Catalysis: A Review of Chemical Literature: Vol 15 (Specialist Periodical Reports)

Catalysis Volume 15

A Specialist Periodical Report

Catalysis Volume 15 A Review of Recent Literature Senior Reporter James J. Spivey, Research Triangle Institute, Research Triangle Park, North Carolina, USA

Reporters F? Aghalayam, University of Massachusetts, Amherst, MA, USA T. Baba, Tokyo Institute of Technology, Japan S. Bordawekar, University of Virginia, Charlottesville, VA, USA B.H. Davis, University of Kentucky, Lexington, KY, USA R.J. Davis, University of Virginia, Charlottesville, VA, USA E.J. Doskocil, University of Virginia, Charlottesville, VA, USA M. Machida, Miyazaki University, Miyazaki, Japan Y. Ono, National lnstitution for Academic Degrees, Yokohama, Japan Y.K. Park, University of Massachusetts, Amherst, MA, USA A. Ueno, Shizuoka University, Shizuoka, Japan D.G. Vlachos, University of Massachusetts, Amherst, MA, USA Y. Zhang, University of Kentucky, Lexington, KY, USA

ROYAL SOCIETY OF CHEMISTRY

ISSN 0 140-0568 ISBN 0-85404-219-9 @] The Royal Society of Chemistry 2000

All rights reserved. Apart from any fair dealing for the purpose of research or private study, or criticism or review us permitted under the terms of the UK Copyright, Designs and Patents Act, 1988, this publication may not be reproduced, stored or trunsmitted, in any form or by m y means, without the prior permission in writing of The Royal Society of Chemistry, or in the cuse of reprogruphic reproduction only in accordance with the terms of the licences issued by the Copyright Licensing Agency in the UK, or in uccordance with the terms of the licences issued by the appropriute Reproduction Rights Organizution outside the UK, Enquiries concerning reproduction outside the terms stuted here should be sent to The Royul Society of Chemistry ut the address printed on this page. Published by The Royal Society of Chemistry, Thomas Graham House, Science Park, Milton Road, Cambridge CB4 OWF, UK For further information see our web site at www.rsc.org Typeset by Computape (Pickering) Ltd, Pickering, North Yorkshire, UK Printed and bound by Athenaeum Press Ltd, Gateshead, Tyne &Wear

Preface

The challenges that the chemical industry and research institutions face are increasing. The cost-efficient and environmentally sound conversion of natural resources into fuels, chemicals, and energy require innovative solutions. Catalysts are essential to these processes, and the subjects of the chapters presented here reflect this. In two separate chapters, Robert Davis, Eric Doskocil, and Shailendra Bordawekar (University of Virginia), and Yoshio Ono (National Institution for Academic Degrees, Japan) and Toshihide Baba (Tokyo Institute of Technology) review catalysis by solid bases. These reviews reflect the interest of catalyst researchers in this area. Although acid catalysis is much more widely studied and applied industrially, there are key reactions that require base catalysts, such as aldol condensations. In some cases, the synthesis of heterogeneous base catalysts to replace aqueous base catalysts can have positive environmental effects by eliminating wastewater streams that are otherwise produced. The challenge is to maintain industrially practical selectivity and activity without significant deactivation. Masato Machida (Miyazaki University) has provided a review of solid sorbents for No, removal. These materials can be used in conjunction with catalytic reduction, especially when the NO, concentrations are low. Carbons, zeolites, and modified aluminas can be used - both in stationary and mobile sources. He shows how the NO, separation and catalytic reduction can be combined into a cyclical process for NO, control. P. Aghalayam, Y.K. Park, and D.G. Vlachos (University of Massachusetts) review progress in catalysts for the partial oxidation of light alkanes. Most of the interest in this area, and the focus of this review, is on the conversion of methane. However, the authors have expanded their review to oxidative dehydrogenation of ethane and other alkanes. In an closely related chapter, Akifumni Ueon (Shizuoka University) describes recent work in the partial oxidation of methane, with a special focus on silocomolybdic acid catalysts. These solid acids are one of several classes of catalysts being investigated for this process, and they are more selective for oxygenates such as formaldehyde and methanol - both industrially important intermediates now produced by other processes. Finally, Burt Davis and Yongquing Zhang (University of Kentucky) have reviewed the status of indirect liquefaction, and the critical role that catalyst development must play in bringing these processes to wider commercial practice. This review clarifies the relationships among the wide range of V

vi

Preface

processes that have been developed and offers insight into the challenges that remain. Volume 16 is underway and I look forward to bringing it to you. As always, comments are welcome. James J. Spivey Research Triangle Institute Research Triangle Park, NC USA Email:[email protected]

Contents

Chapter 1 Strong Solid Bases for Organic Reactions By Yoshio On0 and Toshihide Baba 1 Introduction

1

1

2 Role of Solid Base and Basic Sites as a Catalyst 2.1 Abstraction of Protons 2.2 Activation of Reactants without Proton Abstraction 2.3 Cooperative Action of Acidic and Basic Sites

2 2

2 2

3 Base Strength of Basic Sites 3.1 H - Acidity Function 3.2 Indicator Method 3.3 Other Methods for Determining Basic Strength

4 Base Strength and Catalytic Reactions

5

5 Solid Base Materials 5.1 Alkaline Earth Oxides 5.2 Alkali Metals Supported on Metal Oxides 5.3 Alkali Metals on A1203Treated with Alkali Metal Hydroxides 5.4 KNH2/A1203 5.5 Alkali Metal Compound Supported on Alumina 5.5.1 KNOJA1203, KHCOJA1203, K2C03/A1203 5.5.2 LiOH/A1203, NaOH/A1203,KOH/A1203, RbOH/A1203, CsOH/A1203 5.6 Zeolites 5.7 Modified Zeolites 5.8 E~,Yb/A1203 5.9 KF/A1203 5.10 Mixed Oxides

8 8

14 14 15 16 16 18

6 Reactions Catalysed by Solid Bases

20

~ _ _ _

Catalysis, Volume 15 0The Royal Society of Chemistry, 2000

vii

9 10 10 14 14

...

Contents

Vlll

6.1 Isomerizations 6.1.1 Isomerization of Alkenes 6.1.2 Isomerization of Alkenyl Amines 6.1.3 Isomerization of Ally1 Ethers 6.1.4 Isomerization of Alk- 1-ynes 6.2 Nucleophilic Addition Reactions of Phenylacetylene 6.2.1 Dimerization of Phenylacetylene 6.2.2 Reactions of Alk-1-yne with Ketone or Aldehyde 6.3 Aldol-type Condensations 6.3.1 Aldol Condensations 6.3.2 Claisen-Schmidt Reactions 6.3.3 Knoevenagel Condensations 6.3.4 Michael Additions 6.3.5 Dehydrocondensation with Methanol 6.4 Nucleophilic Substitution at Silicon - Si-C Bond Formation 6.4.1 Metathesis of Trimethylsilylacetylene 6.4.2 Cross-metathesis of Alkynes 6.4.3 Reaction of Alk-1-yne with Silane 6.4.4 Reaction of Alk-1-ene with Diethylsilane 6.4.5 Reaction of Toluene with Diethylsilane 6.5 Miscellaneous Reactions 6.5.1 Tishchenko Reaction 6.5.2 Reactions of Silanes with Ketones or Aldehydes 6.5.3 Reaction of Aniline with Trimethylsilylacetylene 6.5.4 Ring Opening of Epoxides with Me3SiCN 6.5.5 Hydrocyanation 6.5.6 Alkylations 6.5.7 Formation and Ring Transformations of Heterocycles

20 20 20 21 21 21 21

7

Concluding Remarks - Unsolved Problems

35

References

36

Chapter 2 Catalysis by Solid Bases By Eric J. Doskocil, Shailendra Bordawekar and Robert J. Davis

22 25 25 25 27 28 28 28 29 30 30 31 31 32 32 33 33 34 34 34 35

40

1 Introduction

40

2 Types of Basic Catalysts 2.1 Supported Alkali Metal Oxides

41 41

Contents

ix

2.2 2.3 2.4 2.5

Hydrotalcites Zeolites and Mesoporous Oxides Alkali Metals Supported on Nanoporous Carbon Basic Phosphate Oxynitride Catalysts

3 Characterization of Solid Base Catalysts 3.1 Theoretical Ranking of Solid Basicity 3. I. 1 Sanderson Intermediate Electronegativity 3.1.2 Molecular Orbital Calculations 3.2 Experimental Ranking of Solid Basicity 3.2.1 Infrared Spectroscopy of Adsorbed Probe Molecules 3.2.2 Visible Absorption Spectroscopy of Adsorbed Iodine 3.2.3 Microcalorimetry of Adsorbed Probe Molecules 3.2.4 Temperature Programmed Desorption (TPD) of Probe Molecules 3.2.5 X-Ray Absorption Spectroscopy 3.2.6 X-Ray Photoelectron Spectroscopy

41 42 44 45 45 46 46 46 47 47 50 52 54 54 56

4 Probe Reactions over Solid Base Catalysts 57 4.1 Double-bond Isomerization 58 59 4.2 Alcohol Dehydrogenation 4.3 Hydrogenation Reactions 60 4.4 Condensation Reactions 61 4.4.1 Aldol Condensation 61 4.4.2 Knoevenagel Condensation 62 4.5 Alkylation Reactions 63 4.6 Side-chain Alkenylation of o-Xylene with 1,3-Butadiene 64 4.7 Miscellaneous Reactions 65

5 Conclusions

67

Acknowledgements

67

References

67

Chapter 3 Solid Sorbents for Catalytic NO, Removal By Masato Machida

73

1 Introduction

73

2 Materials for NO, Sorption 2.1 Carbonaceous Solids 2.2 Zeolites

74 74

75

Conten ts

X

2.3 2.4 2.5 2.6 2.7 2.8

Alumina Alkaline Solids Transition Metal Oxides containing Alkaline Earths NO, Intercalation Compounds Non-alkaline Solid Oxides Heteropoly Compounds

76 77 78 82 83 87

3 Regeneration of NO, Sorbent 3.1 Pressure Swing Process 3.2 Thermal (Temperature) Swing Process 3.3 Reduction-Oxidation Cycles

88 89 90 90

4 Practical Applications of NO, Sorption 4.1 Flue Gas Treatment for Stationary Sources 4.2 Automobile NO, Storage Catalysts

90 90 92

5 Conclusion

93

References Chapter 4 Partial Oxidation of Light Alkanes in Short Contact Time Microreactors By P. Aghalayam, Y.K. Park and D. G. Vlachos

1 Introduction

93

98

98

2 CH4 Partial Oxidation on Platinum and Rhodium Catalysts 2.1 Unsupported Catalysts 2.2 Foam and Extruded Monoliths 2.3 Fluidized Beds 2.4 The Effect of the Support and Pore Diameter

99 101 103 103 104

3 Influence of Operating Conditions 3.1 Temperature 3.2 Flow Velocity and Contact Time 3.3 Pressure 3.4 Dilution 3.5 Inlet Fuel Composition 3.6 Influence of Different Surface Coverages

104 104 104 107 108 108 109

4 Analysis of the Effect of Operating Conditions 4.1 Model 4.2 Flow Velocity 4.3 Temperature

110 110 112 113

xi

Contents

5 Bifurcation of Methane-Oxygen Mixtures Near Catalysts 5.1 Influence of Fuel Composition 5.2 Influence of Pressure and Flow Velocity 5.3 The Inhibiting Role of the Catalyst in Gas-phase Ignition 5.4 Importance of Gas-phase Reactions

118 119

6 Direct versus Indirect Path to Syngas Formation

121

114 114 118

7 A Quantitative Reaction Mechanism for Oxidation of Methane 7.1 Currently Proposed Mechanisms 7.1.1 Selectivity Mechanisms 7.1.2 Catalytic Ingition Mechanisms 7.1.3 Other Proposed Reaction Pathways 7.2 Limitations of the Existing Detailed Surface Reaction Mechanisms 7.2.1 Adsorption of CH4 7.2.2 The Role of Adsorbed Oxygen 7.2.3 Adsorbate-Adsorbate Interactions 7.2.4 Other Limitations

126 127 128 129 129

8 Partial Oxidation of Ethane and Higher Alkanes 8.1 Production of Olefins from Alkanes using Pt and Rh 8.2 The Mechanism of Alkane Dehydrogenation

