Vaidyanathan and G. R. Youngquist, Chem. Eng. Sci., 26 (1971) 147 6. P. N'Gokoli-Kekele, PhD thesis, Paris 7. J. K~irger, H. Pfeifer, F. Stallmach and H. Spindler, Zeolites, 10 (1990) 288 8. P. N'Gokoli-Kekele, M.-A. Springuel-Huet, J.-L. Bonardet and J. Fraissard, Stud. Surf. Sci. Catal., 130 (2000) 2939 9. B.F. Chmelka, J. V. Gillis, E. E. Petersen and C. J. Radke, AIChE J., 36 (1990) 1562 10. W. Heink, J. K~irger and H. Pfeifer, Chem. Eng. Sci., 33 (1978) 1019 11. J. Krank, The Mathematics of diffusion, Clarendon Press, Oxford, U. K., 1956 12. T. Domeniconi, P. N'Gokoli-Kekele, J.-L. Bonardet, M.-A. Springuel-Huet and J. Fraissard, Proc. 12th Inter. Zeolite Conf., M. M. J. Treacy et al. eds, Materials Research Society Publ., Warrendale, PA, USA (1998) 2991

Studies in SurfaceScienceand Catalysis133 G.F. Fromentand K.C.Waugh(Editors) 9 2001 ElsevierScienceB.V. All rightsreserved.

383

Catalytic activity of carbon nanotubes and other carbon materials for oxidative dehydrogenation of ethylbenzene to styrene N. Maksimova, G. Mestl and R. Schlrgl 1

Fritz-Haber-Institut der Max-Planck-GeseUschaft, Faradayweg 4-6, 14195, Berlin, Germany

Abstract

Different carbons, i.e. carbon black, graphite and multiwalled nanotubes were tested as catalysts for oxidative dehydrogenation of ethylbenzene to styrene. The influence of the catalyst structure was determined on the catalytic activity and selectivity, and the catalyst stability in the oxidative dehydrogenation of ethylbenzene. In comparison to carbon black and graphite, multiwalled carbon nanotubes were found to exhibit the highest performance in the reaction. Carbon nanotubes as catalyst gave the highest styrene selectivity at maximum conversion. The characterization of the carbon nanotubes and the other carbon materials before and after the catalytic tests was conducted by several techniques, e.g. TEM, XRD and BET area measurements. This comparative characterization of the fresh and tested catalysts indicates that structural modifications in the nature of the catalysts occurred during the reaction. These modifications could be correlated with the changing catalytic behavior with time-on-stream.

1. INTRODUCTION

The catalytic dehydrogenation (DHYD) of ethylbenzene (EB) to styrene (ST) is the major industrial process for the styrene production [1]. The industrial process is usually realized in the temperature regime between 550-620~ with an excess of overheated water vapor mainly over a potassium promoted iron oxide catalyst [1 ]. Because this process is limited by the thermodynamic equilibrium of the reaction and because it is very energy

ITo whomcorrespondenceshouldbe addressed, e-mail [email protected]

384 intensive, there is a strong incentive for the development of alternative technologies. The oxidative dehydrogenation (ODH) of EB is promising due to the realization of an exothermic reaction. Such a process operates at lower temperatures as compared to DHYD, but the presence of oxygen could lead to lower selectivity and possibly to catalyst degradation, e.g. burn-off. The reported observations that (i) different, mainly oxide based compounds, which tend to coke deposition, are active for ODH of EB to ST, and that (ii) a characteristic induction period of several hours, during which the coke deposition occurs, correlates with an increase of the catalytic activity [2], may indicate that the carbon deposited on the catalyst surface plays an important role in the styrene formation. It was also reported that amorphous carbon activated by oxidation is an active catalyst for the ODH of EB to ST [3]. On the other hand, great interest was paid to carbon nanotubes during last decade due to their stability at high temperatures, severe environments, and the possibility to modify them by metal introduction [4-6]. It may be assumed that pure carbon nanotubes, or nanotubes filled with Fe, could be active and stable catalysts for the ODH of EB to ST. In order to test their catalytic activity and to develop a deeper understanding of the relation between carbon structure and its catalytic activity, different types of corbons were used for the ODH of EB to ST.

2.

EXPERIMENTAL Carbon black (Lamp soot, Degussa), natural g r a p h i t e - "AF special"

(Kropfmiihl), and commercial (Applied Science) as well as multiwalled nanotubes (MWNTs), made by the decomposition of ethylene over Fe or Ni catalysts, were used in this study. The dehydrogenation of EB to ST over these different carbons was conducted in tubular quartz reactors (4 mm i.d. x 200 mm) in the temperature regime between 450 and 550~

0.02 g of the catalyst particles were held in the isothermal zone of the reactor

between two plugs of quartz wool. The EB and He was mixed with 02 and fed to the reactor (EB : He : 02 = 9.10 -3 : 100 : 1), to give a total gas flow of 10 ml/min resultung in a LHSV of 0.5 h-1. The analysis of aromatics were done by one-line gas chromatography

385 using 5%SP-1200/1.75% Bentone 34 packed column and FID detector. The permanent gases were analyzed simultaneously using a Carboxen 1010 PLOT column and a TCD detector. For the comparison the same test experiments were conducted for the dehydrogenation without the presence of oxygen. The physicochemical characterization of the carbon nanotubes and the other carbon materials prior and subsequent to the catalytic tests was done by HRTEM (Phillips CM200, equipped with a field emission gun operated at 200kV), XRD (STOE STADI P with CuKa-radiation, Ge monochromator and a curved PSD) and BET area measurements.

3. RESULTS

Figure 1 shows a comparison of the evolution of the catalytic activity over the different carbon catalysts in the anaerobic dehydrogenation of EB to ST with time on stream at 550~

An induction period was observed of about 1.5 hours, during which all

carbon materials have shown different behavior. The highest activity after having reached steady state was determined for carbon nanotubes as the catalyst with a ST yield of 28%.

Styrene yield 100 90

=

80 70

~

c a r b o n black

9

graphite carbo~ nanotubes

60 % 5o

40 30 L 20 10 0

i

,..~/ 0

I

~

20

1, I

~

40

I

~ & ~ &

.

~

=~

60

~ I

80

I

I

&

&~&~&__& ===in

I

I

&__&

. ~ l p = = =ll ~ai ~ , I

I

I

I

I

100 120 140 160 180 200 220 240 260

Time on stream min Fig. l. Comparison of the evolution of the ST yields with time on stream over carbon black m, graphite ST at 550~

, and carbon nanotubes

in the anaerobic dehydrogenation of EB to

386 Figure 2 shows a comparison of the evolution of the catalytic activity over the different carbon catalysts in the ODH of EB to ST with time on stream at 550~

Again an

induction period of about 1.5 hours was observed during which all carbon materials exhibited different behaviors. The highest activity was found again for carbon nanotubes as the catalyst, after having reached steady state operation in ODH, with a ST yield of 58%.

loo~

:Oto

--=--- carbon black --=--- graphite - - A - - carbon nanotubes

60

.A_._.---&~A~&.._----&---._._&_._...-A~_&

.

%

50

i ~ l l - - - - - . ~ . l l ~ ii_._.__i._-----II~ll~ll

10" O"

20

40

60

80

100 120 140 160 180 200 220 240 260

Time on stream, rain

Fig. 2. Comparison of the evolution of the ST yields with time on stream over carbon black i , graphite

, and carbon nanotubes

in the ODH of EB to ST at 550~

Table 1 shows the dependence of the catalytic activity of the carbon nanotube catalyst with the reaction temperature at steady state operation. Catalytic activity for the anaerobic dehydrogenation of EB was observed for carbon nanotubes in the temperature range between 450 and 550~

The highest activity for oxidative dehydrogenation of EB to ST

was observed at 550~ . With increasing temperature, the ST yield increased, but the selectivity to ST decreased as can be expected from thermodynamics (Tab. 1). Figure 3 shows the evolution of the catalytic activity of carbon nanotubes for ODH of EB to ST with time on stream at 550~

After an induction period, during which the

387 activity increased, an ST yield was reached of 52+5%. The EB conversion was at 64+5%, with a selectivity to ST of 78+5% (Fig. 3).

Test reaction

Temperature,~

Conversion of EB, %

ST yield,%

450

1.1

0.86

500

21.3

21

550

12

11.4

450

58.3

57.25

500

75.7

70

550

67

63.42

Dehydrogenation

Oxidative dehydrogenation

Table 1. Catalytic activity of carbon nanotubes in anaerobic dehydrogenation and ODH of EB to ST at different temperatures.

80t~&/ ~n[

l~n--m~m--m--m--m--m--m

4O 2O |

0

.

i

30

9

I

,

i

.

i

.

i

.

| . l

m

60 90 120 150 180 210 Reaction time. min

Fig.3. Evolution of EB conversi0nl, ST yield

, and ST selectivity

i

with time on stream

for the ODH of EB to ST over carbon nanotubes at 550~ Fig. 4 shows an HRTEM overview of the fresh nanotube material. The comparative HRTEM characterization of the fresh and tested catalysts, shown in Fig. 5, indicates that some modifications of the carbon nanotube surface occurred during the reaction (Fig. 5, a and b). Fig. 5, a shows that the walls of the carbon nanotubes before the reaction consisted of two layers, an inner layer formed by conical graphene sheets and a thick outer layer of amorphous carbon parallel to the tube axis. After 20 hours time on stream in the ODH, the stacks of the inner, conical graphene layers can now be recognized well. The thick, outer

388 amorphous layer has almost disappeared (Fig. 5, b) but, the walls of the nanotubes were still covered by a thin amorphous carbon layer.

Fig. 5. TEM image of the wall of carbon nanotube prior (a) and subsequent (b) to the ODH reaction: dashed lines in (a) indicate the directions of the carbon layers; black arrows in (b) indicate the amorphous carbon layers. BET analysis showed that the specific surface area of the initial carbon nanotube material was 26 m2/g, while it was 47 m2/g subsequent to the ODH reaction. This increase in the specific surface area has to be related to the HRTEM observation of an altered surface layer of the nanotubes. Therefore it is proposed that the observed increase of the

389 catalytic activity of the carbon nanotube catalyst with time on stream is correlated with this changed surface properties.

4. DISCUSSION AND CONCLUSIONS

From the comparison of the evolution of the catalytic activity and the results obtained by the TEM, XRD, and BET analyses of the carbon samples prior and subsequent to the ODH of EB, it can be stated that the catalytically active carbon species are formed during the reaction on the carbon surface. Carbon black is very active from the beginning of the reaction. During the induction period, a part of the active sites is removed as indicated by the decreasing catalyst activity. This decrease could be correlated with the bum-off of amorphous carbon. The activity of graphite was almost constant during the reaction. The active sites on the graphite surface, hence, are stable under reaction conditions. In case of carbon nanotubes for anaerobic and ODH of EB to ST, the activity of the sample was initially low, but increased during the induction period. This was correlated with the increase of the BET surface area. Carbon nanotubes have shown the highest ST yields at the highest EB conversions as compared to carbon black and graphite. The nature of the active carbon species is unclear to date, but the structure of the underlying carbon has to play an important role in its formation and its stability, as proven by the catalytic tests. In the catalytic reaction, the carbon surface most probably acts as a donor of the active oxygen species. To date unidentified surface oxygen groups, e.g. carbonyl, carboxylic, or quinone groups, may react with EB under the abstraction of hydrogen and the formation of ST and water.

REFERENCES

1. 2. 3. 4. 5.

F. Cavani, F. Trifiro. Appl. Catal. A: General, 133 (1995) 219. T.G. Alkhazov, Kinet. Katal., 13 (1972) 509. M.F.R. Pereira, J.J.M. Orfao, J.L. Figueiredo, Appl. Catal. A: General, 184 (1999) 153. S. Iijima, Nature, 354 (1991) 56. K. Hemadi, A. Fonseca, P. Piedigrosso, M. Delvaux, J.B. Nagy, D. Bemaerts, J. Riga, Catal. Lea., 48 (1997) 229. 6. O.P. Krivoruchko, N.I. Maksimova, V.I. Zaikovskii, A.N. Salanov, Carbon, 38 (2000) 1075. 7. G. Emig, H. Hofmann, J. Catal., 84 (1983) 15.



Heterogeneous applications

catalysis

391

at

high

temperature

for

space

M. Balat-Pichelin and J.M. Badie Institut de science et de g~nie des Mat~riaux et Proc~d~s, CNRS-IMP, UPR 8521, BP 5, 66125 Font-Romeu Odeillo (France) Ph. : 33 4 68 30 77 68 Fax : 33 4 68 30 29 40 e-mail : [email protected] During the atmospheric entry phase of space vehicles, one of the main physicochemical phenomena acting on some hot parts of the vehicle walls is the atomic oxygen recombination leading to possible important excess of heating and damage of the protective materials. The aim of this work is the evaluation of the catalytic recombination of atomic oxygen under conditions nearby Earth atmospheric entry using thermal and chemical approaches in order to obtain the weakest catalytic material. 1. INTRODUCTION The most important conditions for the simulation of atmospheric entry of space vehicles (high temperature, low pressure plasmas) have been realized in the MESOX set-up associating a solar radiation concentrator and a microwave generator. On one hand, the study of the atomic oxygen recombination on partially catalytic based-silicon or-aluminum ceramic materials, at high temperature has been performed at different pressures by a thermal approach (macroscopic) and leads to a catalytic scale of materials. On the other hand, a chemical approach (microscopic) is developed for the evaluation of the recombination coefficient 7 using atomic emission spectroscopy (actinometry) on the same device and in the same temperature range but only at 200 Pa total air pressure. Finally, the activation energy of atomic oxygen recombination can be reached for all the ceramics. 2. EXPERIMENTAL SET-UP The MESOX set-up which associates a solar radiation concentrator for the sample heating (up to 2300 K under pressures comprised between 102 and 105 Pa) and a microwave generator (2450 MHz, 1200 W max), originally developed for oxidation studies has been adapted to allow catalytic recombination measurements. Atmospheric entry conditions can be partially simulated, pressure and temperature can be reproduced independently with a high accuracy.

392 The experimental device is placed at the focus of a solar furnace equipped with a variable opening shutter. It can be moved away from the focus to be replaced by a calorimeter to measure the incident solar flux. The available incident concentrated solar flux can reach 4.5 MW.m -2. The temperature measurements on front (Tf) and back (Tb) faces of the sample are realized using a single optical pyrometer (5 ~tm) with a system of one rotating mirror and two stationary mirrors. The experimental reactor consists of a silica tube, 50 cm length and 5 cm diameter with CaF2 viewports. This reactor crossing the refrigerated wave guide contains the sample (25 mm diameter and 3 mm height) placed on a sampleholder in stagnation point position at the center of the discharge. The microwave generator works at a constant power of 300 W. A regulator, a gauge and a vacuum pump are used to control precisely the total pressure during the experiment. 2.1. Recombination thermal flux We have developed a thermal approach based on a heat balance on a reference cylinder volume in the sample to evaluate the atomic oxygen recombination. The surface of this cylinder, 6 mm diameter and 3 mm height, represents the measurement area by pyrometry. Convective phenomena are neglected (rarefied gas flows) compared to the radiative fluxes to establish the equations for steady state heat transfer under different environments. The thermal balance is established under each atmosphere " standard air, air plasma, standard argon and argon plasma, the determination of the recombination flux being done by difference between the experiments under air plasma (reactive) and argon plasma (inert) [1, 2]. Under some assumptions precisely described in [2], a like-one dimensional equation is obtained " (1) [~.q~r = qrad,f air - qrad, farg + qrad,b air - qrad,barg equivalent to : ~ . q r e c air " e O [(TP~) 4 - (T~rg)4 + (Tbair) 4 - ( T b a r g ) 4] (2) The absorbed recombination flux ~ q~ec (~ is the accomodation coefficient) can be calculated from the following parameters : e the total hemispherical emissivity and Tf, Tb the front and back face temperatures under air and argon plasmas. The uncertainties A ~qreJ~ q~ec have been calculated taking into account the errors on temperature measurements due to the accuracy of the optical pyrometer (0.5%), on the spectral emissivity (at 5 ~tm, 1 to 2% depending on the materials) and on the total hemispherical emissivity (1%).

2.2. Experimental results The pressure and the flow rate are fixed at the beginning of each experiment. Three total gas pressures have been applied 9200, 1000 and 2000 Pa, for five temperature levels between1000 and 1800 K. The surface temperature depends on the incident solar flux controlled by the opening of the shutter. This study has been realized on several sintered ceramic materials (SIC, Si3N4, AIN, A1203) or on samples oxidized at 1300 K during 24 hours (SiC + Si02, AIN + A1203) and on composite ceramic material. The determination of the recombination flux confirms the tendencies obtained with the temperature increase. Silicon-based

393 ceramic materials have similar and low catalytic activities. Sintered SiC has the weakest recombination flux which is nearly constant between 1000 and 1600 K (around 30 kW.m-2), except under active oxidation at 1800 K (13 kW.m -2) [3]. This confirms its very low catalytic activity, the higher values obtained in the case of the silica layer (SiC + Si02) between 1000 and 1600 K being due to the presence of this oxide layer (~ cristobalite). Aluminum-based ceramic materials have more pronounced catalytic activity and sintered alumina is by far the more catalytic material with an increasing recombination flux which reaches 151 kW.m -2 at 1400 K, before decreasing at 1800 K (65 kW.m2).

Table 1 Recombination fluxes (experimental data and modeling) on some ceramic materials at high temperature for 200 Pa total air pressure. SiC

T (K)

S i C + SiO~

~.q~r (kW.m -2)

T (K)

~.q~r (kW.m-2)

Al20~

T (K)

Exp.

~.qrec(kW.m -2)

Exp.

Mod.

Exp.

Mod.

984 + 3

30 + 3

34

966 9 3

39 + 4

977 + 6

97 + 16

109

1191 + 5

29 + 3

31

1182 + 5

35 + 3

1167 + 9

139 + 24

153

1377 + 7

30 9 3

32

1379 • 7

33 • 3

1401 • 11

151 • 29

166

1590 + 9

25 + 3

26

1566 • 9

38 + 4

1577 + 16

108 + 23

120

1780+11

13+2

15

1770+11

39•

1710+21

65+15

-

2.3. B i - d i m e n s i o n a l m o d e l A bi-dimensional model has been developed in order to estimate the radial thermal losses by conduction at the sample edges which have been neglected in eq. (1). This model is based on the simulation of axial and radial heat transfers in material at high temperature. The bi-directional heat equation is solved with a finite difference scheme applied to a cylindrical shape [4]. The resolution of the equations system leads to the recombination flux determination. Modeling values are given in Table I for sintered SiC and A1203. The results of the modeling agree with the experimental values within the measurement accuracy ; the modeling values are a little higher due to the difference between conductive radial losses which is a positive value. Moreover, the discrepancy between modeling and experimental values for the recombination flux is less important at higher temperature levels (1600-1800 K). Finally, the studied materials have been classified according to a catalytic scale (for thermal flux transferred to the surface) available between 1000 and 1800 K 9 SiC ~ SiC+SiO2 ~ Si3N4 ~ A1N ~ A1N+A1203 -

0~, 0

=

1

I

I

2 Time

A

5

(h)

4-20 (mi,ma)

6-20 (mi,ma)

O mi (KOH) V mi (NaOH) r-! ma (KOH) OH)

t-

tO

I

4

2-20 (mi,ma)

2

N t.O

I

3

1

O t,-

a

0

I

I

I

,,

I

I

I

I

I

I

2 c~

o

o .c; ._o) 11) I

1

I

3

I

5

b

I

I

1

I

3

I

5

I

1

I

3

I

5

Base (M)

Fig 3 Effect of salt concentration for the imbibed weight of chlorobenzene and water for various resin: chlorobenzene = 25 cm 3, aqueous solution = 25 cm 3, resin (45-60 mesh) = 1g, 250.


Studies in Surface Science and Catalysis 133 G.F. Fromentand K.C. Waugh (Editors) 9 2001 Elsevier Science B.V. All rights reserved.

435

Catalyst Development for Methanol and Dimethyl Ether Production from Blast Furnace Off Gas Jun-ichiro Yagia, Tomohiro Akiyama b and Atsushi Muramatsu a alnstitute for Advanced Materials Processing, Tohoku University 2-1-1, Katahira, Aoba-ku, Sendai, 980-8577 Japan. bDept. Chemical Engineering, Osaka Prefecture University 1-1, Gakuen-cho, Sakai, Osaka-fu, 599-8531 Japan For C u - Z n O - A 1 2 0 3 catalyst, compositional research was on two parameters, namely Cu/ZnO ratio (3/7-5/5) and content of A1203 (0-33.0 mol%). The influence of each parameter was estimated by DME and MeOH yields for the same catalyst mass. Some Cu-ZnO-AI203 catalysts synthesized DME more than MeOH, in which the DME activity was related to specific surface area and existence of broadened ZnO peaks in XRD pattems. In contrast, three other Cu-ZnO catalysts including another oxide (Cr203, ZrO2 or Ga203) synthesized only MeOH without DME. Content of A1203 was more influential on DME synthesis than Cu/ZnO ratio. Physically-admixed (hybrid) catalysts of Cu-ZnO-Ga203 for CH3OH synthesis and y-AI203 for its dehydration were experimentally studied, in which the influence of mixing ratio of the two catalysts on both yield and selectivity of DME was mainly examined for the same catalyst mass. The results showed that the developed hybrid catalyst is very effective in producing DME directly from BFG without equilibrium limit of methanol. Interestingly, the yield of DME had a significant dependence of mixing ratio, and the hybrid catalyst with only 5 mass% ~/-A1203 showed the highest yield with 99.3%selectivity of DME + methanol. This implies that methanol formation governs the rate of this series reaction (BFG-->CH3OH-->DME) due to fast dehydration of methanol to DME. In conclusion, the most active composition of Cu-ZnO-A1203 catalyst for DME synthesis was Cu/ZnO=4/6 with 14.3mo1% A1203. However catalysts mixture of C u - Z n O - G a 2 0 3 ( 9 5 % ) and ~,-A1203 (5%) showed higher activity.

Key Words: methanol, dimethyl ether, catalyst, blast furnace, energy, industrial waste gas, CO-CO2-H2, Cu-ZnO-A1203, 7-A1203. 1. INTRODUCTION Methanol (CH3OH, MeOH) is now attracting worldwide attention because it is not only an important industrial raw material for chemicals such as formaldehyde 1), but is also a clean fuel which burns without emission of NOx and SOx. Previously, Feasibility studies 2-5) of MeOH production from blast furnace offgas (BFG) have been carried out theoretically and experimentally for energy saving and environmental protection due to reduction of carbon dioxide emission. Since composition of BFG is quite different in comparison to industrial

436 synthesis gas (syngas) for MeOH synthesis, it leads to serious thermodynamic restriction of MeOH yield due to less hydrogen concentration and higher carbon dioxide concentration. Therefore, we need to overcome this problem to promote the effective use of BFG. Dimethyl ether ((CH3)20, DME) is also expected to be a clean energy source with large calorific value and excellent transportation properties, almost same as LPG 6). Industrially, DME is generally produced in a two-step process 6), namely, MeOH formation and its dehydration. It should be pointed out that its equilibrium yield is far beyond that of MeOH. Therefore the use of a bifunctional catalytic system7), that is, a combination of a MeOH synthesis component with a dehydration partner s 12) can avoid the equilibrium limit of MeOH. Direct synthesis of DME from syngas (CO+H2) has been reported in the literature. 7' 12, 13) The method has been directed so far, to only the typical syngas; carbon monoxide/hydrogen with or without a small content of carbon dioxide, but not yet to BFG with a large carbon dioxide content. In addition, influence of the mixing ratio in the hybrid catalyst has not been well investigated in spite of its importance for the process design and operation. Therefore, the aim of this study is to determine an optimum composition of Cu, ZnO and A1203 as high activity catalysts for high yield of oxygenates (MeOH + DME) from BFG, in which Cr203, Zr02 or Ga203 were also examined instead of A1203 and to investigate catalytic activity and mixing ratio of the hybrid catalyst which is a physical mixture of Cu-ZnO-Ga203 and y-A1203. 2. MATERIALS For synthesizing DME from COCOa-Ha mixtures such as BF gas, we prepared catalysts of Cu-ZnO-X (X= A1203, Cr203, ZrO2, Ga203) by the coprecipitation method. Chemical composition of the catalysts prepared is given in Table 1. Twelve catalysts of the Cu-ZnO-A1203 system were prepared, in which additive content of alumina was systematically changed under three different ratios of Cu/ZnO. In addition, three catalysts with different oxides (Cr203, ZrO2, Ga203) instead of alumina were also prepared for comparison. The preparation procedure of the catalysts was reported in detail elsewhere. 3' 5,6,) For the Cu-ZnO-A1203 catalyst, the procedure framework is as follows; (1) Three commercial chemicals of Cu(NO3)2"3H20, Zn(NO3)2"6H20 and Al(NO3)3"9H20 are first dissolved in distilled water at room temperature, to

Table

1. Chemical composition of fifteen catalysts used.

Cu/ZnO/A 1203 system Cu/ZnO A 1203 mol 6.7 14.3 23. 1 33.0

Composition, mol% 30.0/70.0/0.0 28.0/65.3/6.7 25.7/60.0/14. 3 23.1/53.8/23.1 20. 0/47.0/33. 0

4/6

0.0 6.7 14.3 23. 1

40. O/GO.0/0. 0 37. 3/56. 0/6.7 34. 3/51.4/14. 3 30. 8/46.1/23. 1

5/5

6.7 14.3 23. 1

46. 7/46. 7/6.7 42.9/42.9/14.3 38. 5/38.5/23. 1

3/7

0.0

Cu/ZnO/Oxide (Cr203, ZrO2, 6a203) system Cu/ZnO Oxide mol~ ] Composit ion, mol~ 3/7 14. 3 I 25. 7/60.0/14. 3

437 obtain aqueous solutions of 1 mol/dm 3 with the desired molecular ratio of the metallic elements. (2) The obtained solution and 0.7 mol/dm3 sodium carbonate (Na2CO3-10H20) solution are mixed at 353 K all at once, to precipitate the carbonates of Cu, Zn and A1. The precipitate, recovered by filtration, is compressed to from a cylindrical sample and then sintered at 673 K after sufficient drying at 393 K. (3) The sinter is crushed in a mortar and particles of size 0.35-0.84 mm are separated by sieving for packing into a reaction tube. In preparing hybrid catalyst, two catalysts of CuZnO-Ga203 and y-A1203 are admixed at different mass ratios and then fixed into a reaction tube. (4) The packed catalyst is finally reduced by pure hydrogen at 653 K just before the experiments. 3. METHOD

Figure 1 shows a schematic diagram of an experimental system for DME synthesis, consisting of three parts: gas supplier, fixed-bed reactor and gas analyzer. The molar ratio of feed gas excluding nitrogen, CO/CO2/H2=4/4/1 (molar basis), corresponds to operating data 2) of a BF with natural gas injection of 50 Nm 3 per ton-hot-metal. Nitrogen was added to the feed gas to avoid CO2 liquefaction in the cylinder so the composition of the actual feedgas is CO/CO2/H2~2 =35/35/8.7/21.3. After the feedgas was introduced, outlet gas from the fixedbed reactor was analyzed by two kinds of gas chromatography: FID (Flame Ionization Detector) for DME, MeOH and hydrocarbon (HC, mainly methane), and TCD (Thermal Conductivity Detector) for CO, CO2, H2 and N2.

438 An enlargement of the reactor is shown in Figure 2 The reactor employed was made of SUS316 stainless steel tube with 10 mm outer diameter and 8 mm inner diameter. The prepared catalyst was fixed at the middle position of the uniform temperature zone, on which glass beads were placed for preheating the feed gas and for obtaining radically uniform flow. To measure temperature changes of the catalyst during the experiment, a CA type thermocouple, 1 mm in outer diameter, was inserted into the bed. The experimental conditions of this study were 523 K, 1.0 MPa inlet pressure and 1.0xl0 -6 m3/sec (STP) flow rate of the feed gas. These were determined based on our previous studies 3' 4). The measurements were continued until outlet gas composition reached a steady state. In the experiment for finding optimal mixing ratio of 7-A1203 to the hybrid catalyst, mass ratio of ~/-A1203 was changed from 0 to 100% keeping the hybrid catalyst mass 2.0g. Since the densities of 7- AlaO3 and Cu-ZnO-GazO3 are considerably different, the height of the hybrid-catalyst bed depends on their mixing ratio, resulting in much different residence times. Therefore, not only the catalyst mass(2.0g) but also the residence time have to be considered in the unit of C-mol%/g-cat s, for quantitatively evaluating yields of DME, MeOH and HC. They are defined as follows: 2CI:,MV. CO + c co2

1 I00

CMr

:(c0 co

1

0 + Cco2 CHC

-

(1)

W x (L / u )

I

(C~ + C~ )/ lo0 • w • (L / u)

(a) (3)

4. RESULTS AND DISCUSSION All monitored components of the outlet gas (CO, CO2, He, MeOH, DME, HC, N2) reached constant values with negligible further temperature changes after approximately 6 hours. The obtained concentrations of MeOH and DME for each catalyst were then defined as their final yields. Figure 3 shows bar graphs of evaluated yields of DME and MeOH, for all of the catalysts prepared. Here, Cu/ZnO ratio in the catalysts in the horizontal direction of the graph is constant and the three rows of the graph are for different Cu/ZnO ratios: 3/7, 4/6 and 5/5. The bar height indicates the DME and MeOH yields. Obviously, some catalysts of Cu-ZnOA1203 gave good yields of DME with little MeOH synthesis although only MeOH is normally synthesized from CO-H2 gas as mentioned later. The results suggest significant influence of A1203 content on DME yields, in comparison to Cu/ZnO ratio. The highest yield of DME+MeOH was given by the catalyst of 14.3 % A1203 with Cu/ZnO=4/6. In contrast, addition of oxides of Cr203, ZrO2 and Ga203 instead of A1203 did not cause DME generation, although Cu-ZnO-Ga203 showed the highest yield of MeOH among all of the catalysts. Strangely, Cu-ZnO catalysts without A1203 showed little activity for DME and MeOH syntheses, in spite of experimental reports 8) claiming that Cu-ZnO is effective for synthesizing MeOH from mixtures of CO, H2 and a little CO2. In addition, they claimed no effect of Cu-

439 0. 2 0.15

Cu/Zn0=5/5 '

0.1 "-" ~

00.05 0.15

Cu/Zn0

O. 002

-o 0 ~9

_4/; ~

,n

n,~

m

DME

[-~

MeOH

-

Cu/Zn0=3/T

0.012

Cr203

0. 15

Zr02

0.1 0.05 0

O. 004

0

r

O. 040

o o,o] 6. 7

f

O. 002

14. 3

tool%

AI203

23. 1

33. 0

O.

l

[

Ga203

055

14. 3mo1% Ox i de

Fig.3 Measured yields of methanol and dimethyl ether for all catalysts used. ZnO-A1203 catalyst on DME synthesis at all, although the chemical composition of industrial catalysts (Cu: 12--66 mol%, ZnO: 17-62 mol%; A1203:4-38 mol%) covers this study's best catalyst (34.3/51.4/14.3). This is mainly due to syngas composition (CO-CO2-H2). Regarding the relationship between syngas composition and catalyst composition, a noteworthy result was recently reported by Fujita and coworkers a2). They examined the catalyst mechanism of Cu-ZnO under CO-H2 and CO2-H2 by means of diffuse reflectance FT-IR spectroscopy. The result demonstrated that Cu promotes CO2 conversion to MeOH and ZnO promotes CO conversion. This suggests that the CO/CO2 ratio of syngas would have significant influence on the optimum catalyst composition in the Cu/ZnO system. Therefore a similar study should be extended to CO-CO2-H2 mixtures for constructing an advanced catalyst theory. To make clear the effects of A1203 content and Cu/ZnO ratio in the Cu-ZnO-A1203 system, we redrew data in the following 0.1 :-. . . . . . . . .-.-. -. .h. . . . . . . . . . . . . . . . . I 9 DME i figure. Figure 4 shows the effect of A1203 content on DME and MeOH yields with Cu/ZnO=3/7. The sharp rise in DME yield indicates strong sensitivity to A1203 content, in comparison to the MeOH yield. Both peaks of DME and MeOH curves are identically located at 14.3 mol% A1203. Similar tendencies were also observed for the other Cu/ZnO ratios of 4/6 and 5/5. This is explained well by the concept 15) that DME is produced by a series reaction via MeOH synthesis (BF gas --> MeOH ~ DME).

/ \

----

/

o MeOH

o.o. I.............................. ./ ,.:,E 0.06

........ /

~= o.o4 il. . . .

......

-/-- . . . . . .

0 02 i....... _ .O. :::--9 . / ~ 0

~

10

........ \ .....

-o o

':":-""O ".. . . . . . . . .

...................... 5

[

..9 \

.... . ............ ............ .....

15 20 25 30 35Cr~Oa ZrO~ Ga2Oa MoI% AI2Oa 14.3 mol%-Oxide

Fig.4 Effect of alumina content on dimethyl ether and methanol yields over catalysts of Cu/ZnO=3/7, together with other oxides

440 Figure 5 shows measured surface 01 . . . . I . . . . i . . . . i . . . . ~. . . . ~. . . . ! . . . . area of the catalysts of Cu/ZnO=3/7. I The measured surface area showed ,-. 501............................................................................................................................................ , .......... I strong A1203-content dependence. I Ga203 Interestingly, this result can explain o~ 401........................................................................................................................................................ well the catalyst property of the CuZnO-A1203 system (see Fig.3). That is, we can conclude that the catalysts with large surface area cause high yields of 201....................../ .................. \ ............l..........'-..................-'-............................i DME. The surface area of 14.3 and 23.1 mol% A1203 catalysts (approximately 33 m2/g) are more three times higher in comparison to the 0 catalyst without A1203. Similarly, the 0 5 10 15 20 25 30 35 14. 3 mol~ Oxide 14.3 mol% Ga203 catalyst showed mo196 AI203 large surface area and good catalyst Fig.5 Effect of A1203 content and oxides on properties. However, this applied to specific surface area of catalysts with only MeOH synthesis, not DME molar ratio of Cu/ZnO=3/7 synthesis. Figure 6 shows XRD (X-ray diffraction) patterns of the Cu-ZnOA1203 catalysts used, in which three ~ ~~ II . . . . . o ZaO substances of Cu, ZnO and ZnA1204 m 9 ^ 33.0 mol% AI203 were detected as main peaks. Peaks of A1 or A1203 never appeared in spite of O ,, J ~ 23.1 tool% AI.O. our expectation. With increasing o o'" m.~, A V 9 9 . ," ",~.,,~~,~,.~. A1203 from 0 to 14.3 mol%, peaks of o Cu and ZnO were broadened and their r l 14.3 tool% A120:3 []~9 ~ o D o o heights became smaller. In this range, the catalysts became more active for [] I o , [] l ' 6.7 mol% AJ20. DME and MeOH synthesis with l,l~ l, n o g um o o " increasing A1203 (see Fig. 4). In contrast, 23.1 and 33.0 mol% A1203 ? I o catalysts showed ZnAl204 peaks o f , instead of ZnO peaks, however its 'D' t Ill 'i o o El 0rn~ l peak width was very broad. In i lU II~, i ' O O O I addition, the Cu peaks of the 33.0 mol% catalyst became very weak. 20 40 60 80 100 According to these results, we can 2 t9 ( d e g r e e , CuK a ) expect that the optimum composition of the catalyst would range between Fig.6 XRD patterns of catalysts with molar ratio of Cu/ZnO=3/7 14.3 and 23.1 mol% A1203. As for the mechanism of DME direct synthesis from CO-CO2-H2, A1203 microstructure must be a key point as mentioned above. Microscopic research of catalyst structure is also needed in the future. Figure 7 shows an example of changes in gas concentration with time, where the hybrid catalyst with 5mass% - A1203 was employed. It is clear that, DME is synthesized. This

1~

i i ;i; . . . . . . . . . . . . . . . . .,. . . . . . . . . . . . . . . .

441 demonstrates that DME synthesis from BFG over the hybrid catalyst is a typical series reaction: BFG

Cu-Zn~

>MeOH

>D M E

y-AI203

Concentration of BFG decreases with time at the initial stage, then, it remains steady after aboutl0 ks. Conversely, DME concentration increases and shows a peak, corresponding to the initial decrease of BFG concentration. In contrast, such a peak is not observed in MeOH concentration. It takes about 15 ks for the hybrid catalysts to reach steady state. The same tendency to form a DME peak in an initial stage is found in other hybrid catalysts. This is probably caused by first conversion of DME from MeOH, in comparison to MeOH synthesis from BFG.

0.15

~ S

--

:

f

o

BFG(CO+CO2,H2)

o0,

~

M

.~

,.--

..

~

"~ ..m

60

"~ -=

40

m

20 0.1

E

:_~_~

~

77

5

10

15

20

.

.

. . . . ! . .......... . . .i . i...........

'

25

.

.

.

i

.

'

i-

............. i

"~ ~

0

.

-,

!Moor

0

.

,~M,~ ooH| o.o6 ~t ....... ~

....... ::. . . . . .

:

~

t~

0.04

Time / ks

Fig.7 Concentration-time curves for DME, MeOH and C0+CO2+H2 in outlet gas. 5mass% y-A12o3 + 95mass % _ Cu_ZnO_Ga203 _ _ _ _

00

2

~

o 0

l_

too

10

20

30

T -A1203 / mass%

,Jo

~o

,

"

40

Cu-ZnO-Ga203 / mass%

40

I

60

50

~o I

Effect of mixing ratio in the hybrid Fig.8 Effect of hybrid catalyst on DME, catalyst on yields of DME, MeOH and HC MeOH and HC. is shown in Figure 8, together with changes (Cu-ZnO-Ga2Q3)+ y-A1203, 2.0g in bed height. Yield of DME has very strong W/F=3.44x 10 -~g-cat's/mo l, dependence on the mixing ratio, showing CO/CO2/HE/N2=35/35/8.7/21.3, peak of DME yield at 5 mass% y-AI203 1.0Mpa, 523K addition. In contrast, HC yield is negligible in all cases. The selectivity of DME+MeOH is 99.3% on carbon mole basis. The MeOH yield is almost constant in the region 50 to 90 mass% of Cu-ZnO-Ga203 in spite of an increase of the methanol-active catalyst. This was probably caused by the fact that DME synthesis is a series reaction via MeOH from BFG and intermediately synthesized MeOH is very quickly converted to DME. The reaction of MeOH synthesis from BFG was the rate controlling step under this experimental condition. It was, therefore, concluded that the hybrid catalyst

442 developed is quite promising for the direct conversion of BFG to DME, and very small additions of T-A1203 to a methanol-active catalyst of Cu-ZnO-Ga203 is enough because the rate of dehydration due to T-AhO3 is much larger. These results appealed to the possibility to promote effective use of BFG by overcoming the thermodynamic limitation of MeOH yield.

5. CONCLUSION Fifteen kinds of single component catalysts in Cu-ZnO-A1203 and Cu-ZnO-Oxide (Cr203, ZrO2, Ga203) systems and the hybrid catalyst of Cu-ZnO-Ga203 and y-A1203 were systematically studied for direct synthesis of DME from CO- CO2-H2 mixture. The experimental conditions were 523 K, 1.0 MPa pressure and 2.0 g catalyst weight. The following conclusions were drawn: (1)The most noteworthy finding is the significant catalytic activity of Cu-ZnO-AI203 in DME synthesis, not MeOH, from CO-CO2-H2 mixture. The catalyst of 34.3Cu-51.4ZnO-14.3 A1203 (Cu/ZnO=4/6 (mole basis)) showed the best activity for DME synthesis. (2)The catalytic mechanism of the system of Cu-ZnO-AI203 is still not well explained. However, catalysts showed strong relationship between DME activity and surface area, and XRD peaks of ZnO in active catalysts tend to be broadened without Zn and ZnA1204 peaks. (3)Maximum yield of DME+MeOH was given by the hybrid catalyst with 5mass% y-AI203 and 95mass%Cu-ZnO-Ga203. As for the BFG conversion to MeOH, the single catalyst of Cu-ZnO-Ga203 reached 75% of equilibrium conversion, while for the DME conversion, the hybrid catalyst of 5mass% y-A1203 had only 4% in spite of large yield. (4)Rate controlling step of DME synthesis in the hybrid catalyst system was MeOH formation. In other words, y-A1203 is better catalyst for the reaction from MeOH to DME. For enhancing the DME conversion, development of more active catalyst for MeOH synthesis is expected.

REFERENCES

1. Y. Ohno: Personal communication (1996). 2. T. Akiyama, H. Sato, A. Muramatsu and J. Yagi: ISIJ Int., 33 (1993), 1136. 3. A. Muramatsu, H. Sato, T. Akiyama and J. Yagi: ISIJ Int., 33 (1993), 1144. 4. H. Sato, T. Akiyama, A. Muramatsu, T. Sugimoto and J. Yagi: J. Jpn. Pet. Inst., 38 (1995), 390. 5. H. Sato: Ph. D. Thesis, Tohoku University, (1996). 6. H. Hover: Ullmann's Encyclopedia of Industrial Chemistry, A8 (1989), 541. 7. A.C. Sofianos and M.S. Scurrell: Ind. Eng. Chem. Res, 30 (1991), 2372. 8. R.G. Harman, K. Klier, G.W. Simmons, B.P. Finn, J.B. Bulko and T.P. Kobylinski: J. Catal., 56, (1979), 407. 9. B. Peplinski, W.E.S. Unger and I. Grohmann: Appl. Sur. Sci., 62 (1992), 115. 10. R. Burch, S.E. Golunski and M.S. Spencer: J. Chem Soc. Faraday Trans., 86 (15), (1990), 2683. 11. I. Nakamura, T. Fujitani, T. Uchijima and J. Nakamura: J. Vac. Sci. Technol., A14 (3),

443 (1996), 1464. 12. S. Fujita, M. Usui, H. Ito and N. Takezawa: J. Catal., 157, (1995), 403. 13. Y. Han, K. Asami and K. Fujimoto: J. Japan Inst. of Energy, 75, (1996), 42. 14. K. Fujimoto, K. Asami, H. Saima, T. Shikada and H. Tominaga: Ind. Eng. Chem. Prod. Res. Dev., 25, (1986), 262. 15. S. Machida, T. Akiyama, A. Muramatsu and J. Yagi: ISIJ Int., 37 (1997), 536.



445

Heterogeneous Photo- and Thermal Catalytic Oxidation of CO: Effects of Metal Deposition I. Ozen and D. Uner Chemical Engineering Department, Middle East Technical University, Ankara 06531, Ttirkiye [email protected]

ABSTRACT In an effort to elucidate the mechanism of rate enhancement of photocatalytic reactions over TiO2 upon metal doping, CO oxidation reaction was studied under dark and illuminated conditions. It was observed that the rate enhancement due to the presence of Pt was merely due to the thermal reactions with no contribution from UV illumination under the sensitivity conditions of the experiments. Oxygen adsorption studies performed separately indicated that the oxygen adsorption capacity of the TiO2 catalysts decreased after wetting and drying. Among differently pre-treated Pt containing catalysts, all of the samples exhibited a higher oxygen adsorption capacity than the TiO2 substrate, except a sample for which the last treatment step was calcination. The results of this study indicate that the thermal catalytic effect of Pt is the only source of rate enhancement of CO oxidation, within the specified conditions.

1. INTRODUCTION It is a well-known fact that the presence of some metal additives such as Pt, W, Au or Ag enhances the rate of low temperature or photo catalytic oxidation of the organic matter over TiO2 surfaces. However, the mechanism of the rate enhancement is still a subject of debate in the scientific community. The photo catalytic oxidation reactions require catalytic materials that can be activated by the proper wavelength electromagnetic radiation, i.e., the wavelength should be coherent with the band gap of the catalyst. TiO2 for example has a band gap of around 3.2 eV, which is appropriately activated by electromagnetic radiation in the UV-A (-350 nm) to UV-C (-200 nm) range. The wide ranges of the UV wavelengths are tolerable in the TiO2 framework when the active catalyst is used in the powder form. The particle size effects, the lack of long range structural order in the bulk and the surface irregularities all contribute to variation of the Fermi level in the bulk structure, therefore the material can be activated by a wide range of wavelengths of electromagnetic radiation. The

446 band gap of TiO2 allows us to consider it as a semi-conducting material. Thus, when a semiconducting material is brought into a contact with a metal, a charge transfer is expected to occur from one side to the other depending on their respective bulk Fermi levels. As a result of this charge transfer around the contact area, a charge accumulation around the interface surrounded by a charge depletion layer occurs. This phenomenon was thought to influence the catalytic processes during the photo excitation of the catalytic material. It is argued that due to the presence of the depletion layers, the exothermic recombination of photogenerated electron-hole pairs (e- - h +) was retarded, allowing them more time to interact with the adsorbed organic material to react. However, one should also look at the chemical and catalytic phenomenon taking place as a result of this newly added catalytic material to the system. The intrinsic activity of the catalyst can play important roles in increasing the rates of these extremely slow reactions. Vorontsov et al. investigated the thermal and photocatalytic oxidation of CO [1 ] and resulted in an increase in the percent quantum yield upon Pt photodeposition onto the TiO2 support in the photocatalytic oxidation reaction. In their study, among samples with different oxidation state of Pt, the highest photocatalytic activity was observed in Pt~ and this phenomenon was attributed either to a better photogenerated charge separation with a decrease in recombination rate of the e- - h pairs, or to Pt atoms being reaction sites themselves. Parmon et aL [2] observed no dark activity of CO oxidation on the 0.4 % photodeposited Pt/TiO2, however they do not clearly report the reaction temperatures. This lack of information about the reaction conditions prevents the differentiation between photoand thermal activities under the same conditions. In a previous study, the deactivation of TiO2 during the photocatalytic decomposition of benzene was reported [3]. The reaction studies along with supporting evidence from the literature indicated that deactivation during this reaction occurred as a result of a carbonaceous material deposited on the surface of the catalyst during the reaction. Further evidence from the literature suggested that the presence of Pt inhibited the deactivation [4-6]. It was postulated that the spillover of oxygen in the presence of Pt might enhance the final oxidation step [3,7]. In this article, the results of the adsorption experiments performed to test the spillover postulate will be presented. Furthermore, by assuming that the final oxidation step in a thermal or photo-oxidation reaction of a volatile organic carbon (VOC) is the CO oxidation reaction, the results on effect of Pt on the activity of TiO2 during this reaction will also be presented.