130 131 131

9 Conclusions and Outlook

133

125 125 125 126 126

Acknowledgement

134

References

134

Chapter 5 Indirect Liquefaction - Where Do We Stand? By Yongqing Zhang and Burtron H. Davis

138

1 Introduction

138

2 Today’s Commercial Operations 2.1 South Africa 2.2 Sasol 2.3 Mossgas 2.4 Shell

139 140 140 145 145

3 Large Pilot/Demonstration Plant Operations 3.1 Rheinpresussen-Koppers 3.2 British Fuel Research Station

148 148 153

Contents

xii

3.3 US Bureau of Mines 3.4 Stanolind/Carthage Hydrocol

153 155

4 Standard Oil Co. (New Jersey) - Exxon

156

5 Pilot Scale Operations 5.1 Syntroleum 5.2 Gulf Oil 5.3 Rentech 5.4 Chinese Studies 5.5 Mobil Oil

161 161 170 171 172 172

6 Process and Economic Evaluations

174

7 Potential Commercial Operations

175

8 Summary of Current Status

180

Acknowledgement

182

References

182

Chapter 6 Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts By Akifumi Ueno

185

1 Introduction

185

2 Partial Oxidation of Methane on Mo031Si02 and Alkali Metal-doped Mo03/Si02

186

3 Active Sites on Mo03/Si02and Reaction Mechanism of Selective Oxidation of Methane

192

4 Characterization of Surface Species Generated on Mo03/Si02 by Heat Treatments

197

5 Catalytic Activity of SMA Generated On or Impregnated on Si02

204

6 Other Topics Concerning Heteropoly Acid Catalysts or Partial Oxidation of Methane

208

7 Summary of the Selective Oxidation of Methane on Mo03/Si02 Catalysts

21 1

References

212

1 Strong Solid Bases for Organic Reactions BY YOSHIO O N 0 AND TOSHlHlDE BABA

1

Introduction

Carbanions are important intermediates in many organic reactions such as isomerizations, additions, alkylations, and cyclizations. They are formed by abstraction of a proton from a C-H bond of an organic molecule by a base. These organic reactions often require a stoichiometric amount of liquid base to generate carbanions and produce a stoichiometric amount of metal salts as a by-product. For example, the methylation of phenylacetonitrile with methyl iodide proceeds in the presence of base under a phase-transfer condition. PhCH2CN

+ CH31 + NaOH

-

PhCH(CH3)CN

+

Nal

+

H20

(1.1)

In this case, more than a stoichiometric amount of sodium hydroxide is required to neutralize the hydrogen iodide produced and to keep the system basic. Furthermore, a stoichiometric amount of sodium iodide is inevitably formed and has to be disposed of in an appropriate manner. Organometallic compounds such as Grignard reagents and alkyl lithium serve as donors of carbanion-like species. Here, again, a stoichiometric use of these reagents is required. Therefore, there is a need to develop solid bases to avoid these problems. Solid base catalysts have many advantages over liquid bases. They are noncorrosive and environmentally benign, presenting fewer disposal problems, while allowing easier separation and recovery of the products, catalysts and the solvent. Thus, solid base catalysis is one of the economically and ecologically important fields in catalysis and the replacement of liquid bases with heterogeneous catalysts is becoming more and more important in the chemical industry. * Furthermore, high activities and selectivities are often obtained only by solid base catalysts for various kinds of reaction. Since the ability of bases to abstract a proton from a C-H bond is directly connected to the base strength, stronger bases are in general more effective in forming carbanions. Alkaline earth oxides such as magnesium oxide are strongly basic when properly pretreated. Extensive works by Tanabe, Hattori and their co-workers have been carried out using these material^.^.^ Recently, other strong base catalysts have been reported. Potassium amide supported on alumina (KNH2/A1203) is effective for a number of baseCatalysis, Volume 15 0The Royal Society of Chemistry, 2000 1

2

Catalysis

catalysed Even toluene is activated to react with silanes at 329 K,8 and the isomerization of 2,3-dimethylbut- 1-ene proceeds even at 20 1 K.495 Potassium fluoride supported on alumina (KF/A1203) has been used by organic chemists for a long time,10-12but Tsuji and Hattori revealed that this catalyst becomes much more active when pretreated at 573-673 K under vacuum. l 3 Yamaguchi et al. reported that catalysts prepared by loading alkalimetal salts such as KN03, followed by heating at 773-873 K, were very strongly basic and active for the isomerization of cis-but-2-ene at 273 K.14 Fu et al. used alkali-metal compounds supported on alumina for the reaction of catechol with dimethyl carbonate and found that the rate and selectivity depended strongly on the catalyst used.I5- Furthermore, modified zeolites and calcined hydrotalcites are often reported as strong bases. In this review, we will describe the preparation and characterization of these strong bases. Then, application of these catalysts to a variety of catalytic reactions is described. The reactions include the isomerization of alkenes and alkynes, the dimerization of alkynes, aldol reactions, and the formation of Si-C, Si-N and Si-0 bonds. 2

Role of Solid Base and Basic Sites as a Catalyst

2.1 Abstraction of Protons - On the surface of solid bases, there are specific sites or centers, which function as a base. Basic sites (centers) abstract protons from the reactant molecules (AH) to form carbanions (A-),

Here, the basic site B- on the solid surface acts as a Bronsted base. Stronger bases can abstract a proton with molecules with higher pK, values.

2.2 Activation of Reactants without Proton Abstraction - Reactants such as ketones and aldehydes are often activated by bases without proton transfer, as expressed by the following equation.

I

R2

B

Here, the basic sites B - act as a Lewis base. It should be noted that a same surface site can serve as a Bronsted base as well as a Lewis base, depending on the nature of the adsorbate.

2.3 Cooperative Action of Acidic and Basic Sites - Magnesium oxide is active for the hydrogenation of 1,3-butadiene. It is assumed that hydrogen heterolytically dissociates in the presence of a pair of a coordinatively unsaturated Mg2+ and an oxide Hydrogen adsorption is schematically expressed as:

1: Strong Solid Bases for Organic Reactions

3

(1.4)

3

Base Strength of Basic Sites

3.1 H- Acidity Function - The H - acidity function is defined as a measure of the ability of the basic solution to abstract a proton from an acidic neutral solute.

To determine the H - value of a solution, the concentrations of AH and A have to be measured accurately. When half of a solute AH is deprotonated in the solution, i.e., [A-] = [AH], the H - value of the solution is equal to the pK, value of AH. The basic strength of a solution is stronger when a neutral molecule of larger pK, value is deprotonated. Tanabe proposed transferring this concept to solid bases as a measure of their strength.2 The base strength of solid bases is expressed by means of the H - value, equated to the highest among the pK, values of the adsorbates from which the basic site is able to abstract a proton. Tanabe defined solid superbases as materials with H - values higher than 26. This value, like that of superacids ( H - < -12), is 19 units from a neutral solution of 7. In the use of this concept for solid bases, two important points should be noted: (a) In the discussion of solid bases, the H - value is treated as a parameter to describe the nature of individual basic sites. It is often assumed that there are a certain number of basic sites on solid surfaces and that each of the sites has its own basic strength. In the original definition, H - scale is used to describe basic property of the solution, not that of individual basic molecules (or ions) in the solution. (b) In principle, the idea of the H - scale is only applicable to the Bronsted base. It is not, at least directly, related to the ability of the sites to function as Lewis bases, as shown in eqn. (1 3. 3.2 Indicator Method - The H - values of basic solutions are determined by using indicator molecules.18If the pK, value of the indicator AH is known, the H - value can be calculated by determining the ratio of [AH]/[A-I. To cover a wide range of the H - scale, a series of indicators with different pK, values have been selected to obtain the accurate value of [AH]/[A-1. In the case of solid bases, the color change of indicator molecules upon adsorption is taken as a measure of basic strength.* If the color change of the indicator is observed, the H - value of the basic sites on the solid is higher than the pK, value of the indicator. Similarly, if the indicator does not change color

4

Catalysis

upon adsorption, the H- value of the sites is judged to be less than the pK, value of the indicator. By using indicators of different pK, values, the H value of the basic sites can be determined. It is important to know that the color change is due to proton abstraction by basic sites and not due to other types of interactions such as charge-transfer between the adsorbate and the surface.

3.3 Other Methods for Determining Basic Strength - Temperature programmed desorption of carbon dioxide is often used. When the interaction between basic sites and carbon dioxide is stronger, the molecule desorbs at higher t e r n p e r a t ~ r e . ~One . ' ~ disadvantage of using carbon dioxide is that this molecule adsorbs on solid surfaces in several different forms. For example, as revealed by infrared spectroscopy, carbon dioxide is adsorbed on alkaline earth oxides to form a unidentate complex as well as a bidentate complex.20 Since the interaction of carbon dioxide with the surface does not involve a proton-transfer process, the result may not be directly related to the Bronsted basicity of the sites. The XPS binding energy value of elements depends on the charge carried by the atom. The binding energy of 0 1 , is then expected to decrease with increasing negative charge on the oxygen. The 0 1 , binding energy of X- and Y-type zeolites decreases with increasing Si/Al ratio and decreasing electronegativity of the counter cation.21'22Since the XPS binding energy is measured as the average of those for all of the 0 atoms in the material, the method is not applicable for the materials where only a fraction of the oxygen ions are active as basic sites, as in most alkaline earth oxides. Infrared spectroscopy of adsorbed molecules is often used for characterizing surface bacisity. Pyrrole is an amphoteric molecule. The basic strength may be estimated from the shift of NH vibration upon its interaction with basic sites through hydrogen bonding. For example, when pyrrole interacts with framework oxygen ions in zeolites, the NH vibration is shifted to low wavenumbers from 3430 cm- in the pure liquid to around 3200 cm- .23 The extent of the shift increases with the negative charge on oxygen, which is calculated from Sanderson's electronegativities. Though pyrrole adsorption has become a popular technique, the spectra are rather complex and the molecule is not always stable on the surface; it polymerizes or dissociates on some oxides.24 The shift of C-D stretching mode of adsorbed CDC13 is also a measure of the basic strength. Berteau et al. examined the base properties of modified aluminas by IR spectroscopy of probe molecules and C 0 2 TPD.25 Acetylene and substituted acetylenes have also been used as probe molecules for surface basicity.26 Bosacek proposed the use of 13C NMR of adsorbed methyl iodide for the basicity of zeolites.27Methyl iodide heterolytically dissociates and the methyl group attaches to the lattice oxygen. The chemical shift of the carbon, therefore, reflects the actual electronegativity of the oxygen. With this method, Hunger et al. confirmed that strongly basic sites were created by incorporating alkali metal oxides into the zeolite pores.28

1: Strong Solid Bases for Organic Reactions

4

5

Base Strength and Catalytic Reactions

Catalytic reactions provide an accurate measure of basic strength, especially when the reaction starts by formation of carbanions by abstraction of a proton from the reactant, since the ease of carbanion formation depends on the PKa value of reactant. In Table 1.1, PKa values of various compounds are listed. Isomerization of alkenes such as but-1 -ene proceeds through the formation of allylic anions, which are formed by abstraction of a proton from but-1-ene, as shown in reaction ( 1.6).29

Since the PKa value of alkenes is high (PKa of C3H6=35), strong bases are required to activate the alkene molecule. Thus, alkene isomerization is an appropriate test reaction for strong solid bases. Moreover, the reaction is mechanistically simple. This makes the interpretation of experimental results straight forward. Table 1.2 shows the catalytic activities of various solid bases for the isomerization of 2,3-dimethylbut- 1-ene (DB-1) to 2,3-dimethylbut-2-ene after 20 h.30The activities vary significantly from one catalyst to the other, reflecting a wide variety of the base strength and the number of the basic sites. KY (K+exchanged Y zeolite) has no activity, indicating that no strong basic sites exist on KY. On the other hand, there are groups of catalysts that have very high activities: alkali amides on A1203,alkali compounds on A1203,CaO and MgO. Since the conversion over these catalysts is close to the equilibrium value at 3 13 K, it is hard to know the relative ranking of activities for these materials from Table 1.2. Table 1.3 shows the results when the isomerization was carried out at much lower temperature, i.e. 201 K, and a shorter reaction time on these materials, RbNH2/A1203, a series of alkali hydroxide supported on A1203,and K loaded on A1203 prepared by the deposition of K vapor (K/A1203), being added to the list. The reaction is very fast over RbNH2/A1203, KNH2/A1203, CsOH/A1203 and CaO even at 201 K. The order of the activities for the most active class of solid base catalysts is as follows: RbNH2/A1203> KNH2/A1203 > CaO > MgO > CsOH/A1203 > KN03/A1203> RbOH/A1203> KOH/A1203> K/A1203 The PKa value of propene is 35 and we can assume that the pK, value of DB-1 is not far from this value. Therefore, all of these catalysts can be classified as solid-superbases. CaO was reported to have basic sites stronger than H - = 26 by an indicator method.2 As mentioned above, KY does not show the catalytic activity for the isomerization of DB-1, whereas it catalyses a Knoevenagel reaction of benzaldehyde with ethyl ~yanoacetate.~' O