2. EXPERIMENTAL

2.1 Catalyst Preparation The TiO2 powder obtained from Degussa (P-25) was used as is, or water treated before the adsorption or reaction studies. The Pt/TiO2 catalysts were prepared by incipient wetness technique. Appropriate amount of the metal salt, Pt (NH3)aC12.H20(Johnson Matthey) was dissolved in 2-4 ml H20/g support to bring about incipient wetness. The solution was impregnated to the support and the final mixture was dried overnight at room temperature and for 4 h at 110 ~ The final catalyst was ground, and subjected to various calcination and reduction treatments. The catalyst was reduced in H2 under static conditions. The static

447 reduction started by dehydrating the sample under He at 150 ~ for 15 min followed by evacuation, and at 350 ~ for another 15 min, again followed by evacuation. H2 reduction was carried out at 350 ~ by introducing 750 Torr of hydrogen for 30 min intervals each followed by evacuation, for a total duration of two hours. The catalyst was calcined prior to or after reduction at 450 ~ for five hours. In order to eliminate the C1- contamination from the catalyst, the reduced samples were washed in hot water several times. Hot water elutriation continued until the filtered solution was proven to be free of CI ions by testing with AgNO3 solution. Various catalyst pre-treatments are summarised in Tablel. Table 1a Step codes for pre-treatment Treatment step Code incipient wetting 1 (and Pt impregnation) drying 2 dehydration 3 washing 4 reduction 5 calcination 6

Table lb Catalyst pre-treatment sequences Catalyst Catalyst content code TiO2 A TiO2 B 0.5% (w/w) Pt/TiO2 C 0.5% (w/w) Pt/TiO2 D 0.5% (w/w) Pt/TiO2 E 0.5% (w/w) Pt] TiO2 F 0.5% (w/w) Pt/TiO2 G H 0.5% (w/w) Pt/TiO2

Pre-treatment sequence untreated 1,2 1,2,3 1,2,3,5,3 1,2,3,5,4,2,3 1,2,3,6,5,3 1,2,3,5,6,5,3 1,2,6

2.2 Oxygen Adsorption The oxygen adsorption characteristics of both pure and 0.5% (w/w) Pt doped TiO2 were examined by measuring the oxygen adsorption isotherms. The experiments were run in a home built pyrex manifold o f 203 ml volume, coupled with Baratron gauges (Varian) and a Turbo-molecular pump backed up by a mechanical pump (Varian TurboV). A schematic drawing of the system is given in Fig. 1. The background pressure of the manifold was less than 10-6 tort after bakeout. Effects of calcination, in-situ reduction and washing on the

I 6-portvalve

pressureread-out

coldcathodegauge baratrongauges Pyrexmanifold

l

I

vacuum pumpstation

turbo

}

~

U~ globeVaslvte0pcock~

catalyst containerand r j-I-] heatingmantle

Fig. 1. Experimental set-up for volumetric oxygen chemisorption measurements.

,,c

448 oxygen adsorption capacity were investigated.Adsorption measurements were performed with 99% pure oxygen at room temperature. Isotherms were obtained for the 0 - 600 torr pressure range in about 20 torr pressure increments. 2.3 CO Oxidation

The oxidation experiments were performed in a packed bed cylindrical Pyrex reactor with 0.5 g catalyst samples (see Fig. 2). The feed mixture contained 8 ml/min of CO and 50 ml/min of dry air, adjusted by mass flow controllers (MKS 1179), giving a space time of 1.5"105 g cat-s/mol. The catalyst samples were activated under UV irradiation of 365 nm and 100 W. The produced gases were separated in an HP 4890 gas chromatograph equipped with a Porapaq-Q column and analyzed via a TCD. The reaction environment was heated to 110 ~ by the UV lamp.

Fig. 2. Experimental set-up for CO oxidation reaction.

449 3. RESULTS AND DISCUSSION The oxygen adsorption isotherms measured at various stages of catalyst preparation are presented in Fig. 3. The data presented in Fig. 3 indicated that the oxygen adsorption capacity of an untreated 0.5% Pt/TiO2 (C) is the highest for the whole pressure range -except for the twice reduced catalyst (G)-, probably due to the oxidative interaction between the catalyst precursor and the adsorbate. Adsorption decreases upon reduction (D) and starts to approach back to the original level upon hot water elutriation (E), which is performed in order to clean the catalyst from C1- ions coming from the metal precursor salt. However, the increase in oxygen adsorption capacity after hot water elutriation of C1- ions is only observable at the high-pressure range of the experiment. Both of samples D and E posses higher oxygen adsorption capacity than the pure support only at this high-pressure range, i.e. above 350 Torr. Adsorption on the reduced sample decreases drastically upon calcination and becomes lower than that of the untreated pure support. This is thought to be due to the presence of strongly bounded oxygen on the support surface upon calcination. However the adsorption capacity reaches back to the initial value upon a second reduction (G). Wetting and drying of the TiO2 support reduces adsorption capacity, due to an increase in the bulk density and an accompanying decrease in the external surface area (B) compared to the untreated pure sample.

390

-

340

-

v

290

-

z =-

240

-

190

-

140

-

0

1D

--o-E ,i.

F

a

G

-*--B A

900

40-10

! 0

100

200

300 Pressure,

400

500

600

P (torr)

Fig. 3. Oxygen adsorption isotherms. See Tables 1a and lb for the pre-treatment conditions. The adsorption isotherms were fit to dissociative and non-dissociative Langmuir models. Both models fit equally well for Pt/TiO2 (C and D) while none of them gave a reasonable correlation for the undoped supports. A comparison between Pt/TiO2 (E) and water treated pure TiOa (B) indicated that oxygen adsorption capacity of TiO2 increased by 0.93 mol O2/mol Pt upon Pt addition. Although the Pt dispersion of this catalyst is not known, a typical supported Pt catalyst

450 prepared by incipient wetness technique can have dispersions of about 30% [8]. Therefore, close to 1:1 O2: Pt stoichiometry with the total metal loading can increase to 3:1when the true dispersion is taken into account. Such a high O2: Pt stoichiometry is indicative of O2 spillover although the actual evidence must be provided by spectroscopic methods. The results of the CO oxidation reaction carried out under dark and illuminated conditions are given in Fig. 4. Neither non-catalytic reaction nor deactivation was observed within the duration of the experiments. Essentially the same behaviour was observed for dark and illuminated conditions at the same temperature, indicating that the CO oxidation reaction was merely due to the thermal activity of the catalyst at the specified reaction conditions. Meanwhile, the overall conversion of the reaction did not depend on the pre-treatment methods of the sample. Pure TiO2 samples showed no measurable activity for neither of the media, which indicates that the rate enhancement is, dominantly, via the intrinsic catalytic 100 90A

80

x o

70

o ~

60

>0

-

'- t9 8~

o~ E~

50

40

k

30

-B-- H2 / UV

[

---z~ F

20 10 0

m

0

i

i

i

i

i

i

10

20

30

40

50

60

time,

t (min)

activity of Pt. This result supports the proposed role of Pt as a spillover media for oxygen, thus increasing the rate of reaction.

Fig. 4. Steady-state conversions for various catalyst samples in CO oxidation reaction. Considering the overall mechanism of VOC oxidation, it could be agreed that Pt addition slows down deactivation of the catalyst in by speeding up the final oxidation step, if this step is CO oxidation as proposed in this study. As also stated in literature, oxygen spillover could be an important step in the removal of the surface deposited carbon causing deactivation [9].

4. CONCLUSIONS Deposition of Pt onto the TiO2 support was observed to significantly increase the rate of CO oxidation reaction, via its intrinsic catalytic activity. Under the same experimental conditions, the pure support, TiO2, was totally inactive. No evidence for an electronic effect

451 due to UV illumination was observed. Conversions up to 80-90% could be achieved with Pt/TiO2 in both dark and illuminated media, independent of the pre-treatment method of the catalyst. Oxygen adsorption tests indicated that Pt deposition onto the support resulted in a significant increase in the oxygen adsorption capacity, which in turn could be acting in the rate enhancing mechanism of the CO oxidation reaction.

ACKNOWLDEGEMENT

The funds for this project was provided by TUBITAK (Turkish Scientific and Technical research Council) under research grant no MISAG-96. REFERENCES

1. Vorontsov, A. V., Savinov, E. N., Zhensberg, J., J. Photochem. Photobio. A: Chem., 125 (1999) 113-117 2. Parmon, V.N. et.al., J., Cat. Today, 39 (1997) 207-218 3. Ozbek,, S., Uner, D.O., Stud. in Surf Sci. and Catal., 126 (1999) 411-414 4. Sauer, M., Ollis, D. F.,J.CataL, 158 (1996) 570-582 5. Fu, X., Zeltner, W., Anderson, M.A.,App. Cat. B: Env., 6 (1995) 209-224 6. Papaefthimoiu,P., Ioannides, T., Verykios, X.E.,Appl. Cat. B: Env., 15 (1998)75-92 7. Falconer, J.L., Magrini-Bair, K.A., J. CataL, 179 (1998) 171-178 8. Bhatia, S., Engelke, F., Pruski, M., Gerstein, B. C., King, T. S., J. Chem. Phys., 147 (1995) 129-140 9. Conner, W. C., Falconer, J. L., Chem. Rev., 95 (1995) 759-788


Studies in Surface Science and Catalysis 133 G.F. Froment and K.C. Waugh (Editors) (r 2001 Elsevier Science B.V. All rights reserved.

453

Harmonization of the empirical information on the de-NOx catalysts for the comparison of the catalyst performance through reaction modeling D. Uner and S. Kaya Chemical Engineering, Middle East Technical University, Ankara 06531, Turkey [email protected]

In an effort to rationally design a monolith catalytic converter, data harmonization and converter analyses were performed on model reactors. The empirical data presented in the literature on CO and hydrocarbon oxidation over various catalysts were harmonized for space times through a simple model reaction. The model parameters on CO and hydrocarbon oxidation were used along with the reported temperature and velocity profiles to analyze the axial and radial conversion behavior in a cylindrical monolith geometry. The axial conversion behavior was investigated via a multiple CSTR reactor model. In order to investigate the radial effects, the reactor was divided into one central cylinder and three doughnut shaped subsections. The division points were based on the shape of the velocity profile. During this analysis t-he diffusional and mass transfer effects were neglected. Analyses were performed for net oxidizing and net reducing exhaust atmospheres. The temperature and flow non-uniformities present in the catalytic converters resulted in unique conversion profiles: The velocity maximum in the central part of the reactor resulted in lower conversions than the intermediate parts. The lowest conversions occurred at the peripheral annular region. With this information, it will be possible to tailor the metal loading as a function of radial coordinate to improve the conversion non-uniformities and the precious metal economy. 1.

INTRODUCTION

The design of catalysts for automotive applications is still an active field, changing its dimensions and directions simultaneous with the changed regulations on the environmental issues, the engine design and also the changes as a result of the fuel characteristics based on the demands of the automotive industry and the environmental regulations. Therefore, rational design of the exhaust emission catalysts is an economical as well as an environmental imperative. Precious metals, such as platinum and palladium, are well-known complete oxidation catalysts with high activity and stability, and are widely used for control of exhaust gas and automotive exhaust (1). Most of the unwanted exhaust emissions occur within 70 seconds of engine ignition, before the catalyst system has reached operating temperature. As a result, emission system engineers have focused on ways of reducing the time it takes the catalytic converter to "light-off," or become operational. A number of options have been tried: close-coupled converters to take advantage of engine heat, increased idle speed, dual-wall exhaust pipe technology and high-technology engine manifold designs. Ceria coverage of Pt strongly promotes CO oxidation, suggesting that new sites at the metal/oxide interface become available for the reaction (2). The oxidation of hydrocarbons near the stoichiometric point is the primary process for hydrocarbon removal for the modem three-way catalyst equipped light duty vehicles. When oxidation of propane propylene composition over

454 palladium catalyst was compared, it is observed that propylene was considerably easier to oxidize than propane. The easier to oxidize species are converted before for those more difficult to oxidize. The ease of paraffin oxidation is a function of the length of the hydrocarbon chain (3). Automobile converters are characterized by non-uniform flow distributions that can give rise to severe radial gradients even when the reactor operates under adiabatic conditions. It is evident from the geometry of the inlet section that the flow rate in each channel of the monolith will be very strong function of the channel's radial position. Fluid velocities are much higher near the center of the reactor (r=0) and fall the very low values at its periphery (r=R). Non-uniform flow distributions cause significant radial temperature and concentration gradients, and can severely degrade the performance of a converter (4). Catalytic converters are exposed to various thermal and concentration fields depending on airfuel ratio control system and configuration of the exhaust system like the volume/location of catalysts, the size of exhaust pipes, and so on. The central region near the inlet is heated. In this portion, CO is oxidized actively due to high wall temperature, but passes through the channels without being oxidized in the outer region because the wall temperature rises very slowly (5). In an effort to find the optimum loading and performance of the monolith, the temperature and flow non-uniformities must be taken in to account. Therefore, a study was carried out firstly to harmonize the data reported in the literature for comparing the catalyst performances based on similar space times. Selected data from literature on the precious metals for CO oxidation reactions were harmonized and Arrhenius parameters were obtained (1-3,6-9). When the bed porosity data was not available an average estimate of the bed porosity was used to uniformly report the space times in terms of h -~ units. The same analysis was applied to oxidation reactions of hydrocarbons present in the exhaust gas such as benzene, toluene, hexane and octane (3, 9-11). After the performance harmonization of the rate data on selected catalysts, the reaction model was used to estimate conversion profiles in a typical monolithic reactor. 2.

THE M O D E L

2.1. Data harmonization model:

In order to extract the rate parameters such as the order, the activation energy and Arrhenius pre-exponential factor of the reaction from the light off-curves, available empirical data from the literature were collected from the references mentioned in the introduction section. Considering the fact that most of the reactions were carried out in packed bed reactors, a plug flow reactor (PFR) design equation was used to model the light-off process. The experimental data was compared to the model predictions and the rate parameters were extracted. It was observed that a first order kinetics in CO and hydrocarbon concentrations was sufficient to have a reasonable fit (a correlation better than 90%) to represent the oxidation light-off data for these substances. In a previous study, the surface concentrations, orders and activation energies of a NO+CO reaction were determined by a micro-kinetic analysis, and it was observed that the reaction orders of the NO+CO reaction with respect to NO and CO partial pressures were temperature dependent, and for the temperatures dictated by the experimental conditions, the reaction is first order with respect to CO partial pressure (12, 13). The experimental conditions of the data treated in this study were such that oxygen was in some cases in excess, thus the order of the reaction with respect to oxygen partial pressures were not sought. In addition, the conversions presented in the light off curves were

455 reported mostly in terms of the oxidizing species giving no information about the oxidant behavior. The details of the data used for this study are collected in Table 1. Table 1 The catalysts, reactive species and the reaction conditions in terms of the amount of the catalyst and the space velocity and their respective references of the data used for this work

Inlet gas mixture

Space velocity, -1 1; ,

Catalys t weight (mg)

Gas flow rate (cm3/min)

Pt/SiO2

200

40

4 % CO, 4 % O2/He

200

Pt/CoOx/SiO2

200

40

4 % CO, 4 % O2/He

200

Pt/MnOx/SiO2

200

40

4 % CO, 4 % O2/He

200

Catalyst

(cm3/g cat.min)

1%CO, 0.5%02/N2 1%CO, 0.5%02/N2 1% CH4, 2.2-2.5 % 02 1% CH4, 2.2-2.5 % 02 1% CH4, 2.2-2.5 % 02 1.26 % N20 in helium, 150 100 1 or 5 % CO in helium 1.26 % N20 in helium, Rh/A1203 CO+NO 150 100 1 or 5 % CO in helium 1.26 % N20 in helium, Rh/A1203 O+N2 0 150 100 1 or 5 % CO in helium 200 100 1% CO, 0.6 % 02, 25 Pt/A1203 ppm 1-hexene, and benzene 200 100 1% CO, 0.6 % 02, 25 Rh/A1203 ppm 1-hexene, and benzene 200 100 Pd/A1203 1% CO, 0.6 % 02, 25 ppm 1-hexene, and benzene ! Rh/A1203/M 200 100 1% CO, 0.6 % 02 Calcination temperaturesare 1000, 1300,and 1400K, respectively. Pt/AI203 Pt/CeO2/A1203 Pd/Sro.8Lao.2A12019* Pd/Sro.8Lao.2A12Ol9 Pd/Sr0.8Lao.2Al2Ol9 Rh/A1203 CO+O2

300 300 200 200 200

50 50 50 50 50

Ref.

i 2

166.7 166.7 250 250 250 1500

8

1500

8

1500

8

500

11

500

11

500

11

500

10

2.2. Modeling the monolithic reactor as a combination of parallel and series reactors: After extracting the kinetic parameters, selected results for CO oxidation over were used to analyze the effect of non-uniform temperature and velocity distributions on the conversion of CO. In order to determine the optimum number of multiple CSTR's to capture the behavior of a PFR, the rate law of Oh and Carpenter (14) for the NO+CO reaction was used to model a monolith channel as a CSTR in series. The results indicated that it was sufficient to use 5 reactors in series to capture the performance of the PFR behavior in the NO+CO reaction The cells of a monolith reactor were taken as independent parallel reactors ignoring the mass transfer and diffusion through the ceramic pores. The axial and radial temperature and velocity profiles collected from the literature(4,5) are used to calculate the

456 individual average space time and space velocity of the independent CSTRs connected in series and in parallel. The axial and radial conversion profiles are obtained by calculating the conversions in each reactor cell defined by its unique velocity and the temperature profile. Radial and axial divisions of a cylindirical monolith used in the model are shown in Figure 1.

radlus

Figure 1. Subsections of the quarter monolith. The radial divisions were decided based on the shape of the velocity profile (4). The azimuthal dimension was treated as uniform. 3. RESULTS AND DISCUSSION 3.1. The harmonization of the empirical information: The data harmonization was performed on several hydrocarbon species and CO oxidation as well as on N O x - C O reaction. A small review of some of the data treated in this study is presented in Table 1. The first four columns in Table 1 indicate the experimental variables such as.the catalyst type, the sample weight, the gas volumetric flow rate and inlet gas mixture composition. The fifth column reports the calculated space velocity based on the given data, and in the last column the reference from which the data were collected is given. There are several factors to note in the information presented in Table 1. First of all, the sample sizes and the gas 1 ................................... ~ - , : ~ . , . - q y . ~ ....................... flow rates are not uniform even in the limited number of the studies collected. 0,7 r The space time values are calculated by the authors based on the information presented in the source articles. Most of the studies 0,4 , 0,3 did not report the catalyst porosity and the 0,2 bulk density, therefore the space time units o~ ?~." --'" ! had to be determined in the gcat.min/cm 3 . . . . . o ,-~V:"; :-.-':"--"-" 250 300 350 400 450 500 550 600 650 700 750 units. These units make it difficult for Temperature, K using the resulting kinetic information in monoliths where it is more meaningful to ~/CeO2/AI203 (6) ---~--- *PJCeO21AI203(6) J- R/CoOx!902(2) use proper time units for the space time -..A..-*R/CoOxlSiO2(2} ~ R/AI203(10) ...o... 'P,JAI203(10) due to the open geometry used in the monoliths. The adaptation procedure will Figure 2. A comparison of the be discussed in the forthcoming section. catalyst performances, solid lines represent Finally, the gas compositions varied the data as reported in the literature, between the source articles, another factor dashed lines represent the data after the contributing to the difficulty in the data space velocity correction to 300 h ~. The comparison. experimental space times are given in Table 1.

o,

,: ,...

457 Decreasing the space velocity used in the experimental studies can shift the light-off curves towards right (lower temperatures). In other words, a lower space velocity can show a lower light-off temperature for the same catalyst. When the empirical data was corrected for space velocities, the performance order of some of the catalysts was changed. The comparison is given in Figure 2 for a selected number of catalysts. Depending on whether the space velocity had to be increased or decreased in order to bring the data to a single space velocity of 300 h~, the light off curves switched positions, as indicated by dashed lines in Fig. 2. 3.2. The effect of the velocity and temperature non-uniformities on the chemical conversion:

The effect of the flow and temperature non-uniformities on the conversion was investigated on a cylindrical monolith geometry. The CO oxidation reaction kinetic parameters obtained in the previous section were used. An average porosity of 30% and bulk density of 0.42 g/cm 3 of y-A1203 measured by mercury porosimetry in our laboratory (Micromeritics Pore Sizer 9310) were used to obtain the space velocities in h 1 units. The gas composition was selected as 0.4% CO, 1.26% 02 and the rest as inerts for the net oxidizing conditions, as 1.6% CO, 0.465% 02 and the rest as inerts for the net reducing conditions. The results for the net oxidizing conditions are presented in Figure 3. In Figure 3.a the distribution of the space velocity in the monolith as a function of axial and radial distance, and in Figure 3b the CO molar flow rate as a function of axial and radial coordinates were shown. As seen in Figure 3.b, most of the chemical conversion takes place in the first quarter of the reactor, and the conversion non-uniformity along the radial coordinate is obviously seen. The non-uniform conversion profiles along the radial direction arise due to the temperature, but most importantly velocity distributions. As one can see from Figure 3.a, the space velocity increases as one approaches towards the center. In the boundary layers near the wall, lower space velocities serve as a means of increasing the chemical conversion. But on the other hand, the monolith is colder near the wall, thus reducing the overall rate of the reaction. Whereas the fastest flow takes place in the central region, giving rise to a decrease in conversions despite high temperatures.

Figure 3. (a) The space velocity distribution and (b) the axial and radial conversion profiles in a cylindirical monolith. The conversion results presented here are determined with a feed composition representative of net oxidizing conditions.

458 In summary, the flow and temperature non-uniformities can give rise to non-uniform conversion profiles in a monolith reactor. As long as the desired conversion levels are reached, these non-uniformities are tolerable. However, tightening emission standards require improved catalyst performances while the market for the converters require cheaper units. By carefully analyzing the reactor effluents and adjusting the metal or active component loading radially, it can be possible to obtain close to uniform conversion profiles, and improve the metal economy. 4. CONCLUSIONS In this study, performance comparisons of various catalytic materials were performed. In order to compare the performances, data from different sources were modeled, their kinetic parameters were extracted, and for comparison, all of the data used for this study were brought to a single space time and the results were compared. Some drastic changes in the catalyst performances were observed when the space time corrections were properly done. The results of this study clearly demonstrated that, it is imperative to reach a consensus about the catalyst testing procedures in order to be able to compare the performances of the catalysts tested in different laboratories. Selected kinetic information was used to analyze the reactor performance in a monolith. The temperature and velocity distributions in a monolith geometry leads to non uniform conversion profiles. The present model results can be used to tailor the metal loading as a function of radial coordinate to improve the conversion non-uniformities and the precious metal economy. ACKNOWLDEGEMENT

The funds for this project was provided by METU under research grant no AFP-2000-03-0404. REFERENCES

1. Luo, Meng-Fei, Zheng, Xiao-Ming, Appl. Catal. A: Gen., 189 (1999) 15. 2. Mergler, Y. J., van Aalst, A., van Delft, J. and Nieuwenhuys, B. E . , Appl. Catal B: Env., 10 (1996) 245. 3. Summers, J. C., Silver, R. G., SAE Technical Paper no 902072 (1990). 4. Zygourakis, K., Chem. Eng. Sci. 44 (1989) 2075 5. Baba, N., Ohsawa, K., SAE Technical Paper no 962076 (1996). 6. Martinez-Arias, A., Coronado, J. M., J. Phys. Chem. B, 102, (1998) 4357. 7. Sekizava, K., Machida, M., Eguchi, K., Arai, H.,J. Catal. 142 (1993) 655. 8. McCabe, R. W., Wong, C., J. Catal., 121 (1990)422. 9. Efstathiou, A. M., Papageorgiou, D. and Verykios, X. E., J. Catal., 141, (1993) 612. 10. Patterson, M.J., Angove, D.E.,Cant, N.E., and Nelson P.F., Appl. Catal B: Env., 20 (1999) 123. 11. Cant, N.W., Angove, D.E., and Patterson, M.J, Catal. Today, 44 (1998) 93. 12. Ernur, D. and Uner, D.O., Turk J. Engin. Environ. Sci, 24 (2000) 277. 13. Ernur, D. MSc. Thesis, Middle East Technical University, Ankara, 1997. 14. Oh, S. H. and Carpenter, J. E., J. Catal., 98 (1986) 178.

Studies in Surface Science and Catalysis 133 G.F. Frornent and K.C. Waugh (Editors) 9 2001 Elsevier Science B.V. All rights reserved.

459

Synthesis of Cumene (Isopropylbenzene) from Diisopropylbenzenes in the presence of Benzene using Triflic acid as catalyst at room temperature M. C. A1-Kinany, S. H. A1-Khowaiter and F. H. A1-Malki

Petroleum & Petrochemical Research Institute, King Abdul Aziz City for Science and Technology, P.O. Box 6086, Riyadh 11442, Saudi Arabia, e-mail : [email protected]. ABSTRACT Cumene (isopropylbenzene) is presently produced from benzene and propylene using either solid phosphoric acid or anhydrous aluminium chloride or zeolite as catalyst. Large amounts of m- and p- diisopropylbenzenes (DIPB) were produced as by-products from the above processes. Therefore, in order to study the activity of trifluoromethane-sulphonic acid (triflic acid) as catalyst to produce higher yield of cumene from DIPB isomers a series of isomerization and transalkylation reactions, of m- and p-isomers in benzene with different molar ratios 1:1:1, 1:3:1 and 1:6:1 of isomer to benzene to catalyst respectively with each isomer using triflic acid as catalyst, were carried out in liquid phase at room temperature and under nitrogen atmospheric pressure. In the case of 1:1:1 molar ratio of m-isomer or p-isomer to benzene to catalyst respectively transalkylation reaction gave a higher yield of cumene (ca. 46 mol%) based on benzene and isomer. It was found that when the molar ratio of isomer to benzene decreases the yield of cumene decreases. In the case Of 1:6:1 molar ratio of either m- or p-isomer to benzene to catalyst, yield of cumene decreases from (46 mol%) to about (20 mol%). The conclusion that can be drawn from these series of reactions is that, in all cases the rate of transalkylation of DIPB isomers is faster than the rate of isomerization leading to cumene with a higher yield of (46 mol%). Therefore triflic acid as a catalyst is efficient and effective in producing more yield of cumene from the by-products of isopropylation reactions at room temperature and atmospheric pressure. It is more superior than the other catalysts such as supported phosphoric acid aluminium chloride and zeolite, which are currently used in industries at higher pressures and temperatures. 1. INTRODUCTION Cumene (isopropylbenzene) is an important intermediate product, in petrochemical industries, for the major route of phenol and acetone[ 1]. Further downstream derivatives are caprolactame, bisphenol A and others producing end-products. Cumene is presently produced by alkylation of benzene with propene using either solid phosphoric acid or anhydrous aluminium chloride or zeolite as catalyst[2-5].

460 In most of the above processes, several competing side reactions take place to greater or lesser extent depending on the conditions of alkylation, and the greater difficulty of transalkylating of higher substituted isopropylbenzenes such as diisopropylbenzene (DIPB) isomers, triisopropylbenzene (TIPB) isomers and tetraisopropylbenzene (TetlPB) isomers. During the alkylation, about 7-10% of the above polyisopropylbenzenes referred to as high-boiling point residue is unavoidably obtained as byproduct which mainly contains 5060% DIPB (ortho-, meta-, and para-). It is not possible to recycle these products to the alkylation reactors[ 1]. Transalkylation of DIPB isomers with benzene is an effective way for utilizing this byproduct and increasing the cumene yield again. Therefore, it was of some interest to investigate the efficiency and catalytic behaviour of triflic acid as catalyst for cumene synthesis with a higher yield from DIPB isomers (p-, and m-) by isomerization and transalkylation reactions at room temperature. 2. EXPERIMENTAL 2.1 Materials p-DIPB and m-DIPB (Fluka Chemie 97%) were used directly without further purification. Benzene (Sigma Chemicals Ltd., 99.9%) was purified using the standard method, and dried prior to use over sodium wire. Hexane (Fluka Chemie 99.9%), which was used as a solvent for GLC analysis was high grade and spectroscopically pure. Its purity and benzene purity was checked by GLC. 2.2 Catalyst Triflic acid was a commercial sample (Fluka Chemie, d=1.696. 98%) purified by double distillation (b.p. 68-70~ under dry nitrogen at reduced pressure immediately prior to use.

2.3 General Procedures Isomerization and transalkylation reactions were carried out in a closed system with continuous stirring under nitrogen atmospheric pressure and room temperature. The reactions were monitored by GLC analysis using a Varian 3400 instrument fitted with a 60m x 0.32 mm I.D. capillary column phase DB 1, film thick 1.0 micron.

3. RESULTS AND DISCUSSION Transalkylation reaction of DIPB (p-, and m- isomers) to cumene is of industrial interest as some low value products can be converted to cumene, which is in higher demand and higher value. While the aluminium chloride catalyst enables the reaction to be carried out at 80100~ temperature, there are inherent problems with this catalyst such as, complex formation with polyisopropylated products, short catalyst life, require large amount of catalyst, it is not recycled, problem in disposing the waste catalyst complex, and polluting problems. Using solid catalysts there are also inherent problems such as rapid catalyst deactivation, poor transalkylation capabilities, energy consumption is high since the transalkylation rection carried out at 200-240~ and high pressure.

461

A lot of research work has been carried out on transalkylation of DIPB isomers with benzene using many kinds of zeolites at high temperatures and pressures[3-5], but commercial use of these zeolites could not be made, mainly due to its fast deactivation. However no reports were available in the literature, for the use of triflic acid as catalyst for the transalkylation of DIPB isomers with benzene. Therefore, the objective of the study reported in this work was to test the activity and the efficiency of triflic acid as catalyst for the transalkylation of DIPB (p-, and m- isomers) with benzene in order to obtain higher yield of cumene. The activity of the catalyst was evaluated in terms of mole percent of DIPB isomers converted and mole percent yield of cumene formed. We have investigated the effect of various molar ratios of isomer to benzene to catalyst (i.e. 1:1:1, 1:3:1, and 1:6:1 respectively). Isomerization and transalkylation of p-DIPB isomer with benzene Reactions of p - D ~ B isomer, benzene and triflic acid with different molar ratios of 1 :I: I,I :3:1 and 1:6:1 respectively at room temperature caused mainly rapid isomerization as well as transalkylation to cumene at the first half an hour. Then there was rapid transalkylation and little disproportionation to mainly cumene and TIPB as shown in figures (1,2, and 3). When the mole percentage values of benzene and p-DIPB were normalized, it was observed that the conversion of p-isomer (77.5 mol%) and the mole percent of benzene remained (22.5 mol%) after 0.5 h with 1:1 molar ratio. The conversion of p-isomer gradually increased until approximately 4 h to reach about (82 mol %). The same trend was found with 1:3 and 1:6 molar ratio, but the rates of isomerization and transalkylation were slower. Figure (4) shows that the yield of cumene resulted from 1:1 molar ratio is only (18 mol%) after 0.5 h, and gradually increased until approximately 6 h to reach about (46 mol%), and it decreases as the ratio of isomer to benzene decreases. I

~ cuB~mlne - - u - umene

m.Diisol~q)ylbenzerm

/

-.-p-oi~,nzene

I --x- p-Diisopropylbenzene I-u-1,3,S-Trlisopropylbenzene

/I

4O

,~,

10,

o

o.s

1

1.s

2

2.s

3

3.s

4

4.s

s

s.s

s

s.s

o

o.s

1

Tlme-hr F i g . (1) Product distribution from isomerizatlon and transalkylation of p-DIPB Isomer with benzene ( 1 : 1 m o l a r ratio ) in the presence of anhydrous triflic acid as catalyst at

room temperature.

1.s

2

2.s

3

3~

4

Time-hr

45

s

s.s

s

Fig. ( 2 ) Product distribution from isomerJzstion and transalkylstion of p - D I P B !"s omer with benzene ( 1:3 molar ratio ) in the presence of anhydrous triflic acid as catalyst at room temperature.

so so

l:3~-Tdisopi'op~bcmzene

7o

e0

4O ~0 2O 10

ss Fig. ( 3 ) Product dlsUtbodonfrom IsomerizaUonand transalkylalionof p-DiPO isomer with benzene ( 1:6 motar retlo)lnthepresonceofanhydroustdfllcactdascatalystat room temperature.

e

0

os

1

15

2

25

3 Tim~ s hr

4

45

5

s5

e

Fig. ( 4 ) Yield o f c u m ~ e from Isomerizatlonand tranulkylaUon of p-DIPB isomer with benzene ( 1:1 0 1:3 and t:6 molar retlo ) in lhe presence of anhydrous trlflk: acld as

catalyst at room temperature.

ss

s.s

462 It can be seen from the figures (1,2 and 3) that p-isomer gave predominantly m-isomer during the early stages of the reaction as expected of 1,2-shift mechanism scheme(I).

_. + H +

1,2-Shift

~_ I-I+._

p-DIPB

m-DIPB (Schemel)

It is interesting that m-isomer is the only product of isomerization reaction with a very little trace of o-isomer. The absence of o-DIPB, it might have been predicted on steric grounds. The rate of transalkylation, after 0.5 h is faster than the rate of isomerization. The formation of cumene is possible either by transalkylation reaction between DIPB isomers and benzene or dealkylation of DIPB isomers as shown in reactions (1) and (2) respectively.

..... (1) p or m-isomer

Benzene

Cumene

+ fro----( p or m-isomer

Cumene

(2)

Isopropyltriflate

The formation of TIPB is possible by very little disproportionation of DIPB isomers as shown in reaction (3).

2

-..

.+

"-

)

+

2 p or m-isomer

Cumene

TIPB

..... (3)

463

Isomerization and transalkylation of m-DIPB isomer with benzene The same trends were observed when m-DIPB isomer used under similar conditions as shown in figures (5,6 and 7). It can be seen that the rate of isomerization and transalkylation of m-isomer are slower than that found with p-isomer, but both gave mainly cumene. This was expected because of the thermodynamically more stable m-isomer. When the mole percentage values of benzene and m-DIPB normalized, the conversion of m-isomer(67 mol%) with 1:1 molar ratio after 0.5 h. This value continues to rise through out the reaction to reach (70 tool% ) after 4 h. It can also be seen from figures that there is a marked increase in the conversion to cumene with an increase in the molar ratio of isomer to benzene. The maximum yield of about (46 mol%) was observed at 1:1 molar ratio while (18 tool%) with 1 "6 molar ratio after 6 h of the reaction as shown in fig. (8). 100

100 --*-klm. -4,--C.m~* - - . -

OO aO

go ~

1.3,S-Tdtsopr

7O

--I-c.ame -

oo

-

*

-

~

7o

OO

!-

eo

4O 30

1o

s T~n~e3.~r F ~ . ( S ) Product distaUutk~ ~ m

with benzene (

~omer~t~,nd

t r e n s ~ k y ~ k m or r r ~ ' P 8 ~omer

1:1 mo4ar ratio ) In tim presence of awdwdrous trtfllr ackl u r room temperature.

OS

65

at

1

15

2

25

3

Tlme3.Shr

+

,5

s

55

e

e5

Fig. ( I ) Product distrlbutkm from I

s m and tramsalkyhltk~ of m-DIP8 1:3 molar nltto ) In the presence ~ anhydretm trmk: ackl ill catalyst at room t m p e r m r e .

Isomer with benzene (

100

o

o

~

70

OO

oo,

so,

3O 20

11 05

1

15

2

ZS

T I ~ s- hr

Fig. ( 7 ) Product distribution from Isomerlzation and tmnsalkylation of m-DIPB

isomer with benzene (

t:6 molar retio ) in Ule presence of anhydrous tdflic acid as catalyst at room temperature.

0

os

1

ls

2

25

3 Tim~S hr

4

45

5

55

e

Fig. ( 8)Yield of cumene from Lsomedzation and b a n s a l W I d o n of m4)iPB ~

benzene(

wnh 1:1.1:3 and l:6 molar ratio ) ln the presmlce of anhydrous tdflk: acid as catalyst at room temperature.

The activity of triflic acid was similar to that found in our previous studies such as alkylation of benzene with ethene[6], isopropylation of benzene with propene[7], isomerization and disproportionation of diethylbenzene isomers[8], and isomerization and transalkylation of diethylbenzene isomers [9]. 3. CONCLUSIONS The following conclusions can be drawn from this work; The isomerization and transalkylation reactions gave higher yield of cumene (46 mol%) at 1:1 molar ratio with both isomers. When the molar ratio of isomer to benzene increases, the conversion of DIPB isomers and the yield of cumene increases.

464 Based on the results, the triflic acid is a promising active acid catalyst for isomerization and transalkylation of DIPB isomers in order to produce higher yield of cumene. Its activity is comparable with solid acid catalysts and other Lewis and Bronsted acid catalysts at room temperature. ACKNOWLEGEMENTS The authors would like to thank Mr. Rasheed Al-Rasheed and Mr. Ali Al-Kharji for their assistance in conducting the experiments. REFERENCES

1. Ullman's Encyclopedia of Industrial Chemistry, 1980, vol. A13, p.257. 2. G.R. Meima, M.J.M. Van der Aalst, M.S.U. Sanson, J.M. Charles, and J. G. Lee, ERDOL ERDGAS KOHLE, 112, Jahrgang, Heft 7/8, Juli/Aug. 1996, p. 315-18. 3. H. Mikki, U.S. Patent 4,347,393 (1983). 4. A.R. Pradhan, B.S. Rao, Appl. Catal., 106, 143 (1993). 5. K.S.N. Reddy, B.S. Rao, V.P. Shiralkar, Appl. Cata!., 95, 53(1993). 6. B.L. Booth, M. A1-Kinany, and (in part) Khosrow Laali, J. of Chem. Soc. Perkin. Trans., 1, 1987, 2049. 7. M.C. A1-Kinany and S.H. A1-Khowaiter, Proc. of the 15th World Petroleum Congress, 1997, John Wiley and Sons, 8. M.C. A1-Kinany and S.H. A1-Khowaiter, Proc. of 2nd Middle East Refining and Petrochemicals Conf. and Exhibition, PETROTECH'98, Bahrain, September 1998, vol. 1, p. 523-42. 9. M.C. A1-Kinany and S.H. A1-Khowaiter, Proc. of the International Symposium on Reaction Kinetics and Development of Catalytic Processes, Brugge, Belgium, April 1921, 1999, vol. 122, p. 375-84.

Studies in Surface Science and Catalysis 133 G.F. Fromentand K.C. Waugh (Editors) ~) 2001 Elsevier Science B.V. All rights reserved.

465

Diffusion Analysis of Cumene Cracking over ZSM5 using a Jetloop Reactor P. Schwan, P.J. Henry and K.P. Mrller* Department of Chemical Engineering, University of Cape Town, Private Bag, Rondebosch, 7701, South Africa. Cumene is cracked in a recycle reactor over commercial H-ZSM5 extrudates. A Thiele modulus approach is used to determine the diffusion coefficient and the intrinsic rate constant. The results are compared to those obtained from pulse experiments. A linear model for diffusion, adsorption and reaction rate is applied for reactants and products. In contrast to literature it is argued that if the Thiele modulus is greater than five, the system becomes over parameterised. If additionally adsorption dynamics are negligible, only one lumped parameter can be extracted, which is the apparent reaction constant found from steady state experiments. The pulse experiment of cumene is strongly diffusion limited showing no adsorption dynamics of cumene. However, benzene adsorbed strongly on the zeolite and could be used to extract transient model parameters which are compared to steady state parameters. 1. INTRODUCTION In reactions occurring within porous catalysts the reactants have to adsorb and diffuse to and from the active centres. Typically this diffusion-sorption behaviour is measured under inert conditions at temperatures well below that at which reactions occur. It is preferable to study catalysts under reaction conditions. Such data can be interpreted by reaction-diffusionadsorption models, the so-called "Thiele modulus" approach [ 1]. The model parameters can be determined by using steady state or transient measurements. Thiele's analysis [1], has been applied under steady state conditions by a number of workers (Haag et al. [2], Weisz [3], Voogd and van Bekkum [4], Garcia and Weisz [5]) to estimate diffusion in zeolite systems under reaction conditions. However, it is necessary to vary crystallite or pellet size in order to estimate the diffusion coefficient. In the case of zeolite catalysts, this is extremely difficult to accomplish without changing the intrinsic reaction behaviour. Furthermore this method cannot distinguish between adsorption and intrinsic reaction constants. Transient techniques have been developed [6-9] which allow the estimation of rate parameters using gradientless reactors. A number of workers [7,8] claim that all model parameters can be derived from transient experiments by zero, first and second moment analysis. This paper investigates the simultaneous estimation of reaction, adsorption and diffusion coefficients for cumene cracking over ZSM-5. Transient and steady state data are compared. 2. EXPERIMENTAL The Jet-Loop is an internal recirculation reactor (volume = 50ml), which approximates CSTR behaviour [10]. Argon (Ar) (>99.995%) was used as carrier gas flowing through the jet at 175 to 500 ml/min (STP). H-ZSM5 extrudates (Stidchemie T4480, Si/Al=25, 50 vol%

Corresponding author" [email protected]

466 binder) were used. Reactions were carried out at atmospheric pressure, n-decane (20ml/min, 0.1 kPa in argon) was fed to the effluent as internal standard.

2.1

Steady State Cumene (0.25 kPa) diluted in Argon was fed with a constant flow of 20 ml/min (STP) into the reactor, while the flow rate of the jet stream was varied. Samples were taken with a syringe and analysed in a GC. Carbon mass balances were better than 95%. The extrudates were crushed to three different sizes (diameter = 0.15, 0.085, 0.065 cm). The reactions were carried out for 3 hrs. The extrudates were calcined at 480~ in air for 12h between experiments. Kinetic studies with the powder form (0.5 gm) were measured in a plug flow reactor at a constant contact time (WHSV=0.7 g cat/(g hr) ) 2.2

Pulse Experiments

0.5gl of liquid cumene was injected into an injector port heated to X X 300~ and flushed with 20ml/min Ar. For the 0.1 o sorption studies of o benzene, 150 gl of 0.01 ! vapour was injected. The effluent gas was 0.001 analysed with an 0 10 20 30 40 50 60 automated rotating t[s] Multi-Ampoule-Sampler Fig 1 9Benzene blank response curves at a residence time of system (MAS), where 2.2s (o) and 5.8s (A,x) measured by the Multi-Ampoulepre-evacuated ampoules Sampler technique compared to online FID ( ). were broken, sealed and later analysed in a GC. The fastest sampling time was ls. Fig 1 shows that the online FID and MAS are in excellent agreement. Thus the MAS can be used with confidence to analyse reaction data. Further details regarding the MAS is in preparation. 3. MODEL Model assumptions include the following : (i) The adsorbent has an uniform bidisperse pore structure, (ii) The pellets have spherical geometry, (iii) The reactor behaves like a CSTR, (iv) Ideal pulse input, (v) Macropore diffusion is Fickian, (vi) No external film resistance, (vii) Linear equilibrium, (viii) First order irreversible reaction, (ix) The crystals are small ( HSO: (1)

SO; + 1/2 O= -+ SO 4 (2) When the pH is relatively low, as it occurs in flue gas desufurization processes, the prevailing sulfurous species in solution is bisulfite ion (HSO3-) and the oxidation may be described by equation (1) with the following rate expression: kt"1/2p 31D r . . . . M "~s(w) (3) where r is the reaction rate expressed as mole of SO4= produced per unit time and volume, k is the kinetic constant, CM the catalyst concentration, and Csav) the total sulfite concentration. In the case of the uncatalyzed oxidation of calcium bisulfite, the reaction rate can be expressed as in equation (3) by neglecting the catalyst concentration and substituting the concentration of the total sulfite with the concentration of bisulfite ions [10]. Experimental results considered in the present work refer to the uncatalyzed oxidation of calcium bisulfite performed in a laboratory scale stirred reactor with lines for continuous feeding and discharging of both gas and liquid phase [ 10]. The gas was pure oxygen or mixture of oxygen and nitrogen with oxygen concentration of 40% or 21% volume. The liquid phase was a clear solution prepared by dissolving calcium hydroxide into sulfur dioxide and bidistilled water. 4. RESULTS AND DISCUSSION As discussed above, the computer code developed by Markos et al. [6] on the basis of previous results [7-9] is capable of handling any type of kinetic scheme and reaction rate expression, including the specific case of zero-th order with respect to the gaseous reactant. In this case, which applies to both the p-xylene and the bisulfite oxidation, the computer code compares the oxygen mass flux entering the liquid bulk with the maximum possible rate of oxygen consumption calculated from the reaction rate expression. On the basis of this comparison, the computer code distinguishes between chemically and mass transfer controlled regime. Thus, if the oxygen flux entering the liquid bulk is larger than the maximum possible rate of oxygen consumption, the reactor is recognized to operate under chemically controlled regime. On the contrary, if the flux entering the liquid bulk is lower than the maximum possible rate of oxygen consumption, the reactor is recognized to operate under the mass transfer controlled regime, and the reaction rates become proportional to the ratio of the corresponding oxygen mass flux entering the liquid bulk and the maximum possible rate of oxygen consumption. Figure 1 shows a comparison between computer code results and experimental data [8] in the case of p-xylene oxidation, where the reaction is considered of zero-th order with respect

475 to the gaseous reactant and of first order with respect to the liquid reactants. The overall rate of oxygen uptake is reported in an Arrhenius-like plot as a function of temperature for different values of the gas feed compositions (i.e. 21% and 100% volume of oxygen). At low temperature the reactor operates under chemically controlled regime (straight line) and the rate of oxygen uptake is the same for pure oxygen and air. On the other hand, at higher temperature values (i.e. T>105~ diffusional limitations become important. These arise for the oxidation performed with air as indicates the deviation from the linearity (mass transfer controlled regime), while when a gas feed with pure oxygen is used data lie on the same straight line for all range of temperature (chemically controlled regime). Model parameters used for simulation are reported elsewhere [8,9].

10 .2

0 2 100% vol

0 2 21% vol

ext) o 10.3 O

A

exp. data

model results

0.0024

0.0025

0.0026

0.0027

0.0028

0.0029

1/T [ K " ]

Figure 1. Oxidation of p-xylene. Comparison between computer code results and experimental data [8] for two gas feed composition (100% and 21% volume of oxygen). In Figure 2 the comparison between computed results and experimental data [ 10] is shown for the uncatalyzed oxidation of calcium bisulfite where the reaction is considered of zero-th order with respect to the oxygen and of three halves order with respect to the bisulfite ion. The reaction rate is reported as a function of the bisulfite ion concentration for different composition of gas feed (i.e. 21%, 40% and 100% volume of oxygen). Model parameters used for simulation are reported elsewhere [10]. Fol; relatively low values of bisulfite ion concentration experimental data and computed results lie in a straight line of slope equal to 1.5 (chemically controlled regime). By increasing the concentration of bisulfite ion (up to about 27 mol/m 3 when the gas feed contains 21% volume of oxygen, up to about 46 mol/m 3 for oxygen at 40% volume and up to 100 mol/m 3 for pure oxygen) the reactor starts to operate under mass transfer controlled regime. Under these conditions, a further increase of the bisulfite ion concentration does not produce any effect on the oxidation rate, which displays an upper limit. Figures 1 and 2 clearly show that the computer code properly describes the

476 transition from kinetic to mass transfer regime induced by an increase of reaction temperature for the case of p-xylene oxidation (cf. Figure 1) and of HSO3 concentration for the case of bisulfite oxidation (cf. Figure 2). It should be noted that these results are accomplished neither by changing the reaction order with respect to the gaseous reactant nor defining specific diffusional rate equations where the product distribution is neglected.

02 100% vol

T = 45 ~ 10q

02 40% vol

A--A

r--..I

~

_A

02 21% vol

r~ 10: o

10:

10 4

10~

o: i

.

o ,

,

i

,

, , I

.

.

.

.

I0'

E!

/k

exp.

data

model predictions

.

,

, , I

,

,

.