C

H

-

O + NCCH2COOEt

O

CN C

H

=

(1.7)

& COOEt

6

Catalysis

Table 1.1 Approximate acidities of various common reagents PKl

4A

9

PK,

ref:

NH3

35

a

H2C=CH2

36.5

b

37

b

37

a

CH4

40

b

H3C-CH3

42

b

H3C-CH2-CH3

44

b

45

b

ref: a

H H

OH

9

11

a

13

a

15

C

15.75

a

16.7

C

y

3

H H

CH30H CH3CH20H

O

C

E

C

H

17 18

a

18.5

b

a

0 II

CrCH3

19

a

25 25

a,b

CH3CN CH3C02C2H5

25

a

25

a

HCGCH

a

H.O. House, Modern Synthetic Reuction, W.A. Benjamin, INC., Menlo Park, CA, USA, 1972, p. 494. D.J. Cram, Fundamentuls of Curbanion Chemistry, Academic Press, New York and London, 1965.

a

This indicates that basic sites of KY are able to abstract a proton from the latter, which has a pK, value of 8.6. Therefore, when this reaction proceeds over a solid catalyst, it is judged that the solid has basic sites stronger than H- =8.6. Though the H - values of a base catalyst can be used to decide whether a

1: Strong Solid Basesfor Organic Reactions

7

Table 1.2 The catalytic activities of solid base catalysts for the isomerization of 2,3-dimethylbut-l-ene at 313 K after 20 h 303a

Catalyst

Metals supported from liq. NH3b*C KNH2/A1203(2.6) K(N H 3)/AI2O3(2.O) Na(NH3)/A1203(3.5) E~(NH3)/A1203(0.5) Y b(N H3)/A1203(0.5) E~(NH3)/A1203(0.5) Y b( N H3)/A1203(0.5) Alkali metals/A1203 KN03/A1203(5.O) KOH/A1203(5.0) KF/A1203(5.0) Metal oxides CaO MgO SrO BaO Sm203 EU203 Yb203 A1203 Mixed oxidesd 4 CaO, A1203 4 MgO, A1203 Zeolites KY

Pretreatment TemperaturelK Timelh

573 423 423 523 473 423 473 873 673 623

Conversion (%)

91.3 89.4 89.3 83.4 29.4 2.3 1 .o 3 3 3

90.7 86.0 25.6 89.8 84.9 1.6 0.4 1.2 1.o 0.9 3.9

1073 873 1023 1073 773 923 923 773 1073 773

1 4

11.9 1.3

773

3

60%) and conversion (> 60%) have been obtained with contact times of the order of rnilliseconds'O to ca. 50 microseconds.Iol Use of Pt-Sn alloys as catalysts leads to improvements in both selectivity and conversion. Comparable conversions and selectivities to olefins have been achieved for higher alkane oxidation over Pt-based catalysts, such as propane,lo4 pentane and hexane. Io2 However, in contrast to partial oxidation of CH4 to syngas, alloying Rh with Pt leads to decreasing conversion and selectivities, regardless of the fue1.99~101~102 Use of a combined separationreactor for the dehydrogenation of i-C4H 10 has also shown potential for improved conversion and selectivities by shifting equilibrium limitations. Io5 Addition of H2 to the feed has a positive effect on performance.ll-lO1 Regarding the role of H2 in improving selectivities to olefins, it has been proposed that H2 is preferentially oxidized to H20 on the catalyst surface. This has two performance-promoting effects: first, preferential consumption of 0 2 that can lead to deep alkane oxidation; and, second, an increase in reactor temperature, due to the exothermicity of H20 formation, that promotes alkane dehydrogenation. Overall, Pt is an active catalyst for the production of olefins in oxidative alkane dehydrogenation in short-contact time reactors. Among the other catalysts studied, Rh is found to be non-selective to olefins, whereas Pt-Sn alloys gave higher selectivities than Pt. Aside from the catalyst, reactor design (such as heat transfer, distributed feed of reactants, etc.) plays a key role in performance as the experiments with co-feeding of H2 indicate.

8.2 The Mechanism of Alkane Dehydrogenation - The short-contact times in alkane dehydrogenation processes suggest a purely heterogeneous reaction mechanism. In fact, a detailed surface reaction scheme has been proposed and found to give reasonable predictions of data.9 However, various experiments without the catalyst have been conducted for ethanelO' and propane oxidationIoOrevealing that, even in the absence of a catalyst, high conversions of the fuel and selectivity to olefins may be obtained. In fact, the presence of a

Contact time

1-10ms 1-10 ps 50 ps 50 ps 1 ms 1 ms 1-10 ps ca. 1 ms ca. 5 ms ca. 5 ms ca. 5 ms

Catalyst

Pt/A1203 Pt/Rh gauze Pt gauze Pt/Rh gauze Pt/A1203 Pt-SdAl203 Pt/Rh gauze Pt/A1203 Pt/Al203 Rh/A1203 Pt/A1203

Reactant

!% Conv.

70 62 65 60 64 83 58 60 60 40 60

!% S(o1eJins)

Schmidt and Goetsch and Schmidt99 Lodeng et al."' Lodeng et al."' Bodke et al." Bodke et al." Goetsch and Schmidt99 Beretta et Dietz et a1.Io2 Dietz et al. Io2 Dietz et a1.Io2

Reference

Table 4.5 Selected summary of the experimental data for catalytic oxidation of higher alkanes at short-contact times

4: Partial Oxidation of Light Alkanes in Short Contact Time Microreactors

133

catalyst seems to be detrimental to olefin production at low conversions, as compared to a non-catalytic system.lO*On the other hand, in the absence of the catalyst, the induction times are found to be high (of the order of seconds), as compared with milliseconds for Pt.lm+lol Calculations were recently carried out with a detailed gas-phase reaction mechanisrn,lo6 and the selectivity to olefins vs. conversion of ethane was compared to experimental results' for various inlet ethane/oxygen/hydrogen compositions. The agreement was poor for all systems studied. The improvement in selectivity due to the addition of hydrogen in the feed was underpredicted, whereas the induction times were overpredicted. Selective alkane dehydrogenation in short-contact time catalytic reactors has thus been proposed to be a homogeneous-heterogeneous process, involving intimate coupling between gas-phase and surface-phase chemistry. Reliable detailed reaction schemes for both gas- and surface-phase chemistry are necessary in order to model this system.

'

9

Conclusions and Outlook

The experimental efforts of various researchers have shown that partial oxidation of CH4 over Rh-catalysts and the oxidative dehydrogenation of C2H6 over Pt-catalysts are viable and are potentially more attractive methods to syngas and ethylene production over the traditional methods of steam reforming and thermal cracking. Major advantages of short-contact time oxidation reactors include autothermal operation with significant economic and environmental impact, high yield, large throughput, small reactor sizes, a more suitable H2 :CO ratio for downstream utilization of syngas, and easy scale-up by increasing the reactor diameter (although better heat and mass transfer management would also be essential for this). These reactors run under rather severe conditions of high pressure and temperature, short-contact time, complex fluid flow, and with strong coupling of heat and mass transfer with surface- and possibly gas-phase kinetics. Such conditions impose serious challenges to both experiments and simulations. Despite the extensive experimental and simulations efforts, our understanding of such processes is far from being complete. In order to realize the potential of partial oxidation in short-contact time reactors, several factors, outlined above, need to be further studied. Some critical aspects include: 0

0

Kinetic studies of very fast, exothermic reactions - Novel reactor configurations are needed which approach realistic operating conditions while allowing deconvolution of gas-phase from surface chemistry and chemistry from heat and mass transfer. Reactor scale-up - A combined effort of experiments and modeling is needed to study reactor behavior at commercial scale operating conditions, including high pressures and large reactor sizes. The engineering aspects of temperature control and catalyst management also need to be considered.

134

Catalysis

Safety - The bifurcation behavior should be elucidated at high pressures for safe and optimal reactor operation. Again, features such as flammability limits and homogeneous ignition are sensitive to reactor geometry, and need to be carefully considered during scale-up to industrial conditions. Surface reaction mechanisms - The major theoretical impediment is systematic and rapid development of elementary surface reaction mechanisms that can be extrapolated, are thermodynamically consistent, and are validated with a wide range of types of experiments and conditions. The outcome of this effort will pave the way for efficient catalyst screening with minimal experimental information and rapid commercialization of such processes. Gas-phase reaction mechanisms - While the gas-phase reactions of small alkanes have been well studied for fuel-lean mixtures, the understanding of the chemistry of fuel-rich mixtures, typically of interest in partial oxidation, is at an embryonic stage. Knowledge and techniques developed in the combustion community can be an asset in overcoming this problem. Computational efficiency - The computational demand of high dimensionality fluid mechanics, multicomponent transport, and detailed chemistry codes prohibits parametric studies and reactor optimization. Techniques need to be developed to accelerate codes by up to two orders of magnitude in order for simulations to be an integral part of experimentation and optimization.

Acknowledgment

Acknowledgment is made to the Office of Naval Research with Dr G.D. Roy through a Young Investigator Award under contract number N00014-96-10786 and to the National Science Foundation (CAREER CTS-9702615) for support of this work.

References 1 2 3 4 5 6 7 8 9 10

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4: Partial Oxidation of Light Alkanes in Short Contact Time Microreactors

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5

Indirect Liquefaction - Where Do We Stand? BY YONGQING ZHANG AND BURTRON H. DAVIS

1

Introduction

Following its discovery by Fischer and Tropsch,’ the desirability of converting synthesis gas to hydrocarbons has undergone many cycles ranging from the view of an immediate, urgent need for commercialization, to periods of benign neglect. In the 1930s, because Germany did not have an internal source of petroleum, but did have a vast reserve of coal, interest in Fischer-Tropsch Synthesis (FTS) was great. This interest is illustrated by the visit in 1937 of Chancellor Adolf Hitler and Albert Speer to the leading coal research center in Germany and their discussions with its leader, F. Fischer.2 A feeling for the perceived concern for the petroleum supply situation at that time is represented in the f~llowing:~ ‘Jersey (now Exxon) sought still other methods for making motor fuels of high quality. For example, it investigated the Fischer-Tropsch hydrocarbon synthesis process, developed by Ruhrchemie, A.G., which converted brown coal into liquid fuel. In 1938 and 1939, patents for this process outside Germany were transferred by Ruhrchemie to Hydrocarbon Synthesis Corporation, in which Standard Oil Development (now Exxon) took 680 shares, Shell and Kellogg 425 each, and I.G. Farben 170. Both Great Britain and France considered the building of plants using this synthetic process for providing aviation gasoline, but they had been unable to accomplish anything definite by the time the war broke out in Europe in 1939.’ Germany developed a significant synfuels industry prior to and during WWII. However, their total peak production was small-scale in comparison to the US consumption of petroleum today. Furthermore, much of the German transportation fuel was derived from coal tars and direct coal liquefaction rather than from FTS. During WWII, the German FTS production was mainly conducted at atmospheric pressure using cobalt catalysts in fixed-bed reactors. Following WWII there was a perceived shortage of petroleum crude. In many countries, crash programs to develop a synfuels industry were initiated. US officials were among those who recognized the potential limits of petroleum crude. Major efforts were directed toward developing the FischerTropsch (F-T) synthesis processes in the US. One of these efforts was led by Catalysis, Volume 15 0The Royal Society of Chemistry, 2000 138

5: Indirect Liquefaction - Where Do We Stand?

139

Dobbie Keith, the engineer responsible for building the Oak Ridge Weapons facility, and his efforts led to the construction and operation of a commercialscale plant at Brownsville, T e ~ a s . While ~ - ~ the subsequent discovery of major petroleum reserves in the Middle East prevented this plant from being a commercial success, it introduced the use of large-scale oxygen plants for the production of syngas. The Brownsville plant utilized a fixed fluid-bed catalyst technology. At the time of the shutdown, the Brownsville plant was approaching a technically viable operation. On a smaller scale, pilot plants for both direct and indirect coal liquefaction were constructed and operated by the US Bureau of Mines in Louisiana, Missouri in 1 9 5 0 ~ The . ~ Louisiana Fischer-Tropsch plant could utilize a version of a slurry-bubble column reactor that was being tested in Germany, primarily in work led by Herbert Kolbel.*

2

Today's Commercial Operations

In general, three operations are needed for a commercial Fischer-Tropsch plant to produce transportation fuels: syngas generation, Fischer-Tropsch synthesis and conversion of light- (C,-C,) and heavy wax (about C18+) products to transportation fuel range or developing markets for these materials. For a process based on natural gas and heavy-product F-T synthesis, the capital cost is considered to be approximately in the range: syngas generation :F-T synthesis : hydrocracking = 50 : 35 : 15. For operation with coal as the carbon source, the capital cost is divided approximately: syngas generation : F-T synthesis = 67 :33. A recent Arthur D. Little, Inc. study provides estimated cost components of a gas-to-liquids unit (Figure 5. 1).9 Natural gas Pdudion

Oxidant

I

Syngas Conversion

Pmducl Upgrading

,

I

I

I Convemion &.20hbl

Upgrading

Total = $17.5Wbbl

-I

$3.mbl

Feedstock cost

I

Figure 5.1 Typical gas-to-liquids (GTL) product cost breakdown for a 100000 bbl day-'plant in North Field, Qatar (from ref: 9).

Catalysis

140

Currently there are three major commercial operating groups. Two of these are located in South Africa: the Sasol operation based on coal as the source of syngas and the Mossgas operation that is based on natural gas. The third commercial plant is operated by Shell in Bintulu, Malaysia and is based on natural gas. Rentech, Inc. claims to have operated '. . . at the commercial scale in 1992 and 1993'.'* However, the Rentech commercial operation at Pueblo, Colorado was to produce approximately 235 barrels of liquid hydrocarbons per day, and the failure to produce sufficient methane limited the operation to a few months. 2.1 South Africa - In South Africa the fear of being boycotted due to apartheid led to major efforts to develop for the country an independence from crude imports. The method used by the South Africans was the Fischer-Tropsch synthesis. This led first to the creation of Sasol and, more recently, Mossgas, and the introduction of a government subsidy. However, since crude is now readily available to South Africa, political and economic considerations have forced the South African Government to begin phasing out the subsidy for transportation fuels. With vision, Sasol anticipated the eventuality of losing its subsidies and has gradually shifted its emphasis from producing only transportation fuels to deriving a significant fraction of its operating profits from the sale of chemicals and petrochemical feedstocks. Today Sasol and Mossgas produce about 130 000 bbl day- corresponding to about 35% of the transportation fuels used in South Africa. While the South African production is impressive, its total output would only satisfy the needs of a medium-sized US petroleum refinery. Put in another perspective, the South African total production could replace only about 1% of the US consumption of crudes.

',

Sasol - Initially, Sasol utilized Arge fixed-bed reactors and iron catalysts. Subsequently they employed circulating fluidized bed reactors. Today the circulating fluidized bed reactors, combined with the fixed fluidized bed reactor that has recently been developed, account for about 90% of the Sasol production." The Arge reactors are being replaced by slurry phase bubble column reactors, and a 2400 bbl day - slurry reactor has operated successfully for almost ten years.12 Sasol still employs iron-based catalysts in all of its reactors. The production at Sasol is approximately 100000 bbl day-'. Of this total, about 10% is produced in the Arge and slurry reactors. Because of the very exothermic nature of the F-T synthesis (FTS), the introduction of the slurry reactor, with its ability to allow rapid heat transfer and isothermal operation, offers a decided advantage. Sasol features two types of product selectivities. Because the FTS is a 'polymerization reaction' using 'C1 monomers', the products fit an AndersonSchulz-Flory distribution (Figure 5.2). This distribution means that only methane can be produced as a pure product, with the ratio of chain termination to chain growth, a,being zero; all other values of a produce a mixture of products (Figure 5.3). The fluid-bed reactors must operate at an a value that is 2.2

5: Indirect Liquefaction - Where Do We Stand?

141

I 0'

10

10

ld

a

4

8

12

16

.

.

.

a

.

20

24

P(= carbon number) Figure 5.2 Theoretical Anderson-Schulz-Flory plots for U-values relevant for the F-T synthesis.

0.2

0.4

0.6

0.8

1.o

a

Figure 5.3 The Fischer-Tropsch reaction leads to a product distribution common to polymerizations from Cl monomers. The value of u (the probability for chain growth) depends on the relative rates of the propagation and termination steps.

142

Catalysis

Figure 5.4 Schematic showing sizes of Sasol fluid-bed reactors (compiled from T. Shingles and D. H. Jones, ChemSA, August 1986, 179- 182 and B. Jager, M.E. Dry, T. Shing1esandA.P. Steynberg, Catal. Lett., 1990, 7, 293).

small enough so that heavy waxes are not produced. Non-volatile waxes would condense on the catalyst particles, causing agglomeration so that the fluid state of the catalyst bed would be destroyed. Thus, the Sasol fluid-bed reactors are operated to produce 'chain-limiting' products. In general, the high temperature operation produces little wax, whereas the low temperature operation produces high molecular weight wax as the dominant product. Sasol operates four types of reactors at the commercial scale: tubular fixedbed (TFB), Synthol circulating fluid bed (CFB), Sasol Advanced Synthol (SAS; fixed fluid bed) and slurry phase (SP). During more than 40 years of operations, the Synthol circulating fluid-bed reactors have grown in size (Figure 5.4) until it would be difficult to scale up to a larger size. Sasol has recently introduced an advanced Synthol reactor (SAS) that makes use of conventional solid-gas fluidization that offers the advantage of requiring only about half the size of the CFB for the same production rate. Because of the decrease in reactor size and the elimination of the need to recirculate large masses of catalyst, the SAS reactor requires only about half the capital cost. The SAS is a much simpler reactor making it easier to operate, lowering operating cost by eliminating the need for catalyst recycle, reducing catalyst consumption by about 40%and reducing maintenance to only 15% of the CFB system. The SAS reactor accommodates higher gas loads and more cooling coils; this allows for larger capacity equipment with the advantages of economy of scale. Whereas the capacity of the CFB reactor in use at Sasol today is 6500 bbl day-', the SAS reactor has a capacity of 11 000 bbl day-' and it is anticipated that it has the potential for producing 20000 bbl day-' (Table 5.1)."

5: Indirect Liquefaction - Where Do We Stand?

143

Table 5.1 Sasol Fischer-Tropsch commercial reactors - capacities (bbl day- ') (from ref: 1 1 ) ~~~

Total installed capacity Capacity per reactor Potential per reactor

CFB

SAS

TFB

SP

1 10 000 6500 7500

1 1 000 1 1 000 20 000

3200 500-700 1550

2500 2500 10000

10000

Slurryqhase

Tubular fixed bed

0

195s

1987

1991

1993

Future

Figure 5.5 Reactor capacity increasefor LTFT reactor (from reJ 13).

A similar increase in size occurred with the tubular fixed-bed reactors as seen with SAS reactors (Figure 5.5) where the output increased from 500 bbl day-' reactor-' in 1955 to 1200 bbl day-' reactor-' in 1991.13 These increases in output may be viewed as evolutionary advances by operators of a commercial plant. In 1993, Sasol introduced, for them, a revolutionary advance in switching from the tubular fixed-bed reactor to a slurry bubble column reactor (SP) with 2400 bbl day-' reactor-' output. It is anticipated that the SP reactor can be scaled to 10000 bbl day-' reactor-', and even more. In a commercial operation, economic incentive drives operators to make scientific and engineering advances that lead to improvements in both the scale and the efficiencies of the operation. Sasol has made significant improvements in the energy efficiency of the Lurgi gasifiers during many years of operating them for the production of syngas from coal. These improvements have in turn allowed for the production of the products of the F-T process at a lower cost. The 16 conventional CFB Synthol reactors have been replaced with eight SAS reactors: four with a diameter of 8 m and four with a diameter of 10.5 m. The size differences were determined by the plant configuration. The first SAS reactor was commissioned in 1996. The Synthol reactors will be decommis-