,

10 2

CHSO; [ m o l / m 3]

.Figure 2. Uncatalyzed oxidation of calcium bisulfite. Comparison between computer code results and experimental data [ 10] for different compositions of the gas feed (100%, 40% and 21% volume of oxygen). REFERENCES

1. 2. 3. 4.

N.M Emanuel and D. Gal, Akademiai Kiado, Budapest, 1986. P. Cavalieri D'Oro, E. Danoczy and P. Roffia, Oxidation Communications, 1 (1980) 153. M. Morbidelli, R. Paludetto and S. CarrOt, Chem. Eng. Sci., 41 (1986) 2299. A.Z. Kryzstoforski, Z. Wooichik, R. Poharecki and J. Baldyga, Ind. Eng. Chem. Proc. Des. Dev., 25 (1986) 894. 5. G. Astarita, Mass transfer and Chemical Reaction, Elsevier, Amsterdam, 1967. 6. J. Markos, M. Pisu and M. Morbidelli, Comp. & Chem. Eng., 22 (1998) 627. 7. J. Markos, M. Pisu and M. Morbidelli, in METECC 94, Vol. C, Ed. E. Clementi, Stef, Cagliari, (1993) 167. 8. G. Cao, A. Servida, M. Pisu and M. Morbidelli, AIChE Journal, 40 (1994) 1156. 9. G. Cao, M. Pisu and M. Morbidelli, Chem. Eng. Sci., 49 (1994) 5775. 10. A.Lancia, D. Musmarra and F. Pepe, Chem. Eng. Sci., 51 (1996) 3889. 11. C.H. Barron and H.A. O'Hem, Chem. Eng. Sci., 21 (1966) 397. 12. S. Bengtsson and I. Bjerle, Chem. Eng. Sci., 30 (1975) 1429. 13. G.C. Mishra and R.D. Srivastava, Chem. Eng. Sci., 31 (1976) 969. 14. W. Pasiuk-Bronikowska and T. Bronikowski, Chem. Eng. Sci., 44 (1989) 1361. 15. A.Lancia, D. Musmarra, F. Pepe and M. Prisciandaro, Chem. Eng. Journal, 66 (1997) 123.


477

Evaluation of external diffusional effects in a microreactor for S R G O hydrotreating L.C. Castafieda-L6pez a,*, F. Alonso-Martinez

a, j. Ancheyta-Jtuirez a,b

a Instituto Mexicano del Petr61eo, Eje Central L~izaro Cfirdenas 152, 07730 M~xico, D.F., MEXICO b Instituto Polit6cnico Nacional, ESIQIE, 07738 M6xico D.F., MEXICO In this work a study about the effect of external diffusion on the hydrodesulfurization reaction of straight run gas oil is presented. The experimental study was carried out at typical lIDS operating conditions over a commercial catalyst using an upflow and downflow tricklebed microreactor (TBMR). The experimental results showed that upflow is the best system for conducting kinetic studies in order to avoid interphase mass transfer limitations. 1. INTRODUCTION Trickle-bed reactors (TBR) are commonly used for hydrotreating of crude oil fractions, such as straight run gas oil (SRGO), vacuum gas oil, and residuals. Experimental information obtained in small-scale TBR is very important for reactor design, for catalyst screening, for evaluating alternative feedstocks, for estimating kinetic parameters, for studying the effect of operating conditions on product selectivities, etc. Microreactors are preferred for conducting such investigations mainly due to economic benefits associated with them. In trickle-bed microreactors (TBMR), it is frequently used a commercial catalyst sample and real feedstocks, however, the length of the catalyst bed and hence the reactor length to catalyst particle diameter ratio are low as compared to commercial reactors. In addition, low liquid velocities are used in order to match the liquid hourly space velocities (LHSV) of industrial units. These differences cause number of problems in testing catalyst having commercially applied size and shape, such as poor wetting of catalyst, wall effect, axial dispersion, maldistribution, and the data obtained in such TBMR may not be reliable [ 1,2]. The use of small size of catalyst particles is helpful for reducing these phenomena, this means that the commercial catalyst sample has to be crushed in order to test it in the form of fine particles in TBMR. However, the data generated with crushed catalyst may not be representative for industrial practice, since pore diffusion limitation, which may occur with commercial catalyst size, will exhibit a higher activity with crushed sample [3]. Upflow and downflow reactors are often used in laboratory-scale studies, however, few comparative studies between these two systems are reported in the literature. Dudukovic et al [ 1] summarized some of these reports. They have found that the upflow reactor behavior can be better than downflow reactor and viceversa depending on gas and liquid velocities, level of Author to whom correspondenceshouldbe addressed, fax: (+52-5) 333 8429, e-mail: [email protected]

478 conversion and the type of reaction system. They also have concluded that there is no clear guidance as to which reactor will perform better for a given reaction system. In order to understand the differences between upflow and downflow reactors for a given reaction, more systematic studies need to be done. In this work we present a comparative study of a TBMR with downflow and upflow, using the hydrodesulfurization of SRGO as test reaction. 2. EXPERIMENTAL The feedstock used in this study was a SRGO recovered from an HDS industrial unit, and its properties are shown in Table 1. The catalyst used was a trilobe extrudate commercial available Ni-Moh/-Al203 sample (surface area of 204 mZ/g, pore volume of cm3/g, 9.5 wt% MOO3, 2.4 wt% NiO). Table 1. Characterization of the feedstock (SRGO) Property Specific Gravity 20/4~ Elemental Analysis, wt% C H O N S Aromatics, wt% ASTM Distillation, ~ IBP/10 vol% 30/50 70/90 FBP

Value 0.8687 85.08 13.00 0.250 0.054 1.616 39.71 153/297 317/327 337/353 365

Prior to processing, the catalyst was in-situ sulfided with a desulfurized naphtha contaminated with 0.6 wt% CS2 at 54 kg/cm 2, a hydrogen-to-oil ratio of 2000 fi3/bbl, at 230~ liquid hourly space velocity of 3 h l , 12 h, to ensure complete catalyst presulfiding. When the presulfiding was completed, the feedstock was introduced and presulfiding feed was stopped without stopping the hydrogen flow, and the reactor temperature and other conditions were adjusted to the desired start-of-run conditions. The hydrotreating was performed in a fixed-bed upflow and downflow reactor, which was operated in isothermal mode by independent temperature control of a three-zone electric fumace. The SRGO hydrotreating was carried out at constants reaction pressure and hydrogen-tooil ratio without hydrogen recycle (54 kg/cm a and 2000 tt3/bbl, respectively). Reaction temperature was studied in the range of 340-380~ To study the effect of external diffusional effects different amount of catalyst and liquid flow were used.

479 3. RESULTS AND DISCUSSION In order to compare the behavior of both reactions systems, upflow and downflow, and verify the presence of interphase limitations, the experiments were conducted with different amounts of catalyst and reaction temperature as can be seen in Figure 1. These experiments are based on the principle that in the absence of interphase transport limitations, the conversion at any LHSV must be independent of the linear velocity through the bed [4]. Two series of tests were run, first using the minimum amount of catalyst in the TBMR (14 ml), and second, With the maximum amount of catalyst (29 ml). In each series, LHSV is varied and two curves are traced for conversion vs. 1/LHSV. If the two curves overlap there are no interphase limitations [4]. According to Figure 1, the curves of liDS vs. 1/LHSV never overlap for downflow system. The opposite behavior was found when the upflow system was used, since the curves for maximum and minimum amounts of catalyst loaded in the reactor are very close. The same behaviour was found at low and high reaction temperatures. This is mainly due to the better wetting of the catalyst and the solvent vaporization in upflow systems [5]. It can also be observed from Figure 1 that when the LHSV is decreased, and the reaction temperature is increased, both systems give almost the same results. These experimental data indicate that there are no external diffusional effects in upflow system.

T=340~

T=380~

100

100

95

~

95

90 85

90

80 85 75 70 0.3

,

,

.

0.4

0.5

0.6

.

. 0.7

1/LHSV, h

. 0.8

.

80 0.9

1

1.1

. 0.3

0.4

. 0.5

.

. 0.6

.

.

0.7

. 0.8

0.9

1

1.1

1/LI-LqV, h

Figure 1. Effect of LHSV and temperature on HDS in upflow (dotted lines) and downflow (solid lines) systems (circles: minimum amount of catalyst, squares: maximum amount of catalyst). Figure 2 shows some experimental results about the effect of both reaction temperatures and LHSV on SRGO hydrodesulfurization using upflow and downflow systems. It can be clearly seen the differences in sulfur removal between these two systems, specially at low temperature and high space velocity. This observation is very important because kinetic studies should de carried out preferably in upflow system in order to avoid mass transfer limitations (interphase). Of course, additional experiments have to be performed to study intraphase gradients consists of determining the conversion for particles of different size at constant space velocity [4]. If conversion is constant indicates that the system is under chemical kinetic control.

480

100 9590858075 706560 330

LHSV

~..,.~~--~ I

340

I

I

350

360

I

370

1.0 2.5

I

380

390

Temperature, ~ Figure 2. HDS at different LHSV and reaction temperature for upflow (dotted lines) and downfiow (solid lines) systems.

4. CONCLUSIONS Experimental results about the hydrodesulfurization of straight run gas oil over a commercial Ni-Mo/AI203 catalyst are presented. The tests were conducted in an upflow and downflow microreactor at pressure of 54 kg/cm2, hydrogen-to-oil ratio of 2000 ft3/bbl. LHSV and reaction temperature were studied on the renges of 1-2.5 and 340-380~ respectively. External diffusional effects were examined by varying the amount of the catalyst loaded in the reactor and the space-velocity. It was found that the best system to minimize interphase gradients is upflow since very similar sulfur removals were obtained when the amount of catalyst in the reactor was changed. It was found that both systems, upflow and downflow, give similar results at high temperature and low space velocity, which was attibuted to wetting of catalyst and higher solvent vaporization at high temperature. REFERENCES

1. Y. Wu, M.R. Khadilkar, M.H. A1-Dahhan, M.P. Dudukovic, Ind. Eng. Chem. Res., 35 (1996) 397 2. S.K. Bej, R.P. Dabral, P.C. Gupta, K.K. Mittal, G.S. Sen, V.K. Kapoor, A. K. Dalai, Energy & Fuels, 14 (2000) 701 3. S.T. Sie, AIChE J., 42, 12 (1996) 3498 4. C. Perego, S. Peratello, Catal. Today, 52 (1999) 133 5. D. Letourneur, R. Bacaud, M. Vrinat, Ind. Eng. Chem. Res., 37 (1998) 2662


481

A Study of a Phase Transfer Catalytic Reaction between N-Butyl Bromide and Sodium Phenolate in an Oscillatory Baffled Reactor B. Wilson 1, X.Ni 1. and D.C.Sherrington 2 1Department of Mechanical and Chemical Engineering, Heriot-Watt University, Edinburgh, EH14 4AS, Scotland, UK. 1Department of Applied Chemistry, Strathclyde University, Glasgow, G1 1XJ, Scotland, UK. Abstract: The most basic form of a Phase Transfer Catalytic (PTC) reaction consists of

two immiscible liquid phases each with a reagent that, until the introduction of a Phase Transfer Catalyst (PT-Cat), cannot react with each other. The PT-Cat manages to transfer one of the reagents from one phase to the other phase whereupon the second reagent can react with the catalyst-reagent transient. This paper presents our current investigation into the use of an Oscillatory Baffled Reactor (OBR) for a PTC reaction between n-butyl bromide (BuBr, with Toluene as the organic solvent) and sodium phenolate (NaOPh, dissolved in water) with tetrabutylammonium bromide (QBr) as the catalyst. The OBR consists of a cylindrical column containing periodically placed orifice baffles together with superimposing fluid oscillation. The OBR is of 25 mm in diameter and 140 mm in length, and is operated vertically. The oscillating fluid motion in the OBR interacts with each baffle to form vortices, which gives efficient and uniform mixing between each baffled cavity. It has been known that PTC process development remains - almost exclusively- a chemistry based field, and has not yet been examined with chemical engineering principles. This work is the first of its kind. Industrial applications of PTC are concentrated upon the manufacture of Fine Chemicals and Organic Intermediates. Recent reviews identify the potential of PTC in these fields as 'almost unlimited'. The reasons for such an endorsement are the advantages of cost and environmental friendliness that PTC enjoys over existing technologies. In this paper, we examined the effects of oscillation amplitude and frequency on the reaction rate and conversion of the PTC reaction, and compared the results with those carried out in a Continuously Stirred Tank Reactor (STR). Our preliminary results show that the OBR promotes improved reaction rates and conversion at lower power requirements. 1.INTRODUCTION Oscillatory Baffled Reactors (OBRs) are based on the concept of superimposing fluid oscillation onto a cylindrical column containing periodically placed orifice baffles, and are a relatively recent innovation to reactor technology. The key feature of OBR is that mixing can be controlled to a very high degree of precision, either in a dynamic sense by altering frequency and amplitude, or by customizing the baffle geometry. This allows a wide range of mixing conditions to be achieved, from "soft" mixing, exhibiting plug flow characteristics, to the most intense, corresponding to the mixed flow (a single STR) The corresponding author. Tel: 441314513781, Fax: 441314513077, email: [email protected]

482 (Brunold et al. 1989, Mackley and Ni, 1991, 93, Ni, 1995, Ni and Pereira, 2000). The mechanism of mixing in OBC can be understood with the help of Figure 1.

Figure 1 Mixing mechanism of Oscillatory Baffled Column The essential feature is that sharp edges (provided by the baffles) are presented transverse to an oscillating, fully reversing flow. Because the motion is periodic and fully reversing, there are two half cycles, each containing flow acceleration and deceleration corresponding to a sinusoidal velocity-time function. On each flow acceleration, vortex tings form downstream of the baffles. A peak velocity is reached, and then as the flow decelerates, the vortices are swept into the bulk, and subsequently unravel with the bulk flow acceleration in the opposite (axial) direction. It is the strong radial velocities, arising from the repeating cycles of vortex formation, and of similar magnitude to the axial velocities, that gives uniform mixing in each inter-baffle zone and cumulatively along the length of the column. The OBR has been applied to various chemical engineering systems (Ni et al. 1999, 2000a,b), this paper deals with the first example of OBR in conjunction with Phase Transfer Catalysis (PTC) reaction systems. The most basic form of PTC consists of two immiscible liquid phases each with a reagent that, until the introduction of a Phase Transfer Catalyst (PT-Cat) cannot react with each other. The PT-Cat manages to transfer one of the reagents from one phase to the other phase whereupon the second reagent can react with the catalyst-reagent transient. In our study, we elect the reaction between n-butyl bromide and the phenoxide ion (of sodium phenolate trihydrate) as a model liquid-liquid PTC reaction scheme. This system has previously been investigated (Wang & Weng, 1996a) in a continuously stirred flask, as a typical reaction system reported for PTC scheme. They also identified tetrabutylammonium salts as the most efficient phase transfer catalyst. Our investigation is to perform the similar PTC reaction in an OBR and compare the conversion achieved in both reactors.

483 2.EXPERIMENTAL PROCEDURE 2 . 1 0 B R and STR

A stainless steel OBR was used in the investigation. The column is approximately 48 mm in diameter with an external-heating jacket supplied with hot water. Mixing is provided by oscillating the baffles at the top of the column. A motor with a linking cam powers the baffle set. The orifice baffles have a 20 % free baffle area and placed at an equal distance of 1.5 times the column diameter along the column. Oscillation frequencies ranged from 2.5 to 6.5 Hz and oscillation amplitudes from 5 to 20 mm (centre-to-peak) can be achieved. Sampling during reaction is achieved through means of a 150mm long needle syringe at 2 ports fixed along the column. The STR consists of a glass column of 50 mm in diameter with a 6 bladed agitator of diameter 36mm. The column is mounted with 4 baffles equally spaced round the circumference and along the height of the column. The rotation of the agitator is controlled by an IKA Eurostar Basic mixer unit. An external heating jacket supplies the energy of activation, with sampling achieved through the use of a pipette through lid ports. 2.2 Chemicals

All chemicals used in these trials; toluene, N-butyl bromide, sodium phenolate trihydrate (NaOPh.3H20), tetrabutylammonium bromide ((C4H9)4 NBr)(TBAB) and butyl phenyl ether, were purchased from Sigma-Aldrich chemicals (U.K.). 2.3 Methods and Procedures

The reaction mixture is prepared with known quantities of Water, Toluene, Sodium phenolate trihydrate, and QBr (TBAB) all introduced into the column. The reactor contents are thermostated at 65~ (+/- 0.1 C) by the external heating jacket. The reaction was then initiated by the addition of a known amount of N-BuBr, also thermostated at 65~ (+/- 0.1 C) and agitation started. Sample ports allow the extraction of 2ml of the reaction mixture at predetermined intervals. These samples were then evacuated into 3 ml of cooling water and the two liquid phases (organic and aqueous) allowed to separate in a sealed vial. A sample of the organic phase (extracted immediately thereafter) from the separate layer was then analysed using gas chromatography (Perkin Elmer GC-2500 with a 2m long column and 2 mm internal diameter, using packing of SP 2100 of Supelcoport 80-100 Mesh). Thus the concentration of the reagent (N-BuBr) in the organic phase can be measured and the reaction rate determined. 3.REACTION M E C H A N I S M S This PTC matrix has been used to investigate the special case of TLPTC (Tri-Liquid Phase Transfer Catalysis), and previous studies from associated literature [9-14,16] show a variety of mechanisms available when conditions inducing a third liquid phase to form are invoked. However, in this elementary study only a two-phase system was investigated. This leaves only two mechanisms available through which the desired intrinsic reaction can proceed. The first takes place predominantly in the organic phase and could be considered the catalysed system:

484

Org

N-BuBr

+

QOPh

~

BuOPh

+

QBr

[AI Interface

Aq

NaBr

+

QOPh

-~

NaOPh

+

QBr

In contrast to the mechanism [A], the mechanism [B] shown below is non-catalysed and takes place at the organic-aqueous phase interface only.

N-BuBr

BuOPh

[B]

Org N-B~a-B r - - - + - - - N a O P t ~ -

.... ,-

-BtagPh--- §

. . . . . . . . . haterfaee---

Aq NaOPh

NaBr

The mechanism [B] is assumed to be significantly slower than that of the catalysed system. 3.1 R a t e E x p r e s s i o n

The intrinsic reaction- that which occurs in the organic phase in the mechanism [A] has been shown by Starks (1971) to be an example of a bimolecular SN2 substitution reaction. Treated as such the reaction can be considered a second order with respect to the concentrations of either of the two organic phase reactants (N-BuBr and QOPh). This implies that the substrate (N-BuBr) and the nucleophile (NaOPh) react directly via a transition state to produce the product (Butyl Phenyl Ether). Expressed with respect to the concentration of the organic phase reactant the reaction rate in the organic phase is:

d [N - BuBr ]o =k[QOPh]o[N-BuBr]o dt

(1)

Where the quantity of NaOPh is in excess with respect to that of N-BuBr for the system, the concentration of QOPh can be assumed constant throughout, thus

d [N - BuBr ]o dt

=k'[N-BuBr]o

(2)

485 The other rate expression to be taken into consideration concerns reaction the mechanism [B]. In this situation the nucleophile (NaOPh) exists in the aqueous phase whilst the substrate (N-BuBr) remains in the organic phase only. Thus

4 d[N-dtBUBr]o = ki' IN- BuBr]o [NaOPh], "-Vo

(3)

Again some simplifications can be made to this expression where the values of [NaOPh]a, Ai, and Vo are assumed constant throughout the reaction. Thus:

d[N-BuBr]o dt =ki[N-BuBr]~

(4)

Where

4 ki = k; [NaOPh la "-Vo

(5)

Both rate expressions are pseudo first order terms, and thus the results are tested by comparison with a 1st order kinetic model. 4.RESULTS AND DISCUSSION The 1st order kinetic model used to test the order of the reaction is, as shown in Figure 2 below, a plot of-In(l-X) versus time (where X represents the percentage conversion of Butyl bromide). The plot shown compares a STR and OBR trial, with the value of the rate constant given by the slope of the line. 2.5

-

9 STR

_

o OBR ~

~'/1.5~

= !

1~

0.5 |

0

I

'

'

50

100

150

' Time

(mins)

200

Figure 2: Comparison of the effect of reactor type on the reaction rate. It is clear that the data suggests pseudo 1st order behaviour of the system. However, an unusual feature of both plots is the non-zero origin. This is due to a significant

486 conversion of N-BuBr in the first few moments of the reaction and is discussed in more detail later. The percentage conversion of Butyl Bromide (the substrate in the organic phase), with respect to the energy dissipation per unit mass of the system, e, are evaluated for the two reactor types. The expression for e are given below

S=~Vp

p3D3

p

6.

.

3

.

1~ ( oo,/3 ct 2 It

(W/kg)

(OBR)

(6)

(W/kg)

(STR)

(7)

5

PoN Ds .

Figure 3 shows the conversion of N-BuBr only one minute aider the reaction has started. The results show a significant substrate conversion occurs in both reactors with the higher conversion values occurring in the OBR. 60o ,m

r~

==, o

~Z

sSTR o OBR

40 20

, w (W/kg)

0

0.5

1

Figure 3: Comparison of N-BuBr conversion 1 minute after reaction is initiated. It can be seen that both systems have shown the early conversion, instead of zero. This is similar to the around about 25% conversion by Hsiao & Wang (1999) in their investigation of TLPTC for a very similar system. The enhanced reaction rates of the TLPTC system are reportedly due to the concentration of the PTC in the third liquid phase formed. So far we are unable to explain the phenomena satisfactorily in our two phase system. Further work is being carried out to further examine the very first few seconds of the reaction. 5.CONCLUSION We have reported the PTC reaction system in an OBR. The preliminary results show that higher initial conversion rates are seen in OBRs than in STRs, although the same reaction kinetics are found post the initial conversion. It is envisaged that the on-going research

487 on this subject will provide more comprehensive results and a probable hypothesis for this finding. 6.NOMENCLATURE ct Area ratio Ai Interfacial Area between organic and aqueous phases. D,Ds,Dv Diameter of Oscillatory Column, STR agitator, and STR. ko 2 no order reaction rate constant in the organic phase. ko' Pseudo 1st order reaction rate constant in the organic phase ki' 2nd order reaction rate constant of the uncatalysed reaction taking place on the organic-aqueous phase interface. ki Pseudo 1st order reaction rate constant of the uncatalysed reaction taking place on the organic-aqueous phase interface. L Height of Liquid in STR (m). n Number of Baffles per unit length of OBC N Rotational Speed of Agitator P Power requirement for mixing Po Power number 9 Density of mixture Vo,V Organic phase volume, Total Volume of liquids in reactor. X Fractional conversion of N-BuBr Xo,C0 Centre to peak amplitude (m) and angular frequency (rad/s) of oscillation. ~t Dynamic Viscosity of reaction mixture. [] Concentration of species. []a in the aqueous phase []o in the organic phase () Species. 7.LITERATURE CITED 1. Brunold, C.R.; Hunns, J.C.B.; Mackley, M.R.; Thompson, J.W. Chem Eng Sci (1989), 44, 1227. 2. Mackley, M.R.; Ni X. Chem Eng Sci (1991), 46, 3139. 3. Mackley, M.R.; Ni X. Chem Eng Sci (1993), 48, 3293. 4. Ni X. J.Chem.Tech. Biotechnol., (1995), 64, 165. 5. Ni X; Pereira N.E AIChE Journal. (2000) 46(1), 37. 6. Ni X.; Cosgrove J.A.; Arnott A.D.; Greated C.A. Cumming R.H. (2000a) Chem Eng Sci, 55, 3195. 7. Ni X; Bennett D.C.; Symes K.C.; Grey B.D. (2000b) Journal of Applied Polymer Science, 76, 1669. 8. Ni X; Zang Y and Mustafa I, (1999) Chem. Eng.Sci., 54, 841. 9. Starks C.M.; J.Am. Chem. Soc. (1971), 93, 195. 10. Starks C.M.; Liotta C.; Halpern M; Phase Transfer Catalysis, Fundamentals, Applications and Industrial Perspectives, Chapman and hall, New York (1994) 11. Wang D.H.; Weng H.S.J.Chin.Inst.Chem.Eng (1995a), 26, 147 12. Wang D.H.; Weng H.S. Chem.Eng.Sci. (1995b), 50, 3477 13. Wang D.H.; Weng H.S.J.Chin.Inst.Chem.Eng (1996a), 27, 129 14. Wang D.H.; Weng H.S.J.Chin.Inst.Chem.Eng (1996b), 27, 419 15. Weng H.S.; Wang C.M; Wang D.H.; Ind. Eng. Chem. Res.(1997), 36, 3613 16. Hsiao H.C.; Weng H.S; Ind. Eng. Chem. Res.(1999), 38, 2911



489

KINETICS OF THE MAIN AND SIDE REACTIONS OF THE METHANOL OXIDATION OVER IRON MOLYBDATES A. P. Vieira

S o a r e s a,

M. Farinha Portelaa*, A. Kiennemann b

aGRECAT-Grupo de Estudos de Catfilise Heterog6nea, Universidade T6cnica de Lisboa, Instituto Superior T6cnico, Avenida Rovisco Pais, 1049-001 Lisboa, Portugal, Fax: 351-218477695, e-mail: [email protected]; [email protected] bECPM-LERSCI, UMR CNRS 7515, 25, Rue Becquerel, BP 08, 67087 Strasbourg Cedex 2, France, e-mail: [email protected] A kinetic study of methanol oxidation over stoichiometric iron molybdate catalyst was performed in a fixed-bed integral reactor showing kinetic influences of reaction products. In the temperature range of 548-618 K it was not possible to fit the formation rate data to a single power rate law. Dimethyl ether formation presents only a second order dependence with respect to methanol. CO formation seems to be inhibited by water and formaldehyde and rate data fit well to the power rate law:

Such law seems to point out that in tested conditions methanol direct overoxidation to CO is more important than formaldehyde oxidation. INTRODUCTION Iron molybdates are industrial catalysts for selective oxidation of methanol to formaldehyde [I]. Since the pioneer work of Atkins et al [2] in 1931, a large number of papers and patents has been published and the catalytic systems his relatively well characterised. As recently stated by Tatibou~t [3], methanol oxidation can lead to several products depending on the catalyst, reaction temperature, contact time and reactants (O2 and methanol) partial pressure. In spite of the high selectivity towards formaldehyde displayed by iron molybdate catalysts, dimethylether (DME), methylformiate (MF) and dimetoxymethane (DMM) are often referred as reaction by-products [3,4]. Formic acid is rarely observed. Deep oxidation products CO and CO2 are produced in severe reactional conditions. Cheng [4] has shown that in analogous conditions, iron molybdates and MoO3 present similar selectivities toward formaldehyde. In a general way all researchers agree with the fact that by-products are due to long residence time of intermediate species on catalyst surface and its formation is favoured by low reaction temperature, low conversion and long contact time. Several reaction schemes have been proposed to explain the formation of all byproducts during methanol oxidation over Mo-Fe catalysts. Edwards et al [5] and Machiels [6] have suggested reactional mechanisms for the formation of formaldehyde and by-products. However intermediate species proposed in such mechanisms have not been identified by any spectroscopic or other techniques. More recently Busca [7] on basis of infrared studies of surface intermediates species has proposed a rake-type mechanism for methanol oxidation over oxide catalysts. This mechanism account for the formation of formaldehyde and byproducts.

* Corresponding author

490 Many kinetics studies have been performed for methanol to formaldehyde oxidation over iron molybdates. The majority of authors accept a Mars van Krevelen mechanism. However a great number of kinetic laws have been presented for this reaction. In fact the authors do not agree with respect to the kinetic orders of methanol and oxygen. This apparent disagreement is probably due to the fact that kinetic studies of each research group were performed in different conditions and over catalysts with different Mo/Fe atomic ratios. Additionally product effects, like water and formaldehyde, on the oxidation rate of methanol remain ambiguous. For the above-mentioned by-products only the CO formation has been object of kinetics studies, what can be attributed to the fact that this product is the major by-product usually formed. EXPERIMENTAL

Catalyst preparation and characterisation Stoichiometric iron molybdate (Fe2(MoO4)3) was prepared by coprecipitation from aqueous solutions of iron (III) nitrate and ammonium heptamolybdate; it was characterised by BET analysis, XRD, atomic absorption spectrophotometry, XPS, EDX, SEM, FTIR and Mfssbauer spectroscopy. Surface acidity was characterised by NH3 TPD (temperature programmed desorption). Preparation and characterisation procedures have been described elsewhere [8].

Catalytic m e a s u r e m e n t s Methanol oxidation was carried out in a conventional flow apparatus at atmospheric pressure, using a tubular pyrex reactor. The feed mixtures were prepared by injecting liquid methanol into airflow using a Gilson 302 precision pump. The catalyst bed (0.5 g) was diluted (1:7 w/w) with inert SiC to avoid adverse thermal effects. A thermocouple was placed in the centre of catalyst bed for measuring the reaction temperature. Kinetic measurements, in steady state conditions, were performed for temperatures (TR) and contact times (x) in the 2 .h. molM~ -1 H respectively. Methanol concentration, in range of 548-618K and 28-98 mca reactor feed, was kept constant (4% molar in air). The reactor outlet was heated (413 K) to avoid liquid products condensation and formaldehyde polymerisation. Reaction products were analysed by GC using a TCD detector. RESULTS AND DISCUSSION

Catalyst characterisation Detailed results of the catalyst characterisation were reported elsewhere [8]. Mo-Fe mixed oxide prepared by coprecipitation with bulk Mo/Fe=l.5 atomic ratio (A.A. analyse) had a surface area of 2.8 m~/g and a XRD pattern attributable to Fe2(MoO4)3 pure phase. Phase purity was confirmed by Mfssbuaer spectroscopy. EDX and XPS evidenced a slight Mo enrichment of superficial layers of catalyst.

Catalytic m e a s u r e m e n t s Due to the high activity of tested catalyst, conversion was always higher than 10%. Thus integral reactor was considered. In tested conditions only appreciable amounts of formaldehyde, DME and CO were detected. The product yields (Yi) were plotted (Fig. 1) versus contact time (x) for each studied reaction temperature (j). It must be noted that used contact time was computed with respect to surface area and not to catalyst weight. In such circumstances the products rates of formation were calculated by differentiation of the experimental curves Yij=fi(xj)

491

0f(r) r,j-( Or )j

(I)

In Table 1 are presented the computed rates of products formation and the corresponding reaction mixture compositions. Table 1 Rates of CO, HCHO DME formation and mixture compositions. T(K) rco r H C H O r D M E CMeOH , (~mol.sl.m "2) , " 3.4 3.2 0.6 0.51 4.3 3.3 0.8 0.59 5.1 3.4 0.9 0.65 oO 5.8 3.5 1.1 0.69 6.5 3.5 1.2 0.72

OO r,-. tt%

oO

corresponding reaction Co 2

CH20

CHCHO

4.15 4.19 4.24 4.24 4.26

(m01.m-3) 0.57 0.48 0.41 0.36 0.34

0.10 0.07 0.06 0.05 0.04

0.9 2.4 3.8 4.9 6.0

2.7 3.1 3.3 3.6 3.8

0.8 1.1 1.4 1.6 1.8

0.42 0.49 0.52 0.57 0.60

3.91 3.93 3.95 3.95 3.97

0.61 0.53 0.52 0.47 0.43

0.11 0.09 0.07 0.06 0.06

1.5 3.0 4.3 5.5 6.5

3.5 4.0 4.4 4.7 5.1

1.2 1.6 1.9 2.1 2.3

0.35 0.41 0.46 0.52 0.55

3.77 3.81 3.83 3.85 3.85

0.66 0.56 0.56 0.49 0.46

0.15 0.11 0.10 0.09 0.08

4.7 6.3 7.8 9.0 10.2

0.3 1.7 2.9 4.0 5.0

0.7 1.1 1.4 1.7 1.9

0.21 0.30 0.36 0.36 0.40

3.50 3.56 3.62 3.60 3.62

0.86 0.73 0.65 0.67 0.61

0.17 0.17 0.16 0.15 0.13

Fit of the rate data of Table 1 to single power rate laws of the type ~n,

~

r~

r = ,....MeOH._.H20._.HCHO

(II)

was tried. In fact the amount of the available experimental data does not recommended to test more complex kinetic laws. For formaldehyde it was not possible to fred an expression of the above mentioned type providing good fit for all the studied reaction conditions. This is not surprising because for the most part of the rate equations for formaldehyde reported in the literature are based upon redox and Langmuir-Hinshelwood [12] models. Plotting the rate of formaldehyde formation, reported in Table 1, versus methanol concentration, it is observed for 578 and 593 K an

492 apparent first order with respect to methanol (Fig. 2). Some authors report a water inhibition effect for law water concentrations [13] and a formaldehyde inhibition [4]. The DME formation rate data fits well to a simple power second order rate law: 2

FDME - - k C M e O H

(III)

with an activation energy of 69 kJ/mol and no products inhibition effect was observed (Fig. 2). Such kinetic law agrees with the SN2 mechanism generally proposed for this reaction. No published data have been found for this reaction over Mo-Fe-O catalysts. Plotting CO formation rate versus the formaldehyde concentration of the reaction mixture an irregular behaviour was observed. In fact CO can be formed via H C H O + ~1 0 2 --->CO + HEO C H 3 0 H + 02 --->CO + 2 H 2 0 1 0 z --->CO + 2 H 2 0 H C O O H + -~ H C O O C H 3 + 02 --~ 2 C 0 + 2 H z O

The fit of the CO rate data of Table 1 to a power rate law of the type (II) provide the rate law t,c,05 c~-~ c,-~

F "- o,,,..,MeOH "-" HCHO"" H20

(IV)

evidencing inhibition effects of water and formaldehyde. Such rate law does not fit well to the lowest temperature data. For the other temperatures a good fit was obtained and an activation energy of 151 kJ/mol was computed (Fig. 2). Bibin and Popov [14] reported that oxidation of formaldehyde to CO, over Fig. 1 - Formaldehyde, DME and CO yields the catalyst under consideration, could versus contact time. be described by a first-order kinetic law,

493 but Wilson [15] found a much better fit with the equation

ke~Lo ~CO -

-

(v)

~

1 + P~t4o

The importance of the above mentioned pathways for CO formation depends upon operation conditions. It seems that for the conditions of our data CO is formed mainly by direct oxidation of methanol instead formaldehyde oxidation. CONCLUSIONS The mechanism of the catalytic oxidation of methanol to formaldehyde seems to be dependent on the reaction conditions, namely on temperature. The same occurs for CO formation. Such formation seems to take place not only by formaldehyde oxidation, as is commonly accepted, but also by direct methanol overoxidation and is inhibited by water and even by formaldehyde. Dimethylether formation exhibits a second order kinetics. SYMBOLS

Ci- concentration of product i, mol.rn -3

DME - dimethyl ether DMM - dimetoxymethane icomponent (MeOH, DME, O2, H20, CO) M e O H - methanol MF - methylformiate

HCHO,

nl, n2, n3 - kinetic orders Pi - partial pressure of component i rco - rate of CO formation, grnolco Fig. 2 - Formation rate of HCHO, DME and CO.

-2t . S . m ca

-1

494

rDME

rHCHO

rate of DME formation, /.lmOlDME.mc2at.S -1

--

--

rate of formaldehyde formation, ~tmoli4ci~o.inca t-2.s-1

rij - rate of formation of product i at temperature j X - methanol conversion Yi- yield of component i x -- contact time,

2t . m O l M~eOH -1 inca .S

REFERENCES

1. A.B. Stiles and T.A.Koch, "Oxidation Catalysts", in Catalyst Manufacture, Editor Marcel Dekker, 2 na Edition, New York Chap.20 (1995) 197 2. H. Adkins and W. R. Peterson, J. Am. Chem. Soc., 53 (1931) 1512. 3. J.M. Tatibou~t, Appl. Cat. A: General, 148 (1997) 213 4. W.-H. Cheng, J. Catal., 158 (1996) 477. 5. J. Edwards, J. Nicolaidis, M. B. Cutlip and C. O. Bennett, J. Catal., 50 (1977) 24. 6. C.J. Machiels and A. W. Sleight, J. Catal., 76 (1982).238 7. G. Busca, Cat. Today, 27 (1996) 457. 8. A . P . V . Soares, M. Farinha Portela, A. Kiennemann, L. Hilaire and J. M. M. Millet, accepted for publication in Appl. Cat. A (2000). 9. P. Jiru, B.Wichterlova and J.Tichy, in Proceedings of 3 rd International Congress on Catalysis, 1 (1965) 199. 10. P. Jiru, J.Tichy and B.Wichterlova, Collection Czechoslov Chem. Commum., 31 (1966) 674. 11. M. Dente, R. Poppi and J. Pasquon, La Chimia e L "Industria, 46(11) (1964) 1326. 12. N. Pemicone, F. Lazzerin and G. Lanzavecchia, J. Cat, 10 (1968) 83. 13. N.P. Evmenenko and Ya. B. Gorokhovatskii, Kinet. Cat. (Engl. Transl..), 10(.6) (1969) 1071. 14. V. N. Bibin and B.I.Popov, Kinet. Cat. (Engl. Transl.), 10(6) (1969) 1091. 15. J. H. Wilson, PhD Thesis, University of Wisconsin, Madison (1986).


495

MECHANISTIC INVESTIGATIONS OF OXIDATION OF PURINE AND P Y R I M I D I N E BASE C O M P O N E N T S O F N U C L E I C ACIDS BY B R O M A M I N E - B IN A Q U E O U S A L K A L I N E M E D I U M : A K I N E T I C APPROACH Nirmala Vaz* and Puttaswamy Department of Chemistry, Central College, Bangalore University, Bangalore- 560 001, India. ABSTRACT Mechanism of oxidation of purine bases (adenine and guanine) and pyrimidine bases (uracil, thymine and cytosine) in presence of NaOH by bromamine-B(BAB) has been investigated. The reactions follow identical kinetics for all the bases, being first order dependence on [BAB]o and fractional order each in [substrate]o and [NaOH]. Addition of the reaction product retards the rate and the dielectric effect is positive. Variation of ionic strength and addition of halide ions had no effect on the rate. Proton inventory studies were made in H20-D20 mixtures for adenine and cytosine. Oxidation products were identified and activation parameters were evaluated. An isokinetic relationship is observed with 13= 336 K indicated that enthalpy factors control the rate. The rate of oxidation of purines is in the order: guanine > adenine while in case of pyrimidines the order is thymine > uracil > cytosine. A suitable mechanism is proposed and discussed. 1. INTRODUCTION Considerable attention has centred around the chemistry of N-metallo-Narylhalosulfonamides, generally known as organic haloamines, because of their versatility in behaving as mild oxidants, halogenating agents and N-anions, which act both as bases and nucleophiles. The prominent members of this class of compounds are chloramine-T (pCH3C6H4SOzNC1Na.3H20 or CAT), chloramine-B (C6HsSOzNC1Nal.5H20 or CAB) and the corresponding bromine analogues, bromamine-T (BAT) and bromamine-B (BAB). Since these oxidants react with a wide variety of functional groups, they are used as reagents for analytical and kinetic investigations[I-3]. Reports on the bromine analogues particularly on BAB are sparse. Bromamine-B is gaining importance as mild oxidant and is found to be a better oxidizing agent than its chloro analogue. Loss of biological activity by radiations and chemicals in living cells is believed to be the result of damage to DNA. Many chemical modifications of nucleic acids have been elucidated[4,5]. The damage by chemicals on nucleic acids is generally attributed to the changes in base components of nucleic acids and is generally due to oxygen radicals. There is hardly any reference to the kinetics of oxidation of purine and pyrimidine bases in the literature. In this background and in continuation with our studies on the kinetics of oxidation of biological compounds by N-haloamines, we report here the kinetics of oxidation of purine bases (adenine and guanine)and pyrimidine bases (uracil, thymine and cytosine) by bromamine-B in aqueous alkaline medium with a

496 view to elucidate the mechanism of oxidation of these bases in solution and also to assess the relative rates of oxidation. 2. EXPERIMENTAL Bromamine-B was prepared by the reported procedure[6]. The purity of BAB was checked iodometrically through its active bromine content and also by UV, IR and NMR spectra. An aqueous solution of BAB was standardized iodometrically and stored in brown bottles to arrest photochemical deterioration. Purine and pyrimidine bases were of Merck make and were used as such. All other chemicals were of analytical grades of purity and double distilled water was used throughout. Solvent isotope studies were made with DEO (99.4%) supplied by BARC, India. The ionic strength of the system was maintained at a constant high value (I = 0.5 M) using a concentrated solution of NaC104. The reactions were studied under pseudo-first order conditions by keeping an excess of the substrate over the oxidant. The reactions were studied at constant temperature (+ 0.1K), and were followed by monitoring the decrease in [BAB] upto 75% reaction. Pseudo-first order rate constants, k/, evaluated were reproducible within + 3-5%. Regression coefficients 'r' were evaluated. Oxidation products of purine bases were found to be their 8-hydroxy derivatives where as in case of pyrimidine bases corresponding 5,6-dihydroxy-5,6-dihydropyrimidines were identified. 3. RESULTS In case of guanine, the reactions were very fast compared to other bases and hence we could not establish the similar experimental conditions. However, similar oxidation behaviour was observed for all the bases irrespective of experimental conditions. With the substrate in excess, at constant [NaOH] and temperature, plots of log[BAB] vs time were linear (r>0.9908) indicating a first order dependence of rate on [BAB]o. The pseudo-first order rate constants (k/) were not affected by a change in [BAB]o (4.0 x 10-4 1.4 x 10-3 M). Values of k / increased with increase in [substrate]o (Table 1) and plots of log k/ vs log[substrate]o were linear (r > 0.9989) with fractional slopes (0.48-0.72), thus indicating a fractional order dependence on [substrate]o. The rate increased with increase in [NaOH] (Table 1) and plots of log k/ vs log[NaOH] were linear (r > 0.9976) with fractional slopes (0.52-0.80). Addition ofbezenesulfonamide (BSA; 5.0 x 10-3- 3.0 x 10-2 M) retarded the rate and the log-log plots were linear (r> 0.9973) with negative fractional slopes (0.39 0.61). Rates were unaffected by the addition of halide ions or varying the ionic strength of the medium, but solvent composition studies using methanol showed that plots of logk/ vs 1/D were linear (r > 0.9968) with positive slopes. The reaction was studied at different temperatures (283-298 K), and from the Arrhenius plots (r > 0.9903), values of the activation parameters for the overall reaction calculated for adenine, guanine, thymine, uracil and cytosine were found to be Ea = 97.1, 31.6, 50.0, 60.1,77.4 kJ moll; AH# = 94.7, 29.2, 47.6, 57.7, 75.0 kJ mol l ; AG#= 90.8, 82.9, 84.6, 84.7, 87.4 kJ mo1-1 and AS#= 13.7, -187.9, -127.6, -93.0, -42.5 JK -1 mo1-1 respectively. Solvent isotope studies in D20 showed that the rates were further increased in D20 medium and the solvent isotope effect k/(H20) / k/(D20) = 0.51 and 0.56 for adenine and cytosine, respectively. Proton inventory

497 studies were made using different atom fractions of deuterium n (0.00, 0.25, 0.50, 0.74 and 0.98) for both the bases. Addition of reaction mixture to acrylamide did not initiate polymerization showing the absence of free radicals. Table 1 Effect of varying [substrate] and [NaOH] on the rate of reaction 10 3 [substrate] a

102[NaOH] a

(103[substrate]) b

(102[NaOH]) b

M

10 3 k / Adenine

s -1

Guanine

Thymine

Uracil

Cytosine

M

2.0(0.5)

1.6(0.8)

0.13

3.95

2.05

1.50

0.58

4.0(1.0)

1.6(0.8)

0.20

6.35

2.92

2.42

0.97

8.0(2.0)

1.6(0.8)

0.34

9.20

3.98

3.65

1.53

14.0(4.0)

1.6(0.8)

0.49

14.9

5.15

4.90

2.40

20.0(8.0)

1.6(0.8)

0.64

23.8

6.40

6.25

3.10

4.0(1.0)

1.6(0.8)

0.20

6.35

2.92

2.42

0.97

4.0(1.0)

2.5(1.2)

0.26

8.00

3.70

3.55

1.35

4.0(1.0)

4.0(2.0)

0.33

10.7

4.85

4.90

1.80

4.0(1.0)

6.0(3.0)

0.41

12.9

6.00

6.95

2.25

4.0(1.0)

8.0(4.0)

0.48

15.0

7.05

9.50

2.70

values in parenthesis refer to guanine. a [BAB]o = 8.0 X 10-3M ; I = 0.5 M ; T = 288 K b

[BAB]o = 6.0 X 10 -4 M ; I = 0.5 M ; T= 283 K

4. DISCUSSION

B r o m a m i n e - B, like chloramine- T and chloramine- B, behaves as a strong electrolyte in aqueous solutions forming different species [7-9] : PhSO2NBrNa

~

PhSO2NBr- + H + --. 2PhSO2NHBr

-..

PhSOzNHBr + H20 ---

PhSO2NBr" + Na + "- PhSO2NHBr ~-

PhSO2NBr2 + PhSO2NH2 ~

PhSOzNH2 + HOBr

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

498 PhSO2NBr2 + H20 --HOBr

--.