144

Catalysis

Table 5.2 Selectivity (carbon basis) of Sasolprocesses (fromref 11) Product ~

LTFT

HTFT

~~~~~~

CH4 C2 to C4 olefins C2 to C4 paraffins Gasoline Middle distillate Heavy oils and waxes Water soluble oxygenates

4 4 4 18 19 48 3

7 24

6 36 12

9 6

sioned and left on-site for the immediate future.14 The use of SAS reactors rather than slurry reactors was also determined by the in-place downstream equipment. In general, Sasol operates their reactors in two modes: low-temperature (LTFT), high-alpha; and high-temperature (HTFT), low-alpha. The typical products from the LTFT are about 50% heavy oils and waxes (Table 5.2). The product make shown in Table 5.2 can be fitted to a two-alpha plot with the higher-alpha, heavier product polymerization accounting for about 75% of the total F-T products. The products from the high temperature operation are representative of a low-alpha product distribution, producing more of the light CI-C4 products (about 33% of the total hydrocarbon products) with the dominant fraction of the products in the transportation range. The ideal F-T situation is one in which the Cl-C4 products are decreased significantly (to less than 10%) and, at the same time, have the remaining products in the transportation fuel range. To date, this has not been accomplished in reasonable sized reactors that are operated at steady-state conditions. A significant fraction of the products from the fluid-bed reactors at Sasol are in the C I - C ~range and these are not in the boiling range of transportation fuels. In many locations this would be a severe disadvantage but apparently this is not the case in South Africa today. Since South Africa did not have abundant reserves of natural gas, methane and ethane were salable as town gas. The alkenes in the C2-C4 fraction have been utilized by Sasol to develop a strong position as a supplier of feedstocks for the petrochemical industry. In fact, many of the compounds being utilized as petrochemical feedstocks are valued at 2-4 times their value as transportation fuels. This has led Sasol to reserve Sasol I, using the slurry bubble column reactor and Arge fixed-bed reactors, for the production of chemical feedstocks, including paraffin wax. Until sufficient plants are on-stream to saturate the market, petrochemicals will allow Sasol to reap significant benefits from this part of their operation. Sasol has entered the petrochemical business with great determination and has been able to develop significant new business areas. For example, the Sasol

5: Indirect Liquefaction - Where Do We Stand?

145

Synthetic Fuels operating profit increased 28% for the year ending June 1997 over that of the previous year. Even without including the tax subsidy from the government, Sasol now is able to operate at a profit.

2.3 Mossgas - Mossgas, established in 1989, is a subsidiary of South African Parastatal Central Energy Fund.15 The life of one of the two initial gas fields (FA) used as the source of their feedstock was found to be significantly shorter than originally expected, and was expected to run out in 1997. Consultants were used to advise on the potential sale of Mossgas but, after reviewing the bids, the government decided against privatization. Subsequent efforts to use satellite facilities to tap nearby pockets of gas and improvements in compression in the field are expected to extend the life of FA by 3-5 years. Current estimates indicate that there will be sufficient gas from FA and its satellites to maintain synthetic fuels production at 30200 barrels of refined product per day until 2001. Production from another gas field is expected to provide gas to operate Mossgas until the end of 2005. The onshore Mossgas plant receives more than 5 million m3 day-' gas and associated condensate of more than 9000 barrels a day. The methane reforming plant is considered to be the world's largest. The reactors are of the circulating fluidized bed-type utilized by Sasol. The iron catalyst is manufactured on site. The plant consists of three identical trains. Since an iron catalyst is utilized, some water-gas shift will occur and, even with carbon dioxide recycle, it would appear that the plant would have a surplus of hydrogen. Mossgas has had an operating surplus after the first year of operation even before taking credit for the synlevy and tarriff protection. However, the stateowned plant is uneconomic. l 6 Long-term operations at Mossgas are dependent upon the development of economically recoverable gas reserves in addition to those now available to the company, the development of another source of the synthesis gas, or conversion to another operating process. 2.4 Shell - Shell has developed a process to produce middle distillate (SMDS). 7*1* The process consists of syngas generation, Fischer-Tropsch synthesis, the heart of the SMDS, and selective hydrocracking (Figure 5.6). The synthesis gas is generated using a version of the Shell methane reformer (SMR) that utilizes autothermal partial oxidation, first utilized by Shell in 1956 for the gasification of heavy resid. SMR produces a Hz :CO ratio that is close to 2 : 1, the consumption ratio in the F-T synthesis step using a cobaltbased catalyst. However, the complete balance requires that some hydrogen in excess of this ratio be generated, and this is done using conventional catalytic steam reforming of methane. The F-T synthesis is carried out in fixed-bed reactors configured similar to the multi-tubular Arge reactors that are utilized in commercial operations at Sasol. Shell acquired the F-T technology developed by Gulf-Badger that included work with promoted cobalt catalysts and both fixed-bed and fluid phase reactors. Gulf workers had shown the beneficial effects of the incorpora-

Catalysis

146 SYNTHESlS

SYNGAS

CONVERSION

MANUFACTURE

I

SGP

+ I

HEAVY

NA 0, E PLANTLJ GAS

PARAFFIN CONVERSON

HMU"

CH, t 10 2 - t C 0 2

+ 2H2 --c (-CH2-) + H 2 0

t

DISTILLATES

HEAVY PARAFFINS

Figure 5.6 Shell Middle Distillate Synthesis (SMDS): basic concept (HMU* =hydrogen manufacturing unit) (from ref. 17).

tion of ruthenium in a cobalt-based catalyst and the utilization of Group IIIB and IVB metal oxides. The Gulf workers also showed that an activation technique consisting of a reduction, re-oxidation and re-reduction (ROR) procedure for the promoted cobalt catalysts provided a higher conversion than just a single reduction. The catalyst utilized by Shell in its commercial operation is apparently a silica-supported cobalt catalyst promoted with ruthenium and zirconium oxide, a catalyst composition which presumably falls in the range covered by patents assigned to the former Gulf Oil Corporation. Shell workers provide a comparison of the products obtained from the classical catalyst formulations, those obtained as a result of new catalyst development, and the Shell catalysts. The data indicate that the Shell catalyst has an U-value in the 0.90-0.95 range, indicating a 5-10% probability for chain termination at any carbon number (Figure 5.7). Based upon the product distributions reported by Sasol workers, we calculate that Sasol, operating with a 'classical precipitated iron catalyst' in their slurry reactor which is designed to produce the same product distribution as they obtain in the Arge fixed-bed reactors, produces heavy products with an a-value of 0.95, and this high-alpha pathway produces about 75% of the products. Thus, it appears that even the 'classical iron catalysts' can be modified to produce very high avalues, and one should not limit their view to the need to utilize a cobalt catalyst to produce heavy wax F-T products. The distribution of the products from F-T synthesis is normally considered to follow an Anderson-Schulz-Flory distribution. However, the products produced in the early large-scale plants exhibited a 'two alpha' distribution (Figure 5.8). l 9 With the advent of more sophisticated analytical instruments, plots like those shown in Figure 5.8 have been extended to higher molecular weight products. Most authors show data that support the two (or more)

5: Indirect Liquefaction - Where Do We Stand?

147

PRODUCT. %w

0.75

0.85 0.90 0.95 PROBABILITY OF CHAIN GROWTH CLASSICAL CATALYST 0.80

o

-- ---

N E W CATALYST

SHELL

DEVELOPMENT

CATALYST

-----------t

Figure 5.7 Product distribution in Fischer-Tropsch synthesis (from reJ 17).

alpha product distribution although it is the paraffin fraction that is responsible for this (for example, see refs. 20 and 21). Shell workers,I7 on the other hand, report that, ‘in a few hundred FT synthesis experiments, conducted with various catalyst formulations and a range of operating conditions, the products follow a single ASF distribution with U-values in the range of 0.7 to 0.95’. However, the plot (Figure 5.9) that the authors provide to support this conclusion does not provide a full range of product distributions to cover both sides of the carbon number region where the shift from one U-value to the second occurs. The hydrocracking of F-T products can be effected using conventional hydrocracking processes and catalysts. Typical product distributions of a Fischer-Tropsch product and the products obtained following medium- or high-severity hydrocracking are illustrated in Figure 5.10.l 8 While the data in Figure 5. 10l8indicate that little wax needs to be converted, operations at alpha about 0.95 will produce up to about 50 wt% of product that needs to be hydrocracked to produce transportation fuels. The Shell plant in Malaysia was started up in 1993 and, despite encountering start-up problems, was operated successfully. The process is at the point where it could, in 1994, be ‘. . . considered as technically proved and, subject to local circumstances, commercially viable’.18 The plant was designed to have a production of about 10000 bbl day-’; however, the actual production from the plant over its lifetime has not been widely publicized.

Catalysis

148 100

10

I

0.1

0.02

I

1

5

I 9

I

I

13

17

I 21

CARBON NUMBER

Figure 5.8 Logarithmic plots of moles against carbon number. Hydrocarbons from the Schwarsheide tests compared with those from fluidized synthesis (from re$ 19).

A fire/explosion in the air separation plant damaged the SMDS plant and caused the plant to be shut down since December 25, 1997. The fire/explosion was due to accumulation of debris in the vaporizer of the air separation plant, presumably as a result of the severe forest fires in the vicinity of the plant. Current plans are to repair and modify the plant so that it is back in operation in 2000.22

3

Large Pilot/DemonstrationPlant Operations

3.1 Rheinpresussen-Koppers - A large (1 1.5 tons day- synthesis products; about 110 bbl day-') demonstration plant, which featured a slurry phase reactor, was operated at Meerbeck, Germany during 1952-1953 by Kolbel and co-workers (Figure 5.11).* The feedstock for this work was German brown coal. Work with laboratory and demonstration plants provided the most complete relations for estimation of the hydrodynamic properties of the slurry reactor at that time. For the demonstration plant, the synthesis gas was produced using a

149

5: Indirect Liquefaction - Where Do We Stand? MOLES IN PRODUCT, ARBITRARY UNITS

0

Figure 5.9

5

10

15

20

25

30

35 40 CARBON NUMBER

Typical carbon number distributions (from ref: 17).

PROOUCT CORWOSITION

10r

t h

%w

AFTER IGH-SEVERITY Y OROCRACKING

MEDIUM-SEW R ITY HYDROCRACKING

FISCHER-TROPSCH

0

10

20

30 40 CARBON NUMBER

Figure 5.10 Cmbon number distribution of Q Fischer-Tropsch product before and after selective hydrocracking (from ref: 18).

Catalysis

150

Froduet-water

A

Figure 5.1 1 Flow diagram of the large-scale demonstration plant (Rheinpreussen Process): A = freely separated primary products; B =final gas for recovery of low-boiling hydrocarbons; C = high-boiling primary products; a = compressor; b = gas meter; c = orifice plate; d = sampling intake; e = bubble volume reactor; f = steam collector; g = heat exchanger; h = separator; i = container for separated substances; k =pump; 1 = condenser; rn = CO2 expulsion unit; n =pressure cfilter; o = contact stirring container and container for suspension removed from reactor; p = centrifuge; q = mash oil (from re$ 8).

Koppers water gas generator, with some of the CO2 generated during synthesis being recycled. The reactor was 1.55 m in diameter and 8.6 m high (Figure 5.12). The reactor temperature was held constant by controlling the saturated steam pressure in the cooling system.23The height of the catalyst suspension (about 8 m) was maintained at a constant level either by collecting wax using an external, high-pressure filter (high alpha operation) or by adding higherboiling synthesis products to the reactor (low alpha operation). The synthesis gas enters the bottom of the reactor through a gas distributor with jets about 2-3 mm in diameter. The catalyst concentration, because of the micron particle size, was essentially constant from the bottom to the top of the reactor, and the optimum catalyst loading was reported to be about 10 wt% based upon iron. Complete removal of the heat of reaction of 1000 m3 synthesis gas per hour required less than 50 m3 of cooling pipe surface area. The space-time yield of C3+ products was 38.75 kg m-3 reactor volume h- I . The operators of the plant believed that the ultimate yield of a plant running at higher temperatures may be more than 125 kg m-3 reactor volume h-'.24 When operating in the gasoline (low-alpha) mode, it was reported that C3+ products accounted for 85% of the total products. This is surprisingly high

5: Indirect Liquefaction - Where Do We Stand?

151

Gas exit

Slurry

Gas distribution

Gas inlet

Figure 5.12 Reactor for liquid-phase synthesis (Rheinpreussen Process): 1 = reactor cylinder; 2 = cooling pipe register; 3 = liquid level regulator; 4 = steam drum (from ref: 8).

since most reports indicate that 85% C3+ synthesis products are only obtained with high-alpha operations. The work of Kolbel and co-workers was conducted using primarily unsupported iron catalysts. However, Kolbel and Ralek reported that, 'Basically all the catalysts which are suitable for the F-T synthesis can also be used in the liquid phase.' They also reported that, in contrast to other situations, it is desirable for the catalyst particles to break up during use since this leads to a better dispersion of the catalyst. Catalysts with high content of support are less suitable since they lead to unusually high viscosities. The deactivated catalyst recovered following filtration or centrifugation was reported to be regenerated repeatedly simply by oxidation after which the catalyst regained its initial activity. Furthermore, the regenerated catalyst led to the formation of only about half as much carbon as the fresh catalyst, thereby enhancing the catalyst lifetime and its activity. The formation of methane was also decreased with the reactivated catalyst. About 400 kg of hydrocarbon per kg iron was produced up to the time of catalyst regeneration. Because of limited compressor output at the plant, the gas flow for complete synthesis was limited to 3 m3 kg-' Feoh; however, the authors believed that this could be doubled with no problem. Even when operating at the high temperature needed for the gasoline mode of synthesis, less than 4% of the hydrocarbon product was methane plus ethane. This is a surprisingly low amount of methane and ethane. Based upon reported results and their data, Mobil workers25 reported a correlation between the amount of heavy wax and methane that are produced; the lower the heavy wax production the higher the methane make (Figure 5.13). It is noted that the