(5)

"- PhSO2NHBr + HOBr

(6)

"- H + + O B r

The possible oxidizing species in alkaline BAB solutions are PhSO2NBr , PhSO2NHBr, HOBr and OBr-. Hardy and Johnston's calculations[7] on alkaline bromamine-B solutions indicated that there could be considerable concentration of PhSO2NHBr even in alkaline medium. Further as the concentration of alkali increases, there is also increase in the concentration of hypobromite ion. Hence, in the present investigations, a fractional order dependence on [substrate] and [OH] and the observed retardation of rate by the reaction product (PhSO2NH2) can be explained by the two pathway mechanism shown in schemes 1 and 2: kl ~ PhSO2NH2 + HOBr PhSO2NHBr + H 2 0 k2 HOBr + S ~ X k3 X ~ products

(i) slow & rds (ii) fast (iii)fast

Scheme 1

PhSO2NHBr + OH- -..

k4

-~ PhSO2NH2 + O B r

(i)slow

k-4 OBr ~+ S k6

Y

k5

~ Y

~ products

(i)slowest & rds (iii)fast

Scheme 2

Here X and Y are complex intermediate species and S is substrate. From Scheme 1, rate = kl [PhSO2NHBr] [H20]

(7)

From Scheme 2, assuming steady state conditions for [OBr-] d [OBr-]

k4 [PhSO2NHBr] [OH- ] - k_4 [PhSO2NH2] [OBr-] - k5 [OBr ] [S] = 0

dt Rate Rate

=

k5 [OBr-] [S]

-

k4k5 [PhSO2NHBr] [OH-] [S,]. k .4[PhSO2NH2] + k5 [S]

The combined rate law for the disappearance of BAB is given by

(8)

499 -d [BAB] dt

=

k4 ks [BAB] [OH] [S]

+ kl [BAB] [H20]

(9)

k-4 [PhSOENH2] + k5 [S]

Rate Law (9) is in agreement with experimental results. The proposed mechanism is also supported by the fact that there is enhancement of rate in D20 medium since OD- is a stronger base than OH- ion[10]. Proton inventory studies in H20-D20 mixtures could throw some light on the nature of the transition state[11,12]. The dependence of rate constant k/n on n in a solvent mixture of H20-D20 is given by eq.(10) k/n

TS x (1-n + ng}i)

k/o

RS ~(1-n + n~j)

(10)

Here (~i and ~j are isotopic fractionation factors for isotopically exchangeable hydrogen sites in the transition state (TS) and reactant state (RS) respectively. Plots ofk/n vs n are curves in the present case, which clearly shows that the process involves a single proton of H-D exchange in the reaction sequence from the hydroxide ion. This proton exchange can be shown by a comparison with standard proton inventory plots reported in the literature[ 13]. Hence, the participation of hydroxide ion in the formation of trasition state is inferred. The rate increases with decrease in dielectric constant of the medium indicating a charge dispersal in the transition state[14], which is less polar than the reactants. The change in ionic strength of the medium does not alter the reaction rate suggestive of the involvement of non-ionic species in the rds. Addition of halide ions has no effect on the rate indicating that no interhalogen or free bromine is formed. All these observations confirm the proposed mechanism. The proposed mechanism is also consistent with the moderate values of energy of activation and the thermodynamic parameters. The activation energy is highest for the slowest reaction and vice-versa, indicating that it is enthalpy controlled. This is verified by calculating the isokinetic temperature (13) from the slope of a linear plot of AH# vs AS# (r = 0.9992) for each base. The 13 value of 330 K, which is higher than the experimental range used in the study, this implies that the substrate oxidation is enthalpy controlled. A further confirmation of the existence of an isokinetic relationship was inferred from the Exner criterion[ 15] where a plot o f 1ogk/(283K) vs 1ogk/(293K) showed a linearity (r = 0.9988). From the Exner slope 13 was found to be 342 K. The near constancy of AG# values indicates that all the five bases react with BAB via the same mechanism. The rate of oxidation of purines is guanine > adenine while in the case of pyrimidines the order is thymine > uracil > cytosine. Similar observations were made by Freeman et al.[16] concerning the oxidation of these bases by potassium permanganate. ACKNOWLEDGEMENTS One of the authors (NV) thanks the University Grant Commission, India for the award of Research Fellowship under the Faculty Improvement Programme and also the Principal and Management of Jyothi Nivas College, Bangalore for the encouragement.

500 REFERENCES

1. D.H. Bremner, Synthetic Reagents, 6 (1986) 9. 2. K.K. Banerji, B. Jayaram and D.S. Mahadevappa, J. Sci. Ind. Res., 46 (1987) 65. 3. Puttaswamy, T.M. Anuradha, R. Ramachandrappa and N.M.M. Gowda. Int. J. Chem. Kinet., 32 (2000) 221. 4. A.M. Burkhoffand T.D. Tullius, Nature (London), 331 (1988) 455. 5. C.G. Riordan and P.wei, J. Am. Chem. Soc., 116 (1994) 2189. 6. M.S. Ahmed and D.S. Mahadevappa, Talanta, 27 (1980) 669. 7. F.F. Hardy and J. P. Johnston, J. Chem. Soc. Perkin Trans II (1973) 742. 8. E. Bishop and V.J. Jennings, Talanta, 1 (1958) 197. 9. J.C. Morris, J. A. Salazar and M.A. Wineman, J. Am. Chem. Soc., 70(1948) 2036. 10. C. J. Collins and N. S. Bowman, Isotope Effects in Chemical Reaction, Van Nostrand, Reinhold, New York (1970) p. 267. 11. W. J. Albery and M. H. Davies, J. Chem. Soc. Faraday Trans., 68 (1972) 167. 12. G. Gopalakrishnan and J. L. Hogg, J. Org. Chem., 50 (1985) 1206. 13. N. S. Isaacs, Physical Organic Chemistry, Wiley, New York (1987) p. 275. 14. E. S. Amis, Solvent Effects on Reaction Rates and Mechanisms, Academic Press, New York (1966). 15. O. Exner, Collect Czech. Chem. Commun., 29 (1964) 1094. 16. F. Freeman, C. O. Fuselier, C. R. Amasted, C. E. Dalton, P. A. Davidson,. E. M. Karchesfski, D. E. Krochman, M. N. Johnson and N. K. Jones, J. Am. Chem. Soc., 103 (1981) 1154.


501

ACTIVITY-ACIDITY RELATIONSHIPS IN SOLID ACID CATALYSIS- A QUANTUM CHEMICAL STUDY X. Wang a' b, M.A.N.D.A. Lemos a, F. Lemos a, and F. Ram6a Ribeiro a aDepartamento de Engenharia Quimica, Instituto Superior Trcnico, Av. Rovisco Pais, 1049-001 Lisboa, Portugal. bInstitute of Industrial Catalyst, Dalian University of Technology, 158 Zhongshan Road, 116012 Dalian, China. In the present work the influence of the acidity of silanol-type sites on the transformation of propane and n-hexene-3 is inspected by resorting to ab initio and semi-empirical methods. The present work corresponds to the preliminary results of a systematic investigation, both experimentally and resorting to quantum mechanical methods, on the relations between the acidity of a site and the activity that site presents in relation to the transformation of hydrocarbons. 1. INTRODUCTION Catalysis by solid acids is of paramount importance in industrial chemistry, namely due to its application in catalytic cracking, one of the most important processes in the world. However, despite its enormous importance, only recently have practical and quantitative relationship between the acidity of the catalyst and its catalytic activity began to appear, unlike homogeneous acid catalysis, which has made use of the Brrnsted relations for many years. The difficulties to be overcome are of various nature but it was found by some of the authors that Brrnsted-like relationships also apply to solid acid catalysts [ 1]. The clarification of this aspect may have a very significant beating in catalyst developing, since it is envisaged that it will be possible to predict the behaviour of a particular catalyst by the characterisation of its acid site distribution, which can be obtained using as a practical acidity scale the activation energy distribution obtained by temperature-programmed desorption of bases. The objective of the study presented in this paper is to inspect the nature of the relation between the acidity and the activity of a given site towards the transformation of hydrocarbons over zeolites and to compare the relation derived from first-principles modelling of the catalytic mechanism with the type of correlations that are experimentally obtained and which can be viewed as an application of the BeU-Evans-Polanyi principle. These studies are rather uncommon, although, for instances, Van Santen et al [2], presented results of a computational study on the effect of differences in the strength of the acid site on the activation energy of the reaction steps for hydrocarbon conversion over a zeolite model cluster. In order to change the acid strength of the cluster, which was used in the calculations to rn~odel the acid site, the terminal Si-H bonds were constrained to various lengths. Their

502 results, however, indicated that there should be only a light effect of acid strength variation on the activation energy for D/H exchange in methane and for ethylene chemisorption, which have covalent transition states. 2. COMPUTATIONAL METHODS All computations were carried out using either PC Spartan Plus or PC Spartan Pro (from Wavefunction inc.) and both DFT and a semi-empirical method, PM3, were used. All DFT calculation were done using the non-local BP86 functional with DN** numerical polarization basis set. Since calculations using PM3 are much faster than by DFT, and since this semiempirical method supplies good estimates of the equilibrium geometries, all systems were first optimised using PM3 and then the final values for the energies were obtained by DFT. Only relatively simple models for the acid sites were used, mainly based on a single silanol or on sites with only two T atoms (one silicon and one aluminium) and the acidity of the sites was changed by replacing some of the OH groups by electroactracting halogen elements. Ammonia was used as an "acidity probe", since this molecule is the one that has been used for the establishment of the practical acidity scale in the experimental correlation [3]. Thus, the interaction of ammonia with the various acid sites has been inspected and the activation energy for its desorption has been computed. The deprotonation energy of each site was also computed so that a more conventional acidity measurement can be used for comparison For the hydrocarbon transformation two reactions were selected, a relatively difficult one, the transformation of propane, and a relatively easy one, the cracking of n-hexene. The interaction of these molecules and the expectable intermediates in their transformation has been computed. Transition states for the interconversion between the species were obtained, using PC Spartan's facilities for this purpose. The vibration frequencies for all transition states were computed so as to ascertain that only one of the vibrations corresponded to an imaginary frequency (along the line of the reaction pathway) and that this vibration was indeed along the required pathway. Some of the possible pathways for the transformation of these hydrocarbons have also been inspected and the influence of the acid strength of the site on the activation energy of the various steps involved in the transformations was analysed. 3. RESULTS AND DISCUSSION 3.1 Interaction with ammonia

Since, in the experimental studies, the temperature programmed desorption of ammonia was used to obtained the acid site distribution for the catalysts, the acidity of a site being measured by the activation energy for the desorption of ammonia from that site, the first part of this study was to inspect the interaction of ammonia with the various sites and to compare the activation energies that were obtained with the deprotonation energy for that same site. From PM3 calculations, it was observed that the adsorption of ammonia has virtually no activation barrier, so that the activation energy for desorption corresponds to the heat of adsorption of ammonia. This can be seen in figure 1, were the energy of the system is plotted as a function of the distance between the acidic proton and the nitrogen atom.

503

-850 -855 -860

e

/ -865

~

j

~; -870 r,r

-875 0

i

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i

i

2

4

6

8

......................... i

10

12

dn.N (A)

Figure 1. Energy (computed by PM3) of the C13SiOH---NH3 system as a function of the H---N distance, showing that the adsorption of ammonia has neglectable activation energy. Comparing the activation energies for the desorption of ammonia, obtained for each site, with the values of the deprotonat!on energy for the same site, we can see that there is a linear relationship (figure 2), clearly indicating that the heat of adsorption of ammonia on the acid site (or the activation energy for the desorption of ammonia, to which it is numerically equal in most cases) constitutes, indeed, a good acidity scale, which can be easily obtained experimentally.

50 45 40

~' 35

3o 25 0

......

1460

I

I

I

1480

1500

1520

1540

EDeprotonation(kJ/mol) Figure 2. Linear relation obtained between the heat of adsorption of ammonia and the deprotonation energy of the cluster.

504

3.2 Propane transformation The transformation of propane over the acidic sites can occur by various routes, namely by breaking of a C-C bond, resulting in the production of methane and ethylene, or by breaking of two C-H bonds, resulting in propene as a dehydrogenation product. The detailed scheme for the transformation of this hydrocarbon is out of the scope of this paper and will be the subject of a separate publication. It was observed that the energy of the various species bonded to the acidic OH is strongly dependent on the acidity of the site were we consider the transformation to be taking place and that the activation energies vary in a linear way with the acidity of the site. As examples we can look at the activation energies for the C-C bond breaking and for the dehydrogenation step, which are depicted in figure 3; as the acidity of the site increases, the activation energy for these steps decrease, as expected for acid-catalysed reactions.

500-

--,.....

450400350 300 250

"~ 200-

.........

20

I

I

I

I

I

25

30

35

40

45

50

-A HAmmonia(kJ/mol)

Figure 3. Relation between the acidity of the site and the activation energy for C-C bond cracking (~), and for dehydrogenation (~). All energies were computed by DFT (see text). In figure 4 we present, as an example, the structure of the transition states, obtained by PM3 transition-state search calculations, for the breaking of the C-C (figure 4a) and for the dehydrogenation (figure 4b). As it can be seen, in the first case the acidic proton (HI) inserts into de C1-C2 bond, and C1 leaves the propane molecule in the form of methane, leaving an ethoxide species directly bonded to the O atom in OH group. This ethoxide species will, upon desorption, produce an ethylene molecule in gas phase. In the second case, a bond is formed between the acidic proton (HI) and one of the hydrogen atoms (H2) in the propane molecule. The net result will be the formation of an hydrogen molecule and the production of an adsorbed propoxide species, which will eventually desorb as an propylene molecule.

505

Figure 4. Transition-state structures for the breaking of the C-C bond (a) and for the dehydrogenation step (b) in the transformation of propane over acidic sites.

3.3 n-Hexene transformation

The interaction of n-hexene-3 with the various acid sites has been inspected as well as the pathways to transform this molecule into an adsorbed ethylene moiety and a butene-1 molecule.

Figure 5. Two possible pathways for the transformation ofhexene-3 into ethylene and butene-1 and the corresponding transition-state structures for the relevant steps. A stable adsorbed state was obtained for n-hexene where the integrity of the molecule is not essentially changed. This adsorbed hexene molecule can then react by two possible pathways: one corresponding to the formation of an adsorbed carbocation, with an hexoxide

506 structure, which can then crack by 13-scission to yield an adsorbed etoxide species and a butene-1 molecule and another one which corresponds to the direct breaking of the C-C bond straight from the adsorbed hexene molecule. In figure 5 we present two possible pathways and the corresponding structures for transition-states.

500 450 400 350

300 250 200 20

i

i

i

i

25

30

35

40

.... i ...................

45

50

-A HAmmonia(kJ/mol) Figure 6. Relation between the activation energy for the cracking of the C-C bond in n-hexene and the acidity of the site measured by the heat of adsorption of ammonia. In the case of these reactions, however, a different effect was observed, as it can be seen in figure 6. Since this reaction is much easier than the cracking of propane, the acidity of the sites seems not to have any influence on the activation energy of the transformation, for sites with low acidity. Only when the acidity of the site exceeds a certain value does the activation energy start to decrease with the increase of acidity. 4. CONCLUSIONS As one would expect, there seems to be a relation between the acidity of a site, as measured by the heat of adsorption of ammonia, for instances, and the activation energy for the various steps in the transformation of hydrocarbons over these sites. The use of computational models, which are increasingly available, can help in the design of helpful and practical relationships that can be used to estimate catalytic behaviour based on acid-site distribution measurements. REFERENCES 1. C. Costa, J.M. Lopes, F. Lemos, F. Ram6a Ribeiro, J. Mol. Catal. A: Chemical, 144 (1999) 233. 2. A.M. Rigby, G. J. Kramer and R. A. van Santen, J. Catal., 170 (1997) 1. 3. C. Costa, J.M. Lopes, F. Lemos, F. Ram6a Ribeiro, J. Mol. Catal. A: Chemical, 144 (1999)221.


507

CATALYTIC REACTIONS MECHANISMS WITH SOME NONLINEAR CONSERVATION LAWS E. S. Patmar and N.I.Koltsov Chuvash State University, Cheboksary, Russia, 428015 Cheboksary, Moskovskii pr.15, Fax:(+78352)428090, E-mail:[email protected]

The common approach to downturn a dimension of system of the equations of any nature is the method of unknowns exclusion. With a reference to chemical systems, the simple example of its use is the concept of so-called "dependent" and "independent" substances, relations between them beeing defined by the laws of conservation of atoms quantity. For catalytic reactions always there is one such law - law of conservation of amount of atoms of the catalyst in each of elementary stages. This law is linear and allows easily to exclude one of unknowns, that is to express quantity of any one of intermediate substances through others. Generally in chemical systems can exist both linear and nonlinear conservation laws (CL). For the first time the linear CL Z PiXi -- c o n s t i

(1)

( x i are substance concentrations, Pi are some constants) were investigated in paper [ 1], then in [2]. In [3] the computer program for searching linear CL was developed. For some classes of reaction stage schemes the method of nonlinear CL (NCL) determination is described in paper [4]. They are given as elementary functions of exponential form [5] exp(f0)l-I fPi

= const ,

(2)

i

where f is a vector of rational functions of substance concentrations p is a vector of constant parameters which depend on the order and form of the right partsof the differential equation system describing the dynamic behaviour of an investigated reaction. In this paper the new results are given. They widen the classes of catalytic reactions mechanisms having some NCL under isothermal conditions of a gradientless differential reactor. Quasistationarity via basic substances is supposed to be. In table 1 the mechanisms in which Xi are intermediate substances on the catalyst surface are given.

508

1

Schemes

Unlinear CL

XI :=> X 2

ln(x2 + q~ ) + q2" arctg(q3t) + q4 ln(q5 + t) + q6 ln(q7 - t) = const

x,+&=&+x4

t = (x 2 + q~)/(x 3 + qs)

Xl +X3 => 2X 1

ln(x2 + Ul ) + U2" arctg(u3g) + q4 In(us + g) + q6 ln(u7

g) = const

g = ( x 2 + u 1) / ( x 4 + u 8)

X 4 -[-X 1 ~

2

2X~

X1:::>X~, Xl +X~ ~ X 3 +X4 "

x, +iv.=2x; x. +x, = ~

ln(x2 + ql) + q2" arctg(q3t) + q4 ln(q5 + t) + q6 ln(q7 t) = const

ln(x2 + ul) + u2" arctg(u3g) + q4 ln(u5 + g) + q6 ln(u7 - g) = COl g = ( x 2 + 121)/ (X4 -!-~ ) , t =(X 2 +ql )/(X3 +q8), h=(xa +V1)/(Xs+Vs)

x,+x= =2& X 1 + X 5 ~ 2X~ 3

XI :=> X 2

ln(x2 + q~) + q2 " arctg(q3t) + q4 ln(q5 + t) + q6 ln(q7 t = (x 2 + q~)/(x 3 + qs)

& +x~ ~ y~ +X. X~ + X3 =z. 2 X ~

t) = cons1

ln(x2 +ul)+u2 " a r c t g ( u 3 g ) + q 4 ln(us + g ) + q 6 ln(u7 - g ) = c o n s g = (x 2 + u 1)/(x4 + u s)

X . + X , =a, 2 X s

X 5 :z~ X 1 * q i , Ui are functions of rates constants ( i = l ..... 8).

The schemes which are given in table 1 can be applied to concrete reactions. For a catalytic reaction of hydrogen oxidation scheme 1 of table 1 is the following 1.K+O2 ~ K O z 2Jf+K(~ ~ K O + K H 2 0 3 J f + K t l z O ~ 2 K + H 2 0 4 J f + K O + H

2 ~2K+H20

The obtained NCL give the possibility to lower differential equations system dimentions. In this way they help to simplify qualitative and numerical analysis of these systems and that of kinetic regularities of corresponding catalytic reactions. REFERENCIES 1. Korzukhin M. D. Zhurn. Phys. Chem., 46, 1845, 1972. 2. Alekseev B. V., Koltsov N.I. and Fedotov V.Kh., Ibid., 66, 3219, 1992. 3. Koltsov N.I., Alekseev B. V. and Kozhevnikov I. V. Mathematical Methods in Science. Modelling Critical Phenomena in Catalytic Reactions. Cheboksary, Publ. House of Chuvash State Univ., 1998. 4. Alekseev B. V., Koltsov N.I. and Fedotov V.Kh., Zhurn. Phys. Chem., 62, 3069, 1988. 5. Prelle M.J., Proceedings of the 1981 ACM Symposium of Symbolic and Algebraic Computation, Snowbird Utah, 30, 1981.

Studies in Surface Science and Catalysis 133 G.F. Froment and K.C. Waugh (Editors) (_c~2001 Elsevier Science B.V. All rights reserved.

509

Modelling Diffusion, Cracking Reactions and Deactivation in FCC Catalysts F. L6pez-Isunza a, N. Moreno-Montiel a, R. Quintana-Sol6rzano b, J. C. Moreno-Mayorga b and F, Hemfindez-Bekrfin b aDepartamento de Ingenieria de Procesos e Hidrfiulica Universidad Aut6noma Metropolitana-Iztapalapa, M6xico 09340, D.F., MEXICO b Area de Investigaci6n en FCC, Subdirecci6n de Transformaci6n Industrial Instituto Mexicano del Petr61eo, M6xico 07730, D.F., MEXICO

A series of experiments varying temperature, micro-sphere size and time on stream have been performed in a fixed fluidised bed microactivity reactor to study the role of intraparticle diffusion in commercial fluid catalytic cracking (FCC) catalysts, particularly on gasoline yield and catalyst deactivation by coke deposition, for the cracking of a vacuum gas oil. Additionally, a mechanistic model that describes interface and intraparticle mass transfer interactions with the cracking reactions, has been used to study the combined influence of pore size and intraparticle mass diffusion on the deactivation of FCC catalysts and the gasoline yield.

1. INTRODUCTION First studies on the influence of intraparticle diffusional mass transfer on catalytic reactions, and about deactivation of cracking catalysts by coke deposition, started with Thiele [1] and Voorhies [2], respectively. To-date, Thiele's analysis remains valid, however the approach followed by Voorhies in which coke formation is expressed as a function of time on stream (tos), although still used in many studies [3], is not adequate. It has been stated that deactivation, due to the coverage of active sites by coke deposition and to pore blockage by coke growth [3, 4], should be directly related to coke itself and not to tos [5]. In this way, coke formation is linked to the operating conditions, the nature of the feedstock and the type of catalyst. Modelling diffusion, reaction and deactivation in FCC catalysts becomes important as new trends in vacuum gas oil processing point to short catalyst/oil contact times (less than 1 sec) in industrial riser reactors. Under short contact times catalysts particles may experience diffusion limitations [6], and thus catalyst pore dimensions become very important in determining conversion and selectivity. FCC catalysts are made of several different components to produce the necessary activity and mechanical properties. As the accessibility of large molecules to the active sites within the catalyst might be constrained,

510 the matrix becomes a key component in FCC catalyst.activity and selectivity with regular micro-porous structure given by the zeolite component, but unknown overall pore size distribution. Typically, FCC catalyst particles are 70 ~tm average diameter while the matrix average pore size very often range from 30 to 90 A. Accordingly, the purpose of this work is to study, both experimentally and theoretically, the role of intraparticle diffusion on the gasoline yield and catalyst deactivation by coke deposition during the cracking of a vacuum gas oil using an isothermal ideally mixed continuous fluidised bed micro-reactor, varying the reaction temperature, the particle size and tos. The study employs a mechanistic model [6] that describes the mass balances in an isothermal ideally-mixed fluidised bed micro-reactor, where a set of surface reactions occur inside a cylindrical pore in a single pellet cracking catalyst, in which all reactions and deactivation by coke deposition are described in terms of Langmuir-Hinshelwood expressions. The kinetics considers the cracking of a vacuum gas oil in terms of 5 lumps (gas oil, ligth cycle oil, gasoline, ligth olefins, and ligth gases), coke, adsorbed pseudo-species, and the unoccupied fraction of active sites. Mass diffusion of all lumps is assumed to take place in the Knudsen regime.

2. EXPERIMENTAL

Experiments were conducted on an ACE-R T M fixed fluidised bed microreactor, using a typical vacuum gas-oil. The study considers three sets of catalyst micro-spheres, after being steamed (hydrotermically deactivated) for 8 h at 788 ~ After steaming catalyst particles were dryed at 150 ~ for 2h, then sieved to particle sizes in the range: F3=44-74, F2=74-150, and F1=150-250 Bm. Texture analysis, as shown in table 1, was performed by N2 adsorption (ASAP-2000) following ASTM-D-3663 and ASTM-4222 (t-plot) methods, which allowed to estimate the total BET surface area and the micropore (zeolite) area, respectively. Table 1. Texture properties of catalyst fractions Properties Ref. Case F1 Range in diameter, lam 44-250 44-74

F2 74-150

F3 150-250

Surface area, mZ/g Zeolita area, mZ/g Area ratio Zeolite/Matrix Pore Volume, cm3/g

104.7 55.2 1.12 0.2173

104.1 54.8 1.11 0.2315

105.8 50.9 0.94 0.2343

107.6 58.7 1.20 0.2180

In all experiments, the amount of catalyst employed was 9 g, and vacuum gas oil feed rate was 72 g/h, for times on stream:60, 75, 100 and 150 s. The reactions were carried out at temperatures of: 480, 500 and 520 ~ Gas products (dry gas and LP gas) were analyzed by cromatography. The liquid product (gasoline, light and heavy cycle oils) was analyzed after simulated distillation (ASTM-D-2887). The amount of coke on the catalyst was determined through the IR analysis of CO2 produced by the coke burning in air at 695 ~ The mass balance calculated from product recovery was in the range: 97.2 to 100.9 %.

511 3. RESULTS AND DISCUSSION To assess the role of intraparticle diffusion on the cracking reactions experiments were performed at three temperatures and three sets of particle sizes. The model employed [6] to analyze the behaviour of the reaction-diffusion phenomena under these conditions, required the estimation of all transport and kinetic parameters. However, at this stage of our study we only estimated the order of magnitude of these parameters for one set of conditions. Otherwise, a fairly large amount of experimental data would be required. All predictions were performed using an average particle size for F2 and T-520 ~ and as it will be shown afterwards, predictions obtained were very good. However no model fitting was attempted beyond these conditions. On the other hand, although surface areas of all fractions are practically the same, the particle size distribution within each batch was unknown. In turn, since this determines the magnitude of effective diffusivities for all lumps, an average particle size was assumed for each batch for model simulations. Following Thiele [1 ], if intraparticle mass diffusion is not intruding in the cracking reactions, then the yield of gasoline should vary only with temperature and not with particle size. However, experiments shown in Figure 1 indicate that at 480 ~ these differences are smaller than those at 520 ~ suggesting that the particle size has some effect on the gasoline yield. To verify this, model simulations were performed using effective diffusivity values one order of magnitud larger (-10 .7 m2/s) than those employed in the predictions shown in Figure 1. It can be observed that if diffusion limitations decrease, the gasoline yield is entirely determined by the reaction temperature as shown in Figure 2.

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Temperature (C)

Fig. 1. Comparison between experimental and model predictions at different temperatures and particle sizes, for effective diffusivities Deft -10 -8m2/s A

512

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The amount of coke deposited on the catalyst as a function of temperature and time of stream (tos) is shown in Figure 3 for the two-extreme temperature conditions. It can be observed that the strongest effect on coke deposition, for tos < 100 s, is due to temperature rather to particle size. As tos increases deactivation becomes the controlling mechanism. Product distribution in terms of the lumps used in the model at T=520 ~ for an average F2 particle, is compared with the experimental data in Figure 4; whereas predictions obtained for different values of tos are given in Figure 5. It can be observed from these figures that predictions for gasoline yield are better than those for coke on catalyst. This had to be expected, as the complexity of the mechanism for coke deposition, due to both the coverage of active sites by coke deposition and to pore blockage by coke growth, are not contained in the model [6], which describes this process as due to the adsorption of most lumps (gas oil, light cycle oil, gasoline and light olefins) during cracking. 4. CONCLUSIONS A mechanistic model has been developed to study deactivation of FCC catalysts, based on the description of interactions between mass transfer and a set of heterogeneous

513

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,,,

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60

80

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Time on Stream (sec) Fig. 3. Amount o f coke deposited on catalyst as a function o f temperature and time on stream

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Product distribution predicted by the Model (wt %)

Fig. 4. Model predictions versus experimental values obtained at T=520~ for F2 particle size distribution.

514 50

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i,,, 100

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i,,, 120

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(sec)

Fig. 5. Comparison between model predictions and experimental values for F2 particles at T=520~

reactions occurring inside a cylindrical pore in cracking catalyst particles, in which coke deposition is given in terms of a Langmuir-Hinshelwood expression. This approach could be used in future work to transfer information between microactivity and riser reactors, and to evaluate the performance of cracking catalysts.

Acknowledgements" The authors gratefully acknowledge Instituto Mexicano del Petr61eo (projects DO 1024 and D00291) and Consejo Nacional de Ciencia y Tecnologia (M6xico) for supporting this research. 5. REFERENCES 1. E.W. Thiele, Ind. Eng. Chem. 31 (1939) 916. 2. A. Voorhies, Jr., Ind. Eng. Chem. 37 (1945) 318. 3. J.B. Butt and E.E. Petersen, Activation, Deactivation and Poisoning of Catalysts, Academic Press, San Diego, 1988. 4. J.W. Beeckman and G.F. Froment, Chem. Eng. Sci. 35 (1980) 805. 5. G.F. Froment, in G. Bond, P. Wells and F.C. Yompkins (eds.) Proc. Sixth Int. Cong. Catal. 1976. The Chem. Soc., London, 1977, p. 10. 6. F. L6pez-Isunza, Fifth Int. Syrup. Adv. Fluid Catal. Cracking, Preprints 218 th Nat. Meet. ACS, 1999, p. 540-544.


515

Oxidehyrogenation of ethylbenzene to styrene on P-O-Ni-Mn/Alumina catalysts D. Ardissone a, A. Bachiller a, M. Ponzi a, J. Orejasb alNTEQUI- CONICET- UNSL. Facultad de Ingenieria y Ciencias Econ6mico-Sociales. Avda. 25 de Mayo 384. (5730) Villa Mercedes (San Luis). Argentina. bUniversidad Nacional de Rio Cuarto, Facultad de Ingenieria, Ruta Nac. 36 Km 601. (5800) Rio Cuarto. Argentina The oxidehydrogenation of ethylbenzene to styrene on a P-O-Ni-Mn/A1203 catalyst was studied. Kinetic analysis of reaction data was performed by complementing parameter estimation and model discrimination techniques with experimental design strategies. Both ihe principal oxidehydrogenation reaction and the side reaction of complete oxidation of styrene follow kinetics of the Langmuir- Hinshelwood - Hougen- Watson (L-H-H-W) type. The selected model shows acceptable goodness of fit. 1. INTRODUCTION Styrene is of commercial interest due to its extensive application in the manufacturing of plastics as well as in the production of synthetic rubber. As for many petrochemicals, the world demand for styrene is in continuous expansion. An annual increase of 11.4 million tons in production capacity has been estimated for the period 1995-2005, which means that by the year 2005 the world annual production of styrene will reach 29.5 million tons [ 1]. Several technological alternatives have been implemented for the commercial production of styrene. The one with the biggest capacity of installed production consists in making the synthesis of styrene by catalytic dehydrogenation of ethylbenzene in vapor phase. This synthesis route, that requires ethylene and benzene as raw materials, amounted to 53% of the world benzene demand in 1996. The main reaction in this process is strongly endothermic, with the conversion being limited by thermodynamic equilibrium. High temperatures and low pressures favor high conversions. The high energy demand of these adiabatic processes are met by combining superheated vapor (800 - 950 ~ with preheated ethylbenzene [2]. Therefore, this stage of the process is very expensive from the point of view of energy consumption. Nowadays, there is intense activity in the development of new techniques to overcome the above mentioned limitation [3]. Some of the most advanced methodologies include the oxidehydrogenation of ethylbenzene, followed by hydrogen oxidation and the use of membrane reactors. In this work, the kinetics of the oxidehydrogenation reaction from ethylbenzene to styrene on a catalyst of P-O-Ni-Mn supported on alumina is studied. Unlike the traditional process of catalytic dehydrogenation of ethylbenzene, catalytic oxidehydrogenation, does not present thermodynamic limitations. Its reaction:

516 C6HsCzH s +0.502 ~

is

(1)

C6HsC2H 3 + H 2 0

exothermic (AH~

29,7 kcal/mol). As it is shown in Table 1, the performance of the catalyst used in the present work is similar to that of the one employed in the traditional catalytic dehydrogenation. Table 1 Catalyst performance comparison

Technology

Conv. (%)

Dehydrogenation Oxidehydrogenation

60 - 65 50

Sel.(%) 90 90

Temp.(~ 600-650 4 5 0 - 550

Catalyst Fe/KJCr/O P-O-Ni-Mn/A1203

In the last two decades, numerous works related to the oxidehydrogenation of ethylbenzene to styrene have been published [4 - 5]. Among the catalysts described in the literature, metallic phosphates have been reported to exhibit high activity and selectivity [6 7], particularly when a transition metal is added. Ponzi et al. (1986) [8] report ethylbenzene conversions of 50 % with styrene selectivities of 90% for the massic P-O-Ni-Mn catalyst. These authors have also reported good catalyst stability for 50 days of continuous operation, no effects on selectivity and only a slight decrease in conversion were found. Ardissone et al. [9] reported the results obtained with an alumina supported P-O-Ni-Mn catalyst. Unlike the results obtained with massic catalysts, better levels of ethylbenzene conversion were obtained. The results obtained from studies of the combustion reaction of styrene and ethylbenzene have showed that the combustion of the former predominates over that of ethylbenzene [ 10]. C6HsC2H 3 +10 02 ---> 8C02 + 4 H 2 0

(2)

2. EXPERIMENTAL The catalyst was prepared by impregnating alumina with metallic P, Ni and Mn salts with the adequate proportion of phosphoric acid, followed by in situ precipitation by the addition of ammonium hydroxide. Conventional equipment for flow under dynamic regime was used. The reagents were fed into the reaction by bubbling the gaseous current (Air + N2) in a bubbler train containing either ethylbenzene or ethylbenzene and styrene, submerged in a thermostatized bath. A 20-mm ID reactor built in Pirex glass, operated in descending flow, was employed. The catalyst was located in the area of the oven where the temperature profile was fiat. The reactor consists of three sections independently controlled by means of PID controllers. The temperature was measured using a thermocouple located axially in the center of the reactor. The product analysis was performed by on-line gas chromatography. The hydrocarbonated products were analyzed by FID using a 3 m column packed with Carbowax 20 M on 20 % Chromosorb W in weight. The gaseous products were analyzed by TCD using columns packed with Porapak Q and 5 A, molecular sieves. In all cases, the major products detected were styrene and carbon dioxide. Other byproducts, such as carbon monoxide, benzoic acid and benzene were only detected at concentrations below 1%.

517 In order to prevent hot spots in the catalytic bed, 3 g of catalyst diluted in Pyrex glass of identical granulometry in a 1:1 ratio were used. No intra- or inter-particle resistance to mass transfer was detected. The independent variables investigated were reactor temperature, molar fraction of ethylbenzene, oxygen and styrene in the feed, and residence time. A three-level factorial experimental design [11] was used to investigate the spectrum of possible experimental conditions. 3. KINETIC ANALYSIS AND PARAMETER ADJUSTMENT A marked influence of the product on the adsorption term was observed in the (L-H-HW)-type models in a previous work [9], where only the oxidehydrogenation reaction was considered. An identical effect was detected in a recent kinetic model of the same reaction on a P-O-Ni catalyst [ 12]. Different mechanisms were proposed and developed for the reactions, considering the following three possibilities: all the catalyst sites are equal; there are two kinds of sites; the reaction mechanism is of the redox type. The parameters were obtained by applying the integral method [13-14] by means of nonlinear regression using a combination of both the direct search method OPTNOV [15] and the Marquardt method [ 16]. The rate equations were integrated by means of a fourth order Runge-Kutta routine. The parameter estimation was performed by taking simultaneously the data obtained at all the temperatures. In this way, frequency factors, activation energies, preexponential factors and adsorption heats were all estimated in a single stage. Discrimination between rival models was performed by taking into account the goodness of the fit, the physicochemical meaning of the estimated parameters [ 17] and their statistical validity. The proposed models are shown on Table 2. The table also shows the reasons why each model was accepted or rejected. The parameters estimates and the confidence limits of the estimates at 95 % probability for de chosen model (model 14 in Table 2), are reported in Table 3. As can be seen, the signs of the obtained activation energies (Ea) have physicochemical significance. The equilibrium constants of adsorption K which appear in the rate expressions can be written as: K=A'-e

an~ Rr withA ~

As__~_~.

(3)

where zU-/~ is the standard enthalpy of adsorption and 115'0 is the standard entropy of adsorption. Because adsorption is exothermic, the adsorption enthalpy has to satisfy the inequality: - zM-/"~ > 0

(4)

and the adsorption entropy has to satisfy As ~ < 0

(5)

then, the preexponential factor (A*) of the equilibrium constants K must be 0 < A*350 ~ The most active metal for these reactions was found to be palladium [6]; consequently, this was the catalyst selected in our experiments. Most studies on hydrodechlorination reactions in presence of an organic solvent have been devoted to hydrotreatment catalysts, being very scarce the studies using palladium [5, 7]. On the other hand, to the best of our knowledge, there are no works published dealing with mixture effects when different organochlorinated compounds are processed together, even though inhibition effects are important in similar processes such as hydrodesulfurization [8]. The main aim of this work is to study the kinetics of the hydrodechlorination of TTCE and CBZ, both alone and in mixtures, in an organic matrix, over a commercial palladium catalyst. A model based on the postulates of LangmuirHinselwood was used to model the mixture effects 2. E X P E R I M E N T A L S E C T I O N The chemicals used in this work, TTCE, CBZ, benzene, toluene and decahydronaphtalene were supplied by Panreac and Merck, with a minimum purity of 99%. The commercial palladium catalyst used was ESCAT -16, 0.5% Pd/A1203, Engelhard (specific surface:103.35 m2/g). Hydrogen was supplied by Air Products with a minimum purity of 99.999% and a-alumina was supplied by Acros. Reactions were carried out in an 11 mm internal diameter fixed-bed reactor containing 0.25 or 0.5 g of catalyst. Solutions of TTCE and CBZ in toluene in the range 0- 1 molfl were used as reactor feed. A hydrogen excess of 10:1 over the stoichiometric required was fed co-currently to the reactor. All the experiments were carried out at 0.5 MPa pressure. Operation temperature (250~ was fixed in order to make sure that all reactants were present in the reactor as gases, and to ensure the stability of the catalyst during the experiments (higher temperatures lead to fast deactivation of the catalyst). Space times of 0-3 min.g/mmol of reactant) were reached changing the flow rate of liquid feed, the concentrations in the feed and the amount of catalyst, keeping constant the H2/organochlorinated compound ratio. In previous experiments it was demonstrated t h a t the solvent (aliphatic or aromatic) did not have relevant influence on the catalyst activity. A more detailed description of the experimental set-up and further studies about the selection of the operation conditions are given in reference [9]. Liquid reaction products were analysed by gas chromatography in a HewlettPackard 5890A apparatus equipped with a HP-1 30-m capillary column and a FID detector. Decahydronaphtalene was used as internal standard. The response factors were determined using standard calibration mixtures.

523 3. R E S U L T S AND D I S C U S S I O N 3.1. R e a c t i o n s t u d i e s The main reaction products were ethane (from TTCE) and benzene (from CBZ) (i.e. total dechlorination), in both cases with selectivities higher than 99 %. This result is very important considering that the final aim of this work is the development of a clean technology for the treatment of wastes. Very little amount of trichloroethylene were found in the hydrodechlorination of TTCE, whereas in the case of CBZ hydrodechlorination neither chlorocyclohexane nor cyclohexane were found. Likewise, the solvent (toluene) did not react in appreciable extension, being observed only small amounts of methyl-cyclohexane. TTCE was found to be more reactive t h a n CBZ. Experiments with TTCE and CBZ mixtures were carried out at 250~ and 0.5 MPa (5 bar), with 10% (w/w) concentration of CBZ and CBZ/TTCE molar ratios of 1:1, 1:0.5, 1:0.25 and 1:0, and with 10% (w/w) concentration of TTCE and TTCE/CBZ molar ratios of 1:1, 1:0.5, 1:0.25 and 1:0. Important mixture effects were observed when the two compounds were reacted together: the reactivity of CBZ is drastically reduced in presence of TTCE (Fig.l), whereas the effect of CBZ in the hydrodechlorination of TTCE is not so marked (Fig.2).

.2

0.9

o

0.8

o

0.7

O

0.6

0.8

"-----.o

ra~

0.6 0

r N

0.4 0.2

0.5 0.4

' 0

'

,

,

,

,

0.2

0.4

0.6

0.8

Inlet CBZ concentration

1

(mol/1)

Fig.1. Influence of CBZ concentration in TTCE conversion (0.8 molfl TTCE). Space time 0.6 (O), 0.8 (~) and 1 (A) min'g/mmol

0

0.2

0.4

0.6

0.8

1

Inlet TTCE concentration(mol/1) Fig.2. Influence of TTCE concentration in CBZ conversion (0.8 molfl CBZ). Space time and symbols: see Fig. 1

This result can be explained assuming that the number of chlorine atoms attached to the organic structure has a stronger influence on the adsorption strength than the type of organic structure (TTCE has four chlorine atoms and CBZ only one, while aromatic structures are considered to hold stronger interaction with the metallic surface). The large importance of the chlorine atoms in the metal-catalysed hydrodechlorination reactions and their substantial role in

524 the adsorption of the molecule was stated by some authors [10, 11]. This hypothesis is also consistent with the observed negligible effect of the solvent in the catalyst performance.

3.2. Kinetic m o d e l l i n g of m i x t u r e s Internal and external mass transfer limitations were estimated not to be important, according to standard correlations [12]. In the same way, a PFR-like behaviour was also demonstrated. When mixture effects on hydrogenolysis or hydrogenation reactions are considered, the most successful models are often the based on the LangmuirHinselwood (LH) mechanisms [8]. The main assumption of this model is that the reaction occurs between two adsorbed species. Considering the high affinity of hydrogen for the palladium surface, the chemisorption of hydrogen is usually considered as dissociative. Concerning to the adsorption of the organochlorinated compound, two possibilities can be considered: chemisorption on the same active sites than hydrogen, or chemisorption of hydrogen and organochlorinated compound on different adsorption sites. According to the literature [8], the first supposition is more accurate for hydrogenolysis reactions (such as hydrodechlorination) whereas the second is considered more accurate for hydrogenation reactions. In the case of CBZ, the only reaction is the hydrogenolysis of a C-C1 bond, whereas in the case of TTCE two reactions occur: hydrogenation of the double-bond and hydrogenolysis of C-C1 bonds. Even some authors state that the mechanism of TTCE hydrodechlorination is a succession of catalytic hydrogenation of C-C double bond and thermal elimination of HC1 [6,13]. Kinetic expressions were derived for both cases. In the case of chemisorption of chlorinated and hydrogen over the analogous sites (LHA), the reaction rate is given by the following equation:

J,4KuKiPu:P, (-r,.) = (1 + ~KMp M + Kip, + Xjpj) z

(1)

When the hydrogen and the chlorinated are adsorbed on different active sites (LHNA), the equation is: 0.5

/,~KHKipH~pi (-~) = (1 + ~/KH PH + Ki p, + K i P i)(1 + ~KH PH )

(2)

In these equations ji is the intrinsic kinetic constant and Ki are the adsorption constant of component i. Considering that the hydrogen is present in large excess and its partial pressure is almost constant in all the experiments (0.43 MPa), equations [1] and [2] can be rearranged:

525

(-r~) =

J" p' (1 + K'~ p~ + K'j pj )"

(3)

The new constants ji' and Ki' are defined in Table 1 for both models, and n is equal to 1 for LHNA model and n=2 for LHA model. Experimental data were fitted to these rate expressions. Considering the high conversion attained in some experiments, the reactor was considered as integral.

Table 1

Summary of rate models with estimated parameters and correlation coefficient

j'i (MPa:l"mm01/min'g)

K'i (MPa -1)

j;= J'K'K~PH~.2=. (l+K~P~)

K,'= K~ (l+K~P~)

TTCE

2.374" 1012

1.896" 10 ~

CBZ

3.873.109

1.615-104

j,K,K~P~ J;= (1+ K~P~)

K I = K,

TTCE

6.527" 1011

2.075" 107

CBZ

2.214"1010

5.351"106

Model LHA

Parameter definition

LHNA Parameter definition

r

0.995

0.964

The kinetic parameters for the two models were estimated by fitting the rate expressions to the experimental data by means of a non-linear least squares minimisation of the error in the prediction of conversions, with a simplex algorithm followed by a Powell minimisation algorithm. The differential equations have been integrated using the EPISODE package. These mathematical tools are implemented i n t h e commercial programme Scientist. The goodness of the fit is quantified using the correlation coefficient (r). Results of the fit and parameter estimation are shown in Table 1. The correlation coefficients reveal that LHA model gives a better fit. The tendency of the values of the constant of the LHA model are in good agreement with the experimental observations (TTCE is more reactive and has the higher inhibition capacity).

526

~9 0.8 0.6 0.4 "~ o

,Jb~"A

,~~

0.2 0

'pTl /0 /

"i 0.8 ' 0"61]/i*~

S ~9 o ~

.

0

.

.

0.2

.

.

0.4

0.4 r~

.

0.6

0.8

1

Conversion, Experimental Fig. 3. Parity plot comparing the experimentally measured conversion of CBZ ( i ) and TTCE (0) with the prediction of LHA model.

o ~ 0

2

4

6

8

10

Space time (g-rain/retoolCBZ) Fig. 4. Experimental (points) and predicted (line) conversion of CBZ at the following TTCE/CBZ molar ratios: 0 (F1), 1 (*), 2 (O) and 4 (A). LHA model.

Experimental and predicted values are compared in Fig. 3 for all the experimental points considered in this work. With more than 70 points for each species, it can be observed that the fitting for the proposed model is fairly good. As an example, Fig. 4 shows that the LHA model predicts CBZ conversion (the compound that is affected in higher extension by the mixture effects).with reasonable accuracy.

REFERENCES 1. E. Goldberg, Sci. Total Environ. 100 (1991) 17. 2. M. Tancrede, R. Wilson, L. Zeise, E. A. C. Crouch, Atmos. Environ. 21 (1987) 2187. 3. T. N. Kalnes, R. B. James, Environ. Prog. 7 (1988) 185. 4. R. C. Dampsey, T. Oppelt, T. Air and Waste 43 (1993) 25. 5. D.I. Kim, D. T. Allen, Ind. Eng. Chem. Res. 36 (1997) 3019. 6. S. Ord6fiez, H. Sastre, F. V. Diez, Appl. Catal. B 20 (1999) 309. 7. R. J. Meyer, D. I. Kim, D. T. Allen, J. H. Jo, Chem. Eng. Sci., 54 (1999) 3627. 8. M.J. Girgis, B.C. Gates, Ind. Eng. Chem. Res., 30 (1991) 2021. 9. S. Ord6fiez, Ph.D. Thesis, University of Oviedo, Oviedo (1999) 10.A.H. Weiss, B. S. Gambhir, R. B. Leon, J. Catal., 22 (1971) 245. l l . A . H . Weiss, K. A. Krieger, J. Catal., 6 (1966) 167. 12.J.A. Moulijn, P.W.N.M. van Leeeuwen, R.A. van Santen (eds.), Catalysis: an integrated approach to homogeneous, heterogeneous and industrial catalysis, Elsevier, Amsterdam 1993. 13.A.R. Pinder, Synthesis, 79 (1980) 425.


527

Kinetic study of the liquid-phase hydrogenation of 1,3-butadiene and nbutenes on a commercial Pd/AI203 catalyst N.O. Ardiaca, S.P. Bressa, J.A. Alves, O.M. Martinez and G.F. Barreto*. Centro de Investigaci6n y Desarrollo en Procesos Cataliticos (C1NDECA). Universidad Nacional de La Plata y Consejo Nacional de Investigaciones Cientifica y Trcnicas. Argentina. Programa de Investigacirn y Desarrollo en Ingenieria de Reactores (PROIRQ). Facultad de Ingenieria, Universidad Nacional de La Plata. Argentina. The simultaneous liquid-phase hydrogenation of 1,3-butadiene and 1-butene on a commercial Pd/A1203 catalyst of the "egg-shell" type was studied with the purpose of kinetic identification. Batchwise experiments were carried out at 40~ with initial composition in the range 1%-2% ofbutadiene, up to 15% of 1-butene and at 0.1-0.3 MPa of H2 partial pressure. Although the tested catalyst shows a good intrinsic selectivity for butadiene hydrogenation, the results evidenced the presence of severe diffusion limitations in spite of the thin active shell (230 ~tm). The experimental data were modeled by Langmuir-Hinshelwood kinetic expressions derived from an elementary mechanism. Nine kinetic parameters were reliably estimated by means of a regression analysis and it is concluded that the proposed kinetic model provides a good fitting of the experimental observations. 1. INTRODUCTION The elimination of small amounts of 1,3-butadiene (BD) and acetylenic compounds (e.g. 1butyne) in the purification of Ca cuts rich in olefins, 1-butene (1BE), cis2-butene (cBE) and trans 2-butene (tBE), is conveniently carried out by liquid-phase selective hydrogenation on Pd/A1203 catalysts [ 1]. A scheme of the overall reactions, excluding the hydrogenation of acetylenic compounds, is depicted in Fig. 1. Selectivity in this reaction network relies upon the different adsorption strengths of the unsaturated compounds. Butadiene will cover essentially all active sites blocking the access of n-butenes. Only when BD concentration becomes low enough 1BE will begin to hydrogenate to n-butane (BA) and isomerize to cBE and tBE. When the process is intended for obtaining high purity 1BE the product should not contain diolefins and acetylenic compounds in amounts higher than about 10 ppm. These specifications should be achieved with minimum losses of 1BE. 1-Butyne has not been included in this study as acetylenic compounds adsorbs much strongly than BD. Therefore, the latter is the crucial impurity concerning the selectivity of the process. Kinetic studies available in the literature have been performed either in gaseous-phase [2, 3] or with laboratory catalysts [4]. It is the main objective of this contribution to present the Corresponding author.

Mailing address: CINDECA,C.C. 59, B 1900AJK - La Plata, ARGENTINA e-mail: [email protected]

528

J

~

rgl

B

"N~ A

Fig. 1. Overall reaction network. results of a study intended to characterize kinetically the set of overall reactions depicted in Fig. 1, when carried out on a commercial catalyst of the egg-shell type. 2. EXPERIMENTAL Samples from six different commercial catalysts were preliminarily tested. All of them were of the egg-shell type, with a narrow external layer in the range 100-250 ~tm impregnated with Pd as the main catalytic agent. These preliminary studies indicated similar behaviour, as regards activity and selectivity. One of them was chosen to continue the study. The main characteristics of this catalyst are: Pd at 0.2% w/w (overall) on alumina support, spheres of 2.34 mm in diameter and 230 ~tm active external layer thickness. The experimental reaction system consists of a 100 ml stirred vessel and an external 88 tube holding the catalyst sample. The liquid solution is recirculated between the stirred vessel and the catalytic bed by a gear pump. The recirculating flow rate was high enough to minimize external transport effects in the fixed bed [5]. The stirred vessel serves for maintaining the solution saturated with H2 and for temperature control. Catalytic particles in its original size were packed in the external tube. Some other details concerning the experimental system were described in [5]. Batch type experiments with respect to the hydrocarbon mixture were carried out. The batch volume was around 100 ml and the catalyst weight was about 1 g. n-Hexane (inert) was used as a solvent and some quantities of propane were employed to regulate independently the total pressure and H2 partial pressure (PH2), which was kept constant during the run [5]. A low decay of catalytic activity was observed for a catalyst sample after successive runs. To make sure that the activity level was the same for tests under different experimental conditions the catalyst sample was changed, following the procedure detailed in [5]. Three tests were carried out with BD as the only initial unsaturated species at around 2% and at three levels of pH2, about 0.1, 0.2, 0.3 MPa. A fourth test was performed with an additional initial amount of 1BE at 16% and pro=0.30 MPa. The four tests were made at 40~ Some other experimental details were discussed elsewhere [5]. 3. KINETIC MODEL Intrinsic reaction rates expressions have been developed by considering a mechanism of fifteen elementary surface steps, Table 1. A basic assumption cast in this mechanism is that H2

529 adsorbs dissociatively (step a) on sites (| different to those (*) involved with the adsorption of hydrocarbons species [6]. The steps (b) and (f-i), for the hydrogenation of BD, are taken from models proposed in the literature [2]; while the rest of the steps were previously proposed from an analysis of experimental data in vapour-phase [7]. To derive steady reaction rates expressions from the elementary steps in Table 1, it is additionally assumed that all adsorption steps are equilibrated, and that the amount of sites " * " occupied by the radicals C4 H7 *, C~ H 9 *, and C] H 9 * can be neglected. The reaction rates expressions for the set of overall reactions defined in Fig. 1 are given in Table 2, where xj stands for the mole fraction of species j. All the parameters depend on the elementary coefficients of the fifteen steps defined in Table 1. From these relations (not given here because space limitations), it is concluded that there are some constraints for the values that can take the parameters in Table 2. These constraints can be expressed as k14i k 5 k8. = k9 '

k6 = k 5 k____8,

k7K~ q'

k5 ~/=k9 ( l + k 7 / k 8 ) + k 7 / K eq 7 "

The equilibrium constants K ~q7, K~q and K9 q , evaluated thermodynamically, show that 2butenes are favoured (Keq-2.739, K 8eq =29.29 at 40 ~ 7 In summary, the kinetic model given in Table 2 presents fifteen parameters which can be evaluated by fitting experimental tests: r

the rate coefficients ki (i=1-3, 5, 7-9), kI4,

r

the hydrocarbon and H2 adsorption constants, KBdD K~dE ' K:d E~Kad tBE and K

r

the coefficients ~ and [3 in the H2 inhibition terms.

The kinetic expressions in Table 2 are written in terms of local mole fraction xj in the liquid solution filling the pores of the active layer. As severe diffusion limitations were expected, mass conservation equations inside the catalytic layer should be solved to relate the observable reaction rates to the bulk liquid composition. Details about the formulation and the numerical solution are given in [5]. For specified values of xj,bu~kand assuming that the values of kinetic parameters in Table 2 Table 1 Catalytic mechanism

(a)

(g)

C4H 7

*+H|

--4 1BE * + |

(h )

C4H 7

*+H|

--~ cBE * + |

(i)

C4H 7

*+H|

~ tBE*+|

H 2 + 2| r

2(H |

(j)

(b)

BD+* r

BD*

(c)

1BE +* r

1BE *

(d)

cBE +* r

cBE *

(k) (1) (m)

(e)

tBE + , r

tBE *

(n)

C4Hl9 * + H |

--+ nC4Hlo + * + |

( O)

C4H ~ * + H |

--~ nC4Hlo + , + @

(f)

BD* + H|

r

C4H 7

*+|

1BE 9+ H| C4Hl9 9+ | 1BE * + H| r

C4H ~ * + |

cBE * + H| r

C4H 9

tBE * + H| r

C4H 9

2 ,+|

2 ,+|

530 Table 2. Kinetic expressions

q=

r4 ~

k, KBdD XBD XH2 DEN.c DEN ~

k4K EXBE

r6

Kad cBE

X cBE

k 2 KBd XBD XH2

r3=

ot

DEN .c DEN H2

kKEXBE

DENnc DEN~ ks

"5

rE=

+

DENHc DEN~2

k3

K ad BD

ot

DEN HCDEN 82

k7Kad EBE r7

=

XBD XI4_2

XcBE 1 XmE - K7 q

DENHc DEN~

ks KIBE ~ XtBE r8 : D E I ~ D-E-N~ X'BE- K:q

X H2

ad

DENHc DEN~2

k6 K~E XtBE XH2 r9 "-"

DENHc DEN~2

E I

DEN. c DEN~2

]

XtBE1 K9 q

DENHc : 1 + KBd XBD + K;dBEX,BE + K:dBEX~BE + K~E X~E

-(1 +

)

DEN~ - DENH2 (1 +13~ K ~ x m )

D E N ~ = D E N m ( 1 + 7 x~~m)

are known (or tried, if a regression analysis is being performed), the mole fraction fields inside the layer are obtained and the observable reaction rates Ri c a n be calculated as Ri=

ridz LSp/Vp,

9

i=1,..,9;

(1)

where L is the thickness of the active layer, Sp and Vp are the external surface area and total volume of a catalyst particle. Note that Ri is expressed by unit of total catalyst weight and not catalytic layer weight. 4. EXPERIMENTAL RESULTS The results of the tests at PH2=0.10 and 0.32 MPa and BD as the only initial unsaturated species are presented in Figs. 2a,b. The only species not shown is tBE, which follows a similar pattern as that of cBE. The time indicated by t* in Figs. 2a,b is defined as the reaction time when the mole fraction of BD has dropped to satisfy

XBD,bulk--(DH2fDBD)XH2,bulk~ 4 XHZ,bulk,

(2)

where DH2 and DBD are effective diffusion coefficients inside the active layer. When relation (2) is satisfied, BD and H2 diffusion capabilities inside the catalytic layer are the same. XH2,bulkis the saturation mole fraction, which remains constant during each test. Assuming that strong diffusion limitations exists, at t>t* there will be H2 in excess to the amount needed for hydrogenating BD up to depletion inside the catalytic layer. The n-butenes will start to react irrespective of the intrinsic inhibitory effect of BD. We can appreciate in Figs. 2a,b that BD is consumed following an apparent zero-th order

531 2.5%

Pm :0.10 MPa o BD D IBE A cBE

2.0%

9 BA

1.5%

(a)

2.5%

2.0%

,/4

tXcBE

,//f//

1.0%

0.5%

0.5%

0

___ 50

100

150

_ 200

REACTION TIME [ min]

250

(b) ~~* ./"

1.5%

1.0%

0.0% _..--

pm= 0.32 MPa o BD ra IBE

300

0.0% 20

40

60

80

100

120

R E A C T I O N T I M E [ rain ]

Fig. 2. Experimental (symbols) and model (full lines) results. The vertical dotted line indicates t*.

reaction when tt*. A quantitative analysis will reveals that 1BE hinders the hydrogenation of cBE and tBE, but this effect is much weaker than that of BD. The results from the test with initial mole fractions XBD,bulk--0.9 10-2 and X1BE,bulk--15.5 10-2 are plotted in Fig. 3. This initial composition is representative of an industrial C4 cut for 1BE purification. At the level of HE partial pressure employed in this test (Pn2=0.30 MPa) the initial XBD,bulkis already lower than the value given by Eq. (2). Therefore, 1BE starts to react just at t=0. This test confirms that the qualitative behaviour already discussed extends to much higher levels of n-butenes concentration. 5. REGRESSION ANALYSIS As the catalytic bed operates under essentially uniform bulk liquid composition (due to the high recirculation flow), the conservation equations in the liquid solution for the unsaturated

532

Xj,bulk

Pm = 0.30 MPa 14% o BD

**

1.5%

12% 10% 1.0% 8%

6% 0.5%

4% 2% 0%

0.0% 0

50

100

150 200 250 300 REACTION TIME [ min I

350

400

Fig. 3. Experimental (symbols) and model (full lines) results. species during each test can be written as

(3)

N Tdx j,bulk / dt = M cat R j,

where NT is the total number of moles in the liquid solution and Rj is the net rate of species j production per unit mass of catalyst sample, Meat. Rj can be expressed in terms of Ri values (Eq. 1), according to the stoichiometry of the reaction system (Fig. 1). For trial values of the kinetic parameters in Table 2, Eqs. (3) were numerically integrated from the initial composition. Predicted values of bulk liquid composition can thus be obtained at the same reaction times at which samples were cromatographically analyzed. Predicted and measured values of Xj,bulk allow to undertake the regression analysis for estimating the kinetic parameters in Table 2. This was performed by employing the pack of routines GREGPAK [8] in multiresponse mode. The best estimates of the independent rate coefficients ki (i=1-3, 5, 7-9), k I are given in Table 3 along with their confidence limits, which can be evaluated as being acceptably narrow. Individual values of the hydrocarbon adsorption constants can not be evaluated with statistical significance since at the experimental conditions there was always enough concentration of unsaturated species to cover all the active sites. Longer reaction times are needed to this end. The ratio (K ~d / K~dE ) could be estimated, Table 3, but with a large confidence limit. The ratio (KcB Ead + KtBEad)/K1BEad could be estimated with a reasonable confidence limit. Tests with either cBE or tBE as the only unsaturated species must be performed to Table 3. Optimal values of the kinetic parameters; k [mole/(kgactive layer S)] kl = 6.332 101 + 4.446 k2 = 5.874 + 1.008

kI4 = 1.483 101 + 9.801 k5 = 6.433 101 + 7.035

k3 = 3.253 101 + 2.669

k7 = 2.448 + 3.139 10 -1

KBdD/KadlBE= 6.310 102 _+4.982 102

k8 = 3.246 + 4.082 10 1 k9 = 2.032 + 2.515 10 -1

K ad2BE/K1BEad= 3.350 10 -1 _+5.856 10 -2

533 Xj

Xj

Xj 1.0%

1BE

2.0%

tBE

0.20% 0.8%

-~-----BD

1.5%

0.15% 0.6% tBE

1.0%

0.10%

H2 0.5%

cBE.....

~

~ ~ 0 . 0 5 %

0.4%

cBE

[

0.2% IBE

0.0%

m 0

, 0.2

,

J 0.4

,

, 0.6

0.00% 0.8

1

0.0%

i

0

|

0.2

i

|

0.4

0.6

0.8

1

Fig. 4.Concentration profiles inside the active catalytic layer, z is the dimensionless coordinate inside the layer (z=l corresponds to the external surface). discriminate between the individual ratios

(KcB Ead /K1B Ead ),

(KtB Ead /K1BEad).

A very low estimate and an unacceptable confidence limit were obtained for K~d . Consequently, the H2 inhibition coefficients ct and [3 (see Table 2) could not be estimated either. This result reveals that p.2 levels were not high enough to assess H2 inhibition effects. It can be appreciated in Figs. 2a,b and 3 that the kinetic model allows a good fit of the experimental data. A similar quality of fitness was achieved for the data at pn2=0.20 MPa and for tBE, not shown in Figs. 2 and 3. The trends are very well reproduced by the model and the deviations between predicted and observed values of instantaneous mole fractions are satisfactorily low. The average over all species j and all measures of the individual errors as ej =l x j,bulk - - ( X j,bulk )predicted 1 / X j,bulk is 1 3 % . It should be remarked that in spite of a large difference between the levels of 1BE for the tests depicted in Figs. 2 and 3, the evolution of 1BE is equally well described by the model. The most significant parameter out of those which could not be correctly estimated was the ratio (K ad ad )" The large, but uncertain value obtained for this ratio indicates that the BD / K~BE catalyst tested is intrinsically very selective, but the experiments did not allow to quantify clearly the degree of selectivity. This was because relatively few data at very low BD concentration were available, for its hydrogenation rate becomes very fast. Additional tests initiated with lower BD concentration and lower amount of catalyst sample are needed to evaluate properly the ratio (K adBD / K adlBE )" The kinetic model also allows the prediction of concentration profiles inside the active catalytic layer, as those plotted in Figs. 4a, b. In Fig. 4a, the H2 capability is lower than that of BD. This case corresponds to a point of Fig. 2a and a reaction time tt*. In Fig. 4a the only significant reactions are those of BD hydrogenation. In Fig. 4b, although a net production of 1BE is predicted, the H2 excess causes an important rate of 1BE consumption by the isomerization and hydrogenation reactions. It can be appreciated that all reactions take place under very strong diffusion effects. defined

534 6. CONCLUSIONS AND FINAL REMARKS An experimental procedure to test commercial Pd-based catalysts of the egg-shell type used for selective hydrogenation of butadiene in C4 cuts has been described. The results of four tests on a commercial catalyst at different partial pressure of H2 and initial composition show very strong diffusion limitations and a high intrinsic selectivity. This behaviour has also been observed in preliminary tests of other five different commercial catalysts. Intrinsic kinetic expressions based on an elementary mechanism have been proposed for modelling the data here presented. This kinetic model is coupled to the solution of mass conservation equation inside the active catalytic layer. Fifteen kinetic parameters arise from the model and nine of them could be satisfactorily fitted with the available experimental tests at 40~ The most significant conclusions arising from the analysis of the results are: r r

the model is able to capture all significant trends displayed by the experimental data; the experimental procedure is suitable for the estimation of intrinsic kinetic parameters, in spite of the loss of sensitivity caused by the diffusion limitations.

Some of the parameters which could not be suitably estimated may not be significant from a practical point of view. Nonetheless, the experimental variables that should be modified to complete the model identification could be clearly established. ACKNOWLEDGEMENTS

The authors wish to thank the assistance of the following Argentinean institutions: ANPCyTSECyT (PICT No. 00227), CONICET (PIP96 No.4791), and UNLP (PID No.11/I058, Fellowship to NOA). JAA, OMM and GFB are members of CONICET, SPB is member of CIC PBA. REFERENCES

1. L. M. Derrien; Studies in Surface Science and Catalysis, Vol. 27, Chap. 18; L. t~erven3~ (ed.), Elsevier, Amsterdam, 1986. 2. J. Goetz, D. Yu. Murzin and R. A. Touroude; Ind. Eng. Chem. Res., 35 ( 1996 ) 703. 3. P. Kripylo, F. Turek, K-D. Hempe and H. Kirmse; Chem. Techn, 27 ( 1975 ) 675. 4. J-P. Boitiaux, J. Cosyns and E. Robert; Applied Catalysis, 35 ( 1987 ) 193. 5. N. O. Ardiaca, S. P. Bressa, J. A. Alves, O. M. Martinez and G. F. Barreto; Catalysis Today, in press. 6. A. J. Renouprez, A. G. Clugnet and H. Jobic; J. Catal., 69 ( 1981 ) 180. 7. S. P. Bressa, O. M. Martinez and G. F. Barreto; Proc. XI Jornadas Argentinas de Cat~ilisis, papers 7-8 ( 1999 ). 8. W. E. Stewart, M. Caracotsios and J. P. Sorensen; AIChE J., 38 ( 1992 ) 641.


535

KINETICS AND MECHANISM OF RUTHENIUM(Ill) AND OSMIUM(VIII) CATALYZED OXIDATION OF DOPAMINE WITH BROMAMINE-B IN ACID AND ALKALINE MEDIA Puttaswamy* and Nirmala Vaz Department of chemistry, Central College, Bangalore University, Bangalore - 560 001, India. ABSTRACT The kinetics of oxidation of dopamine (DPM) by bromamine-B (BAB) has been studied in HC1 medium catalyzed by Ru(III) chloride and in NaOH medium with Os(VIII) as catalyst at 298 K. Despite the fact that the stoichiometry and oxidation products are same in both catalyzed reactions, the kinetic patterns are different. In acid medium the rate shows a first order dependence each on BAB, DPM and Ru(III) and fractional order on H + concentrations. In alkaline medium, the rate is first order each on [BAB]o and [DPM]o, a fractional order on [Os(VIII)] and an inverse first order on [OH-]. The effects of addition of benzene sulfonamide and halide ions as well as variation of ionic strengths and dieletric constant of the medium did not show significant effect on the rate of the reaction. Solvent isotope effects has been studied in D20 medium. Activation parameters have been calculated. The catalytic effeciency is Os(VIII) >> Ru(III). The proposed mechanisms and the derived rate laws are in consistent with the observed kinetics.

1. INTRODUCTION Chloramine -T (CAT) and chloramine -B (CAB) are the prominent members of organic haloamines and their chemistry has been reviewed[I,2]. Bromamine-B (C6HsSO2NBrNa 1.5 H20 ; BAB ) the bromine analogue of CAB is found to be a better oxidizing agent than its chloro compound. Although the oxidation of some substrates with BAB has been studied [3,4], a little attention has been focused on BAB's reactions with pharmaceuticals. The substrate, dopamine (3-hydroxytyramine ; DPM) is one of the naturally occuring catecholamines and the lowered level of this in the brain is known to cause the neurological disorder- Parkinson's disease. It finds applications in the pharmaceutical industry and its hydrochloride salt is widely used in the treatment of shock and in acute congestive failure. Hence, it was found interesting to investigate the oxidative behaviour of BAB towards dopamine. The reaction of dopamine and BAB under the present experimental conditions are sluggish in the absence of a catalyst. In recent years, the use of transition metal ions in the oxidation of several redox processes is of considerable interest. We have found that ruthenium(III) chloride and osmium(VIII) are excellent catalysts for the title reaction. Therefore in the present paper, we report the detailed kinetics of oxidation of dopamine by BAB with Ru(III) and Os(VIII) catalysts in acid (HC1) and alkaline (NaOH) media respectively for elucidating the mechanism of oxidation of this drug.

536 2. E X P E R I M E N T A L Bromamine-B was prepared[5] by the partial debromination of dibromamine-B and its purity was checked iodometrically and also by UV, IR and NMR spectra. An aqueous solution of the compound was prepared, standardized iodometrically and preserved in brown bottles to prevent its photochemical degradation. Dopamine hydrochloride (Merck) was used as received. Reagent grade chemicals and doubly distilled water were used throughout. Ru(III) and Os(VIII) (Johnson Mathey, London) were used as catalysts and their solutions were prepared in 0.5M HC1 and 0.5M NaOH respectively. Allowance for the amount of acid / alkali present in the catalyst was made, while preparing reaction mixtures for the kinetic runs. Constant ionic strength (I - 0.5M) of the medium was maintained by using a concentrated solution of NaC104 for both the cases. The kinetics were followed under the condition [DPM]>>[BAB] by measuring the rate of disappearance of BAB, the estimation of which was done iodometrically. Pseudo-first order rate constants (k/) calculated from log [BAB] vs time plots were reproducible to 3-4%. Regression coefficients 'r' were evaluated. The oxidized product 2-(3 / 4/-benzoquinone) ethylamine from DPM in both the cases was detected by IR.

3. RESULTS

Ruthenium(III) catalysis: At constant

[DPM]o (5.0 x 10 .3 M), [HC1] (1.0 x 10 -z M), [Ru(III)] (5.42 x 10 -5 M) and temperature (T = 298 K) a first order dependence on [BAB]o is noted. The pseudo-first order rate constant (k / = 4.80 x 10 .4 s l ) was not affected by change in [BAB]o (2.0 x 10~-12.0 x 10.4 M) confirming the first order dependence on [BAB]o. Values of k / increased with increase in [DPM]o (2.0 x 10 -3 - 1.6 x 10 -z M) and a plot of logk / vs log[DPM]o was linear (r = 0.9996) with unit slope showing a first order dependence on [DPM]o. With increase in [HC1] (5.0 x 10 -3 - 5.0 x 10 2 M) k 9 values increased and a plot of logk / vs log[HC1] gave a straight line (r = 0.9990) of slope 0.58 indicating a fractional order dependence on [HC1]. When [Ru(III)] was varied from 1.90 x 10 .5 - 9.68 x 10SM rate was found to increase and a plot of logk / vs log[Ru(III)] was linear (r = 0.9983) with unit slope, indicating a first order dependence on [Ru(III)]. At constant [H +] = 0.01M maintained with HC1, addition of NaC1 did not affect the rate which reflected the dependence of the rate on [HC1] is only due to [H+]. The rate was unaffected by the addition of benzenesulfonamide or NaBr (5.0 x 10.4 - 4.0 x 10 .3 M). Variation of ionic strength (0.1-0.8 M) and also dielectric constant of medium by adding different proportions (0 - 40% v/v) of methanol to the reaction mixture do not have any significant effect on the rate. The solvent isotope effect was studied in D20 wherein the reaction rate was further increased with k / = 5.12 x 10.4 s ~ in DzO medium and 4.80 x 10-4 s "1 in H20 leading to a solvent isotope effect, k/(H20) / k/(D20) = 0.94. The reaction was studied at different temperatures (293-313 K), and from the linear Arrhenius plot of logk / vs 1/T (r= 0.9985), values of the activation parameters Ea, AH~,AG~and ASs were evaluated and found to be 52.1, 49.6, 92.3 kJ mo1-1 a n d - 1 4 2 JK 1 mol -~, respectively. Addition of reaction mixture to aqueous acrylamide solution failed to initiate polymerization indicating the non-involvement of free radicals.

Osmium(VIII) catalysis: The

plots of log [BAB]o vs time are linear (r> 0.9916) at constant [DPM]o (5.0 x 10 -3 M), [NaOH] (5.0 x 10 -3 M), [Os(VIII)] (4.28 x 10 -6 M) and T (298 K)

537

indicating a first order dependence on [BAB]o. The pseudo-first order rate constant (k/= 2.35 x 10-3 s-1) increased with increase in [DPM]o (1.0 x 10.3 -1.2 x 10.2 M) and a plot of logk/ vs log[DPM]o was linear (r = 0.9997) with unit slope indicating a first order dependence on [DPM)o. The rate decreased with increase in [NaOH] (2.0 x 103-1.2 x 10-2 M) and a plot of logk vs log[NaOH] was linear (r = 0.9987) with a negative slope of unity indicating an inverse first order dependence of rate on [OH-]. The rate of reaction increased linearly (r = 0.9994) with increase in [Os(VIII)] and the order with respect to [Os(VIII)] was found to be 0.66 indicating a fractional order dependence on [Os(VIII)]. Addition of halide ions or reaction product (5.0 x 104-3.0 x 10.3 M) and variation of ionic strength (0.2-0.8 M) or dielectric constant of medium show negligible effects on the reaction rate. Study of rate in D20 medium showed that k/(H20) was 2.35 x 10-3 s-1 while k/(D20) = 1.52 x 10-3 s"1 leading to a solvent isotope effect of k/(H20) / k/(D20) = 1.55. The corresponding proton inventory plot for the rate constant k/n vs the deuterium atom fraction 'n' in the solvent mixture is non-linear. Kinetic runs were made at different temperatures (293-313 K) and the activation parameters were found to be Ea = 34.9, AH~ = 32.3, AG~ = 88.5 kJ mo1-1 and ASs= -187 JK "1mol -~. The test for free radicals was negative. 4. DISCUSSION

Ru(III) catalysis : Bishop and Jennings[6] and Hardy and Johnston[7] have shown the existence of similar equilibria in acid and alkaline solutions of organic haloamines. Bromamine-B like its chlorine analogues CAB and CAT, acts as an oxidizing agent in both acidic and alkaline media. In general BAB undergoes a two electron change in its reactions. The reduction potential of BAB-PhSO2NH2 is pH dependent and decreases with increase in pH of the medium and depending on pH of the medium, BAB furnishes different types of reactive species in solution, such as PhSO2NHBr, PhSO2NBr2 and HOBr in acidic solutions. If PhSO2NBr2 were to be the reactive species, then the rate law predicts a second order dependence of rate on[BAB]o, which is contrary to the experimental observations. If HOBr is primarily involved, a first order retardation of rate by the added benzenesulfonamide is expected. Since no such effect is seen, HOBr can be ruled out as an oxidizing species. Hence PhSO2NHBr is responsible for the oxidation of dopamine in acid medium. Narayanan and Rao[8] reported that monohaloamines can further be protonated at pH < 2. Hence it is reasonable to assume that PhSO2NH2Br+ of BAB is the likely oxidizing species in acid medium. Although Cady and Connick[9] and Connick and Fine[ 10] based on abso~tion spectra have shown that in aqueous media octahedral complexes such as [RuC15(H20)] -, [RuC14(H20)2], [RuC13(H20)3], [RuC12(H20)4]§ and [RuCI(H20)5] 2+, may not exist for RuC13, others[ 11,12] have shown that in acidic solutions the following equilibria exists for ruthenium(III) 9 RuC13 xH20 + HC1 [RuC16] 3- + H 2 0

~-~

[RuC16]3- + xH20 [RuC15( H 2 0 ) ] 2" + CI"

(1)

Singh et al.[ 13] used the above equilibrium in ruthenium(III) chloride catalyzed oxidation of primary alcohols by BAT. In the present study, however, the ineffectiveness of chloride ion on the rate indicates that equilibrium(I) does not play a role in the reaction and hence the

538 complex ion, [RuC15(H20)] 2, is assumed to be the reactive catalyst species that interacts with dopamine to form a complex intermediate. Further, UV spectral studies suggests that a complexation occurs only between the catalyst and the substrate. Based on the preceding discussion and observed kinetic results, a mechanism (Scheme 1) is proposed for the Ru(III) catalyzed oxidation of DPM by BAB in acid medium : PhSO2NHBr + H + --., K 1 K2 DPM + Ru(III) -.~ PhSO2NH2Br+ + X X/

k4

k3

,._

,.~

PhSO2NH2Br+ X ~ X/

~ products

~ (i) fast (ii) fast (iii)slow & rds (iv) fast

Scheme 1

Here X and X / represent the complex intermediate species. Based on the rds of Scheme 1, rate = -d[BAB]t/dr = k3[PhSO2NH2Br +] [X]. Since the total effective concentration of the oxidant BAB is given by eq.(2), the rate law (eq.(3)) can be derived : [BAB]t = [PhSO2NHBr] + [PhSO2NH2Br +]

(2)

rate = K1K2k3 [BAB]t [H+] [DPM] [Ru(III)] / {K~[H+] + 1}

(3)

Rate law (3) is in agreement with experimental results and can be written as, 1 / k / = 1/K2k3 [DPM] [Ru(III)] + 1 / K~KEka[DPM] [Ru(III)] [H+]

(4)

From the slope and intercept of the linear plot of 1/k/ vs 1/[H+] (r = 0.9826),values Of Kl and K2k3 found were 52 M and 5.27 x 103 M "2 s"l respectively. The proposed mechanism is supported by the enhancement of rate in D20 medium. Since D 3 0 + ion is a stronger acid than H30+[ 14] by a factor of 2 to 3, a solvent isotope effect of this magnitude is expected. But the observed inverse solvent isotope effect is k/(D20)/k/(H20) = 1.07. This can be attributed to the fractional order dependence on [H+]. The absence of dielectric effect of solvent on the rate, as was observed in the present investigations, signifying that the transition state is not very different from the reactants[ 15]. The proposed mechanism is also supported by the moderate values of energy of activation and other activation parameters. The fairly high positive values of AG~ and AH~ indicate that the transition state is highly solvated, while the large negative entropy of activation suggests the formation of the compact activated complex with few degrees of freedom. The reduction product does not influence the rate, showing that it is not involved in preequilibrium. Change in ionic strength of the medium does not alter the rate indicating that non-ionic species being involved in the rds. Addition of halide ions has no effect on the rate indicating that no interhalogen or free bromine is formed. All these observations also confirm the proposed mechanism.

539

Osmium(VIII) catalysis: In alkaline solutions, PhSO2NBr2 does not exist and the possible oxidizing species in alkaline BAB solutions[6,7] are PhSOzNBr- and OBr-.In alkaline solutions, however, PhSOzNBr- ion which is in larger concentrations undergoes reactions represented by the equilibria (5) and (6) : PhSO2NBr- + H20

~

""

PhSO2NBr-+ H20

; ~

PhSO2NHBr + OH-

(5)

PhSO2NI-I2+ OBr

(6)

The base retarding equilibrium (eq.(5)) is generally present at low alkaline pH. The reaction (6) occurs at high alkalinities and is retarded by the addition of benzenesulfonamide. According to Hardy and Johnston[7], who have made detailed calculations on the relative concentrations of reactive species in alkaline solutions of BAB, at pH ~ L) 0.2

~176 *'~'-"'~

Ni-L-CS

------v----,,----T

0.0

. 0

, 100

.

,

9

200

, 300

Time

(a~

(rain)

Ni-E-SBA 9

, 400

9

, 500

-~

0.6

O0

0.4

_.r

Ni-E-SBA

:E 0 > 0.2 0.0

"(~. - ~ - ~ " ' - , ~ = _ . ~

Ni-L-CS ' ~ ' ' Ni-L-SBA

. . . . . . . . . .

o

1~o

2~o

3~o

Time(rain)

H P318

4~o

(b)

Fig. 4 TCEa conversion (a) and VCM selectivity (b) of the catalysts at 400 ~

5~o

563 3.3. Catalyst Deactivation Three catalysts prepared in this study showed a relatively severe and fast deactivation than the commercial catalyst. To investigate the reasons of the catalyst deactivation, SEM and EPMA methods were used to characterize the surface property changes after the dechlorination reaction. SEM images of the catalyst before and after the dechlorination reaction at different temperatures are shown in Fig. 5. Significant coke formation was observed only in the sample tested at 500~

The coke was filamentous and its diameter was

approximately 40nm. High reaction temperature of 500~ is considered to be the major reason of the coke formation.

Fig. 5 SEM images of the Ni-E-SBA catalyst before and after 10 hrs of reaction (a) fresh catalyst, (b) after reaction at 300~

(c) 400~ and (d) 500~

Table 2. Ni-E-SBA catalyst composition after 10 hrs of reaction Temperature

C1/Ni ratio

Carbon Contents (wt%)

fresh catalyst

0.00

0.046

300 ~

3.29

2.603

400 ~

3.04

3.665

500 ~

2.63

28.33

564 Table 2 shows the EPMA results for the Ni-E-SBA catalyst. Carbon contents of the samples are in good agreement with the results of the SEM image. Carbon contents increased and C l ~ i ratio decreased with the reaction temperature. The predominant mechanisms for deactivation of the catalysts were thought to be the irreversible adsorption of HC1 for low reaction temperature (300~ temperature (500~

and coke formation on the catalysts for high

respectively. Results of kinetics for dechlorination of TCEa will be

reported. 4. Conclusions Catalysts were prepared by a novel method via silylation, grafting, metal adsorption and calcination steps using the SBA-15 as a support. Pore structures of the support, SBA-15, were maintained throughout the preparation steps. In the TCEa dechlorinating reaction test, catalysts prepared in this study showed a better selectivity for the VCM than the commercial catalyst. Major deactivation mechanism of the catalyst was thought to be a irreversible adsorption of HC1 at low reaction temperature and coke formation at high reaction temperature, respectively. 5. References [ 1] Dae I. Kim, David T. Allen, Ind. Eng. Chem. Res., 36, 3019-3026 (1997) [2] R.J. Meyer, D.I. Kim, D.T. Allen and J.H. Jo, Chem. Eng. Sci., 54, 3627-3634 (1999) [3] Abdelhamid Sayari, Chem. Mater., 8, 1840-1852 (1996) [4] G. Tavaoularis, M.A. Keane, J MoL Catal. A: Chem., 142, 187-199 (1999) [5] Zhaohua Luan, Estelle M. Maes, Paul A. W. van der Heide, Dongyuan Zhao, Roman S. Czernuszewicz and Larry Kevan, Chem. Mater., 11, 3680-3686 (1999) [6] Dongyuan Zhao, Qisheng Huo, Jianglin Feng, Bradley F. Chmelka and Galen D. Stucky, Jr. Am. Chem. Soc., 120, 6024-6036 (1998) [7] Dongyuan Zhao, Jinyu Sun, Quanzhi Li and Galen D. Stucky, Chem. Mater, 12, 275279 (2000)

Studies in Surface Science and Catalysis 133 G.F. Froment and K.C. Waugh (Editors) ~ 2001 Elsevier Science B.V. All rights reserved.

565

Kinetics of copolyesters of poly(butylene terephthalate), hydroquinone diacetate and terephthalic acid: A simple rate model for catalysed synthesis in melt. Amir AI-Haddad* & Johnson Mathew Chemical Engineering Department, Kuwait University, PO Box 5969, 13060 Safat- Kuwait

Abstract Considerable literature exists on the structural organisation of thermotropic liquid crystalline aromatic copolyesters. Majority of these copolyesters are synthesised via the copolymerisation route, which helps in tailoring the characteristic properties to the predecided values. Another major advantage of copolymerisation is that it helps in conferring specific chemical properties to the major component present in the system.

Synthesis of aromatic polyesters using bromo, chloro, methyl and methoxy substituted hydroquinones have been reported, to lower the melting temperatures relative to the unsubstituted polyesters. These polyesters have applications in fabricating thermally stable high strength fibres and moulding resins with unusual properties. An understanding of the melt polyesterification kinetics is a must to economise on the process productivity and to improve the polymer properties. There is no published literature on the melt polyesterification kinetics of PBT, HQDA and TA. Kinetic investigation of copolymerisation between PET and 4-acetoxy benzoic acid (PET/OB) did not reveal precipitation of poly (4-oxybenzoate). Here we explore the kinetics of a three component system, wherein many parallel reactions take place simultaneously. Homopolymerisation between HQDA and TA lead to a rigid rod system. The work was carried out with the following objectives: (1) To identify plausible routes to simplif the kinetics of a three component system; (2) To determine the kinetic order with respect to the homopolymers of HQDA and TA as well as the copolymers of PBT, HQDA and TA; (3) To predictthe rates of acetic acid generation during the homopolymersiation and copolymerisation reactions and (4) To check whether precipitation constitutes an added complication, as observed in the earlier study on one component 4-acetoxy benzoic acid system.

*Author for correspondence (hadad@ kucO 1.kuniv.edu.kw)

566 INTRODUCTION

Copolymerization reactions are often an attempt to obtain properties intermediate between those of the homopolymersl'2. The properties of PBT based copolymers depend on the composition of the reaction mixture. The composition is influenced by the reaction mechanism and kinetics, which are further influenced by the choice, purity, and the concentration of the main reaction components, by the presence and proper choice of the amount of catalyst and by the reaction temperature. The mechanism and kinetics of polyesterification, polyamidation, polyurethanes and many more have been investigated in the past as well as today 3-11. There has been no evidence of kinetic studies of melt transpolyesterification of PBTTA-HQDA ternary system. An earlier study relating to copolyesterification of PET and 4-acetoxy benzoic acid did not manifest precipitation of poly(4-oxybenzoate) 9'~~ This work is an attempt to investigate the melt reaction behaviour of the PBT/HQDA/TA system. The work is presented as follows: A simple reaction sequence is suggested to outline the dominant steps involved. The major reactions, which lead to acetic acid production, are taken into account in the kinetic analysis. Simple second order kinetics is found to be effective for each of the steps significant to the acetic acid production. The correctness of this procedure is justified by experimental observations. Finally, physico-chemical characterization of the copolyesters are considered.

EXPERIMENTAL Poly(butylene terephthalate) (PBT) of inherent viscosity 0.58 and 1600 micron particle size (M/s Century Enka Pvt. Ltd., Pune, India) and purified terephthalic acid (TA) were used as received. Hydroquinone diacetate (HQDA) was prepared by sulfuric acid catalysed reaction of hydroquinone and acetic anhydride and was recrystallised from acetone (m.p. 121~ A 300 mL electrically heated histalloy based reactor ll, was used for the melt tranesterification reactions. Tranesterification were conducted to generate a copolyester of PET 50/(HQDA + TA) 50 mole% composition. The isothermal reaction temperatures chosen for the kinetic estimations were 265, 275 & 285~ Dry nitrogen blanket was maintained through out the experiments to prevent oxidative degradations. The rate of evolution of by-product, acetic acid, was monitored volumetrically to estimate the kinetic parameters. Dibutyl tin oxide and zinc acetate (0.25, 0.5 & 1.0 mole% of HQDA) were used as catalysts for the melt transesterification reactions. Thermal transitions were obtained with a Mettler DSC 30 apparatus under nitrogen atmosphere using a sample size of 10-15 mg. Samples were analysed in the temperature range of 50-250~ in the first and second heating cycles.

567

RESULTS AND DISCUSSION The reactions when PBT, hydroquinone diacetate and terephthalic acid are heated together or maintained isothermally at temperature in excess of isotropic temperature of crystalline PBT may be visualised as follows: i) PBT chain cleavage with terephthalic acid (TA) resulting in the formation of two PBT segments of similar or differing in chain lengths but terminated with carboxylic acid end group, (SpBTT), ii) The segment of (i) react with hydroquinone diacetate (HQDA) to form PBT segment SpBTTH with acetoxy end group, iii) segments SpBTT and segment SpBTTH could then react to reform the PBT chain with insertion of hydroquinone-terephthalate unit, iv) transesterification of hydroquinone diacetate and terephthalic acid leads to form hydroquinone terephthalate (TH) and v) the insertion of hydroquinone-terephthalate unit into PBT chain (PB(TH)T. Thus, one can visualize the existence of many variables (rate constants) for these independent reactions. It is difficult task to analytically solve for these kinetic variables. Hence the following assumptions are made to simplify the kinetic picture: (1) The reaction between HQDA and TA leads to the formation of dimers (TH); (2) An oligomer (dimer) of HQDA and TA can cleave the PBT chain to form PBT segments SpBTT and SpBTTH; (3) These segments then react to reform PBT chain with insertion of TH moiety. Hence, the reactions leading to the copolyester of interest are given by equations (A) and (B). If c denotes the dimer formed by the reaction of HQDA and TA, then the following steps can be presumed for the copolyesterification reaction. H

+

( a - x) PBT ( q - y)

T

kl

> TH

( a - x) +

TH

+

CH3COOH

(c) k2

> PB(TH)T

(c)

(y)

(A)

(x + y) +

CH3COOH

(B)

(x + y)

where, H, T, TH, PBT, PB(TH)T, CH3COOH, a, c, x, y and q denote hydroquinone diacetate, terephthalic acid, dimer of HQDA and TA, poly(butylene terephthalate), copolyester, acetic acid, initial concentration of HQDA and TA, initial concentration of dimer number of moles of H or T converted, number of moles of PBT chains converted and initial concentration of PBT respectively, kl and k2 denote the rate constant of homopolyesterification and copolyesterification respectively. Rate of formation of the dimer can be given as d_}x= -d___HH= - d T = kl (a - x) 2 dt dt dt Rate of formation of the copolyester can be given as

(1)

568 dy dt

de

=

- dPBT ~ = k 2 c ( q - y) dt

(2)

= k 1 ( a - x) 2 - k 2 c ( q - y)

(3)

dt total rate of acetic acid production is given by the algebraic sum of eqn. 1 and 2. dCH3COOH dt

dx dy =~ + ~dt dt

2 k 1 (a - x)

+ k 2 c ( q - y)

(4)

From the conservation of residues of H and T, we have c = x - y. Equation 2 now becomes

d y = k 2 (x - y) (q - y) dt

(5)

Equation 4 can now be written as dCH3COOH dt

= k 1 (a - x) 2 + k 2 (x - y) ( q - y)

(6)

A simple numerical procedure ~1 was applied to equation (6) to calculate theoretically the values of kl and k2. A number of reactions were carried in steps of 10~ from 265 to 285~ Figure 1 indicates the experimental data points at three different temperatures and fitting curves according to equation (6) for PBT 50/(HQDA + TA) 50. The reaction rate is also adequately modeled by second-order kinetics for catalyzed reactions (Figure 2). Figures 3 is a typical figure indicating that the model adheres well to the experimental data analyzed for different catalyst concentrations. A typical Arrhenius plot is shown in Figure 4. As seen in the plot the experimental data can be adequately modeled by equation (6). Table 1 indicates the rate constants (kl & k2) and activation energies (El & E2) for zinc acetate and dibutyl tin oxide catalyzed reactions for PBT 50 and HQDA + TA 50 mole% concentration. The values of kl are found to be the range of 0.03136 to 0.5 kJ/mole while for k2 it is in the range of 0.004 to 0.1 kJ/mole. No major decrease in the value of the rate constants k~ and k2 are noticed. The activation energy values indicate that DBTO at 0.25 mole% concentration is a suitable catalyst for the PBT 50 mole% concentration.

569

I 1 265C I 9 275C I 9 285C I

0.07

Moa~] I

0.06 0.05

v 0.04

~0.03

c3

0.02

0.01

!

20

---

~

!

-

40

i

!

!

i

!

!

!

!

60

80

100

120

140

160

180

200

Time ( min ) Figure. | : Second order plot for uncatalyzed PBT (50%) & (HQDA+TA) (50%) with the model

9 265 C 9 275 C m 285 C Model

0.09 0.08 0.07 0.06 o.o5 ..~

0.04 0.03 g~

0.02

0.01 v

0

v

i

i

i

!

50

100

150

200

Time ( mins ) Figure 2" Second order plot of 0.25 mole% dibutyle tin oxide catalyzed reaction for PBT (30%) and (HQDA+TA) (70%) with the model

570

9 0.25mole % [] 0.5 mole % 9 1 mole % 9 uncatalyzed ----'Model

0.07 -

0.06

"

"~"0.05 o.o4

i

-

0.03 0.02

0.01

0

.

,i,_--

0

20

m,- .

40

.

.

60

.

.

80

.

.

100

120

.

,

140

160

180

200

Time (mins) Figure 3: Second order plot of uncatalyzed and dibutyl tin oxide catalyzed reactions for PBT (50%) & HQDA+TA (50%) at 285 C with the model

9DBTO (0.25 mole%) m DBTO (0.5 mole%) 9DBTO (1.0 mole%) & uncatalyzed 0 -0.5 -1 -1.5 -~

-2 -2.5

-3.5 -4

9 ,

,

i

,

i

,

,

0.00345 0.0035 0.00355 0.0036 0.00365 0.0037 0.00375 0.0038 l I T (K)

Figure. 4"

A r r h e n i u s p l o t s f o r u n c a t a l y z e d a n d d i b u t y l tin o x i d e c a t a l y z e d r e a c t i o n f o r 50 m o l e % P B T

571

Table 1. Table indicating the effect of catalyst type and temperature on the ratconstant and activation energy for PBT ( 50 mole%) : HQDA + TA ( 50 mole% )

Catalyst type

Temperature ( C )

kl

k2

E1

Ez

265

0.03136

0.00424

94.35

100

275

0.18459

0.04532

285

0.6298

0.10079

Zinc acetate

265

0.10337

not found

(---)

(---)

0.25 mole %

275

0.14269

0.00635 39.25

20.19

51.89

(---)

34.11

(---)

48.22

(---)

40.884

36.922

Uncatalyzed

285

0.10362

0.07107

Zinc acetate

265

0.13539

0.00335

0.5 mole%

275

0.14084

0.00425

285

0.47981

0.00609

Zinc acetate

265

0.12109

not found

I mole %

275

0.37561

0.02924

285

0.628

0.3966

DBTO

265

0.04446

not found

0.25 mole %

275

0.09474

0.01406

285

0.13114

not found

DBTO

265

0.04079

not found

0.5 mole%

275

0.14447

0.02804

285

0.18729

0.02776

DBTO

265

0.13093

0.03195

1 mole%

275

0.17568

0.03669

285

0.48601

0.10477

constant for the first stage constant for the second stage. E~ : Activation energy for the first stage. E2 : Activation energy for the second stage. kl

9 rate

k 2 9 rate

(1/concentration (1/concentration ! (k J/mole) (k J/mole)

572

"t

|

9

~

'

S'O.

Figure

I'

1 ~!o.

5.a

"

'

"

'

"'

!

....

) 5~.

"

'

"

'

-'

'!

'

"

'~ ....

200.

DSC heating cycle of tmcatalyzed reaction for PBT (50%)/ (HQDA +TA) (50 mole%) system

ao~. Figure 5.b

DSC cooling cycle ofuncatalyzed reaction for PBT (50%)/ (HQDA +TA) (50 mole%) system

' ..........

~

i'"

QC

573 Characterization of the copolyesters were carried out using differential scanning calorimeter. A typical uncatalyzed runs (Figure 5) is shown. The plot reveal that the polyesters are liquid crystalline in nature.

CONCLUSION Kinetics of a three component (PBT, HQDA and TA) melt polymerization system has been a analyzed. Compared to one and two component systems, the kinetics of this ternary system would be extremely complex due to possibility of several reactions occurring in parallel. However an assumption that HQDA + TA react to produce a dimer and acetic acid was made. This dimer is assumed to react with PBT to form acetic acid and higher oligomers and the product of interest. This step is also assumed to give equimolar amount of acetic acid as in dimerization step. Other processes forming acetic acid is assumed to be slow. A second order kinetic model neatly traces the experimental data. Arrhenius plots reveal that the catalyst has only a marginal role to play in the PBT 50 mole% system. DSC studies reveal that the composition is liquid crystalline in nature.