Catalysis

152

9

-t

60

IK w ( l n e m a n c + - )

A

X

*:

YtF

e

CT-256-7 CT-256 RWIS 1-S &son ct rf.(19!j4}. O t Circulation Chy (1981) Flrley a d Rav (19661 Kunugi ct 3.(ISSa),SIkri t t at- (1974) Schlesinger et al. 119Sll 0

0

2

4

..

6 8 Methane YicM. Wt X HC

1 10

12

Figure 5.13 Reactor wax yield versus methane yield with data from ref: 8 added (from ref: 25).

data point we added to the Mobil plot for the work of Kolbel and coworkers has a very low methane plus ethane yield compared to the work of others. The results obtained in a 6 liter laboratory slurry reactor were in agreement, except for reactor efficiency, with the 10000 liter capacity demonstration plant. Kolbel and Ackermann reported that the formation of higher-molecular weight products is favored by low reaction temperature, a high alkali content and by the chain-extending molecular build-up of the hydrocarbons present in the reactor or recycled into the reactor.26When much of the fraction boiling to 320°C was recycled back to the reactor about 80%of the recycled fraction was involved in molecular enlargement. This appears to be an exceptionally high fraction of incorporation of recycled olefins, based upon our work with adding I4C-labeled alkenes and alcohols to measure directly the extent of reincorporation. 27-37 It appears that the lack of gas chromatography for analysis of the lower molecular weight products and problems with mass balance during recycle of higher carbon number synthesis products caused very large errors in the mass balance for the operations summarized by Kolbel and Ralek. As noted, the ability to obtain at the same time both low methane and low wax production has not been reproduced by other workers. Furthermore, even a native German has, at best, great difficulty in discerning whether the product yields are based on CO+H2 feed or on C O + H 2 converted. The ReinpruessenKoppers work demonstrated the ability to run a large sized slurry reactor. However, the product distribution was not as reliable as desired.

5: Indirect Liquefaction - Where Do We Stand?

153

3.2 British Fuel Research Station - Work was started at the laboratory scale in 1949 and this led to the operation of a large plant with a gas throughput of 70 m3 h-'.38 The slurry reactor was 0.248 m in diameter and 8.5 m tall, with expanded catalyst slurry occupying the lower 6 m of the reactor (276 liter reaction volume). Gas was supplied through a single jet that had a 19 mm diameter. The operators indicated that the iron catalyst employed during much of the work was unsatisfactory. While the operators had solved the operational problems, the catalyst performance was still a problem when funding for the project was terminated in 1961. 3.3 US Bureau of Mines - During the 1940s the Bureau of Mines undertook extensive studies of direct and indirect coal liquefaction with the view of providing a source of transportation fuels that could supplement the anticipated dwindling supply of petr01eu1-n.~~ Surprisingly, these authors dismiss natural gas from consideration, '. .. since the supply of natural gas was subject to essentially the same limitations as crude oil'. The indirect coal liquefaction pilot plant at Louisiana, Missouri began operation in 1949. Much of the work was devoted to making improvements in the gasification process, the major cost associated with transportation fuels production. The reactor (Figure 5.14) operated at 450-525 "F (232-274 "C) and 300-350 psi but could operate to 600 psi. The reactor was loaded with about 7 tons (6356 kg) of iron catalyst placed on top of iron spheres ranging from 3/8 to 3 inch (0.95 to 7.6 cm) in diameter. These spheres supported the catalyst and acted to distribute the feed gas. The syngas and coolant fluid were mixed prior to entering the reactor and the fluid was fed at a rate to limit the temperature rise. Normally, about 50% of the gas entering the reactor was in the dissolved state. The initial catalyst volume [116 ft3 (3285 liter) of 8 to 18 mesh particles] expanded by 20-30% during feed gadcoolant fluid flow. Since coal was expected to be the source of the syngas, work focused on iron catalysts. The catalyst was reduced with hydrogen at high temperatures (up to 1000°C) and thus the precipitated, as well as the fused iron, catalysts had low activity. It was found that the pumps could handle some catalyst fines during circulation between the reactor and external wax recovery unit. Catalyst disintegration occurred in the demonstration plant to a much greater extent than it did in the smaller laboratory reactors. It was concluded that this was a result of the method of addition of the synthesis gas/coolant stream and that modifications to give better dispersion of the feed gas prior to entering the catalyst region of the reactor would have reduced catalyst disintegration. The operations were terminated in 1953 prior to a demonstration of this conclusion. The Bureau of Mines also operated slurry reactors ranging in size up to a 3 inch (0.076 m) diameter, 10 foot (3.05 m) tall reactor of about 15 liters volume that produced about 20 liters of product a day.40After 124 experiments lasting up to 4 months in duration, they concluded that: (1) catalyst concentrations of 50 to 500 grams of iron per liter of slurry are operable; (2) ratios of synthesis gas of 0.7 : 1.3 H2 : CO can be used; (3) pressures of 300 to 400 psi are preferred; (4) the range of space velocities are limited at one end by catalyst settling and

Catalysis

154 ptoducts 4

synthesisgas and cooling oil

=? For activating catalyst

Figure 5.14 Synthesis reactor at Bureau of Mines, Louisiana, Missouri plant (from ref: 39).

,

at the other by frothing; and (5) certain additives aid in keeping catalyst particles suspended. Although many successful runs were made in the slurry reactor, unpredictable erratic behavior was observed in about half of the tests. The erratic behavior usually resulted in low catalytic activity. While the factors determining the low activity were not clearly defined, one factor was due to catalyst settling out of the slurry and this was considered to be the major obstacle to the successful operation of a slurry reactor. While the work at the Bureau of Mines featured iron catalysts, they did also work with cobalt catalysts. They demonstrated that over a range of alpha values, the liquid and solid products obtained in their experiments agreed very well with those predicted from the alpha of the ASF distribution (Figure 5.15). During the 1970s’ energy crisis, Department of Energy (DOE) officials established a bubble column pilot plant at LaPorte, Texas. Initially this plant was used to develop a process to produce methanol, and later, in order to increase productivity, dimethyl ether. This effort was so successful that Air Products and Tennessee-Eastman have recently brought a commercial plant on-stream. In the 1980s, the plant was expanded to include the capability to conduct the Fischer-Tropsch ~ y n t h e s i sThe . ~ ~ facility has a capacity of about 1 ton day- of Fischer-Tropsch products. The reactor is 22.5 inches (0.57 m) in

5: Indirect Liquefaction - Where Do We Stand?

155

Figure 5.15 Correlation of theoretical (curves) and actual (points) product distributions obtained with cobalt Fischer-Tropsch catalysts (from ref: 40).

diameter and is 28.3 feet (8.6 m) tall. The slurry has a height of about 20 feet (6.1 m). The unit, apart from problems with wax-catalyst separation, has performed to expectations during the initial four runs which have included industrial partners together with the DOE. The first two runs used iron catalysts and the last two runs utilized a Shell proprietary catalyst. The first two runs utilized tangential (cross) flow filters and settling ex-situ the reactor to effect wax-slurry separation; this operation was not very successful because of the fines produced by catalyst attrition. Catalyst attrition during the third run generated sufficient fines to plug the proprietary separation system provided by Shell; however, it appears that the problems encountered in separations have been overcome so that the fourth run was completed as scheduled. 3.4 StanolindlCarthage Hydrocol - Work initiated by Dobie Keith at HRI ultimately led to the construction of a commercial-scale plant at Brownsville, Texas. This plant was designed to convert 90 million cubic feet of natural gas daily into 7000 bbl of petroleum products and 300000 pounds of chemicals. The plant was held by nine companies: The Texas Co. owned close to 50% with The Chicago Corp., Forest Oil Corp., Hydrocarbon Research (HRI), La Gloria Corp., Niagara Share Corp., Stone & Weber, Inc., United Gas Corp., and Western Natural Gas Co. holding the remainder. The total investment was about $50 million (early 1950s). The F-T plant utilized circulating fluidized bed technology and an iron

Catalysis

156

catalyst. The plant was the first to introduce large-scale air separation to obtain oxygen for the partial oxidation of natural gas to produce the synthesis gas. Despite many problems with the early operation of the plant, the operators progressed to the point where the plant was technically verified, and the later success of the process at Sasol confirmed this. However, the Hydrocol plant did not run at design capacity for extended lengths of time since the price of petroleum decreased to the point where the F-T product was not competitively priced. The plant was sold to Amoco in the 1950s who operated it as part of their chemicals production and then shut it down.

4

Standard Oil Co. (NewJersey) - Exxon

As indicated above, Exxon (Standard, New Jersey) has been active in the F-T synthesis area for many years. Their most recent activities have produced their Advanced Gas Conversion Technology (AGC-21) process that is based upon the F-T synthesis.42 The AGC-2 1 technology consists of three integrated processes: syngas generation, conversion, and product upgrading. Synthesis gas generation utilizes a novel fluid bed reactor system that combines partial oxidation and steam reforming reactions in a single large reactor that contains a fluidized bed of catalyst particles. The gasification technology builds upon Exxon’s long history of pioneering work in the application of fluid-bed technology, beginning with fluid-bed catalytic cracking in the 1930s. The refractory lined steel gas-generating vessel is operated at conditions which provide high thermal efficiency and offer economy of scale. The F-T synthesis process utilizes an advanced multi-phase reactor and a proprietary high performance catalyst. One of the reactor designs included in Exxon patents is similar to the Sasol Arge reactors except the cooling fluid is on the shell side and the catalyst slurry is contained in multiple tubes that can be operated to approach plug-flow conditions within each reactor tube (Figure 5. 16).43Exxon workers claim to have met the development challenges of heat removal, scaleup of the multi-phase fluid dynamics of the reactor, maintenance of the catalyst performance, and separation of the liquid wax from the catalyst slurry. The third part of the process involves mild hydroisomerization using conventional fixed-bed reactors that contain a proprietary Exxon catalyst. This work was on a massive scale with the cost through 1993 being more than $150 million. During the 1980-1994 period, more than 150 patents were issued to Exxon that cover a range of technology, and a similar pace of patents and publications continues today. The fluidized bed synthesis gas generation (FBSG) unit was constructed during 1989, and the Fischer-Tropsch unit was completed in 1990. The FBSG unit is 120 feet (36.0 m) tall with a reactor vessel that has a 5 inch (0.13 m) diameter. A variety of catalysts, prepared by commercial catalyst manufacturers, were tested as well as gas distributor configuration and materials of construction. During the last two years of a three year test period, the unit produced syngas for the F-T reactor on a sustained and fully

5: Indirect Liquefaction - Where Do We Stand?

157

16

I9

c .

46

Figure 5.16 Schematic of Exxon slurry phase reactor (from ref: 43).

integrated basis. Testing of the F-T hydrocarbon synthesis process was carried out in a reactor with a four foot (1.12 m) diameter. Productivity of the F-T reactor approached 200 bbl day-', a value greater than initially anticipated. This Exxon demonstration unit was at the productivity scale that Rentech has claimed as a commercial operation. Many of the Exxon patented F-T catalysts are based on promoted cobalt supported on titania catalyst, frequently with a second metal in the 15% range to improve and/or retain surface area of the support under reaction conditions. Since the only patent issued to date that provides operating data for this demonstration plant utilized a cobalt on titania catalyst, it is presumed that the proprietary commercial catalyst is some version of this catalyst. To date, the scientific and patent publications by Exxon workers are the most extensive of any group working in the Fischer-Tropsch area. For the work with iron catalysts, Exxon provided data for operations with both fixedbed and slurry (CSTR) reactors. However, with two or three exceptions, Exxon patents on cobalt catalysts present data that were obtained only in fixed-bed reactors.

158

Catalysis