ACKNOWLEDGEMENT The authors would like to thank the research administration of Kuwait University for the funds provided from Project EC 082 for carrying out this work.

574 REFERENCES

12345678910-

11-

Jackson, W.J. Jr. and Kuhfuss, H.F., J. Polym. Sci. Polym. Chem. Ed., 1976, 14, 2043 Roviello, A. and Sirigu, A., J. Polym. Sci., Polym. Letts. Ed., 1975, 13,455 S.-W. Wong and K.C. Frisch, J. Polym. Sci. Part A Polym. Chem. Ed., 24, 2867 (1986) S.-W. Wong and K.C. Frisch, J. Polym. Sci. Part A, Polym. Chem. Ed., 24, 2877 (1986) C.M. Thompson, S.G. Taylor, and W.W. McGee, J. Polym. Sci. Part A, 28, 333 (1990) K. Ch. Park and S. Ch. Kim, J. Appl. Polym. Sci., 39, 639 (1990) M. Sato, J. Am. Chem. Soc., 82, 3893 (1960) S.C. Chapra, R.P. Canale, Numerical Methods for Engineers, 3 rd Edition, 1998, McGraw-Hill Inc. Mathew, J., Bahulekar, R.V., Ghadage, R.S., Rajan, C.R., Ponrathnam, S. and Prasad, S.D., Macromolecules, 1992, 25, 7338 Mathew, J., Ghadage, R.S., Prasad, S.D. and Ponrathnam, S., Macromolecules, 1994, 27, 4021. Hamb, F.L., J. Polym. Sci., Polym. Chem. Ed., 1972, 10, 3217.


575

Catalytic Oxidation of Sulfite/Bisulfite in a Falling-Film Absorption Column M. Vorbach a, R. Marr a and M. Siebenhofer b alnstitute of Chemical Engineering and Environmental Technology, Graz University of Technology, Inffeldgasse 25, A-8010 Graz, Austria*

bVTUEngineering GmbH, Grottenhofstrasse 3, A-8010 Graz, Austria 1. ABSTRACT The oxidation of sulfite/bisulfite is a topic, still of interest in wet flue gas desulfurization. The catalytic effect of several transition metals on this reaction is well known, but little data is available about the activity of iron. Target of the present project has been to examine the role of Fe 2+ and the role of Fe 2+ 9 sulfite/bisulfite oxidation. combined Wl"th Mn 2+ in The results of the present project indicate, that the role of Fe 2+ is rather limited to an enhancement of the oxygen mass transfer, while acceleration of the oxidation reaction is based on the catalytic support of Mn 2§ Catalytic acceleration is based on the activity of Mn 2§ while Fe 2+ mainly acts as an oxygen-carrier, which enhances the rate of mass transfer of oxygen. The results of this investigation confirm, that acceleration of the S(IV) oxidation at pH = 5 will be observed as long as the FeZ+-content is kept below 3. 10.3 mol/l. Additional Fe z+ will result in a rapid break down of S(IV)-oxidation. Reason for this break down is the oxidation of Fe z+ to Fe 3+, followed by an immediate formation of Fe(OH)3 due to the low solubility product of Fe(OH)3. When increasing the pH-value of the absorption liquor, the concentration limit of Fe z+ can be extended to at least twice the value of the limit, observed at pH = 5. The effect is based on a physico-chemical shift of the ratio of bisulfite to sulfite with increasing pH-value, resulting in an increased rate of oxidation of sulfite. As a consequence, oxidation of the oxygen-carrier Fe E+is retarded. 2. INTRODUCTION Several wet flue gas desulfurization (FGD) processes are based on the absorption of sulfur dioxide in caustic suspension or absorption liquor. The dissolved sulfur dioxide, which we will refer to as S(IV), is neutralized by the additive, commonly lime or limestone [1 ]. In aqueous phase several species of S(IV), hydrated sulfur dioxide, bisulfite and sulfite are formed according to the dissociation properties of S(IV)-oxides in aqueous solution. The distribution among the protonated and non protonated species is determined by the pH-value. The authors are pleased to acknowledgethe helpful assistance of D. Kreuzwieserand T. Kircher.

576 Parallel to dissolution and neutralization, S(IV) is oxidized by residual oxygen, absorbed from flue gas. This oxidation implies the addition of dissolved oxygen to an S(IV) species to produce S(VI). S(IV) + 1 / 2 0 z SO2.H20 HSO~ SO32-

Catalyst. ) S(VI)

(1)

SO]-+H +

This reaction is of great interest in the study of wet flue gas desulfurization processes. In the case of the normally used absorption additive lime, the oxidation of S(IV) is an essential step in quantitatively producing the final product, gypsum (CaSO4- 2H20). Beside detoxication of S(IV) the dewatering properties of gypsum are improved when a high oxidation yield is achieved. The conditions of improved oxidation are realized by bubbling air into the sulfurous suspension downstream the absorber [1 ]. Thus S(IV) oxidation takes place in the spray of the scrubber as well as in the downstream part. It is well known that the rate of S(IV) oxidation is strongly influenced by the pH-value, which determines the sulfite/bisulfite-equilibrium [2,3]. Gmelin [3] reports the maximum oxidation rate to be in the pH-range of 8 - 10. Another major parameter is the presence of catalysts, even in traces. Literature reports the catalytic activity of several transition metals (Co 2§ Cu 2+, Mn 2+, Fe 3+) [ 1,4,5]. Sulfite concentration and oxygen concentration, temperature of operation and presence of inhibitors have to be considered [3]. The reaction type of S(IV) oxidation is supposed to be a radical type reaction, involving several steps, according to general agreement in literature [3,5-9]. Many researchers have studied S(IV) oxidation, yet the published data is sometimes inconsistent or not reproducible. The reason for that is supposed to originate in the sensitivity of oxidation kinetics to operation conditions [1] and catalytic activity or inhibiting trace impurities [6]. Concerning the catalytic effect of the metal ions mentioned above, research has been focused on Co 2§ Mn 2§ or Cu 2§ [ 1,5]. Little and contradictory data is available about the effect of Fe 2+ [2,3,6,7,9]. Target of the present project has been, to examine the role of Fe 2+ and the role of Fe 2+ combined with Mn 2+, as well as the influence of the pH-value in sulfite/bisulfite oxidation. Provided a better understanding of this subject, the performance of S(IV) oxidation in FGD could be optimised by choosing absorption additives with appropriate catalyst concentration or by adding catalyst, to achieve the optimum concentration for maximum oxidation yield, as already suggested by Gmelin [3]. Literature data indicate catalytic activity of Fe 3§ and synergism of iron/manganese. Ulrich et al. [6] investigated, among other metal ions, the effect of iron(m) and manganese(H) under FGD conditions (T = 50~ pH = 4 - 6). Manganese showed almost no catalytic activity for concentration below 10-3 mol/L and steadily increasing enhancement of mass transfer for higher concentrations up to 101 mol/L. Iron(l/I), however, yielded distinct enhancement for the whole concentration range (10 -4 - 10-1 mol/L), with a slightly decreasing tendency for increasing concentration. For catalyst concentration below 3. 10-3 mol/L iron seemed to be more active than manganese (and all the other metal ions under investigation). Increasing the pH-value led to higher enhancement. When adding manganese(H) to sulfite solution with constant iron(Ill) content, a distinct increase in activity could be remarked for Mn 2+ concentrations above equimolarity to iron.

577 Trzepierczynska et al. [7] put forward a kinetic model for iron catalysis of sulfite oxidation, assuming a multistep radical mechanism, and he determined the reaction order of sulfite oxidation. Martin [2] reports a literature survey of the atmospheric chemistry approach to S(IV) oxidation and experimental data, indicating a distinct synergism of iron/manganese. Reda [9] observed a steadily increasing oxidation rate for higher iron(m) content (concentration range 10 4 - 10 1 mol/L) and a synergistic effect of iron/manganese. Gmelin [3] gives a survey of a great amount of data up to 1963, implying several contradictions and inconsistencies. 3. E X P E R I M E N T A L PROCEDURE AND RESULTS

The mass transfer experiments were carried out in a falling film absorption column with a mass transfer area of 0,2 m 2. The column and its calibration have been described in detail elsewhere [10]. Temperature and partial pressure of oxygen were held constant at T = 20~ and po2 = 2,1- 104 Pa, respectively. Na2SO3 (CAS-No. 523-21-7) dissolved in deionized water was used to prepare the absorption liquor, for pH-adjustment NaOH and H2SO4 were added, respectively, FeSO4.7H20 (CAS-No. 7782-63-0) and M n S O a . H 2 0 ( C A S - N o . 10034-96-5) were used as catalysts (all chemicals analytical grade quality). The S(IV) concentration in the feed stream was held constant at 4.2- 10 -2 mol/L. The sulfite content in the absorption liquor was determined in the feed stream and at the outlet of the column by iodometry. The pH-values were measured continuously in the feed and outlet streams, pH-control was located at the absorbent outlet. The enhancement of the oxygen transfer by Fe 2+, which is the predominating species in aqueous solution in the presence of S(IV), was investigated in a first step by varying the pHvalue at a fixed Fe2+-concentration (1.13. 10 -3 mol/L). Figure 1 shows the depletion of S(IV) in the absorption column over the logarithmic mean pH-value (feed stream - outlet stream). Figure 2 displays the formation o f H + as a function of depletion of S(IV). 30

30

~O

25

.o

,__., 25

20

o

20

9 I..,q

~ 15 9-,

*-' 10 r~

r~ 5

4.0

5.0

6.0

7.0

8.0

pH-value (log. mean) Fig. 1. Sulfite depletion versus pH; T -- 2 0 oC,. cs(rv) -_ 4.2" 10-2 mol/L, CFe0I)= 1.13" 10.3 mol/L

10 5

0E+00

!

!

2E-05

4E-05

6E-05

Produced H § [mol/L] Fig. 2. Sulfite depletion versus proton formation during oxidation; T = 20~ cs(rv) = 4.2- 10 2 mol/L, eve(n) = 1.13. 10 -3 mol/L

578

In a further step the Fe2+-content (1.2- 10 -5 - 4.36- 10 -2 mol/L) and the pH-value (4.5 - 6 • 0.1 at column outlet stream) were varied. The results are shown in Fig. 3. 30~

25

~

................

9 pH=4.5

. . .

"~

9

9 pH=5.0

".

9 pH=5.5

i.-.a

=o 20

t/

.~,,~

4-a

10

9 pH=6.0

.

..~,,, ....~.

5 0

.

/ .. ...... A . .

"~ 15

~

-.

9 :"

-~

|. . . . '

,

......

"",.,

"m ,

"-...

,

" ......

":-*:~t

0.0E+00

i ...................

............ :::::::. .................... ..*.a ......................................

5.0E-03

.~,i;;;;i

1.0E-02

1.5 E-02

Fe(II)-Coneentration [tool/L]

Fig. 3. S(IV) depletion as a function of the pH-value and Fe 2+-content; T = 20~ indicated (+ 0.1) in the outlet stream

pH-value, as

In a next step the catalytic effect of Mn 2+ was investigated by varying the Fe2+-content and fixed concentration of Mn 2+ (3- 10 -3 mol/L). These experiments were carried out at pH = 5 + 0.1 in the outlet stream. The results are displayed in Figure 4.

259 Fe(II)

..---.~

20 =9q ~

9 Fe(II) + 0.003 mol/L Mn(II)

15 t

~ ( . ......... ~,,

~

10

m

5

[-.~

-, 9

~

9 i

0E+00

!

i

,

.~.....A.

" ' ~ .............. : . . . : I

i

J

5E-03

i

i

1E-02

Fe(II)-Concentration [mol/L]

Fig. 4. S(IV) depletion, observed for different catalysts: Fe(l:l), and Fe(II) with a constant amount of Mn(II) (3- 10 -3 mol/L), T = 20~ pH = 5 + 0.1 in the outlet stream

579 4. DISCUSSION

Evidently for the operational pH-range between pH = 4.5 and pH = 7, HSO3 is the predominating species, with a growing fraction of SO32- with increasing pH-value. Thus the overall S(W) oxidation is based on two parallel reactions: HSO 3 + 1/20 2 --+ SOl- + H +

pK2o"c = -84.0

(2)

SOl- + 1/2 0 2 --~ SO24-

pK20oc =-94.8

(3)

The equilibrium constants and the E-pH-diagram of Fig. 5 were calculated with the software HSC [11 ]. The chart in Figure 5 points out the shift of the redox potential with increasing pH-value, which explains the increase in oxidation rate observed in Figure 1. At higher pH-value the fraction of SO32- is higher according to the dissociation properties of S(IV). As a consequence oxidation without proton formation according to mechanism 3 and shown in Figure 2, predominates. Eh (Volts) 2.0

I

II

I

I

I

I

'

I

I

I

I

'

I

I

I

I

I

1.5 1.0

....

SO4(-_2a_)

0.5 0.0 -0.5 -1.0 HSO3(-a) SO3(-2a)

-1.5 J

-2.0 0

2

I

4

I

I

6

I

I

8

I

I

I

10

12

! 14

pH

Fig. 5. E-pH-Diagram for S(IV)-S(VI), T = 20~ [12] In Figure 3 the rate of oxidation, considering different pH-value and different iron(II)concentration is compared. The increase of the rate of oxidation with increasing pH-value can be explained with the mechanism, mentioned above. The increase of the rate of oxidation with increasing iron(I) concentration is a result of catalyst content and the pH-value. Process control of the mass transfer experiments was based on constant pH-value at the film bottom. Elevated oxygen consumption therefore needed adjustment of the feed pH to a higher value according to the proton formation by oxidation. The decrease of S(IV) oxidation after passing a maximum for increasing iron(R) concentration is reported to be caused by (colloidal) precipitate formation [6]. The higher the oxidation rate the more S(W) is converted in the

580 liquid film, until finally oxidation and therefore deactivation of the catalyst takes place. This interpretation is confirmed in literature [2,3,6]. For higher pH-value the physico-chemical shift of the ratio of bisulfite to sulfite implies that precipitate formation takes place only for higher iron(H) content. The same trend can be observed when investigating the role of Mn(I~ together with Fe(~, as shown in Figure 4. These results indicate, that the role of Fe2+ is rather limited to an enhancement of the oxygen mass transfer, while acceleration of oxidation is based on the catalytic support of Mn 2+. 5. S U M M A R Y The results of the investigation point out the significance of Fe2+-containing caustic additives in flue gas absorption with preferred sulfate formation and the role of the pH-value. Enhancement of the mass transfer of oxygen by Fe 2+ and the catalytic acceleration of the oxidation process by Mn 2+ have been confirmed. As shown by the results the Fe2+-content must not exceed a pH-dependent optimum concentration.

REFERENCES

1. A. Lancia, D. Musmarra, M. Prisciandaro, M. Tammaro, Chem.Eng.Sci. 54, 3019 (1999). 2. L.R. Martin, in: SO2, NO and NO2 Oxidation Mechanisms: Atmospheric Considerations 63-100(1983). 3. Gmelins Handbuch der anorganischen Chemic, Schwefel Teil B - Lieferung 3, Verlag Chemie, Weinheim (1963). 4. V. Linek, V. Vacek, Chem.Eng.Sci. 36, 1747 (1981). 5. W. Pasiuk-Bronikowska, A. Sokolowski, Chem.Eng.Comm. 138, 53 (1995). 6. R.K. Ulrich, G.T. Rochelle, R.E. Prada, Chem.Eng.Sci. 41, 2183 (1986). 7. I. Trzepierczynska, A. Olszowski, Environment Protection Engineering 18, 41 (1992). 8. A. Lancia, D. Musmarra, F. Pepe and M. Prisciandaro, Chemical Engineering Journal 66, 123-129(1997). 9. M. Reda, Wat.Sci.Tech. 20, 45-47 (2000). 10. W. Zapfel, R. Marr, and M. Siebenhofer, Sep. Sci. Technol. 32(1-4), 617 (1997). 11. HSC Chemistry Ver. 3.0, Outokumpu Research Oy, Finland.


581

Influence of fluorination on the kinetics of the HDN of o-toluidine over tungsten sulfide catalysts ex ammonium tetrathiotungstate Mingyong Sun and Roel Prins Laboratory for Technical Chemistry, Swiss Federal Institute of Technology (ETH), 8092 Zurich, Switzerland The kinetics of the hydrodenitrogenation (HDN) of o-toluidine was studied over alumina-supported sulfided tungsten catalysts prepared from tetrathiotungstate. The Langrnuir-Hinshelwood kinetic model was used to fit the data. The kinetic parameters were obtained by varying the initial reactant partial pressure and the reaction temperature. Fluorination of alumina increased the HDN rate constant, decreased the adsorption constant of o-toluidine, and increased the activation energies and the preexponential factors. 1. INTRODUCTION Fluorination of alumina increases the activity of alumina-supported tungsten catalysts for hydrodesulfurization [1] and hydrodenitrogenation (HDN) [1, 2]. These increased activities have been ascribed to changes in the dispersion of the tungsten species in the oxidic precursor of the catalyst [3] and to changes in the morphology of the WS2 crystallites in the sulfided catalyst induced by fluorination [1]. It was also found that fluorination affected the sulfidation process and composition of the sulfided catalyst [4]. Notwithstanding the large number of studies, the origin of the promotional effect of fluorination is still not clear. Investigation of the influence of fluorination on reaction kinetics provides another approach to solve this question. In heterogeneous reactions, reactants must first adsorb on the active sites before chemical reactions can occur. According to the Langmuir-Hinshelwood mechanism, the reaction rate is determined by the intrinsic activity and the number of active sites (rate constant), and the coverage of the reactant on the catalyst surface (adsorption constant and partial pressure of the reactant). An increase in activity may be due to an increase in rate constant and/or to a change in the adsorption constant. Two catalysts with different adsorption constants may exhibit different relative activities when using feeds with different partial pressures of the reactant. A low partial pressure favours the catalyst with a high adsorption constant. The apparent relative activity is sensitive to feed composition. Therefore, to gain insight into the effects of fluorination on the catalyst, it is necessary to perform a kinetic investigation. Kinetic data can only be interpreted meaningfully if the catalyst contains of a single phase. However, classically prepared tungsten catalysts are only partially sulfided and are actually a mixture of W( ) oxysulfides and WS2 [4]. In such a case, it is difficult to isolate

582 the effect of fluorination on the properties of the WS2 active phase. Fully sulfided tungsten catalysts can be obtained by using tetrathiotungstate as the precursor [4]. Therefore, we have in the present study used such catalysts to study the kinetics of the HDN of otoluidine. 2. EXPERIMENTAL The catalysts used in the present study were prepared from ammonium tetrathiotungstate (ATT), containing 10 wt% tungsten. The content of fluorine in the fluorinated catalyst was 1 wt%. Details of the catalyst preparation can be found elsewhere [2]. The HDN reactions were carded out in a continuous-flow microreactor. 0.4 to 1.2 g of the catalyst sample diluted with 8 g SiC was sulfided in situ with a mixture of H2S (10 mol%) and H2 (90 mol%) at 400~ and 1.5 MPa for 4 h. After sulfidation the temperature was cooled to 370~ the pressure was increased to 3.0 MPa, and the liquid feed was introduced to the reactor by means of a high-pressure pump, with n-octane as the solvent and n-heptane as internal standard. Dimethyldisulfide was added to the feed to generate H2S (6 kPa) in the reaction stream. The partial pressure of o-toluidine varied from 1 to 9 kPa. 3. RESULTS AND DISCUSSION 3.1. The influence of initial partial pressure of o-toluidine The conversions of o-toluidine over the F-containing catalyst (ATT/A1203F) at different initial partial pressures of o-toluidine at 370~ and 3.0 MPa are shown in Fig. 1.

0.8

t

1 kPa A~'""'"~ 2 kPa

o ~0.6

4 kPa

~ "2 0.4 O

0.2 0 0

20

40

60

80

weight time [min-g/mol] Fig. 1 Influence of the initial partial pressure of o-toluidine on the conversion of o-toluidine (ATr/A1203F, 3.0 MPa, 370~

100

583 It can be seen that the HDN conversion of o-toluidine is dependent of the initial partial pressure of o-toluidine. At the same weight time, the feed with a higher initial partial pressure gives a lower conversion because of the self-inhibition effect. This indicates that the reaction order of o-toluidine is between 0 and 1. A similar result was obtained for the F-free catalyst (ATT/A1203). The removal of the nitrogen atom from o-toluidine proceeds via two reactions: by direct C-N bond cleavage of toluidine to toluene and ammonia (path 1), and by hydrogenation of o-toluidine to methylcyclohexylamine (path 2), which reacts further to methylcyclohexene, methylcyclohexane and ammonia [2]. Fig. 2 shows the influence of the initial partial pressure of 0-toluidine on the selectivity for the products of these two pathways. Varying the initial partial pressure of o-toluidine in the reactant mixture did not affect the product selectivity. This indicates that either both reactions occur on the same site with the same order of reaction, or that o-toluidine has the same adsorption constant on the sites for path 1 and path 2. 3.2. Kinetic modeling The Langrnuir-Hinshelwood model has been applied in the study of the kinetics of the HDN of o-propylaniline, decahydroquinoline, and quinoline over NiMo(P)/A1203 catalysts by Jian and Prins [5-7]. It has been reported that direct C(sp2)-Nbond cleavage takes place on a different catalytic site than hydrogenation reactions [8]. Nevertheless, Jian and Prins assumed that o-propylaniline has the same adsorption constant on the catalytic sites used in hydrogenation and in C(sp2)-N bond cleavage [5]. The fact that changes in the initial partial pressure of o-toluidine do not affect the ratio of the product of path 1 to the product 0.08 n 4 kPa '~

91 kPa x 9 kPa

0.06

o

=

o

0.04

o l=i

o

0.02

_

o

!

0

0.2

0.4

0.6

0.8

yield of products of path 2 Fig. 2 Influence of the initial partial pressure of o-toluidine on the product selectivities (ATT/A1203F, 3.0 MPa, 370~

584 of path 2 (Fig. 2) shows that this was a good assumption. It allows us to use only one adsorption constant for the two paths in the kinetic modelling. Assuming that the adsorption of nitrogen-free hydrocarbons and ammonia in the reaction stream can be neglected, the HDN reaction rate of o-toluidine (Tol) can be written as dProl dr

m

(kl + kz)KrotPro t 1 +K

(1)

TolPTol

where x is the weight time, kl and k2 are the rate constants for path 1 and path 2 respectively, Kro t is the adsorption constants of o-toluidine, and Prot is the partial pressure of o-toluidine. The sum of the rate constants kl + k2 and the adsorption constant Kro t c a n be determined through non-linear regression of the kinetic data according to Eq. 1. Table 1 gives the values and their standard deviations (in parentheses) of these constants for the ATT/A1203 and ATT/A1203F catalysts. Path 1 and path 2 are parallel reactions, and the ratio of kl to k2 is equal to the ratio of the product yields of the two paths. Thus, the individual rate constants can be obtained from their sum and the ratio. The values and their standard deviations (in parentheses) are also given in Table 1. Table 1 Rate and adsorption constants of the HDN of o-toluidine over WSz/A1203(F) catalysts (370~ 3.0 MPa) Catalyst ATT ATTF kl + k2 (kPa'mol'min-l"g -1)

0.021 (0.001)

0.039 (0.001)

Krot (kPa -1)

1.4 (0.2)

0.9 (0.1)

kl/k2

0.081

0.078

kl (kPa'mol-min-l'g-1)

0.0016 (0.0005)

0.0028 (0.0005)

k2 (kPa'mol'min-l"g -1)

0.019 (0.001)

0.036 (0.001)

Effective rate constant (kPa.mol.min-l.g -1) At partial pressure 1 kPa 9 kPa

0.012 0.0022

0.018 0.0039

Table 1 shows that fluorination increases the reaction rate constants of both paths, but that of path 2 to a greater extent. The adsorption constant decreases upon fluorination. The resulting effective rate constants for o-toluidine conversion, (kl+k2)Krot/(l+KrotProt), are higher over the fluorinated catalyst, but the difference is greater at higher partial pressure.

585 That means that the promotional effect of fluorination is more pronounced for the feed with the higher concentration of o-toluidine. The rate constant in the Langmuir-Hinshelwood equation is the product of the number of active sites and the intrinsic activity of that site. The higher rate constant shows that the number and/or the intrinsic activity of the catalytic sites for the HDN of o-toluidine are higher over the ATT/A1203F catalyst than over the ATT/A1203 catalyst. The lower adsorption constant indicates that the properties of the active sites were modified by the fluorination. The activation energy also reflects the intrinsic properties of the sites, and is independent of the number of sites. This means that it is possible to determine if fluorination modifies the intrinsic properties of the sites by investigating the changes in the activation energies of o-toluidine HDN over the catalysts. The activation energies of the HDN reaction paths 1 and 2 were calculated according to the Arrhenius equation by varying the reaction temperature from 320 to 370~ (Fig. 3). The results are given in Table 2. Table 2 Arrhenius constants in the HDN of o-toluidine at 3.0 MPa, k=A,exp(-E/RT) ATT ATTF Path 1

Path 2

Path 1

Path 2

E (kJ/mol)

101

93

134

115

A ( kPa.mol.min-l.g-1)

3E + 5

2E + 8

9E + 7

10_

7E +5

9 ATT/A1203 9 ATT/A1203F ~

Path 1

_

.

n

t"7"

Path 2 _

_

1.5

I

I

I

I

1.55

1.6

1.65

1.7

1/T [10 3 x K1] Fig.3 Arrhenius plot o f the H D N rate constants o f o-toluidine (3.0 M P a )

586 The activation energies of both path 1 and path 2 are higher over the ATTF catalyst than over the ATT catalyst. Fluorination thus modifies the properties of the active sites for the HDN of o-toluidine. The higher activation energy of the ATTF catalyst indicates that the fluorination exhibits a greater promotional effect at higher temperature. For both catalysts, the activation energy for path 2 (hydrogenation of the aromatic ring) is lower than for path 1 (the C(sp2)-Nbond breaking), which means that high temperature favours the production of aromatics in the HDN of o-toluidine. 4. CONCLUSIONS The HDN of o-toluidine proceeds along two paths, and o-toluidine exhibits the same adsorption constant on the sites for hydrogenation and C(sp2)-N bond breaking. Fluorination increases the rate constants of o-toluidine HDN, decreases the adsorption constant, and increases the activation energies. Fluorination thus modifies the active sites, and promotes the HDN of o-toluidine over the fluorinated WS2/A1203 catalyst. The promotional effect of the fluorination is more significant at high temperature and for a high o-toluidine concentration in the feed. REFERENCES 1. A. Benitez, J. Ramirez, A. Vazquez, D. Acosta, and A. Lopez Agudo, Appl. Catal. A : Gen., 133 (1995) 103. 2. M. Sun and R. Prins, Stud. Surf. Sci. Catal., 127 (1999) 113. 3. R. Lopez Cordero, J. R. Solis, J. V. G. Ramos, A. B. Patricio, and A. Lopez Agudo, Stud. Surf. Scj. Catal., 75 (1993) 1927. 4. M. Sun, Th. Btirgi, R. Cattaneo, and R. Prins, J. Catal. Accepted, (2000). 5. M. Jian, F. Kapteijn, and R. Prins, J. Catal., 168 (1997) 491. 6. M. Jian and R. Prins, Ind. Eng. Chem. Res., 37 (1998) 834. 7. M. Jian and R. Prins, Stud. Surf. Sci. Catal.,113 (1998) 111. 8. M. Jian and R. Prins, Catal. Lett., 50 (1998) 9.

Studies in Surface Science and Catalysis 133 G.F. Froment and K.C. Waugh (Editors) (9 2001 Elsevier Science B.V. All rights reserved.

Kinetics of the hydrodenitrogenation over sulfided NiMo/T-AI203

587

of cyclohexylamine

Fabio Rota and Roel Prins Laboratory for Technical Chemistry, Federal Institute of Technology (ETH), 8092 Zurich, Switzerland The hydrodenitrogenation of cyclohexylamine and the hydrogenation of its product, cyclohexene, were studied over a sulfided NiMo/7-A1203 catalyst. The kinetical networks were studied successively; i.e., each step of the chemical network was studied independently. This enables us to determine the rate and adsorption constants that best fit the kinetic data and which reliably describe the chemistry. A two-site model was used to describe the cyclohexylamine network, one for the removal of ammonia from cyclohexylamine and the other for the hydrogenation of the cyclohexene intermediate. Cyclohexylamine adsorbs stronger than cyclohexene on the hydrogenation site and inhibits its hydrogenation to cyclohexane. 1. INTRODUCTION Hydrodenitrogenation (HDN) networks are usually quite complex because of the large number of reaction steps and catalytic sites involved [1]. The HDN of a nitrogen-containing aromatic heterocycle like quinoline is a good example of such a complicated HDN reaction [2-4]. It proceeds via the hydrogenation of the aromatic heterocycle to 1,2,3,4-tetrahydroquinoline, followed by ring opening and the removal of ammonia by C-N bond breaking, forming o-propylaniline, oPropylaniline is then hydrogenated to propylcyclohexylamine and the C-N bond can open upon the formation of propylcyclohexene [2]. Finally, propylcyclohexene is hydrogenated to propylcyclohexane. All these reactions can occur via different mechanisms on different sites. The number of molecules involved indicates how difficult it can be to accurately calculate the rate and adsorption constants that describe the kinetics of these reactions. Therefore, model reactants are often used in HDN studies. To simplify the quinoline network reaction, we studied the HDN of opropylaniline and o-toluidine, the latter has chemical properties similar to those of o-propylaniline, o-Toluidine has many advantages as a model reactant. Due to its molecular structure, all reactions which take place in an industrial HDN process also occur in the HDN network of o-toluidine; these include hydrogenation of the aromatic ring, C-N bond cleavage (be it C(sp2)-N or C(sp3)N bond cleavage), and alkene hydrogenation [5,6]. However, it is too complicated

588 even for these four reactions to calculate all the adsorption and rate constants accurately. Moreover, the number of different sites on which the different reactions take place may increase with the increasing number of pathways in the network. The large number of sites has a strong influence on the complexity of the Langmuir-Hinshelwood equations [7,8]. The HDN of cyclohexylamine (CHA), in which neither hydrogenation of an aromatic ring nor C(sp2)-N cleavage take place, is more suitable for a kinetic study. Moreover, CHA has chemical properties similar to those of methylcyclohexylamine, which is an important intermediate in the o-toluidine network. The advantage of the CHA network is that the calculation is more precise and that the results obtained can be used to fit a more complicated network such as that of o-toluidine. In the same way, the results of the o-toluidine network can be used for the quinoline network. Therefore, the HDN of cyclohexylamine (CHA) and the reactions of its reaction intermediates cyclohexanethiol (CHT) and cyclohexene (CHE) were studied over sulfided NiMo/7-A120~. This network, with CHT in the middle (Figure 1), can give important information about the substitution mechanism [1,5] and the reaction of the CHT intermediate.

{•,• N

DCHA

CHE

MCPE MCP

kl

CHA

[~ N H 2

. t O CHE

~

k3 k2 ~

Chemical network

[~CH

Kinetical network Figure 1

2. E X P E R I M E N T A L The NiMo/7-A1203 catalyst used in this work contained 8 wt% Mo and 3 wt% Ni and was prepared by successive incipient wetness impregnation of 7-A1203 (CONDEA, pore volume: 0.5 cm3.g -1, specific area: 230 m2.g-1) with an aqueous solution of (NH4)6MoTO24"4H20 (Aldrich) followed by an aqueous solution of Ni(NO3)2-6H20 (Aldrich). The catalyst was dried in air at ambient temperature for 4 h, and then dried in an oven at 393 K for 15 h after each impregnation. Finally, the catalyst was calcined at 773 K for 4 h. The catalyst was crushed and sieved to the desired particle size to avoid diffusion effects on product distribution and conversion [9]. A sample (0.050 g) of the catalyst was diluted with 8 g SiC to achieve plug-flow conditions in the continuous flow fixed bed reactor. The catalyst was sulfided in situ with a mixture of 10% H2S in H2 at 648 K and 1.0 MPa for 4 h. After

589 sulfidation, the pressure was increased to 5.0 MPa, and the liquid reactant was fed to the reactor by means of a high-pressure syringe pump (ISCO 500D). Reactions were performed at 598, 623 and 648 K. A mixture of 10% H2S in H2 was added to keep the H2S partial pressure constant at 20 kPa in all experiments. In the event of a change in the partial pressure of the reactant, the octane (solvent) flow was adapted to maintain the partial pressure of hydrogen, heptane (internal standard) and hydrogen sulfide constant. The reaction products were analysed by on-line gas chromatography using a Varian 3800 GC i n s t r u m e n t equipped with a 30 m DB-5 fused silica capillary column (J & W Scientific, 0.32 mm i.d., 0.25 ~m film thickness). Detection was made with a flame ionisation detector (FID) as well as with a pulsed flame photometric detector (PFPD), which is very useful for detecting small amounts of amine and sulfur compounds. Weight-time was defined as x = Wc / nfeed, where We denotes the catalyst weight and nfeed the total molar flow fed to the reactor. Weight-time (x) was changed by varying the flow rates of the liquid and the gaseous reactant, while keeping their ratio constant. MacroMath Scientist 2.0 software was used for the mathematical modelling of the kinetic parameters. The rate and adsorption constants of CHA and CHE were calculated by fitting the data obtained at four different temperatures (568, 598, 623 and 648 K) and at different initial partial pressures to Langmuir-Hinshelwood equations. A twosite model was used to describe the surface reactions, one site for the reactions of CHA and the other for the reaction of CHE. The usual kinetic approach is to measure the partial pressures of a reactant and its intermediate and final products as a function of time, and to fit the resulting data with model equations. In our experience, this approach often underestimates the dependency of the concentration of the product on the reactant and may, therefore, give results which are very well acceptable from a mathematical but not from a chemical point of view. Thus, we studied each reaction step separately in order to determine its chemistry correctly. Kinetic information on the separate reaction steps is then combined in an overall kinetic scheme. 3. R E S U L T S AND D I S C U S S I O N The HDN of cyclohexylamine (CHA) can be described by two reactions: the elimination of NH3 leading to cyclohexene (CHE) and the nucleophilic substitution by H2S leading to cyclohexanethiol (CHT). The formed CHT can react via hydrogenolysis to cyclohexane (CH) or via elimination to CHE (cf. Figure 1). The reaction of CHT was studied separately so as to monitor these two possible pathways. The CHA reactions can take place at the same site, because studies of methylcyclohexylamine (MCHA) showed that different partial pressures of CHE, NH3 and MCHA always led to the same product distribution. This means t h a t the elimination and substitution sites are the same or at least have the same adsorption constants [6]. Therefore, the elimination and substitution of CHA are assumed to take place on one site only. The hydrogenation of CHE does not take place at the same site as the reactions of

590 CHA because of the different effect of H2S on hydrogenation and C-N bond cleavage in HDN. CHE reactions: The reactions were performed at 10 and 80 kPa of CHE with 20 kPa H2S at 598, 623 and 648 K. CHE hydrogenation gives CH as the main product. Only 2-3 % of CHE isomerized to methylcyclopentene at higher temperature (648 K). A small amount of CHT was detected, which was proportional to the amount of CHE and the CHT/CHE ratio was 0.004, showing that the H2S addition to CHE is minor. The small amounts of CHT and methylcyclopentene formed from CHE were not taken into account in the kinetic calculation. CHA strongly inhibited the hydrogenation of CHE (see the high adsorption constant of CHA, KCHA,CHE, on the site of CHE). CHT reactions: The reactions were performed at 10 and 50 kPa of CHT with 20 k P a H2S at 523, 598, 623 and 648 K. At 598 K, and higher temperatures, and low weight time, the conversion was 100%, showing that the C-S bond breaking is extremely fast. Moreover, CHE converted to CH in a high extend making it impossible to measure the selectivity at low weight time. Therefore, the CHT reaction was performed at the lower temperature of 523 K to study the selectivity as a function of conversion. The consecutive conversion of CHE to CH was still too high; therefore, MCHA was added during the CHT reaction to inhibit the conversion of CHE to CH. The detection of only CH and CHE shows that CHT reacts via two different mechanisms, elimination and hydrogenolysis. The addition of MCHA enabled us to measure the real selectivity of CHE and CH during the CHT reaction. At 523 K, 85% of the CHT reacted to CHE and 15% to CH; at 598 K, 75% of the CHT reacted to CHE and 25% to CH. This means that the C-S hydrogenolysis is not the major route when C-S bond elimination is possible. The selectivity for CH increases with increasing temperature, showing that the C-S bond hydrogenolysis has a higher activation energy t h a n the elimination reaction. For the kinetic calculation of the kinetic network of CHA, CHT was not taken into account because of its low partial pressure, which was always below 2%. CHA reactions: CHA, like CHT, can react via elimination. The direct C-N bond breaking t h a t leads to CH has been studied in the MCHA reaction and was found to be due to a nucleophilic substitution of the NH2 group by an SH group [5]. The C-S bond breaking is easier to achieve t h a n C-N bond breaking because of the lower bond energy of the C-S bond (270 kJ/mol) t h a n of the C-N bond (305 kJ/mol). This leads to a lower conversion of CHA in comparison to CHT under the same conditions. CHA reactions were performed at initial partial pressures of 3, 25 and 40 kPa. At high partial pressure of CHA, the formation of dicyclohexylamine was observed; this was due either to the disproportionation of two CHA molecules to dicyclohexylamine and NH3 or to the reaction of the formed CHE with CHA. The largest amount of this by-product was about 15%

591 and was not taken into account in the kinetic modeling. At 598 K, 80% of CHA reacted to CHE and 20% to CH.

Kinetic modeling: Fitting of the kinetic data with the rate constants (kl + k2 for CHA, k3 for CHE) and adsorption constants (KcHA,CHA for CHA, KCHE,CHE for CHE and KCHA,CHE for adsorption of CHA on the site for hydrogenation of CHE) at different temperatures with a Langmuir-Hinshelwood equation gave good correlations. The calculation was performed using Arrhenius equations for all rate and adsorption constants. The use of adsorption and activation energies enabled us to reduce the number of parameters. The CHE reaction is strongly inhibited during the HDN of CHA, which allows to calculate the adsorption constant KCHA,CHE. The obtained adsorption and activation energies were used to fit all the data points together. The parameters are given in the following table. Standard deviations are given in parentheses.

kl k= k3 KCHA.CHA KCHA,CHE KCHE,CHE

Rate (623 K) (kPa-mol-gl.min 1) 5.9 (0.6) 1.3 (0.2) 51.4 (6)

Adsorption (623 K) (kPa 1)

Activation E AE (kJ. mo1-1) 128 (13) 128 (13) 88 (9)

Adsorption E AQ (kJ- mo1-1)

0.25 (0.05) 1.34 (0.3) 0.02 (0.004)

51 (10) 61 (15)

The equations used are given in the appendix. Figure 2A shows the data points and the fitted equations for the CHA reactions. Figure 2B shows the data points and the fitted equations for the CHE reaction. All obtained parameters were used to fit the whole kinetic scheme for the CHA reaction (25 kPa) at 598 and 623 K (Figure 3). Figure 3 shows that the obtained kinetics parameters give a good fit. The value of KCHA,CHEis very high compared to KCHE,CHE.This means that, during the simultaneous reaction of MCHA and CHE, it is not possible for CHE to react until most of the MCHA has reacted [9]. Figure 2

40

x T= 568K T = 598 K T = 623 K T = 648 K

30 20

PCHA ('kPa)

80

9

9

!'

6O 4O

PCHE (kPa)

10 0

u

n-

0

0 1 2 3 4 5 6 7

A)

Weight time (g'mirgmol)

B)

' 1% I-'

I '

'---

0 1 2 3 4 5 Weight time (g'mirgmol)

592 25

Figure 3 Partial pressure (kPa)

= CHA = CHE = CH

2O 15 10

dashed line = 598 K solid line = 623 K

0

2 4 6 8 10 Weight time (g.min/mol)

12

4. C O N C L U S I O N The approach in which each reaction step is studied in isolation, in order to describe the chemistry involved correctly, and in which the resulting kinetic parameters of each step are then used as kinetic input to describe the overall kinetic scheme, gave good results. The Arrhenius' law was satisfied (between 568 to 648 K) and the complete chemical network was fitted well.

Appendix k 3 9Kcne,cI.i~. "Pcnv.

(kl + k2 )" Kcm,CHA "PCHA rCHA --

rCH - - - - r C H E - -

1 + K CHA,CHA " PCHA

1 + KCHA,CHE "PcHA + KCHE,CHE " PCHE

REFERENCES 1. G. Perot, Catal. Today, 10 (1991) 447. 2. C. N. Satterfield and J. F. Cocchetto, Ind. Eng. Chem. Proc. Des. Dev., 20 (1981) 53. 3. S. Eijsbouts, J. N. M. van Gestel, J. A. R. van Veen, V. H. J. de Beer and R. Prins, J. Catal., 131 (1991)412. 4. M. Jian and R. Prins, Stud. Surf. Sci. Cat., 101 (1996) 87. 5. F. Rota and R. Prins, J. Mol. Catal., to be published. 6. F. Rota and R. Prins, Topics in Catal., 11/12 (2000) 327. 7. M. Jian and R. Prins, Ind. Eng. Chem. Res., 37 (1998) 834. 8. R. Prins, M. Jian and M. Flechsenhar, Polyhedron, 16 (1997) 3235. 9. F. Rota and R. Prins, Stud. Surf. Sci. Catal., 127 (1999) 319.


593

H2-TPR kinetics; Case study on the reduction of a CrOx/Al203 catalyst J.M. Kanervo, A.O.I. Krause Helsinki University of Technology, Department of Chemical Technology, P.O. Box 6100, FIN-02015 HUT, Finland This work deals with the reduction kinetics of a supported chromium oxide catalyst. H2 -TPR measurements were carried out for a CrOx/A1203 catalyst prepared by the atomic layer deposition (ALD). The observed hydrogen consumption data with three different heating rates was examined by various techniques. Several kinetic models were tested using non-linear regression analysis. Nucleation/nuclei growth models were found promising to explain the observed reduction kinetics. These models were considered more closely and re-derived with time-dependent rate constants. The revised two-dimensional nuclei growth model with the activation energy of 92 kJ/mol best described the measured kinetics. 1. INTRODUCTION Temperature programmed reduction (TPR) is a convenient technique to characterise metal oxide catalysts. Generally, TPR provides information on the influence of support materials, pre-treatment procedures and the influence of metal additives on the catalyst reducibility. The TPR technique is intrinsically quantitative and also produces kinetic information. Hurst et al. [ 1] reviewed in 1982 the thermodynamics, kinetics and mechanisms of reduction thoroughly with illustrative examples dealing with the reduction of many supported and unsupported oxides. In literature there are two, in principle, different techniques to determine kinetic parameters from TPR experiments. One requires TPR data collected with different heating rates and utilises only one point from each TPR curve, and the other is based on computer simulated nonlinear regression and exploits the whole experimental TPR-curve/curves. The kinetic modelling of reduction of metal oxides involves gas-solid reaction with possible topochemical i.e. interface advancement features. The reduction degree of solid material, a, is typically modelled in with the equation type [2]: da

= k(T)f(a)f~ (CH~ , Cg, o)

(1)

where k(T) is the temperature dependent rate coefficient, usually of the Arrhenius form, f(a) is a function with respect to reducible solid and fi(C) is a function of gas phase concentrations. Some typicalf(a) functions are represented in Table 1. The diffusion models were excluded due to the nature of the studied system. The models of type (1) can be used in nonlinear regression analysis. The conventional techniques to extract reduction parameters from TPR data, originally developed for other thermal analysis techniques (DTA, TPD, etc), are also based to general form (1).

594 Table 1 Functionsf(a) for typical reduction kinetic models Random nucleation (l-a) (l_a) 1/2 Contracting area (1.a) 2/3 Contracting volume 2-dimensional nuclei growth 2(1-a)(-ln(1-a)) 1/2 (2D Avrami-Erofeev) 3-dimensional nuclei growth (3D Avrami-Erofeev) 3(1-ct)(-ln(1-ct)) 2/3 Light alkenes are important intermediates for petroleum and chemical industry. They are increasingly needed for the production of transportation fuel components with a high octane rating, such as oxygenates and alkylates. The catalytic dehydrogenation is a selective way for the industrial production of light alkenes. The dehydrogenation of light alkanes (C3-C5) is commercially carried out with catalysts based either on supported chromium or platinum [3]. Silica supported chromium oxide (Phillips-type) catalysts have for long been used in ethene polymerisation. Chromium catalysts have been intensively studied over the years, as recently reviewed by Weckhuysen and Schoonheydt [4,5]. The identification of active sites, the influence of preparation and pre-treatment on chromium oxide species and on the catalytic activity/selectivity, still rise substantial interest. After the preparation, CrOx/A1203 -catalysts contain chromium in the oxidation states of Cr+6 and Cr+3 (XPS) [6]. Cr+3 is considered to be the catalytically active site in dehydrogenation [6] and the catalysts are reduced prior to the reaction. The aim of this study was to develop a kinetic model for the reduction of a supported chromia catalyst. The reduction rate data was examined by various methods and typical kinetic models were tested. Based on the information obtained, the nucleation/nuclei growth models were investigated more closely and derived in a revised form. 2. EXPERIMENTAL The catalyst was prepared by atomic layer deposition (ALD), a technique for preparing highly dispersed supported catalysts. The catalyst sample contained 7.5 wt-% chromium. Amount of Cr+6 species was 2.9 wt-% according to UV-VIS spectrophotometry. The XRD analysis showed no crystalline Cr203 phases on the catalyst. H2-TPR measurements were performed with Altamira Instruments AMI-100 catalyst characterisation system. The catalyst samples (30 mg) were stabilised prior to the reduction by heating up to 590 ~ under the flow of 5.0 % O2/Ar. TPR was performed at heating rates of 6, 11, 17 ~ up to 600 ~ under the flow of 11.2 % H2/Ar (50 cm3/min). The hydrogen consumption was monitored using thermal conductivity detector. The selection of experimental conditions was in agreement with the criterion developed by Malet and Caballero [7]. Instantaneous maximum conversion of hydrogen was less than 5.3 %. Thus the differential reactor assumption was valid. The intraparticle diffusion was considered by Weisz-Prater criterion, the value of which indicated that the observed reaction rate was free of intraparticle mass transfer resistance.