The patents on Fischer-Tropsch synthesis obtained by Exxon could serve as a model for study of patent coverage. There must have been considerable technical/attorney cooperation or the patent attorneys were very well founded in the science and engineering aspects of F-T technology. In addition, the patent process must have been developed following a defined business plan for the commercialization of F-T technology. In many instances the open literature utilizes directly material from patents, but in other instances the open literature extends the coverage into areas not claimed in the patents. In the latter instances, it is presumed that patent coverage was sought but was not attainable. The dominant themes of the Exxon open literature summary papers are the impact of diffusion and of alkene reincorporation upon the rate and product selectivity of the F-T synthesis. The authors divide the effects into a kinetic factor and into factors that are controlled by the physical characteristics of the catalyst and/or reactor. The specific examples apply to cobalt and ruthenium catalysts although it is implied that the conclusions are general and would apply to iron and other catalysts. The theory and data in these papers are generated for the case of fixed bed reactors although it is shown or implied that the results are general rather than limited to fixed-bed reactors. The kinetic factor is very simple: all cobalt sites have the same activity and selectivity with perhaps the exception of Co-Ru supported on titania (Figure 5.17).44 The activity improvement by Re is indicated to be due to an increase in cobalt dispersion whereas the improvement by Ru is accomplished without an increase in cobalt dispersion. Thus, Ru is considered to preserve active sites of cobalt to a greater extent than Re or other promoters during the F-T synthesis. The other factor relates to diffusion effects upon product distribution by allowing for readsorption of olefins and their reincorporation into the synthesis. Exxon workers have developed a reaction-transport model wherein the effect of diffusional restrictions on readsorption rate is controlled by a dimensionless parameter an where Y nreflects the ratio of maximum diffusion rates to maximum reaction rates for olefins of carbon number, n, with the rate constant for first-order olefin readsorption turnover rate, kr,n,and the effective diffusivity, D,, of aolefins within the support pellets. The term Y nreflects the molecular properties of olefins of a given size and accounts for the observed effect of a rapid decrease with carbon number of the chain termination probability (Figure 5.18).44 The other term, X, includes the pellet radius, R,, the support void fraction, E , the Co sites per unit area, @co, and the average pore radius, rp. This term involves the structural properties of the support and the number of surface Co atoms that form and readsorb olefins. Diffusion-enhanced readsorption and incorporation of olefins increase with increasing carbon numbers up to about carbon number 15. Beyond carbon number 15, diffusion-inhibited chain growth impacts the product distribution. Thus, for small values of X the

5: Indirect Liquefaction - Where Do We Stand?

159

,' CO-RU

26

w

0

om

0.04

om

0.00

at

0.e

0.u

Cobalt Fractional Dispersion Figure 5.17 Effect of cobalt dispersion (ratio of surface Co to total Co atoms), support and alloying on FTS Co-time yields (mole CO convertedltotal g-atom Co-s). Reaction conditions: 473 K, 2000 kPa, H2 :CO = 2.05, Cf selectivity > 80%, 0.17 m m pellet size; A,titania; a, silica; m, alumina; A, titania; 0, silica (from ref 44).

reincorporation of olefins enhances the C5+ production up to a maximum after which further increases cause a decline in C5+ due to the diffussioninhibition of chain growth (Figure 5.19).44 The reverse trend is obtained for methane production. Exxon has an extensive number of patents covering F-T and their associated processes. Some of these patents make claims that are obvious whereas other patents claim areas where the novel advance or improvement is apparently obvious only to those with expert knowledge in patent law. Many of the earlier catalysts covered a variety of iron-based catalysts and the high productivity of these catalysts was demonstrated in both fixed-bed and stirred tank reactors. With two or three exceptions, patents covering cobalt catalysts provide activity data obtained only for fixed-bed reactors. The patent coverage for cobalt catalysts includes preparation, activation, reactor configurations, catalyst rejuvenation and processes for their use. In conformity with the theoretical developments, they have covered rim-loaded cobalt catalysts wherein techniques are claimed for the deposition of cobalt on the exterior rim of a support pellet, thereby limiting and/or controlling the diffusional effects. Patents cover catalyst activation in the reactor using fluid-bed techniques and then adding the start-up solvent. Another patent covers the reduction of a fresh catalyst with hydrogen or a hydrogen-containing gas in the presence of hydrocarbon liquids. A patent covers a 'super' activation treatment wherein the catalyst is initially activated in a fluid bed, passivated and then reactivated in a slurry

Catalysis

160 02

I

1

1

5

I

a n e a

0.15

C

= 0

c-

C

at

FF

0.0s

c 0

Figure 5.18 Bed residence time and carbon number effects on chain termination probabilities to olefins and paraffins (9.5% CO conversion, 2 s bed residence time; ColTiO2, 1 1.7% Co, 0.0 15 dispersion, 0.17 mm pellet size, 473 K,2000 kPa, H2: C0=2.1) (from ref: 44).

Figure 5.19 The effect of structural parameters (x) on FTS selectivity. Diffusionenhanced readsorption (----) and diffusion-inhibited chain growth (-) simulations and experimental data; a, dispersionisupport effects; A,pellet size variations; 0, eggshell thickness variations; (a) Cg+ selectivity; (b) CH4 selectivity (473 K, 2000 k h , HZ:CO=2.1, 5 5 4 5 % CO conversion) ( x values from eqn. 2 of ref: 44 with rp and R, in m, and Oco in surface Co atoms m-2) (from ref: 44).

5: Indirect Liquefaction - Where Do We Stand?

161

reactor. From a scientific point-of-view, it is difficult to appreciate the improvement of this technique over that of the ‘ROR’process patented by Gulf. Patented techniques for catalyst rejuvenation (regeneration) have been obtained that permit the rejuvenation to be carried out during reactor o p e r a t i ~ n .Another ~ ~ ? ~ ~patent also covers the rejuvenation of a catalyst using an external The operation of the slurry reactor pilot plant at Baton Rouge is described in patents; other patents cover various reactor configurations. K ~ r o des ~ ~ scribes a reactor in which the coolant is on the shell side of the reactor and the catalyst is dispersed in the slurry phase within small diameter tubes, each operated to approach plug-flow, in a configuration that resembles the Arge reactor except that the fixed catalyst bed is replaced by slurry (Figure 5.16).43 Benefits are also claimed for the operation of a reactor that includes one or more vertical down corner^.^^ Kim and fiat^^^ claim a processing scheme in which the synthesis gas contacts an olefin-producing catalyst (e.g. iron) so that conversion is limited and then the remaining syngas and highly olefinic hydrocarbon stream contacts a catalyst (e.g. cobalt) that will effect incorporation of olefins into the product. Obviously, space limits the coverage to only a few of the many patent claims, and those included are biased by the writer’s interest and understanding.

5

Pilot Scale Operations

Syntroleum - Syntroleum has been especially active in the business aspect of Fischer-Tropsch commercialization and they appear to have been a major factor in the renewed interest in this process. The company’s work, started in 1984, includes the operation of a 2 bbl day-’ pilot plant beginning in Tulsa in 1990 and a second-generation of unspecified type reactor that was introduced in 1996. They have announced, at times, fixed-bed, fluid-bed, HMX (hybrid multiphase technology) and horizontal reactors. It has been reported that a fluid-bed reactor utilizes a high-alpha cobalt catalyst that operates at 300-500 psig and 375-450°F (191-232°C) and has been operated in the 2 bbl day-’ pilot plant?* There have been several reports that ‘product yields and quality [have been] confirmed’ at the pilot plant scale but the basis of the confirmation are not reported. Syntroleum utilizes air rather than oxygen in an autothermal reforming step that employs a catalyst and a reactor of proprietary design. They do not recycle an appreciable amount of gases to the reformer. They claim that the reactor is simple enough so that it does not need large scale to be cost effective.s’ Work on chain limiting catalysts began in 1994, with partial funding from three oil companies.52Agee reports that, ‘Recent [1997] multiweek test runs in a fluid-bed reactor at the pilot plant yielded a product profile that indicates success. This catalyst promises several additional efficiencies to the process configuration, including: a lower operating pressure for the process, the use of higher capacity fluidized-bed reactors that cannot be 5.1

Catalysis

162 PRODUCT COMPOSITION, Xw

I O r

0

10

20

30

4

CARBON NUM8ER

Figure 5.20 The data are for Shell process (hydrocracking product) and the line is based on data taken by B. Davis from slide shown during presentation by M. Agee.

effectively used with the high-alpha, wax-producing catalyst, and elimination of a hydrocracking step.’52A plot of their reported chain-limiting product distribution, compared to Shell FT wax and the product produced with heavyhydrocracking, is provided in Figure 5.20. They also offer a proprietary, highalpha, highly-active cobalt catalyst with relatively low yields of methane (below 10%). They report that, ‘Pilot test runs with commercially manufactured batches have demonstrated the viability of the high-alpha catalyst system.’52 The product distribution reported by Syntroleum is, at first glance, indeed unique. Based upon the high-temperature F-T products in Table 5.1, Sasol produces products that meet the cut-off requirement shown for the Syntroleum data in Figure 5.20. However, there is a major difference between the lower carbon number yields reported by Syntroleum and by Sasol. The Syntroleum data show a methane product that is about 6 wt% of the hydrocarbon products, in reasonable agreement with Sasol data. However, Syntroleum shows the amount of C2-C4 products decreasing with carbon number to about 1.5 wt% for C4; this is in marked contrast to the Sasol data where C2-C4 products account for 30 wt% of the product and for Syntroleum this is about 8 wt%. One way Syntroleum could obtain such a distribution would be to oligomerize the 20 wt% C2-C4 olefins to produce a C10-C20hydrocarbon fraction. Mossgass does have a process that utilizes a silicalite catalyst to produce alkylate from the C2-C4 olefins; however, the oligomerization step is conducted in a separate reactor and at temperatures that are different from the F-T synthesis. Theoretically, it is possible to obtain a product distribution that

5: Indirect Liquefaction - Where Do We Stand?

163

Table 5.3 Agreements reported by Syntroleum Company

Date of agrmt. Specific agreement

Ref:

TEXACO

12130196

Nonexclusive ‘Master’ Licence Joint work to accelerate development of Syntroleum Process Build 2500 bbl day-’ plant outside US, site to be released 1st quarter 1998, start-up Fall, 1999

53 54

1 2110197

MARATHON

31 10197

Non-exclusive ‘Master’ Licence Marathon provides Syntroleum nonexclusive rights to proprietary technologies

55 56

ARC0

4110197

Non-exclusive ‘Master’ Licence with rights to use in broad geographic areas, including Alaska Joint development; Arc0 to construct 70 bbl day-’ pilot plant, start-up Fall, 1998 to demonstrate new reactor design

57 58

10124197

ENRON

02/24/98

Agreement to develop 8000 bbl day-’ speciality products plant, Wyoming; construction begins late 1998, operational in 2001 Non-exclusive volume license for Enron to license Syntroleum Process outside North America

59

YFP

08/04/97

Non-exclusive license agreement, rights to use Syntroleum Process in broad geographical locations outside North America

60

KERR-McGEE

02/05/98

Non-exclusive license agreement for KerrMcGee to use Syntroleum Process outside North America Syntroleum receive certain hydrocarbon processing technology and related patents from Kerr-McGee

61

resembles that shown for the Syntroleum data in Figure 5.20 but it requires two catalysts and, most likely, two reactors. Syntroleum routinely refers to a chain-limiting catalyst, implying that two catalyst formulations are not used and their simple process scheme does not indicate a product upgrading step. Syntroleum has announced agreements with several producing companies, and include Texaco, Marathon, Arco, Enron (world’s leading integrated natural gas and electricity company and owner of $23 billion energy-related assets), YFP (a wholly owned subsidiary of YFP Sociedad Anonima, Argentina’s largest oil company), and Kerr-McGee. These agreements are reported as non-exclusive ‘Master’ license agreements or non-exclusive license agreements. A synopsis of these arrangements are outlined in Table 5.3. The Arco-Syntroleum joint development project will include the construc-

164

Catalysis

tion of a 70 bbl day-’ pilot plant at Arco’s Cherry Point Refinery near Bellengham, Washington. ARC0 announced the completion and successful start-up of the pilot plant on July 28, 1999. In addition, they announced the successful integration of a new catalyst system and the advanced reactor design. This operation will demonstrate a new reactor design and will be used to produce synfuels ranging from diesel to heavy waxes. It was announced that Texaco intends to p a r t i ~ i p a t eThe . ~ ~ type of reactor was not announced but one would presume that it would involve the HMX hybrid multiphase technology. Arc0 does not appear to have made public their plans other than arrangements with Syntroleum. Enron and Syntroleum contributed $3 million (share supplied by each company was not announced) to fund a detailed engineering study, the purchase of land and other development costs for an 8000 bbl day-’ gas-toliquids (GTL) specialty product plant in Sweetwater County, Wyoming. Construction was scheduled to begin in late 1998 and the plant will be operational in 2001. The products will include synthetic lubricants, drilling fluids and liquid n-paraffins and the plant will produce 50 MW day-’ electricity for sale. Bateman Engineering is the engineering contractor for the project. In addition, a non-exclusive license was granted to allow Enron affiliates to acquire the right to license the Syntroleum Process outside of North America to produce liquid fuels. It is unlikely that the Sweetwater plant will be built in the near future. Because of more favorable gas prices, Syntroleum is now considering sites in northwestern Australia and South America for this plant. Final negotiations are expected to be completed in late 1999 (private communication). The Sweetwater plant has now been announced to be located in Australia. Criterion Catalyst Company and Syntroleum formed an alliance in which Criterion will manufacture and supply Syntroleum proprietary catalysts, presumably to those companies who operate plants using the Syntroleum Process. Syntroleum also obtained from Lyondell Petrochemical Company the right to license and sub-lease a new synthetic wax isomerization process based on Lyondell’s catalytic dewaxing process. The Lyondell dewaxing process was developed by Lyondell, Criterion Catalysts Co., and Zeolyst Int. Syntroleum has agreements with a number of engineering companies. Apparently the first of these agreements was with Brown & Root, a Colorado-based engineering company. In February, 1997, Brown & Root, a subsidiary of Halliburton Co., and Bateman Engineering, based in Denver, Colorado, were authorized to represent, market and license the Syntroleum Process to approved third parties for the production of synthetic fuels from natural gas.62 Shortly thereafter a project development agreement was announced for Bateman to be part of Syntroleum’s plan to build a series of ‘natural gas refineries’ in North and South America for production of synthetic lubricants, solvents and chemical feedstocks. In late 1997, it was announced that three firms (Brown & Root, Bateman and AMEC Process and Energy Limited), were named ‘Approved Process Design Provider’ to Syn troleum.

5: Indirect Liquefaction - Where Do We Stand?

165

b m m ROD vita Figure 5.21 (Left) Schematic of Sasol Synthol reactor" and (right) schematic of Syntroleum chain-limiting reactor (from reJ 63).