595 3. RESULTS AND DISCUSSION The experimental reduction data of the CrOx/A1203 exhibited clearly a single-peak behaviour. The comparison of the total hydrogen consumption and the analysed amount of reducible chromia strongly supported the presumed reduction stoichiometry from Cr+6 to Cr+3. This was interpreted that only one clearly dominant reduction process was taking place. The apparent activation energy (E) was extracted using the Kissinger method [8] based on the shit~ of the temperature (Tm) of the rate maximum as a function of heating rate (fl). Plotting ln(Tm2/,6) versus 1/(Tm) resulted in a straight line (correlation coefficient > 0.99). According to the assumptions of the Kissinger analysis the slope of this plot equals E/R, yielding in this case the activation energy of 87 kJ/mol. Another related technique that utilises different heating rates to extract activation energy, is the Friedman's [9] method of constant conversion. In this technique ln(da/dt) is plotted against 1/T at fixed conversions, and the slope gives -E/R. The Friedman analysis gave the activation energy values around 90 kJ/mol for the reduction degrees below 0.6 and then higher (100-115 kJ/mol) values for the higher reduction degrees. The results of the Kissinger and the Friedman analyses agreed reasonably well, which is explained by the fact that, even though no specific mechanistic hypothesis (f(a)) is required, the both techniques assume similar general form (1) for the reduction rate. The results of Friedman analysis showed however, that the reduction process was not satisfactorily explained by any expression of type (1) with the same energetics for the reduction degrees below and above 0.6. Kinetic modelling of the TPR pattems was then performed with non-linear regression analysis. Various typical reduction mechanisms (shown in Table 1) like the random nucleation law, the phase boundary controlled reactions (contracting area or volume) and the two- and three-dimensional nuclei growth laws (also known as Avrami-Erofeev models) were tested [2]. When the parameter estimation was carried out based on a single TPR data, all the above mentioned models described the observations equally well. The activation energy estimates varied in a wide range from one model to another. Obviously, a single TPR pattern was no basis for the model discrimination. When fitting the model simultaneously to experimental data sets with three different heating rates, only the nuclei growth models were able to satisfactorily explain the observed TPR pattems. Discrimination between the two- and three-dimensional Avrami-Erofeev models was not, however, clear. A better description for the data was obtained by fitting a 2.4-dimensional generalised Avrami-Erofeev type model. The apparent activation energy estimate became 86 kJ/mol using this model. This value is close to the value received by applying the Kissinger peak method. The conventional Avrami-Erofeev models described the observed data quite satisfactorily. The two-dimensional nuclei growth model is intuitively acceptable due to the generally assumed location of Cr6+ on the catalyst surface. This is not the case with the threedimensional nuclei growth model and the empirical dimensionality 2.4 is hardly conceivable. The nucleation/ nuclei growth models, originally formulated by Avrami [10] for other purposes than TPR, were chosen for a closer consideration. The physical background of the model includes nucleation, which is a dynamic process initiating the phase change, and the actual conversion process by the growth of nuclei. Characteristic to this approach is that possible overlap of the growing nuclei and the ingestion of germ nuclei before activated, are taken into account. The commonly applied forms of the nuclei growth models for TPR were omitted and models were re-derived from first principles by taking into account thewell-

596 defined temperature rise and the consequent time dependence of the chemical rate constants. A general dynamic equation for the interface advancement reaction is: - ln(1 - a(t)) =

i '

E Cu~ (t)A exp R (flr + To)

dN

dy

0

(2)

t=y

This equation is a convolution integral that is composed of a law of nucleation:

and a law of nuclei growth by chemical conversion:

/s'

CH2(t)A exp

-~

t=y

/R(pr+ro)I/" dr

N denotes the number of active (growing) nuclei. The time y represents the time the nucleus got activated. The exponent m gives the dimension of nuclei growth. The law of nucleation can be postulated in various ways, such as unimolecular decay law. The left-hand side of the equation origins from Avrami's treatment for the nuclei overlap. It gives the relation between the extended rate of conversion and the true rate of conversion. The pre-exponential coefficient includes several constants grouped together. It was found out that the form (2) could not be simplified to form (1) when the rate coefficient is time dependent, as is the case during the temperature programming. The form (2) gives a possibility to test the nucleation/nuclei growth mechanism and it includes correctly the dynamics that were originally intended. In theory, with present computing capacity, it gives a straightforward opportunity to estimate separate rate parameters for the nucleation process and for the nuclei growth, if identifiable. The computations required to solve (2) were established in the MATLAB | environment. The model (2) was solved by trapezoidal type integration for a and the solutions were used in non-linear regression. The optimisation of the least squares' criterion was performed by Nelder-Mead simplex type search method. The usual temperature mean-centring was done for the rate parameters. The re-derived nuclei growth model (2) was able to describe the experimental data (rms > k2, the relations (3) and (4) are written: [I] = Xo exp (- k 2 t) [C7H12N2] = x0 [ 1 - e x p ( - k 2 t)]

(5)

In the relation (5), the 3,4-diazabicyclo[4.3.0]non-2-ene appears as the product of a first order reaction of rate constant k 2. The variation [C7H12N2] = f(t) was studied at pH =13 and T = 25~

for a constant content

of NHzC1 (10 x 10 -3 M) and increasing concentrations of C7HlzNNH2 (30 to 600 x 10 -3 M). For [C7H12NNH2]

600 x 10 -3 M, the curves are superimposed and tend to an

equation of the type : [C7H12N2]/[CvH12N2]oo = 1 - exp (-kap p t)

(6)

One can consider that (1) is quai-instantaneous with respect to (2) and (5) can be identified to (6) by considering 9kapp = k 2 and zoo = Xo. Consequently, the rate constant k 2 was determined directly from the slope of the curve : Ln 1/(1 - y/x0) = k2 t At T = 25~

and pH = 13, k 2 = 2.81 x 10 -4 S "1. The influence of the pH on the second step

(2) was studied for 0.1 to 0.6 M NaOH and initial contents of NH2C1 and CTH12NNH2 equal to 10 x 10 -3 and 300 x 10 -3 M (T = 25~ The formation of 3,4-diazabicyclo[4.3.0]non-2-ene increases according to pH and the rate constant k 2 verifies the relation : k 2 = k'2 [OH-], where k'2 = 2.81 x 10 -3 M -1 s -1

609 The influence of the temperature was studied for an interval ranging between 15 and 45~ (pH 13). The enthalpy and entropy of activation are the following" AH~# = 86.7 kJ tool -1, AS~ # = - 22 J mo1-1 K -1. The low value of entropy of activation confirms a first order mechanism and the rate constant k 2 can be expressed by the relation (E 2 in kcal mo1-1) 9 k 2 = 1.18 x 1012 exp (- 21.35/RT) s -1

EXPERIMENTAL It consists of two thermostated vessels, one on top of the other and joined by a conical fitting. Each one contains one of the two reactants. The lower reactor (200 mL) contains a magnetic stirrer, and has inlets to allow the measurement of pH and temperature, influx of circulating nitrogen, and removal of aliquots for analysis. The upper cylindrical vessel (100 mL) is blocked at its base by a solid machined stopper (17 mm i.d.) fastened to a control rod. This setup allows the rapid introduction of the ampoule contents into the reactor, and therefore a precise definition of the initial time. A slightly reduced pressure is maintained throughout the experiment and the temperature is defined to + 0.1 ~ N-amino-3-azabicyclo[3.3.0]octane solution, adjusted to the desired pH, was introduced into the reactor. While the thermal equilibrium was reached, a titrated monochloramine solution was prepared by reacting an aqueous NH3-NH4C1 solution with sodium hypochlorite as previously described [5]. Chloramine shows a UV absorption in water at ~, = 243 nm (e = 458 M -1 cm'l). It was analyzed by ultraviolet and HPLC. The hydrazine and the hydrazone are determined by gas chromatography according to an original method because of the presence of chloramine [6]. 3,4-diazabicyclo[4.3.0]non-2-ene presents a strong absorption in ultraviolet spectrum (~. = 229 nm, e = 2685 L mo1-1 cm -1) and its analyzed at its maximum wavelength.

BIBLIOGRAPHIE 1. M. Elkhatib, Th6se de doctorat ~s Sciences n ~ 89-94, Avril 1994, Universit6 Lyon I. 2. H. Delalu, M. Elkhatib, A. Marchand, Monatshefte far Chemie, 125 (1994) 1113. 3. M. Anbar, G. Yagil, J. Am. Chem. Soc., 84 (1962) 1797. 4. M. Elkhatib, A. Marchand, J.J. Counioux, H. Delalu, Int. J. Chem. Kinet., 27 (1995) 757. 5. H. Delalu et al., 1977, French patent n ~ 77 07584. 6. H. Delalu, A. Marchand, Analusis, 17 (1989) 205.


Studies in Surface Science and Catalysis 133 G.F. Froment and K.C. Waugh (Editors) ~.c~2001 Elsevier Science B.V. All rights reserved.

611

An easy methodology for estimating kinetic constants in complex kinetic models J. Ancheyta-Jufirez * a,b, R. Sotelo-Boyfis

a,b

a Instituto Mexicano del Petr61eo, Eje Central Lfizaro Cfirdenas 152, Mrxico 07730 D.F., MEXICO b Instituto Politrcnico Nacional, ESIQIE-IPN, UPALM, Mrxico 07738 D.F., MEXICO

An easy method to estimate rate constants in complex kinetic models is proposed. This method reduces the number of parameters to be estimated simultaneously. A 5-lump kinetic model for the catalytic cracking process was selected in order to apply the proposed methodology. Experimental data obtained in a MAT using three gas oils and a commercial equilibrium catalyst unit were used to evaluate the rate constants. 1. INTRODUCTION For some complex reactions which involve small number of products, path reactions and kinetic constants, there are numbers of books that report analytical solutions or methodologies to solve the resulting rate reaction differential equations [ 1-3]. However, in many industrial processes a large number of chemical reactions occurs simultaneously. In petroleum refining operations dealing with feeds containing hundreds of components (i.e. gas oil catalytic cracking, naphtha catalytic reforming, middle distillate hydrodesulfurization), where the complete analysis is a problem, the number of reactions becomes formidable and the reaction network may also become very complicated, so that components of the feed can be lumped into a small number of groups. This lumping technique has been widely used where the evolution of the lump is practically identical to that of the sum of its individual components. For parameter estimation in kinetic models derived from these complex reaction networks a non linear regression analysis (NLRA) is commonly used. Whereas, severe convergence problems are likely when the number of parameters to be estimated is large. The initializing of the kinetic parameter values is another problem frequently found in NLRA that may converge to local minimum of the objective function usually given by the sum of square errors between experimental and predicted product yields. In order to solve these problems, in this work we propose an easy sequential method to estimate kinetic constants in complex models that considerably reduces the number of parameter to be estimated simultaneously. The method is applied for the catalytic cracking reactions and the kinetic parameters of a five-lump model are determined by using successively NLRA with various 3- and 4-lump kinetic models. *Author to whom correspondence should be addressed., FAX. (52-5) 333-84-29, E-mail:[email protected]

612 2. THE M E T H O D

In order to estimate the parameters of a complex kinetic model by using the proposed method, it is necessary first to divide the original kinetic model in various models with less number of products, by lumping some of them in order to determine some kinetic constants. Once the parameters of these models are known, they can be correlated to evaluate the kinetic constants of the original model, because some path reactions in these lumped models will be the same than the original one, and hence the kinetic constants in these path reactions will have also similar values. The 5-lump kinetic model [4] used in this work (Fig. 1) could be divided in three models each one with three lumps (Fig. 2).

~-c [ A

kA-B

kB.c

,,- B

kA_z

" kB.D ~,/

/

~ C

kc_o

A: Gas oil B: Gasoline D: Dry gas E: Coke

E D Fig. 1.5-lump kinetic model for FCC process [4]

A~B~C+D+E4--A A---~ B + C---~ D + E ~---A A---~ B + C + D---~ E 4-- A

Fig. 2. FCC kinetic models with 3-lumps In order to determine the other parameters, the original 5-lump model (Fig. 1) could be divided in other three models, in similar way as in Figure 2, each one with 4 lumps. In this case some kinetic constants, which were calculated with the previous models, can be determined again. 3. RESULTS AND DISCUSSION In order to apply the proposed sequential method for kinetic parameter estimation in the FCC process three feeds were used for MAT experiments: GO-l, a typical FCC feedstock, GO-2, Heavy Vacuum Gas Oil and, GO-3, the typical FCC feedstock plus 5 vol% atmospheric residuum. Characterization of theses feeds were presented in a previous work [5].

613 The following mass balance, solved with a Runge Kutta Method, was used to evaluate the product yields from a set of kinetic constants for a given kinetic model and for each feedstock and temperature:

yill ci

-~z - WHS-----~

(r~)

(1)

The minimization of the objective function, based on the sum of square errors between experimental and calculated product yields, was applied to find the best set of kinetic parameters. This objective function was solved using the least squares criterion with a nonlinear regression procedure based on Marquardt's algorithm [6]. The sequential methodology previously described was followed for kinetic constants estimation in order to decrease the number of parameters estimated simultaneously and to avoid convergence problems. The mass balance given in equation (1), was solved numerically with the following boundary condition: at z=O, yA=l, and ys=yc=yo=ye=O, in order to calculate gas oil, gasoline, LPG, dry gas and coke yields for each feedstock.. The kinetic parameters obtained by following the proposed methodology for each feedstock are presented in Table 1. Values of kinetic parameter for each model are presented in Ref. [6]. Average values were calculated for those kinetic constants evaluated with more than one model. The kinetic constants for gasoline cracking (kB-c, kB-o and kB-F~)show that gasoline gives mainly LPG and dry gas, while the coke is produced only by cracking of gas oil, since the gasoline to coke kinetic constant (kB-F.) was many orders of magnitude smaller than the others. Apparent activation energies for each reaction lump involved, which were quite similar for the three feedstocks, are within the range of those reported in the literature (5-20 Kcal/mol). Table 1. Kinetic parameters of the 5-lump model for each feedstock at 500~ Reaction GO-1 GO-2 kA-8 (wt% 1 hr l ) A~ B 3370.60 3171.00 kA_c(wt% -1 hr1) A~ C 510.88 491.74 kA-D(wt% -1 hr 1) A~ D 10.76 75.30 kA-e(wt% -1 hr-~) A~ E 390.03 442.96 kB-c(hr-~) B --~ C 181.80 154.98 ks-D (hr -1) B~ D 20.89 25.46 kB-~(hr l ) B~ E 0.750 0.915 kc-D(hr1) C~ D 286.58 323.83 ka(hr -~) 1481.85 1785.00

GO-3 2907.18 477.02 86.06 540.67 101.33 29.38 1.060 353.92 1905.24

Figure 3 shows a comparison between experimental and predicted gasoline yields as a function of gas oil conversion in the range of 480-520~ for GO-2 feedstock. Conversion is defined as the sum of gasoline, LPG, dry gas and coke yields. The effect of the type of feedstock on gasoline yield at 500~ is shown in Figure 4.

614 It can be observed from these figures that the application of the proposed sequential method to estimate the rate constants of the 5-lump kinetic model predicts sufficiently well the experimental data with average deviations less than 2%.

0.60

0.60 480~

~: 0.55-

~ 0.55

-,~ 9~ 0.50 o

raO

0.45

0.65

!

!

!

!

0.70

0.75

0.80

0.85

0.45 0.90

Conversion, wt. fi'ac. Fig. 3. Gasoline yield prediction at different temperatures

,

0.60

0.70

0.80

0.90

Conversion, wt. fi'ac. Fig. 4. Gasoline yield prediction for different feedstocks

4. CONCLUSIONS An easy method for estimating parameters in complex kinetic models is presented. This method reduces the number of parameters estimated simultaneously, and hence the converge problems are also reduced. A MAT reactor was used to crack three industrial feedstocks over a commercial equilibrium catalyst in the range of reaction temperature of 480-500~ and WHSV of 6-48 h -~. The experimental data were utilized to evaluate the kinetic parameters of a 5-lump model for catalytic cracking process, which includes the unconverted gas oil, gasoline, LPG, dry gas and coke. Product yields calculated with the proposed methodology show a good agreement with experimental data with average deviations less than 2%. REFERENCES

1. Moore, J.W.; Pearson, R.G. Kinetics and Mechanism. John Wiley & Sons, 1981. 2. J.J. Carberry, Chemical and Catalytic Reaction Engineering. Mc.Graw Hill, 1976. 3. G.F. Froment and K.B. Bischoff, Chemical Reactor Analysis and Design, John Wiley & Sons, 2nd Ed., 1990. 4. J. Ancheyta, F. Lrpez, E. Aguilar, App. Catal. A. 177 (1999) 227. 5. J. Ancheyta, R. Sotelo. Energy & Fuels 14 (2000) 1226-1231. 6. D.W. Marquardt, J. Soc. Ind. Appl. Math. 2 (1963) 431.


615

Experimental validation of a kinetic model for naphtha reforming J. Ancheyta-Jufirez * a,b, E. Villafuerte-Macias

a,b

a Instituto Mexicano del Petr61eo, Eje Central Lfizaro Cfirdenas 152, Mrxico 07730 D.F., MEXICO b Instituto Politrcnico Nacional, ESIQIE-IPN, UPALM, Mrxico 07738 D.F., MEXICO

A kinetic model for the naphtha catalytic reforming process, which utilizes lumped mathematical representation of the reactions that take place, is presented. The reaction are written in terms of isomers of the same nature, which range from 1 to 11 atoms of carbon for paraffins, and from 6 to 11 carbon atoms for naphthenes and aromatics. The kinetic parameters values were estimated using experimental information obtained in a fixed-bed pilot plant. The pilot reactor was loaded with different amounts of catalyst in order to simulate a series of three reforming reactors. The reformate composition calculated with the proposed model agrees very well with experimental information. 1. INTRODUCTION Catalytic reforming of straight run naphthas is a very important process for octane improvement and production of aromatic feedstocks for petrochemical industries. Generally, the reforming is carded out in three or four fixed bed reactors which operate adiabatically at temperatures between 450 and 520~ total pressures between 10 and 35 atm and molar hydrogen-to-hydrocarbon ratios between 3 and 8. The feed to the reforming process is a hydrodesulfurized naphtha cut, composed of normal and branched paraffins, five and six membered ring naphthenes and single ring aromatics. The major reactions in the first reactor are endothermic and very fast. As the feedstock passes through the reactors, the reactions become less endothermic and the temperature differential across them decreases. Recently there has been a renewed interest in the reforming process, firstly, because reformate is a major source of aromatics in gasoline, and secondly, because of the new legislation of benzene and aromatics content in commercial gasolines. In this sense, refiners have reduced the severity of the industrial reforming plants in order to decrease the amount of aromatics in gasoline, however it adversely affects the reformate octane [ 1]. Various kinetic models to represent catalytic reforming have been reported in the literature, which have different levels of sophistication [2-4]. The kinetic model ofKrane et al [3] is one of the more elaborate models which considers all possible reactions for each individual hydrocarbon. However, the temperature and pressure dependency on the rate constants was not reported. In addition, this model does not consider the formation of the main benzene Author to whomcorrespondence shouldbe addressed., FAX. (52-5) 333-84-29,E-mail:[email protected]

616 precursor (N6: cyclohexane) via isomerization of methylcyclopentane (MCP), and it does not take into account the reaction rates of hydrocarbons with eleven atoms of carbon because only hydrocarbon up to ten atoms of carbon are considered. In the present paper the Krane et al [3] model is extended in order to consider these deficiencies. 2. THE KINETIC MODEL The Krane et al [3] kinetic model utilizes lumped mathematical representation of the reactions that take place. These representations are written in terms of isomers of the same nature. These groups range from 1 to 10 carbon atoms for paraffins, and from 6 to 10 carbon atoms for naphthenes and aromatics. The Krane model includes 53 chemical reactions. The proposed kinetic model has 18 more reactions compared to Krane model [5]. Four more lumps can be directly predicted with this new model, Pll, Nll, All and MCP, and by equilibrium calculations, six iso-paraffin lumps can also be estimated. 3. EXPERIMENTAL The feedstock used in this study was a hydrodesulfurized straight-run naphtha recovered from an industrial naphtha HDS unit. The catalyst used in this investigation was a commercial available Pt-Re reforming sample (Pt: 0.29 wt%, Re: 0.29 wt%). The tests were performed in a fixed-bed pilot plant with H2 recycle. The unit consists of a stainless-steel reactor (internal diameter of 2.5 cm and length of 25 cm), which was operated in isothermal mode by independent temperature control of a three-zone electric furnace. The tests were carried out at 10.5 kg/cm 2 pressure, molar ratio H2/oil of 6.5 and temperatures of 490, 500 and 510~ To simulate a series of three reforming reactors, the pilot reactor was loaded with different amounts of catalyst, 6, 15 and 30 ml keeping the same naphtha flow at a constant value of 102 ml/h in order to have different WHSV, 17.72, 7.09 and 3.54 h -1, respectively. These amounts of catalyst and WHSV were selected in order to have 20% of the total mass of catalyst in the first reactor, 30 % in the second reactor and 50% in the third reactor. 4. RESULTS AND DISCUSSION The seventy one kinetic parameters of the proposed kinetic model were estimated using the experimental information obtained at different temperatures and WHSV. For each reaction step, a kinetic expression was formulated as a function of product yields and kinetic constants. All reactions are presumed to be pseudo-first order with respect to the hydrocarbon. The equations for all the reaction steps are combined into twenty four simultaneous differential equations, which comprise the kinetic model. To evaluate the product yields as a function of reactor length from a set of kinetic constants a pseudo-homogeneous model [6] was used, which was solved with a Runge Kutta Method. The minimization of the objective function, based on the sum of square errors between experimental and calculated yields, was applied to find the best set of kinetic parameters. The least squares criterion is the basis of the objective function. The parameters are determined using Marquardt's algorithm [7]. Most of the initial values of the kinetic parameters were those reported by Krane et al [3]. The best values of all kinetic constants are

617 presented in Table 1. This set of values could not be unique, and a sensibility analysis would be necessary. However, because of the small amount of experimental data available, this was not possible to perform. Table 1. Kinetic constants Reaction step Pll ~ Nl~ Plo ~ N10 P9 ~ N9 P8 ~ N8 P7 ~ N7 P6 ~ N6 P6 ~ MCP Pll ~ Plo+P1 P~I ~ P9+P2 P11 ~ Ps+P3 Pll ~ P7+P4 P11 ~ P6+P5 P10 ~ P9+P1 Plo ~ Ps+P2 P~0 ~ P7+P3 PlO ~ P6+P4 P10 ~ 2P5 P9 ~ P8+P1 P9 ~ P7+P2 P9 ~ P6+P3 P9 ~ P5+P4 P8 --> P7+P1 P8 ~ P6+P2 P8 ~ P s + P 3

of the proposed model for catalytic reforming k Reaction step k Reaction step 0.0356 P8 --->2P4 0.0070 N8 --+ N7+P1 0.0243 P7 ~ P6+P1 0.0027 Nil --~ All 0.0500 P7 ~ P s + P 2 0.0018 Nlo ~ A10 0.0266 P7 --+ P4+P3 0.0043 N9 ~ A9 0.0076 P6 ~ Ps+P1 0.0018 N8 ~ As 0.0000 P6 ~ P a + P 2 0.0016 Nv ~ A7 0.0042 P6 --4 2P3 0.0025 N6 ~ A6 0.0075 P5 ~ P 4 + P l 0.0018 Azl ~ Pll 0.0100 P5 ~ P3+P2 0.0022 Alo ~ Plo 0.0135 N11 ~ Pll 0.0050 A9 ~ P9 0.0135 N10 --4 P10 0.0054 A8 ~ P8 0.0191 N9 --~ P9 0.0054 A7 ~ P'/ 0.0015 N8 ~ P8 0.0025 A11 --4 AI0+P1 0.0054 N7 ~ P7 0.0019 All ~ A9+P2 0.0160 N6 ~ P6 0.0204 A10 ~ A9+P1 0.0095 MCP ~ P6 0.0008 A1 ~ A8+P2 0.0095 NIl ~ Nl0+Pl 0.0134 A10 ~ AT+P3 0.0030 Nil ~ N9+P2 0.0134 A9 ~ As+P1 0.0039 N~ --+ Ns+P3 0.0080 A9 ~ A7+P2 0.0068 N10 ~ N9+P1 0.0134 A8 ~ A7+Pl 0.0058 N~0 ~ N s + P 2 0.0134 A6 --~ N6 0.0019 Nlo ~ N7+P3 0.0080 MCP ~ N6 0.0056 N9 ~ Ns+P1 0.0127 N6 ~ MCP 0.0034 N9 ~ N7+P2 0.0127

k 0.0007 0.6738 0.3198 0.2205 0.2150 0.0788 0.1368 0.0016 0.0016 0.0016 0.0011 0.0016 0.0006 0.0006 0.0006 0.0006 0.0000 0.0005 0.0005 0.0001 0.0015 0.0238 0.0040

The conversion of some selected hydrocarbon types (n-Ps, i-Ps, P6, P7, MCP, N6, N7 and A6) as a function of position in the catalyst bed is shown in Figure 1. The solid lines represent the values calculated with the proposed kinetic model and the symbols the experimental data. It can be observed that the calculated compositions agree very well with experimental information with average deviation less than 3%. It can also be seen from Figure 1 that as the naphtha passes through the catalyst bed, A6 concentration increases. The same behavior was found with all aromatics compounds. The concentrations of N6 and 177 and heavy paraffins (P7-Pll, only P7 is shown in Figure 1) decrease as they undergo conversion. A high rate of conversion of naphthenes was found in the first 30 percent of the catalyst bed. After 60 percent of the catalyst bed, naphthenes concentration approaches a very low steady-state value.

618 The relative rates of naphthenes and paraffins conversion are very different in the first 2030 % of the catalyst bed. While N6 and Nz are almost totally converted in this section, MCP and paraffins have a low conversion. This means that MCP is much less reactive than N6 and Nz. The A6 composition calculated with the proposed kinetic model matches very well with experimental data with a maximum deviation of 2%.

8

~ 7

7

-

i-Ps

I

6

6

5

5

,, -

~4 3 2

2

0

0

1

0

10

20

30

40

50

60

70

, 80

Fractional c a t a l y s t w e i g h t

, 90 100

0

10

20

30

40

50

60

70

80

90

100

Fractional catalyst weigth

Figure 1. Experimental (points) and calculated (lines) reformate composition at 510~ 5. CONCLUSIONS A new kinetic model for naphtha catalytic reforming, which takes into account the most important reactions of this process in terms of isomers of the same nature, has been developed. The groups range from 1 to 11 carbon atoms for paraffins and from 6 to 11 atoms of carbon for naphthenes and aromatics. Paraffins and MCP isomerization reactions are also included. The proposed kinetic model has twenty four differential equation with seventy one kinetic parameters, which were estimated using experimental information obtained in a fixed-bed pilot plant. The calculated reformate composition agrees very well with experimental data with average deviation less than 3%. REFERENCES

1. G.H. Unzelman, Oil and Gas J. 88 (1990) 43. 2. R.B. Smith, Chem. Eng. Prog. 55 (1959) 76-80. 3. H.G. Krane, A.B. Groh, B.D. Shulman, J.H. Sinfeit, Proceedings of the fifth world petroleum congress, (1959) 39-51. 4. J. Henningsen, N.M. Bundgaard, Chem. Eng. 15 (1990) 1073-1087. 5. 6. G.F. Foment, K.B. Bischoff, Chemical Reactor Analysis and Design, John Wiley & Sons, 1990. 7. D.W. Marquardt, J. Soc. Ind. Appl. Math. 2 (1963) 431-441.

Studies in SurfaceScienceand Catalysis133 G.F. Fromentand K.C.Waugh(Editors) 9 2001 ElsevierScienceB.V.All rightsreserved.

619

Kinetic analysis of enzymatic esterification of fatty acids and ethanol V. U. Miguel C*),G. C. Trubiano, G. P&ez, D. O. Borio and A. F. Errazu Planta Piloto de Ingenieria Quimica (UNS-CONICET) Camino La Carrindanga km 7. (8000) Bahia Blanca, Argentina

Abstract The esterification of oleic acid and ethanol with immobilised lipase in a solvent-free media is presented in this work. The reaction has been carried out at a laboratory scale in a stirred tank reactor operating batchwise. The evolution of oleic acid conversion to ester was determined by using two analysis techniques: measurement of the acid index and gas chromatography. The influence of the operating conditions on the equilibrium conversion and conversion-time profiles was analysed. The amount of enzyme, temperature, initial alcohol/oleic acid molar ratio and initial water content were considered. A kinetic model has been developed for the enzymatic esterification of oleic acid and ethanol catalysed by Candida Antartica immobilised lipase. Some kinetic models were considered and a second order reaction kinetic was selected. The parameter adjustment was carried out by using a computer program based on Marquardt algorithms. The results from the simulated experiments by using the proposed model showed a satisfactory agreement with the experimental data. 1. INTRODUCTION Most esters are easily produced by acid or base catalysed condensation of an alcohol and an acid at elevated temperatures. In many cases these temperatures cause a degradation of the reactants and/or the product itself. Resulting consequences are not only a loss of substrate but also a laborious processing to obtain the desired ester in reasonable purity. Modem enzymology has achieved improvements in the development and application of lipases as catalysts. New immobilisation techniques make possible to use enzymes in industrial processes in a similar way to the classical catalysis for heterogeneous reactions [ 1] [2]. Lipases belong to the enzymatic class of hydrolases and catalyse both the hydrolysis and synthesis of esters at the interface between water and insoluble substrate. This interface, consisting of water and oil, is essential for the activity of lipase. Some of the advantages of lipases are the specificity for different types of substrate, high yields and the possibility of working under mild operating conditions. In the case of the immobilised lipase, the interface is obtained by immobilisation of the necessary water together with the lipase. The main advantages of immobilised lipases are the easy recovery and reusability of enzymes and low contamination levels for both the final product and the environment [3] [4]. In the present work, the enzymatic esterification was investigated in order to analyse the influence of the operating conditions on the reaction yield and develop a kinetic model to describe the reaction rate for different operating conditions. c*)E-mail: [email protected]

620 2. MATERIALS AND METHODS 2.1. Reagents Enzyme: The catalyst used was Novozym 435. It is a Candida Antartica lipase, immobilised on a macroporous acrylic resin. The product consists of bead-shaped particles with the diameter in the range of 0.3-0.9 mm. The product was delivered with the water content of 1-2% w/w. It was kindly donated from Novo Nordisk AS (Copenhagen, Denmark). Lipase Activity: The ester synthesis activity of Novozym 435 is expressed in Propyl Laurate Units (PLU/g). The activity of Novozym 435 used for our synthesis was 7000 PLU/g. Chemicals: Alcohol was purchased from Dorwil (Ethanol 96% v/v and Ethanol 99% v/v) Fatty acids (mainly oleic acid) were kindly donated by Materia Hnos (Mar del Plata, Argentina). Standards and other chemicals were provided by Sigma. 2.2. Procedure Experimental equipment: The experimental apparatus consisted of a well mixed glass reactor operating batchwise (BSTR), provided with a heat transfer jacket and with a blade type mixer (Fig. 1). Reactants were introduced in the reactor where they were heated up to the desired temperature. When the reaction temperature was reached, the catalyst was added. Stirring speed was set at 300 rpm. Samples of 0.1 cm 3 were taken at regular time intervals for analysis. Once the reaction was finished, the catalyst was separated by filtration, washed with a solvent and dried. In this way the lipase was recovered for a next operation. Analytical Methods: The reaction progress was followed by the determination of the acid index, and the product distribution was analysed by means of capillary gas chromatography. The product composition was analysed following the AOCS official method Cd 1 l b-91 based on capillary gas chromatography. 10mg of reaction sample were dissolved Figure 1: Experimental set up: in pyridine (containing 0.1mg/ml of n-tetradecane as 1,jacket;2, reaetor;3, pressure internal standard), before being silylated with sensor;4, temperature sensor; 5, BSFTA/TMCS, allowing the resulting mixture to react at stirrer; 6, sampling; 7 enzyme 70~ for 20 minutes. Finally, about 1 ~1 was injected on a loadinz Varian 3700 chromatograph equipped with a split/splitless capillary injector, a flame ionisation detector and a SPB 1 (30m x 0.32mm x 0.25~tm) Supelco column. Helium was used as carrier gas, the oven temperature was programmed from 80~ to 290~ with a constant ramp of 10~ The response time and response factors were calculated by comparison with Sigma reference standards. 3. KINETIC MODEL DEVELOPMENT The esterification of fatty acids could be represented by the following global equation: RCOOH

+ R' OH
RCOOR' + H 20

(1)

621 Different kinetic models have been proposed by others authors [5]. A simplified model based on a reversible model with second-order equations, both the forward and the reversal reaction, was considered as a first approach. The reaction rate equation was expressed as: r . . ,~tA~ . dt.

j

[Ac IA~ ] - g 1eq [es

Etk(,

Iw ~

(2)

An Arrhenius-type equation was proposed for the rate constants with activation energy and the pre-exponential for each reaction rate constant. kl =kl exp(_ E1/RT )

(3)

Keq = Keq ooexp (- AH/RT )

(4)

On the other hand, an ordered reaction mechanism (bi-bi model) was also suggested. The reaction sequence is expressed by the following reaction scheme: A c + E < k~,k-~ >A c E

(5)

A c E + AI < k2,k-2 > A c E A l

(6)

AcEAI
E + Es (8) This model was considered taking into account different controlling steps and keeping the remaining steps at equilibrium. The derived kinetic equations are shown in Table 1.

Table 1 Kinetic Models Model

Controlling step

Resulting kinetic equation

2"~ order

....

r = et k, ([Ac Iaz 1- K1 [Es IW b eq

reversible

Et

Ordered mechanism

Eq. (5)

Ordered mechanism

Eq. (6)

r

Ordered mechanism

Eq. (7)

r -

Ordered mechanism

Eq. (8)

c

l -

r = [AI]K2K3K4 + [ w l E s ] + [ W l E s l A t ] g , , ~ + [ E s l A I ] K 2 K 3

Parameters

k, geq

k = k1K4K2K

4-k = k

k_l

= k2K !

=

1.-[-[EsI[W]K3K 4 -~-[Es ]g 4 q-- l i e ]K,

k_k_2g3g 4

t( i cI ,l i.i sl) x, + x,x4 [Ac ]+ [~cIAl]x~x,x4 + [es ]

~-

k ~

k = k4x, x~x, 4--r = [W]+K~K3Kz[AcIAI]+[AclAIlW]K~K, +K~[AeIW] k = k_ 4 e k[acl*]-k

622 4. RESULTS AND DISCUSSION A set of experiments was designed and carried out in order to examine the desired range of process variables. In the following figures the influences of operating variables are shown. These figures contain both the experimental data (as marked points) and the simulation results obtained from the second order reversible model (as continuous lines). 1.0,

1.0J

9

0.8.

9 .& . . . . . . . 94k,,.,, "" " "

0 0.6

0.6-

As I s 0.4-

o.o

s

. *

/~

_. . . . . . . . . .

..........

5~)

9 ....

9

.9. . . . . . . . . 9

9

11

9. . . . . . ... . . . .

ql.

9seso

100

15()

200

0.4-

. . . . T ' - 45 C N = 10

~1)'" .-*" . . . . . . "e E = 10 %0 // ..-'"" 9 E = 5 % .... /'v;'~"'" * E = 1% ...... 0

I,, ~'"

.

d,

/./

L) 0.2,

s s

, s &s s ." '' ~

9

'~

0.8-

w=4%

Simulated data

P'

Simulated data

2,~:)

300

Time (minutes)

Figure 2: Influence of enzyme content

E = N=I 9 W = 9 W =

k

Simulated data

350

0.0

0

5~1

140

1,50

10% 0 0.2 % - 4% ..... 240

Simulated data Simulated

250

300

data

350

Time (minutes)

Figure 3 9Influence of initial water content

4.1. Influence of the enzyme content The influence of weight of catalyst (E) on the reaction progress can be seen in figure 2. As the value of E increases, the reaction rate rises at the same proportion. It is also demonstrated that the initial reaction rate presents a lineal relationship with the enzyme content. The conversion value reached at equilibrium is not shown for all the experiences but it was not affected by the enzyme weight. 4.2. Influence of the initial water content Water is a product of esterification reaction and a reactant for the hydrolysis (reverse reaction). Therefore, the initial water content (W) affects the equilibrium. For this enzymatic reaction, the equilibrium conversion increases as the initial water concentration decreases (figure 3). The water content affects the reaction in two ways: the reaction rate is decreased along the course of reaction and the final conversion is lower because water concentration shit, s the equilibrium towards the reactants. This effect has also been reported by other authors [6]. 4.3.

Influence of temperature Esterification of oleic acid with ethanol has been carried out at three different temperatures (T) as it is shown in figure 4. As expected, a temperature rise increases the reaction rate. The simulated results from reversible kinetic model were according to this behaviour and from the adjusted parameters, an endothermic reaction was validated. 4.4. Influence of initial molar ratio of alcohol/fatty acid The initial molar ratio of alcohol/fatty acid (N) has an important effect on the equilibrium conversion, which increases with alcohol excess (figure 5). As expected, the increment in the

623 amount of one of the reactants shifts the chemical equilibrium towards the product side. On other hand, alcohol excess has an unfavourable effect on the reaction rate. The reversible kinetic model offered a good fit all experiences at different N, although for N=I (stoichiometric value) higher differences between calculated and experimental values were found.

1.0.. ~. . . . .

~...:....

A.. ~t. - _ ~ ._._ ._,

0.8, "g_ o oo

._o _.r

0.6,

.o

0.4,

i/

r o O

0.2,

o

.. -- ~ " "

w_,%

9

T = 35 C

Simulated Data

9

T = 45 C ....

Simulated Data

T = 55 C ......

Simulated Data

" ~o " 1 ~ o ' 1 ~ o ' ~ o ' 2 ~ o ' ~ o ' ~ ' , o o

o,tl t,"

I I ~:,' =" /1~.',' 8 o~qlt;, ~, oo

E= 10%

.

_. . . . Simul~teO n=a

w=o2%

9 N--10

1~;

~

0

Time (minutes)

50

~

100

9 N = 7,5 ~ - S i m u l a t e d data N=5_''''.'" Simulated data N = 1 Ex erimental data 150

200

250

300

350

"Time (minutes)

Figure 4: Influence of temperature on the conversion

4.5.

-

O.ll :~,

"i t

o o c

'~t

0.8

Figure 5: Influence of initial molar ratio on the conversion

Model parameters adjustment

Kinetic model parameters were adjusted to find the best values fitting experimental data with simulation results. A multiparametric, non-linear regression based on Marquardt algorithm was employed. The agreement of adjustment for each model can be compared in Table 2 by means of the variance values: Table 2 Kinetic parameters

Models (controlling step) 2 nd order model (--) Ordered model (5) Ordered model (6) Ordered model (7) Ordered model (8)

Variance 1.18 x 10 -02 4.93 x 10-~ 1.40 x 10 -~ 1.35 x 10 -~ 1.20 x 10 -~

The lowest value of variance from ordered model (5) indicates the best fit. However, the second order model was selected because the variance is similar and the equation representing the model is simpler. The resulting parameters for this model are shown in Table 3. The positive value of AH obtained from the adjustment is in concordance with an endothermic reversible reaction. According to the variance values, the second order model was selected to represent the reaction rate as shown in the previous Figs. 2-5.

624 Table 3 Kinetic parameters for the second order model

Kinetic parameter kloo E1

Keqoo AH

Fitted value 5.39x10 "~ 3447 4.69 918

5. CONCLUSION The solvent-free esterification of oleic acid and ethanol was carried out using a commercial enzyme as catalyst. The enzymatic esterification seems to be an interesting way to obtain fatty acid ester due to several advantages such as the mild operating conditions, the product quality and energetic saving, among others. The experimental results allowed knowing the effect of reaction variables on the reaction rate and the ester yield. The esterification rate can be described adequately by the selected model and the simulated results showed a good prediction of the reaction rate for the different operating conditions. Therefore, this model and its parameters can be considered valid either for reactor design or for simulation purposes. NOMENCLATURE Ac: Oleic acid AI: Ethanol E: immobilised lipase Es: Ester W: water AH = heat of reaction [cal/mol] [i] = concentration of component i [mol/l] El = activation energy Ofkl [cal/mol] Et = enzyme/substrate ratio [%w enzyme/acid]

k_i = rate constant for equation i (reverse) kloo = pre-exponential factor Ofkl I~q = equilibrium constant ki = rate constant for equation i (forward) K i = desorption/absorption constant for equation i N = initial molar ratio (alcohol/fatty acids) r = reaction rate [mol/(1 goat min)] R = gas constant [cal/mol k] T = temperature [K]

REFERENCES 1. E. Lie, G. Molin, Journal of Chemical Technology and Biotechnology 50 (1991) 549-553 2. L. Mojovick, S. Siler-Marinkovic, G. Kukic, G. Vunjak-Novakovic. Enzymes in Microbial Technology 15 (5) (1993) 438-443 3. L.H. Posorske, Journal of the American Oil Chemists Society 61 (11) (1984) 38-42 4. G. Langrand, C. Triantaphylides, J. Baratti, Biotechnology Letters 10(8) (1988) 549-554 5. T. Garcia, M. Martinez, D. Garcia and J. Aracil, Reactions kinetics and the Development of Catalytic Processes, G.F. Froment and K. C. Waugh (eds.) (1999) 6. M. Leitgeb, Z. Knez, JAOCS Vol.67 (1990) nl 1

Studies in Surface Science and Catalysis 133 G.F. Fromentand K.C. Waugh (Editors) (_c92001 Elsevier Science B.V. All rights reserved.

625

Catalytic Combustion of Methane over Pd/'I,-AI203 in a Monolithic Reactor: Kinetic Study Eduardo L6pez, Carlos Gigola, Daniel O. Borio, Ver6nica Bucal/t* PLAPIQUI (UNS - CONICET), Camino La Carrindanga Km 7. (8000) Bahia Blanca, Argentina. e-mail: elopez@plapiqui, edu.ar

ABSTRACT This paper focuses on the kinetic study of the methane combustion over a Pd/~'-A1203 catalyst deposited on a cordierite monolithic structure. The experiments were performed in an integral reactor, and the operating conditions were chosen to adequately represent the operation of a domestic-scale catalytic heater, i.e. relatively high volumetric flowrates and high methane molar ~actions. The experimental data together with an appropriate model for the integral reactor were used to estimate the intrinsic kinetic parameters for a power law type reaction. For the laboratory catalyst, internal mass transfer limitations become appreciable for temperatures higher than 430~ A heterogeneous reactor model, which includes the evaluation of the methane concentration gradient inside the washcoat, allowed to reproduce with high accuracy the laboratory data. 1. INTRODUCTION Catalytic combustion is being adopted for different industrial and domestic applications. This type of combustion is gradually replacing the homogeneous (conventional) one because it offers attractive features, among others, flameless operation, lower emissions of contaminants and lower operating temperatures. The multichannel monolithic reactors are being widely used to carried out catalytic combustion reactions because operate with low pressure drop, high surface/volume ratio and minimum gas channeling due to the uniformity of the matrix (Hayes and Kolaczkowski, 1997). Natural gas is a common fuel for heat generation, and Pd is a widely used catalyst for methane combustion (Ferrauto et al., 1992; Mouaddib et al., 1992; Zwinkels et al., 1993). Anderson et al. (1961), YU Yao (1980), Baldwin and Burch (1990), Briot and Primet (1991), Hoyos et al. (1993), Ribeiro et al. (1994), among others, have reported information about the kinetics of this reaction. The kinetics parameters usually found have been calculated from experimental data obtained for Pd deposited over A1203 powder or pellets. Recently, Kolaczkowski et al. (1996) reported a kinetic expression for methane catalytic combustion in a monolith reactor. Corresponding author

626 This work focuses on the kinetic study of the methane catalytic combustion in a honeycomb monolith wash-coated with Pd/7-A1203 (homemade). The experimental conditions were chosen to adequately represent the operation of a domestic-scale catalytic heater, i.e. relatively high volumetric flowrates and high methane molar fractions (Lrpez et al., 2000). From experimental data (obtained at conditions of negligible mass transfer resistances) and using a mathematical model for the laboratory reactor, the intrinsic kinetic parameters are calculated for a power law type rate expression. 2. E X P E R I M E N T A L

The monoliths (cordierite, 620000 cells/m 2, l x l m m square cross-section) were cut into cylindrical pieces with a length of 0.075 m and a diameter of 0.008 m. These samples were wash-coated following the guidelines given by Skoglundh et al. (1996) and Trrncrona et al. (1997). The amount of alumina deposited on the monoliths was about 0.38 g (-20% of the total weight). The active metal, Pd, was deposited on the wash-coated monoliths by immersion in an benzene-Pd(AcAc)2 solution until equilibrium was reached. Finally, the samples were calcined in air at 600~ during 2 hr. X-ray diffraction studies were performed to verify the complete transformation of boehmite (raw material to prepare the washcoat) to ),-alumina. A Pd content of 0.024 wt% (Pd mass/total mass) was measured by atomic absorption. An average washcoat thickness of 90 ~tm was estimated by SEM. Besides, a BET surface area of 80 m2/gwashcoatwas measured. The monolith reactor was inserted in a quartz tube 0.6 m long, which was placed inside an electrical furnace. The first 0.25 m of the quartz tube was used to preheat the reactive stream up to the desired inlet temperature. This temperature was maintained by means of a digital controller that governed the furnace power. Two mass flow controllers were used to prepare the reactive stream (chromatographic quality air and 99.3 % methane chemically pure). The feed and reactor exit streams were analyzed by GC, using a packed column. Two to three GC analysis were performed for a given set of operating parameters. Methane and CO2 were measured using FID and TCD, respectively. In all the experiments the mass balance was closed with high accuracy. To measure the feed temperature a type K thermocouple was placed at the reactor entrance. Since the laboratory reactor was found to be non-isothermal non-adiabatic, solid axial temperatures were measured every 0.5 cm by a moveable type K thermocouple inserted in one channel (closed at its entrance to prevent flow into it). Temperature differences up to 12 ~ were found in the solid phase. Experiments were carried out at different feed temperatures, total mass flowrates, and inlet methane concentrations. The operating conditions, selected to represent a domestic scale heater (Lrpez et al., 2000), are summarized in Table 1. Table 1 Range of operating conditions used in the experiments

Inlet Temperature, ~ 310- 500

Total Inlet Flowrate (STP), cm3/min 250- 750

Inlet Methane Fraction, % (balance air) 0.8- 1.6

627 3. M A T H E M A T I C A L M O D E L

A steady-state one-dimensional heterogeneous plug-flow model has been used to represent the behavior of the laboratory reactor. The internal and external mass transfer resistances are taken into account. For different operating conditions, very small Prater numbers were calculated. For this reason, the thermal gradients inside the washcoat were not included in the reactor model (Froment and Bischoff, 1990). The axial solid temperature profiles were measured for each experiment, therefore only the mass balances for gas and solid phases are needed to model the reactor. The equations that represent the system are the following:

Boundary conditions

Mass Balances Gas phase: dz

cA =CAo

(5)

(2)

x=O

dCAs : 0 dx

(6)

(3)

x- w Oe,( ,,x

d

Solid phase: km(CA-CSAs):qLwrA(CSAs,Ts) rl=

z=O

rA(CAs Ts)dx/Lw rA(CSAs,Ts)

Deft ~d2CAs =rA dx 2

(CAs ,Ts )

kmtC -

(7)

(4)

For operating regimes where the internal mass transfer resistances can not be neglected (Model A), equations (1) and (4)-(7) have to be solved simultaneously. For conditions of absence of internal mass gradients the system can be modeled using equations (1), (2) and (5) by doing q = 1 (Model B). The mass transfer coefficient between the gas and the channel walls was evaluated using an expression reported by Hawthorn (1974). The pressure drop in the monolith reactor was neglected for modeling purposes. The physical properties of the gas mixture were assumed to be dependent on temperature and composition. To solve the model A, the solid phase mass balance equation was discretized using second order central finite differences. For both models, the differential-algebraic equations were solved using a Backward-Difference Formula (BDF algorithm). 4. RESULTS AND D I S C U S S I O N

Aiming to obtain an intrinsic kinetic expression, experimental data for inlet temperatures between 310-420~ were used. 99 experimental points, the axial solid temperature profiles (measured for each experiment), the mass balances of model B and a numerical nonlinear regression code were simultaneously employed to estimate the kinetic parameters. A power law type expression was used to model the reaction rate behavior: /

\

rA = k~ exp(-E/RTs)( CsAs~ . The the estimate are presented in Table 2.

kinetic parameters that minimize the standard error of

628 Table 2 Kinetics Parameters for Catalytic Combustion of Methane koo,mol~l-n)/m3O-n)seg E, KJ/gmol n Std. error of est. a 8,281,000 85.8 0.30 2.40 [~"~99 )2/(99 3)] '/2 , aDefined as ~.~j--i(X y,c,lc - xj,exp where x represents methane conversion The estimated activation energy is in reasonable agreement with the values reported in the literature for methane combustion over Pd/~-A1203 (see Table 3). Particularly, the activation energy obtained in the present work is lower than that reported by Kolaczkowski et al. (1996) which is the unique kinetic expression found for a monolith reactor. However, the reaction rates calculated using the kinetic parameters showed in Table 2, for different methane concentrations and temperatures, are lower than those estimated through the expression given by Kolaczkowski et al. (1996). The observed differences can not be explained because the catalyst characterization is not given by the authors. Table 3 Reported activation E [kJ/molg] 91.2 71-84 80-160 74-104 112-13 7 75-90 131 85.8

energy for methane combustion Reference Anderson et al. (1961) Yao, Y. Y. (1980) Baldwin and Burch (1990) Briot and Primet (1991) Hoyos et al. (1993) Ribeiro et al. (1994) Kolaczkowski et al. (1996) This work

Type of catalyst Powder Powder Powder Powder Powder Pellets Monolith Structure Monolith Structure

The supposition of negligible internal mass transfer resistances was validated by solving Model A using the obtained kinetics parameters, and evaluating the effectiveness factor (Eq.3). The Defy values were estimated assuming tortuosity factors reported in the literature (Hayes and Kolaczkowski, 1997), i.e. the effective diffusivity was not fitted. The solution of the heterogeneous Model A indicates that the reactor operates with effectiveness factors between 1 and 0.98 for the temperature range 310 - 420 ~ The mass transfer coefficient was set to values one order of magnitude higher and lower than the value predicted by the Hawthorn's correlation. For all the cases, insignificant changes for the calculated outlet conversion were observed. This fact ensures that for the experimental conditions the external mass transfer resistances are also negligible. Therefore, for the selected experimental conditions (310-420 ~ the process can be considered controlled by the chemical reaction (i.e. the fitted kinetic parameters correspond to the intrinsic ones). Figs. 1 and 2 show experimental and calculated conversions vs. inlet temperature for different methane inlet molar fractions and inlet flowrates, respectively. Fig. 1 shows that the higher reactant inlet concentration the lower is the outlet conversion. This is the expected

629 behavior for reactions with orders below the unity. Fig. 2 indicates that higher flows lead to lower outlet conversions, as consequence of the residence time reduction.

Y~o

60 50

,o 11 0,008

o,o12 o,o~6

40

:2 30

9

Q

Calc

Exp 9

~

Exp

40

~ ........

_S

7 .....

........L-,

30

,C'""

,,0

Calc

:

-

9

......

300

, 340

, 360

, 380

, 400

.... j

......~l~,". o..,,"

10 , 320

' _J

.........3,--" ...... :.=,;;L-j

10 0

/,I

, 420

m ....m ~ L ~ - ' " "

0 300

To,*C

i

!

320

340

|

360 To, *C

i

!

i

380

400

420

Figure 2: Methane conversion for different inlet flowrates as function of inlet temperature

Figure 1: Methane conversion for different inlet methane molar fractions as function of inlet temperature

Good fits to the experimental data were obtained for different inlet methane fractions (see curves in Fig. 1). The curves in Fig. 2 represent the calculated conversions for different flowrates. Although the predicted values can follow the general trend, higher deviations than those of Fig. 1 can be observed. In fact for variations in the inlet flowrate, the mathematical model predicts a higher sensitivity than that observed during the experimentation. Besides the set of experimental points used to estimate the kinetic 100 parameters, 43 additional data were 90 YA,0 Exp obtained for inlet temperatures 80 0,008 9 ranging from 460-500 ~ and the 70 0,012 9 z//... previously selected values of yAo 600~ 50 and Q. Fig. 3 shows the influence x 40 of the inlet temperature on the 30 methane conversion for three 20 different yAo values. The solid lines 10 represent the predictions of Model 0 A, which includes the mass transfer 300 350 400 450 500 resistances. This Model predicts the T O , *C experimental behavior with high Figure 3: Methane conversion for different inlet accuracy even for higher inlet methane molar fractions as a function of inlet temperatures than those used to temperature fit the kinetic parameters. Fig. 3 also includes a dotted line that represents, for yAo=O.O16, the prediction of Model B (r/= 1). As it can be seen, internal mass transfer problems become appreciable for inlet temperatures higher than 430 ~ In fact, effectiveness factors lower than 0.90 are found for T0>430~ '

.....'"

,

i

,

,

,

,

!

630

5.

CONCLUSIONS

For the selected set o f operating conditions, the laboratory monolithic catalyst showed good activity to carry out the combustion o f methane. Using a set o f experimental data (methane conversions up to 60% and solid temperature profiles) together with an appropriate model for the integral reactor, the intrinsic kinetic parameters for a p o w e r law type reaction were obtained. For the laboratory catalyst, internal mass transfer limitations b e c o m e appreciable for temperatures higher than 430~ A heterogeneous reactor model, which includes the evaluation o f the methane concentration gradient inside the washcoat, allowed to reproduce with high accuracy the laboratory data for the whole range o f inlet temperatures. NOMENCLATURE C Calc d Deer E Exp F l~

concentration, gmol/m3 model-calculated channel hydraulic diameter, m effective diffusivity, m2/s activation energy, J/gmol experimental data specific molar flowrate, gmol/m2s preexponential factor, gmol/m2 s km maSStransfer coefficient, rn/s Lw washcoat thickness, m n reaction order Q flowrate, cm3/rnin r reaction rate, gmol/m3 s R universal gas constant, J/gmol K T Temperature, K x washcoat coordinate, m

x x y z

conversion washcoat coordinate, m molar fraction axial coordinate, m

Greek letters effectiveness factor ~,eff effective thermal conductivity, W/m K Subscripts 0 at the reactor inlet (z-0) A methane s solid phase w washcoat Superscripts S surface conditions

REFERENCES 9 9 9 9 9 9 9 9 9 9 9 9 9 9 9 9

R.B. Anderson, K.C. Stein, J.J. Freenan and L.E.J. Hofer, Ind. Eng. Chem. 53 (1961) 809. T.R. Baldwin and R. Burch, App. Catal. A. 66 (1990) 337. P. Briot and M. Primet, App. Catal. A, 68 (1991) 301. R.J Ferrauto, M.C. Hobson, T. Kennelly and E.M.Waterman, App. Catal. A, 81 (1992) 227. G.F. Froment and K.B. Bischoff, Chemical Reactor Analysis and Design, Wiley, New York, 1990. R.E. Hayes and S.T. Kolaczkowski, Introduction to Catalytic Combustion, Gordon and Breach Sci. Pub., Amsterdam, 1997. R.D. Hawthorn, AIChE Symp. Ser., 70 (1974) 428. L. Hoyos, H. Praliaud and M Primet, App. Catal. A, 98 (1993) 125. S.T. Kolaczkowski, W.J Thomas, J. Titiloye and D.J Worth, Comb. Sci. Tech., 118 (1996) 79. N. Mouaddib, C. Fermi-Jantou, E. Garbowski and M. Primet, App. Catal A., 87 (1992) 129. E. L6pez, A.F. Errazu, D.O. Borio and V. Bucalfi, Chem. Eng. Sci., 55 (2000) 2143. F.H. Ribeiro, M. Chow and R.A. Dalla Betta, Journal of Catalysis, 146 (1994) 537. M. Skoglundh, H. Johansson, L. Lrwendahl, K. Jansson, L. Dahl and B. Hirschauer, App. Catal. B, 7 (1996) 299. M. Trrncrona, P. Skoglundh, E. Thormahlen, E. Fridell and E. Jobson, App. Catal. B, 14 (1997) 131. Y.Y. Yao, Ind. Eng. Chem. Prod. Res. Dev. 19 (1980) 293. M.F. Zwinkels, S.G. Jaras and P.G. Menon, Catal. Rev.-Sci. Eng., 35 (1993) 319.

Studies in Surface Science and Catalysis 133 G.F. Froment and K.C. Waugh (Editors) (c) 2001 Elsevier Science B.V. All rights reserved.

631

Software Functionality Assessment for Kinetic Reaction Model Development, Model Discrimination, Parameter Estimation and Design of Experiments Rob J. Berger a, Johan Hoorn b, Jan Verstraete c and Jan Willem Verwijs d EUROKIN (http://www.dct.tudelft.nl/eurokin]) aDelft University of Technology, Julianalaan 136, 2628 BL Delft, The Netherlands ([email protected]. nl) DDSM Research, P.O. Box 18, 6160 MD Geleen, The Netherlands cInstitut Franqais du P~trole, CEDI "Ren~ Navarre", Solaize, P.O. Box 3, 69390 Vernaison, France. dDow Benelux N.V., ES/MD, P.O. Box 48, 4530 AA Terneuzen, The Netherlands. An inventory was carried out on the capabilities and user-friendliness of commercially available modeling packages aimed at estimation of (kinetic) parameters and capable to describe two or more dimensional reactor models. Four case studies were developed in order to evaluate these packages in more detail. It appeared that all the packages need improvement in order to become really good and user friendly. Especially the quality of the statistics, the number of useful statistical tools and several user-friendliness aspects need significant improvement. However, discussion of these issues with the software vendors already initiated the developers of the packages to improve the software functionality. This paper is a result of co-operation within Eurokin, a consortium of over 10 European companies and 4 universities. 1. INTRODUCTION A survey, published by Bos et al. [1] indicated the need for improved methods to determine reaction kinetics within the chemical industry. Starting from this survey, the "EUROKIN" consortium has been established in 1998, comprising eleven companies and four universities. Eurokin aims to produce a pre-competitive toolkit for measuring kinetic data and model development. The activities are focused on: 9 Experimental methods to determine reaction kinetics; e.g. investigation of the capabilities of different types of laboratory reactors to measure the kinetics. 9 Development of models for a set of selected laboratory reactor systems, to be used for processing experimental data, and/or the determination of suitable experimental conditions; e.g. assess if the proposed experimental conditions are in the kinetically or in the mass-transfer controlled regime.

632 9 Methods for the determination of kinetic models from experimental data; including model discrimination, parameter estimation and design of experiments. This paper describes the procedure and criteria used to evaluate commercially available software packages for kinetic modeling, and their capabilities for parameter estimation, model discrimination and design of experiments. Also the ease of use and other user-friendliness aspects receive attention. 2. CASE S T U D I E S An initial screening of commercially available modeling packages specifically aimed at estimation of (kinetic) parameters and capable to describe two or more dimensional reactor models yielded about ten packages of interest. In order to evaluate these packages in more detail, four case studies have been developed for testing purposes. The case studies were defined in such a way that most typical parameter estimation problems occurring in the industries and academia were covered with relative small and simple kinetic models. The experimental data for the case studies 1, 2 and 3 were created artificially with addition of realistic noise, the data used case study 4 were taken from a real experiment. The cases are: 1) Estimation of the kinetic parameters in Langmuir - Hinshelwood - Hougen Watson (LHHW) type rate equations, using experimental data sets obtained from a batch and a CSTR reactor. 2) Estimation of the kinetic parameters using experimental data and requiring a mass transfer limited, heterogeneous catalytic liquid phase reaction model, which comprises implicit algebraic and ordinary differential equations. 3) Investigation of the model discrimination and the design-of-experiments capabilities of the various software packages by means of twenty different rate equation models to describe a given set of experimental data. 4) Estimation of the kinetic parameters from experimental data obtained from the start-up of an industrial tubular reactor. This requires dynamic simulation of a tubular reactor, comprising a set of partial differential equations. The four cases were sent to the vendors of the selected software packages, together with a request to use these cases to demonstrate the capabilities of their software packages to the Eurokin consortium. The pre-screening of the software vendors was done using user experiences within the Eurokin consortium and the publicly available information about the software packages (e.g. Internet and journals). Subsequently, the authors used these four cases to test the selected software packages themselves, in order to: 9 assess issues related to ease of use, programming effort required, data handling capabilities, etc., and to 9 exclude as much as possible any differences in e.g. CPU-performance between the packages due to differences in settings of the solution accuracy. This paper will discuss the test cases, the working procedure and evaluation criteria, and the obtained results.

633 3. R E S U L T S A N D D I S C U S S I O N Case study 1

k. 1

k2 ~ C The reaction scheme for which the kinetic k3 ~ kl p a r a m e t e r s had to be estimated is shown in Fig. 1. The pre-exponentials and the a p p a r e n t activation D energies corresponding to the rate coefficients kl, k~ Figure 1: Reaction scheme and k3 had to be estimated from the experimental data sets from four batch reactor and ten CSTR experiments. The initial concentration of reactant A and the t e m p e r a t u r e were varied. The kinetic rate equations of the catalytic reactions can be described by using the following socalled Langmuir-Hinshelwood Hougen-Watson equations. A"

klKA([A] r~' =

~ B

[B])

Keql

1+ KA[A]+KB[B]

r2 =

k2KB[B ]

1+ KA[A]+KB[B]

r3 =

k3K,4[A]

(1-3)

1+ KA[A]+KB[B]

The 'experimental' data and o.7 a typical adequate model fit for ~ ! the first batch experiment is 7 0.6 shown in Fig.2. E o.s Less-experienced research.. ers may start with trying to ~ 0.4 estimate the pre-exponentials ~ 0.3 ko and the activation energies ~ 0.2~ \~ ~ ~ ~/ [D]*! O0 Ea using the standard =~ Arrhenius expression. Our tests o~ 0.1 showed t h a t three out of five 0 software packages had severe 0 100 200 300 400 500 difficulties to obtain a Reaction time Is converged solution for the Fig. 2. Typical fit case 1; symbols: experiparameter estimation run, mental data, lines: simulated curves. which is easily explained by the strong correlation between both parameters. This correlation can be strongly reduced by re-parametrization of the Arrhenius expressions according to:

Original:k=koexp(-R--~Ti Where: k

Re-parametrized: k=krq'exp(;.ET~)

= Reaction rate coefficient

(-Ea) kref- ko exp R'Iref

(4-5)

ko = Pre-exponential factor Tref=Arbitrarilychosenref. temp. within the exp. temp. window

T ~ = Re-parameterized t e m p e r a t u r e [K], obtained from the reactor t e m p e r a t u r e T [K] according to:

1

1

1

T* T TreI

634 After re -p arametrization, all software packages quickly converged to the best possible result. This exercise showed that the optimization algorithms used in some packages need to be improved, as (especially non-frequent) users usually do not like to worry about issues like scaling, equation format, and reparametrization to obtain sufficient performance of the optimizer. Differences between the statistical methods used in the various packages caused that not all the packages gave identical confidence ranges for the estimated parameters. The values obtained varied by more than a factor 2. In order to allow the user to interpret and use the statistical output properly, the software developers are recommended to add more explanation concerning their statistical methods, and/or to add additional methods from which the user can select. In practice, experiments might have been performed in more than one reactor type in order to reduce the probability of systematic experimental errors related to the reactor hardware. This situation was incorporated in this test case by requesting to estimate the parameters, using experimental data from four batch reactor experiments and ten CSTR experiments. It appeared, however, that several packages do not (easily) support estimating parameters from two or more different models simultaneously. It was nice to experience that, after discussing this issue with the software vendors, the developers of the packages made adaptations in order to improve the software functionality to be able to handle this type of problems. Case study 2 In (industrial) practice, many kinetic researchers are not experienced modelers. As a result, they often struggle with models involving sets of implicit algebraic and ordinary differential equations, both for model simulation and parameter estimation. In order to test the capabilities of the various software packages to solve this kind of problems, a case study was defined using a masstransfer limited, heterogeneous catalytic liquid phase reaction model, which comprises implicit algebraic and ordinary differential equations. Since all the selected set of software packages all could solve this problem well, it will not be discussed here in detail. Case study 3 A methanol synthesis problem has been used to investigate the model discrimination and the design-of-experiments capabilities of the various software packages. Twenty different, single rate equations were developed of varying complexity and form. A data set, consisting of 27 data points for the reaction rate at different temperatures and partial pressures of the reactants, was given as well. Duplicate experiments were included in the data set in order to allow an estimate of the variability within the data. The results of this case study clearly showed that none of the packages investigated contained all the advanced statistical tools to perform the requested tasks concerning model selection, model discrimination and experimental design.

635 The ranking of models (best to worse) might be simply based on the sum of (weighted) squares obtained in the parameter estimation run. Preferably, this sum of squares is corrected by dividing this sum of squares by the number of degrees of freedom (calculated as the number of experimental data points minus the number of parameters estimated). However, there are more accurate means available to create a model ranking, e.g. by using F-test values [2] or probability density functions. It appeared that only one of the selected packages was able to do an automatic model ranking. The next step is to perform model selection. Models may be rejected for three different reasons : (i) because the differences between the experimental data and the data calculated with the fitted model are much larger than the measurement error (the model is then qualified as 'inadequate'), (ii) because the fit of the model is significantly worse than an alternative model, and (iii) because one or more parameters in the kinetic model cannot be estimated accurately and independently, which usually indicates that the model contains too many parameters. Although there were large differences between the options available in the packages investigated, none of the packages were capable to perform all these checks. Only one of the packages investigated was able to carry out Design of experiments (DOE), with the goal to facilitate model discrimination. The package selects the most useful experiment out of a set of possible sets of experimental conditions specified by the user. Additional useful features would be the design of a set of experiments instead of only one experiment, and design of experiments to increase the accuracy of specific kinetic parameters. Case study 4 A dynamic tubular reactor model, comprising a set of partial differential equations, has been used to test the computational efficiency and the data handling capabilities of the various software packages. Experimental data of three time-varying model inputs, i.e. the reactor temperature, the fluid velocity and the reactant inlet concentration, are used to estimate the model parameters from experimental data of the reactor temperature measured at several fixed reactor locations as a function of time. This problem was originally published in 1992 [3]. As expected, it appeared that the required CPU time to solve this case varied strongly. This was only to minor part due to differences in time required per reactor simulation, which in particularly depends on the requested accuracy of the integration and the method of integration. Most packages use the backward final differences method for the integration which is a robust but not very fast technique. The differences in CPU time were mainly caused by the different optimization routines used. It appears that so-called indirect search routines, i.e. routines that use the derivatives of the objective function with respect to the parameters to be estimated, are much faster and sufficient robust in comparison with the direct search routines such as the Simplex method [4].

636 User friendliness aspects Generally it is difficult to assess the user-friendliness of a package in an objective way. Most researchers will be already familiar with only one or two packages and will tend to mark these and comparable packages as most userfriendly. The authors tried to avoid this 'trap' by asking someone who was not familiar with any of the selected packages to do the evaluation. The result was that the packages scored differently on different user-friendliness aspects, but that none of the packages was clearly superior on this. The following aspects are considered as most important: 9 Quality of the manual, it should be comprehensible, complete and contain some tutorials. 9 The programming language syntax should be easily understandable. It is also very important that the package contains a large number of useful preprogrammed examples. There should be extensive explanation with these examples. 9 A graphical user interface (GUI) is only useful if it is really user friendly. In some packages the input of experimental data and/or output o f results is difficult and time-consuming. Some GUI's require a (too) large number of mouse clicks to specify and run a problem. 9 Very useful is the availability of a (graphical) output to the screen during the estimation run, so that the user can easily follow the progress of the run. This can vary from simple text output, including the value of the objective function and the parameter values - up to interactive graphs showing the model responses and experimental data for each iteration. 9 Finally, the packages should produce complete and e a s i l y r e a d a b l e report files containing the results, model prediction vs. experimental data, and all the statistically relevant information. 4. C O N C L U S I O N S AND R E C O M M E N D A T I O N The inventory of commercially available software packages aimed at kinetic parameter estimation from experimental data showed that all the packages need improvement in order to become really good and user friendly. Especially the quality of the statistics, the number of useful statistical tools and several userfriendliness aspects need significant improvement.

REFERENCES 1. A.N.R. Bos, L. Lefferts, G.B. Marin, and M.H.G.M. Steijns Appl. Catal. A: General, 160 (1997) 185-190. 2. G.F. Froment and K.B. Bischoff, Chemical reactor analysis and design, 2nd ed., John Wiley & Sons, New York, 1990. 3. J.W. Verwijs, H. van den Berg, and K.R. Westerterp, Startup of an Industrial Adiabatic Tubular Reactor, AIChE J., Vol 38, (1992) 1871-1880. 4. T.F. Edgar and D.M. Himmelblau, Optimization of chemical process, McGraw-Hill, New York, 1989.

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489 333 341 123 375 221 205

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638

Herskowitz M. Hess St. Hoebink J.H.B.J. Hoorn J. Horiuchi J. Hoshi S. Hubaut R. Hudgins R.R.

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STUDIES IN SURFACE SClENCEAND CATALYSIS Advisory Editors: B. Delmon, Universite Catholique de Louvain, Louvain-la-Neuve, Belgium J.T.Yates, University of Pittsburgh, Pittsburgh, PA, U.S.A.

Volume 1

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Volume 6 Volume 7 Volume 8 Volume 9 Volume 10 Volume 11

Volume 12 Volume 13 Volume 14 Volume 15

Preparation of Catalysts I.Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the First International Symposium, Brussels, October 14-17,1975 edited by B. Delmon, RA. Jacobs and G. Poncelet The Control of the Reactivity of Solids. A Critical Survey of the Factors that Influence the Reactivity of Solids, with Special Emphasis on the Control of the Chemical Processes in Relation to Practical Applications by V.V. Boldyrev, M. Bulens and B. Delmon Preparation of Catalysts I1.Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Second International Symposium, Louvain-la-Neuve, September 4-7, 1978 edited by B. Delmon, R Grange, RJaeobs and G. Poncelet Growth and Properties of Metal Clusters. Applications to Catalysis and the Photographic Process. Proceedings of the 32nd International Meeting of the Soci6t~ de Chimie Physique,Villeurbanne, September 24-28, 1979 edited by J. Bourdon Catalysis by Zeolites. Proceedings of an International Symposium, Ecully (Lyon), September 9-11, 1980 edited by B. Imelik, C. Naccache,Y. BenTaarit, J.C.Vedrine, G. Coudurier and H. Praliaud Catalyst Deactivation. Proceedings of an International Symposium, Antwerp, October 13-15,1980 edited by B. Delmon and G.E Froment New Horizons in Catalysis. Proceedings of the 7th International Congress on Catalysis,Tokyo, June 30-July4, 1980. Parts A and B edited by T. Seiyama and K.Tanabe Catalysis by Supported Complexes by Yu.l.Yermakov, B.N. Kuznetsov andV.A. Zakharov Physics of Solid Surfaces. Proceedings of a Symposium, Bechyhe, September 29-October 3,1980 edited by M. L&zniEka Adsorption at the Gas-Solid and Liquid-Solid Interface. Proceedings of an International Symposium, Aix-en-Provence, September 21-23, 1981 edited by J. Rouquerol and K.S.W. Sing Metal-Support and Metal-Additive Effects in Catalysis. Proceedings of an International Symposium, Ecully (Lyon), September 14-16, 1982 edited by B. Imelik, C. Naccache, G. Coudurier, H. Praliaud, R Meriaudeau, R Gallezot, G.A. Martin and J.C.Vedrine Metal Microstructures in Zeolites. Preparation - Properties-Applications. Proceedings of aWorkshop, Bremen, September 22-24, 1982 edited by RA. Jacobs, N.I. Jaeger, R Jin3 and G. Schulz-Ekloff Adsorption on Metal Surfaces. An Integrated Approach edited by J. B~nard Vibrations at Surfaces. Proceedings of theThird International Conference, Asilomar, CA, September 1-4, 1982 edited by C.R. Brundle and H. Morawitz Heterogeneous Catalytic Reactions Involving Molecular Oxygen by G.I. Golodets

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Preparation of Catalysts III. Scientific Basesfor the Preparation of Heterogeneous Catalysts. Proceedings oftheThird International Symposium, Louvain-la-Neuve, September 6-9, 1982 edited by G. Poncelet, R Grange and RA. Jacobs Spillover of Adsorbed Species. Proceedings of an International Symposium, Lyon-Viileurbanne, September 12-16, 1983 edited by G.M. Pajonk, S.J.Teichner and J.E. Germain Structure and Reactivity of Modified Zeolites. Proceedings of an International Conference, Prague, July 9-13, 1984 edited by RA. Jacobs, N.I. Jaeger, R Ji~,V.B. Kazansky and G. Schulz-Ekloff Catalysis on the Energy Scene. Proceedings of the 9th Canadian Symposium on Catalysis, Quebec, RQ., September 30-October 3, 1984 edited by S. Kaliaguine andA. Mahay Catalysis by Acids and Bases. Proceedings of an International Symposium, Villeurbanne (Lyon), September 25-27, 1984 edited by B. Imelik, C. Naccache, G. Coudurier, Y. BenTaarit and J.C.Vedrine Adsorption and Catalysis on Oxide Surfaces. Proceedings of a Symposium, Uxbridge, June 28-29, 1984 edited by M. Che and G.C. Bond Unsteady Processes in Catalytic Reactors by Yu.Sh. Matros Physics of Solid Surfaces 1984 edited by J. Koukal Zeolites: Synthesis, Structure,Technology andApplication. Proceedings of an International Symposium, Portoro~_-Portorose, September 3-8, 1984 edited by B. Drs S. HoEevar and S. Pejovnik Catalytic Polymerization of Olefins, Proceedings of the International Symposium on Future Aspects of Olefin Polymerization,Tokyo, July 4-6, 1985 edited by T. Keii and K. Soga Vibrations at Surfaces 1985. Proceedings of the Fourth International Conference, Bowness-on-Windermere, September 15-19, 1985 edited by D.A. King, N.V. Richardson and S. Holloway Catalytic Hydrogenation edited by L. Cerven~ New Developments in Zeolite Science andTechnology. Proceedings of the 7th International Zeolite Conference,Tokyo, August 17-22, 1986 edited by Y. Murakami,A. lijima and J.W.Ward Metal Clusters in Catalysis edited by B.C. Gates, L. Guczi and H. Kn6zinger Catalysis and Automotive Pollution Control. Proceedings of the First International Symposium, Brussels, September 8-11,1986 edited by A. Crucq andA. Frennet Preparation of Catalysts IV. Scientific Basesfor the Preparation of Heterogeneous Catalysts. Proceedings of the Fourth International Symposium, Louvain-laNeuve, September 1-4, 1986 edited by B. Delmon, R Grange, RA. Jacobs and G. Poncelet Thin Metal Films and Gas Chemisorption edited by RWissmann Synthesis of High-silica AluminosUicate Zeolites edited by RA. Jacobs and J.A. Martens Catalyst Deactivation 1987. Proceedings of the 4th International Symposium, Antwerp, September 29-October 1,1987 edited by B. Delmon and G.E Froment Keynotes in Energy-Related Catalysis edited by S. Kaliaguine

643 Volume 36 Volume 37 Volume 38 Volume 39 Volume 40 Volume 41

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Volume 45 Volume 46



Methane Conversion. Proceedings of a Symposium on the Production of Fuels and Chemicals from Natural Gas, Auckland,April 27-30, 1987 edited by D.M. Bibby, C.D. Chang, R.F.Howe and S.Yurchak Innovation in Zeolite Materials Science. Proceedings of an International Symposium, Nieuwpoort, September 13-17, 1987 edited by RJ. Grobet, W.J. Mortier, E.EVansant and G. Schulz-Ekloff Catalysis 1987.Proceedings of the 10th North American Meeting of the Catalysis Society, San Diego, CA, May 17-22, 1987 edited by J.W.Ward Characterization of Porous Solids. Proceedings of the IUPAC Symposium (COPS I), Bad Soden a.Ts.,Apri126-29,1987 edited by K.K. Unger, J. Rouquerol, K.S.W. Sing and H. Kral Physics of Solid Surfaces 1987. Proceedings of the Fourth Symposium on Surface Physics, Bechyne Castle, September 7-11,1987 edited by J. Koukal Heterogeneous Catalysis and Fine Chemicals. Proceedings of an International Symposium, Poitiers, March 15-17, 1988 edited by M. Guisnet, J. Barrault, C. Bouchoule, D. Duprez, C. Montassier and G. P~rot Laboratory Studies of Heterogeneous Catalytic Processes by E.G. Christoffel, revised and edited by Z. Paal Catalytic Processes under Unsteady-State Conditions by Yu. Sh. Matros Successful Design of Catalysts. Future Requirements and Development. Proceedings of theWorldwide Catalysis Seminars, July, 1988, on the Occasion of the 30th Anniversary of the Catalysis Society of Japan edited by T. Inui Transition Metal Oxides. Surface Chemistry and Catalysis by H.H. Kung Zeolites as Catalysts, Sorbents and Detergent Builders. Applications and Innovations. Proceedings of an International Symposium,Wfirzburg, September 4-8,1988 edited by H.G. Karge and J.Weitkamp Photochemistry on Solid Surfaces edited by M.Anpo andT. Matsuura Structure and Reactivity of Surfaces. Proceedings of a European Conference, Trieste, September 13-16, 1988 edited by C. Morterra,A. Zecchina and G. Costa Zeolites: Facts, Figures, Future. Proceedings of the 8th International Zeolite Conference, Amsterdam, July 10-14, 1989. PartsA and B edited by RA. Jacobs and R.A. van Santen Hydrotreating Catalysts. Preparation, Characterization and Performance. Proceedings of the Annual International AIChE Meeting, Washington, DC, November 27-December 2, 1988 edited by M.L. Occelli and R.G.Anthony New SolidAcids and Bases.Their Catalytic Properties by K.Tanabe, M. Misono,Y. Ono and H. Hattori RecentAdvances in Zeolite Science. Proceedings of the 1989 Meeting of the British Zeolite Association, Cambridge, April 17-19, 1989 edited by J. Klinowsky and RJ. Barrie Catalyst in Petroleum Refining 1989. Proceedings of the First International Conference on Catalysts in Petroleum Refining, Kuwait, March 5-8, 1989 edited by D.L.Trimm, S.Akashah, M.Absi-Halabi andA. Bishara Future Opportunities in Catalytic and SeparationTechnology edited by M. Misono,Y. Moro-oka and S. Kimura

644 New Developments in Selective Oxidation. Proceedings of an International Symposium, Rimini, Italy, September 18-22, 1989 edited by G. Centi and ETrifiro Olefin Polymerization Catalysts. Proceedings of the International Symposium Volume 56 on Recent Developments in Olefin Polymerization Catalysts,Tokyo, October 23-25, 1989 edited by T. Keii and K. Soga Volume 57A SpectroscopicAnalysis of Heterogeneous Catalysts. Part A: Methods of SurfaceAnalysis edited by J.L.G. Fierro Volume 57B SpectroscopicAnalysis of Heterogeneous Catalysts. Part B: Chemisorption of Probe Molecules edited by J.L.G. Fierro Introduction to Zeolite Science and Practice Volume 58 edited by H. van Bekkum, E.M. Flanigen and J.C. Jansen Heterogeneous Catalysis and Fine Chemicals I1. Proceedings of the 2nd Volume 59 International Symposium, Poitiers, October 2-6, 1990 edited by M. Guisnet, J. Barrault, C. Bouchoule, D. Duprez, G. P6rot, R. Maurel and C. Montassier Chemistry of Microporous Crystals. Proceedings of the International Symposium Volume 60 on Chemistry of Microporous Crystals,Tokyo, June 26-29, 1990 edited by T. Inui, S. Namba andT.Tatsumi Natural Gas Conversion. Proceedings ofthe Symposium on Natural Gas Volume 61 Conversion, Oslo, August 12-17, 1990 edited by A. Holmen, K.-J. Jens and S. Kolboe Characterization of Porous Solids II. Proceedings of the IUPAC Symposium Volume 62 (COPS II),Alicante, May 6-9, 1990 edited by E Rodriguez-Reinoso, J. Rouquerol, K.S.W. Sing and K.K. Unger Preparation of CatalystsV. Scientific Bases for the Preparation of Heterogeneous Volume 63 Catalysts. Proceedings of the Fifth International Symposium, Louvain-la-Neuve, September 3-6,1990 edited by G. Poncelet, RA. Jacobs, R Grange and B. Delmon Volume 64 NewTrends in CO Activation edited by L. Guczi Catalysis and Adsorption by Zeolites. Proceedings of ZEOCAT 90, Leipzig, Volume 65 August 20-23, 1990 edited by G. Ohlmann, H. Pfeifer and R. Fricke DioxygenActivation and Homogeneous Catalytic Oxidation. Proceedings of the Volume 66 Fourth International Symposium on Dioxygen Activation and Homogeneous Catalytic Oxidation, Balatonffired, September 10-14, 1990 edited by L.I. Simdndi Structure-Activity and Selectivity Relationships in Heterogeneous Catalysis. Volume 67 Proceedings of the ACS Symposium on Structure-Activity Relationships in Heterogeneous Catalysis, Boston, MA, Apri122-27, 1990 edited by R.K. Grasselli andA.W. Sleight Catalyst Deactivation 1991. Proceedings of the Fifth International Symposium, Volume 68 Evanston, I1_,June 24-26, 1991 edited by C.H. Bartholomew and J.B. Butt Zeolite Chemistry and Catalysis. Proceedings of an InternationalSymposium, Volume 69 Prague, Czechoslovakia, September 8-13, 1991 edited by RA. Jacobs, N.I. Jaeger, L. Kubelkovd and B.Wichterlov& Poisoning and Promotion in Catalysis based on Surface Science Concepts and Volume 70 Experiments by M. Kiskinova Volume 55

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Catalysis andAutomotive Pollution Control II. Proceedings of the 2nd International Symposium (CAPoC 2), Brussels, Belgium, September 10-13, 1990 edited by A. Crucq New Developments in Selective Oxidation by Heterogeneous Catalysis. Proceedings ofthe 3rd EuropeanWorkshop Meeting on New Developments in Selective Oxidation by Heterogeneous Catalysis, Louvain-la-Neuve, Belgium, April 8-10, 1991 edited by R Ruiz and B. Delmon Progress in Catalysis. Proceedings of the 12th Canadian Symposium on Catalysis, Banff, Alberta, Canada, May 25-28, 1992 edited by K.J. Smith and E.C. Sanford Angle-Resolved Photoemission.Theory and Current Applications edited by S.D. Kevan New Frontiers in Catalysis, PartsA-C. Proceedings of the 10th International Congress on Catalysis, Budapest, Hungary, 19-24 July, 1992 edited by L. Guczi, F.Solymosi and RTdtdnyi Fluid Catalytic Cracking: Science andTechnology edited by J.S. Magee and M.M. Mitchell, Jr. NewAspects of Spillover Effect in Catalysis. For Development of HighlyActive Catalysts. Proceedings of theThird International Conference on Spillover, Kyoto, Japan, August 17-20, 1993 edited by T. Inui, K. Fujimoto,T. Uchijima and M. Masai Heterogeneous Catalysis and Fine Chemicals III. Proceedings ofthe 3rd International Symposium, Poitiers, April 5-8, 1993 edited by M. Guisnet, J. Barbier, J. Barrault, C. Bouchoule, D. Duprez, G. Perot and C. Montassier Catalysis: An Integrated Approach to Homogeneous, Heterogeneous and Industrial Catalysis edited by J.A. Moulijn, RW.N.M. van Leeuwen and R.A. van Santen Fundamentals of Adsorption. Proceedings of the Fourth International Conference on Fundamentals ofAdsorption, Kyoto, Japan, May 17-22, 1992 edited by M. Suzuki Natural Gas Conversion I1. Proceedings of theThird Natural Gas Conversion Symposium, Sydney, July 4-9, 1993 edited by H.E. Curry-Hyde and R.E Howe New Developments in Selective Oxidation II. Proceedings of the SecondWorld Congress and Fourth European Workshop Meeting, Benalmadena, Spain, September 20-24, 1993 edited by V. Cortes Corber~n and S.Vic Bellbn Zeolites and Microporous Crystals. Proceedings of the International Symposium on Zeolites and Microporous Crystals, Nagoya, Japan,August 22-25, 1993 edited byT. Hattori andT.Yashima Zeolites and Related Microporous Materials: State of theArt 1994. Proceedings of the 10th International Zeolite Conference, Garmisch-Partenkirchen, Germany, July 17-22, 1994 edited by J.Weitkamp, H.G. Karge, H. Pfeifer andW. H61derich Advanced Zeolite Science and Applications edited by J.C. Jansen, M. St6cker, H.G. Karge and J.Weitkamp Oscillating Heterogeneous Catalytic Systems by M.M. Slin'ko and N.I. Jaeger Characterization of Porous Solids II1. Proceedings of the IUPAC Symposium (COPS III), Marseille, France, May 9-12, 1993 edited by J.Rouquerol, E Rodriguez-Reinoso, K.S.W. Sing and K.K. Unger

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Catalyst Deactivation 1994. Proceedings of the 6th International Symposium, Ostend, Belgium, October 3-5, 1994 edited by B. Delmon and G.F.Froment Catalyst Design forTailor-made Polyolefins. Proceedings of the International Symposium on Catalyst Design forTailor-made Polyolefins, Kanazawa, Japan, March 10-12, 1994 edited by K. Soga and M.Terano Acid-Base Catalysis II. Proceedings of the International Symposium on Acid-Base Catalysis II, Sapporo, Japan, December 2-4, 1993 edited by H. Hattori, M. Misono andY. Ono Preparation of CatalystsVI. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Sixth International Symposium, Louvain-La-Neuve, September 5-8, 1994 edited by G. Poncelet, J. Martens, B. Delmon, RA. Jacobs and R Grange Science andTechnology in Catalysis 1994. Proceedings of the SecondTokyo Conference on Advanced Catalytic Science andTechnology,Tokyo, August 21-26, 1994 edited by Y. Izumi, H.Arai and M. Iwamoto Characterization and Chemical Modification of the Silica Surface by E.RVansant, RVan DerVoort and K.C.Vrancken Catalysis by Microporous Materials. Proceedings of ZFOCAT'95, Szombathely, Hungary, July 9-13, 1995 edited by H.K. Beyer, H.G.Karge, I. Kiricsi and J.B. Nagy Catalysis by Metals and Alloys by V. Ponec and G.C. Bond Catalysis andAutomotive Pollution Control III. Proceedings of theThird International Symposium (CAPoC3), Brussels, Belgium, April 20-22,1994 edited by A. Frennet and J.-M. Bastin Zeolites:A RefinedTool for Designing Catalytic Sites. Proceedings of the International Symposium, Qu@bec, Canada, October 15-20, 1995 edited by L. Bonneviot and S. Kaliaguine Zeolite Science 1994: Recent Progress and Discussions. Supplementary Materials to the 10th International Zeolite Conference, Garmisch-Partenkirchen, Germany, July 17-22, 1994 edited by H.G. Karge and J.Weitkamp Adsorption on New and Modified Inorganic Sorbents edited by A. Dalbrowski andV.A.Tertykh Catalysts in Petroleum Refining and Petrochemical Industries 1995. Proceedings of the 2nd International Conference on Catalysts in Petroleum Refining and Petrochemical Industries, Kuwait, April 22-26, 1995 edited by M.Absi-Halabi, J. Beshara, H. Qabazard andA. Stanislaus 11th International Congress on Catalysis -40th Anniversary. Proceedings ofthe 11th ICC, Baltimore, MD, USA, June 30-July 5, 1996 edited by J.W. Hightower, W.N. Delgass, E. Iglesia andA.T. Bell RecentAdvances and New Horizons in Zeolite Science andTechnology edited by H. Chon, S.I.Woo and S.-E. Park Semiconductor Nanoclusters- Physical, Chemical, and CatalyticAspects edited by RV. Kamat and D. Meisel Equilibria and Dynamics of Gas Adsorption on Heterogeneous Solid Surfaces edited by W. Rudzifiski,W.A. Steele and G. Zgrablich Progress in Zeolite and Microporous Materials Proceedings of the 11th International Zeolite Conference, Seoul, Korea, August 12-17, 1996 edited by H. Chon, S.-K. Ihm andY.S. Uh

647 Hydrotreatment and Hydrocracking of Oil Fractions Proceedings of the 1st International Symposium / 6th European Workshop, Oostende, Belgium, February 17-19, 1997 edited by G.F.Froment, B. Delmon and R Grange Volume 107 Natural Gas Conversion IV Proceedings of the 4th International Natural Gas Conversion Symposium, Kruger Park, South Africa, November 19-23, 1995 edited by M. de Pontes, R.L. Espinoza, C.R Nicolaides, J.H. Scholtz and M.S. Scurrell Volume 108 Heterogeneous Catalysis and Fine Chemicals IV Proceedings of the 4th International Symposium on Heterogeneous Catalysis and Fine Chemicals, Basel, Switzerland, September 8-12, 1996 edited by H.U. Blaser,A. Balker and R. Prins Volume 109 Dynamics of Surfaces and Reaction Kinetics in Heterogeneous Catalysis. Proceedings ofthe International Symposium,Antwerp, Belgium, September 15-17,1997 edited by G.I-. Froment and K.C.Waugh Volume 110 ThirdWorld Congress on Oxidation Catalysis. Proceedings of theThirdWorld Congress on Oxidation Catalysis, San Diego, CA, U.S.A., 21-26 September 1997 edited by R.K. Grasselli, S.T. Oyama,A.M. Gaffney and J.E. Lyons Volume 111 Catalyst Deactivation 1997. Proceedings of the 7th International Symposium, Cancun, Mexico, October 5-8, 1997 edited by C.H. Bartholomew and G.A. Fuentes Volume 112 Spillover and Migration of Surface Species on Catalysts. Proceedings ofthe 4th International Conference on Spillover, Dalian, China, September 15-18, 1997 edited by Can Li and Qin Xin Volume 113 Recent Advances in Basic and Applied Aspects of Industrial Catalysis. Proceedings of the 13th National Symposium and Silver Jubilee Symposium of Catalysis of India, Dehradun, India,April 2-4, 1997 edited by T.S.R. Prasada Rao and G. Murali Dhar Volume 114 Advances in Chemical Conversions for Mitigating Carbon Dioxide. Proceedings of the 4th International Conference on Carbon Dioxide Utilization, Kyoto, Japan, September 7-11,1997 edited by T. Inui, M.Anpo, K. Izui, S.Yanagida andT.Yamaguchi Volume 115 Methods for Monitoring and Diagnosing the Efficiency of Catalytic Converters. A patent-oriented survey by M. Sideris Volume 116 Catalysis and Automotive Pollution Control IV. Proceedings of the 4th International Symposium (CAPoC4), Brussels, Belgium, April 9-11, 1997 edited by N. Kruse, A. Frennet and J.-M. Bastin Volume 117 Mesoporous Molecular Sieves 1998 Proceedings of the 1st International Symposium, Baltimore, MD, U.S.A., July 10-12, 1998 edited by L.Bonneviot, F. Bdland, C. Danumah, S. Giasson and S. Kaliaguine Volume 118 Preparation of Catalysts VII Proceedings of the 7th International Symposium on Scientific Bases for the Preparation of Heterogeneous Catalysts, Louvain-la-Neuve, Belgium, September 1-4, 1998 edited by B. Delmon, RA. Jacobs, R. Maggi, J.A. Martens, R Grange and G. Poncelet Volume 119 Natural Gas ConversionV Proceedings of the 5th International Gas Conversion Symposium, Giardini-Naxos, Taormina, Italy, September 20-25, 1998 edited by A. Parmaliana, D. Sanfilippo, F.Frusteri,A.Vaccari and F.Arena Volume 120A Adsorption and its Applications in Industry and Environmental Protection. Vol I: Applications in Industry edited by A. Dabrowski

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648 Volume 120B Adsorption and its Applications in Industry and Environmental Protection. Vol I1:Applications in Environmental Protection edited byA. Dabrowski Volume 121 Science andTechnology in Catalysis 1998 Proceedings of theThirdTokyo Conference in Advanced Catalytic Science and Technology,Tokyo, July 19-24, 1998 edited by H. Hattori and K. Otsuka Volume 122 Reaction Kinetics and the Development of Catalytic Processes Proceedings ofthe International Symposium, Brugge, Belgium,April 19-21,1999 edited by G.E Froment and K.C.Waugh Volume 123 Catalysis: An Integrated Approach Second, Revised and Enlarged Edition edited by R.A. van Santen, RW.N.M. van Leeuwen, J.A. Moulijn and B.A.Averill Volume 124 Experiments in Catalytic Reaction Engineering by J.M. Berty Volume 125 Porous Materials in Environmentally Friendly Processes Proceedings of the 1st International FEZA Conference, Eger, Hungary, September 1-4, 1999 edited by I. Kiricsi, G. PaI-Borbdly, J.B. Nagy and H.G. Karge Volume 126 Catalyst Deactivation 1999 Proceedings ofthe 8th International Symposium, Brugge, Belgium, October 10-13, 1999 edited by B. Delmon and G.E Froment Volume 127 Hydrotreatment and Hydrocracking of Oil Fractions Proceedings of the 2nd International Symposium/Tth European Workshop, Antwerpen, Belgium, November 14-17, 1999 edited by B. Delmon, G.E Froment and R Grange Volume 128 Characterisation of Porous SolidsV Proceedings of the 5th International Symposium on the Characterisation of Porous Solids (COPS-V), Heidelberg, Germany, May 30- June 2, 1999 edited by K.K. Unger, G. Kreysa and J.R Baselt Volume 129 Nanoporous Materials II Proceedings of the 2nd Conference on Access in Nanoporous Materials, Banff, Alberta, Canada, May 25-30, 2000 edited by A. Sayari, M. Jaroniec andT.J. Pinnavaia Volume 130 12th International Congress on Catalysis Proceedings of the 12th ICC, Granada, Spain, July 9-14, 2000 edited byA. Corma, EV. Melo, S. Mendioroz and J.L.G. Fierro Volume 131 Catalytic Polymerization of Cycloolefins Ionic, Ziegler-Natta and Ring-Opening Metathesis Polymerization byV. Dragutan and R. Streck Volume 132 Proceedings of the International Conference on Colloid and Surface Science, Tokyo, Japan, November 5-8, 2000 25th Anniversary of the Division of Colloid and Surface Chemistry, The Chemical Society of Japan edited byY. Iwasawa, N. Oyama and H. Kunieda Volume 133 Reaction Kinetics and the Development and Operation of Catalytic Processes Proceedings ofthe 3rd International Symposium, Oostende, Belgium,April 22-25, 2001 edited by G.E Froment and K.C.Waugh

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