A recent international patent application, assigned to Syntroleum, describes a gas conversion process using a chain-limiting reactor.63 This patent teaches that, ". . . syngas is passed through a fluidized-bed catalyst in a chain-limiting reactor until the chain length of the hydrocarbon being absorbed [adsorbed?] on said catalyst reaches a chain length which is substantially equal to the chain length of the desired hydrocarbon product. This saturated portion of the catalyst is then removed from the reactor and is replaced with fresh catalyst. The 'saturated' catalyst is regenerated and recycled to the reactor." This patent does not contain examples with experimental data. This is both surprising and disappointing since Syntroleum claimed, in March 1997, to have operated a pilot plant in this chain-limiting mode during multi-week test runs in a fluid-bed reactor. The reactor described in the patent application does not appear to be unique from the circulating fluidized-bed reactor design used for years at Sasol (Figure 5.21). Fresh catalyst is added at the bottom of the reactor, contacts syngas as it rises in the reactor and the length of the chain of the hydrocarbon being absorbed on the catalyst continues to grow. The retention time of the catalyst in the reactor is such that the hydrocarbon chain on the catalyst will grow as the catalyst moves toward the top of the reactor and will attain the maximum desired length as the catalyst reaches the takeout point. It would appear that the gas flow-rate must be sufficient for the reactor to operate with the catalyst particles in, or nearly in, a plug-flow mode. This will require high gas flows and tall reactors. If the catalyst bed is well-mixed,

166

Catalysis

then catalyst particles throughout the bed will have the same average hydrocarbon composition. In the abstract, patent body, and the claims, the author uses absorb, rather than adsorb. As used in chemistrykhemical engineering terminology, adsorption means that the compound is present on the surface of the catalyst where absorbed means that the compound is taken into the bulk. It is not possible for the hydrocarbon product to become a part of the bulk of the cobalt or support; thus, the only way that absorb can have meaning in the context of the patent is that the hydrocarbon is retained within the catalyst porosity. Likewise, the patent does not define ‘saturated’. It appears that there is a contradiction within the description of the claimed operation of the reactor. Claim 1 is ‘A Fischer-Tropsch process carried out in a chain-limiting reactor for converting syngas to a desired hydrocarbon product having a desired chain length, said process comprising: passing said syngas through a fluidized bed of catalyst in said chain-limiting reactor to convert said syngas to said hydrocarbon product; operating said reactor at temperature and pressure at which said hydrocarbon product remains in its vapor phase while in said reactor; removing said hydrocarbon product while in its vapor phase in said reactor; continuously removing catalyst from said chain-limiting reactor as the chain length of any hydrocarbon product absorbed on said catalyst reaches said desired chain length for said hydrocarbon product; and continuously supplying fresh catalyst in said chain-limiting reactor at the same rate as that at which said catalyst is removed.’ The authors claim that the hydrocarbon product remains in the vapor phase while in the reactor. If, as claimed, the hydrocarbon product is in the vapor phase in the reactor, there will therefore be no ‘absorbed’ hydrocarbon on the catalyst. It will therefore be impossible to remove catalyst at the top of the catalyst bed that contains ‘absorbed’ hydrocarbon of the desired chain length. Even if it is granted that some of the hydrocarbon is ‘absorbed’ with the catalyst, the product using a typical F-T catalyst must follow the typical Anderson-Schulz-Flory distribution. This is illustrated in the following. Take the simpler case of plug-flow of catalyst particles with the residence time adequate to make the longest product a C18 hydrocarbon. Chain termination will occur for some of the catalyst particles as they travel the length of the catalyst bed. Furthermore, the chain termination probability will be that which gives the ASF distribution. For example, assume that a catalyst particle that enters the reactor has a chain growing that is terminated at methane (Cl). When this happens, there will be a free site on the catalyst where CO will adsorb, initiate a new chain that will grow as the catalyst particle passes along the catalyst bed, and the second chain will follow the ASF plot. The only way that the products using the chain-limiting reactor can deviate from the ASF plot that would be obtained in a typical reactor is if a catalyst is utilized that has a termination probability of zero. Stated another

5: Indirect Liquefaction - Where Do We Stand?

167

c

Yn m

a

lo-7

2 4

6

(I

10 I2 H

16 I 6 2 0 2 2 2 4

n Figure 5.22 Effect of time on the distribution of Fischer-Tropsch products obtained from a well-stirred slurry reactor when the catalyst is characterized by a single value of a (V,= 15 cm3; V G = 300 cm3; Go=400 STP cm3 min-'; Mw, = 350 g mol- I; PL = 0.7 g ~ r n - ~a= ; 0.7) (from ref: 64).

way, there will be vapor-liquid (vapor-absorbed phase) equilibrium for each carbon number hydrocarbon. Mass balance requires the hydrocarbon product of each carbon number leaving the reactor to be the sum of the vapor and adsorbed product that is produced by the ASF mechanism. Whether the product leaves the reactor in the vapor or absorbed phase is irrelevant as far as the mass balanced product distribution is concerned. The only difference that the reactor could make, whether operated in the backmixed or plug-flow mode, is that which is well defined in any reactor text where at intermediate conversion levels, a higher concentration of an intermediate product (e.g. alkenes in this case) could be higher in the plug-flow reactor than in the backmixed reactor. It appears that the ASF products will not be formed only if the catalyst residence time is sufficiently short or if appreciable hydrocracking occurs. We conclude that the typical Fischer-Tropsch catalyst will produce an ASF product distribution in any reactor; any deviation from this distribution will be due to obtaining product analyses that do not represent a proper mass balance. For example, one can easily obtain a higher fraction of lower hydrocarbons (chain-limiting operation) due to accumulation of heavy hydrocarbon products in the reactor; this has been illustrated by, for example, Dictor and BellM (Figure 5.22). Data obtained at early times on-stream in the runs conducted by Dictor and Bell readily fit the definition of chain-limiting

Catalysis

168

---- -- -

Chaln Llmltlng Hlgh Alpha

Figure 5.23 Product distribution of chain limiting and high alpha product catalyst (redrawn from a Syntroleum handout).

reactor but this is true only if one ignores the heavier products that are temporarily retained within the reactor. The Syntroleum patent application clearly teaches that the reactor is operated so that the hydrocarbon products are in the vapor phase. This requires that the catalyst used in the chain-limiting reactor has a chain termination probability that is greater than zero. It is therefore concluded that the reactors described in the patent cannot be operated to effect an operation with a chain-limiting hydrocarbon product distribution in the sense of deviating from the ASF distribution or that the patent application does not reveal sufficient information for one versed in the art to be able to reproduce the art described in the patent. The products from a Syntroleum chain limiting catalyst are represented in Figure 5.23 together with their representation of a normal high-alpha catalyst. The limitation of the high molecular weight products for the Syntroleum catalyst is typical of a low-alpha catalyst. However, there is a vast difference between the products shown for the Syntroleum chain-limiting catalyst and the products produced by a conventional low-alpha catalyst. Whereas the volume percentage of products decrease from C1 to about C4 for the Syntroleum catalyst before they begin to increase, the volume percent produced by a typical low-alpha catalyst would increase from C1 to about Cq. For example, an average molecular weight of C4 would be obtained for a catalyst with an alpha of about 0.7. At Mossgas, a silicalite catalyst is utilized to carry out oligomerization of the low molecular weight products to produce gasoline and diesel transportation products. Thus, it is anticipated that the Mossgas products, when the alkenes have been converted to liquid transportation fuels,

5: Indirect Liquefaction - Where Do We Stand?

169

will resemble the product distribution shown for the Syntroleum chain limiting catalyst. A problem that one encounters is the definition of a chain-limiting catalyst. To most investigators, a chain limiting catalyst is one that deviates from the normal Anderson-Schulz-Flory distribution. Furthermore, the distribution deviates negatively from the ASF at some carbon number range of products. In other words, the products follow the ASF distribution up to some carbon number, but above this carbon number the products are present at a lower value than that corresponding to ASF, or are zero. This distribution has also been designated as being produced by a catalyst with a ‘cut-off’. It appears that the groups currently using chain-limiting catalysts do not mean a deviation from the ASF distribution. Rather, it appears that a chain-limiting catalyst, as used by Syntroleum and Carbon Resources, is one that produces a product distribution whose vapor pressure is such that essentially all products are in the vapor phase at the reaction temperature. This type of chain-limiting catalyst has been utilized at the commercial scale for about 50 years. For the circulating fluid-bed operation, liquid phase products cannot be tolerated. The presence of liquid causes the catalyst particles to agglomerate and plug up the reactor because they become too large for the gas to maintain them in a fluid phase. Thus, catalyst agglomeration was one of the early problems encountered by the operators of the Brownsville, Texas and the Sasol fluid-bed reactors. By adjustment of process conditions and catalyst compositions, these plants were able to operate their fluid-bed reactors in the absence of a liquid phase. At Brownsville and Sasol, the catalyst was chain-limiting because it produced such a small amount of high molecular weight product. However, the products follow the distribution expected for a ‘normal low-alpha’ catalyst that follows the normal ASF mechanism. Thus, on the basis of limiting the amount of high molecular weight liquid products, Sasol has operated reactors that contain chain-limiting catalysts, and the output from these catalysts accounts for about 90% of their production (that is, 90000+ barrels per day). The distribution shown by Syntroleum is that expected for the Sasol catalyst with the exception of a different distribution at the lower molecular weight products. The lower molecular weight products represented by the Syntroleum curve in Figure 5.23 deviate significantly from that of the Sasol operation, but the higher molecular weight products do not. The higher molecular weight products of the Syntroleum distribution curve are expected if the catalyst produces products representative of a low (e.g. 0.70) alpha catalyst. Furthermore, the Syntroleum curve is consistent with the conversion of the low molecular weight alkene products to liquid products in the range of about CgCI7.This type of process is apparently operated by Mossgas where they use a pentasil catalyst to oligomerize low molecular weight alkenes and, while they may not produce a product distribution curve that is identical to the one shown in Figure 5.23 to represent the Syntroleum products, the Mossgas product distribution is expected to resemble the one shown for the Syntroleum products. Salomon Brothers analyst Paul Ting asserts that, ‘They [Syntroleum] are

170

Catalysis

creating the gas-industry equivalent of the PC standard’. If Mr Ting refers to business operations with the creation of publicity and obtaining agreements for potential commercial operations, Syntroleum indeed is creating a ‘PC standard’. If, on the other hand, Mr Ting is referring to a definition of the scientific and technical aspects of Fischer-Tropsch synthesis, he is certainly incorrect since Exxon and Shell, for example, have provided a much more complete and detailed description of Fischer-Tropsch synthesis in both their scientific publications and in their patents. 5.2 Gulf Oil - The technology developed by Gulf was transferred to Chevron following a merger of the two companies. Later, Chevron transferredsold the technology rights to Shell and Shell appears to have incorporated much of the Gulf technology into the commercial plant in Malaysia. Gulf workers prepared a series of catalysts that demonstrated a beneficial effect of a minor amount of ruthenium on a catalyst containing a major amount of cobalt for the low pressure synthesis of higher hydrocarbons from synthesis gas.65 It was claimed that the incorporation of ruthenium increased not only the activity but also the average molecular weight of the product. However, problems with the analysis of products were probably responsible for the heavier products obtained with the catalyst containing ruthenium. Gulf workers found that a catalyst that was subjected to an initial reduction followed by an oxidation and then a second reduction (ROR process) yielded a catalyst with a higher activity than when the same starting material was activated using only the same initial reduction procedure.66 Gulf workers also patented a two-stage process67 wherein the first stage contained their cobalt catalyst, promoted with a Group IIIB or IVB metal oxide and ruthenium, and was used to generate Fischer-Tropsch synthesis products. The C5-C9 fraction of the F-T products was then converted to high octane gasoline using a typical naphtha reforming catalyst. While it was stated that the reactor type utilized was not a critical factor, a fixed-bed, with gas downflow, was preferred. Another variation of the two-stage process was also patented68 wherein the Fischer-Tropsch products were converted to highly aromatic high-octane gasoline range products using a silicalite ZSM-5 type catalyst in the second stage rather than the naphtha reforming catalyst. Gulf workers demonstrated their process and their promoted catalyst at the commercial scale of operation. With the fixed bed reactor, one must use small diameter tubes in order to effect sufficient heat transfer to maintain reasonable temperatures along the length of the reactor tube. Thus, Gulf and Badger operated a 40 foot tube reactor (two 20 foot sections in series) for extended periods of time and obtained data under conditions where the exotherm was less than 5 O F (2.8 “C) along the reactor (Figure 5.24).69These data verified the Gulf-Badger process and catalyst at the commercial level since a bundle of identical tubes would be utilized in a commercial plant. The takeover of Gulf by Chevron eliminated the urgency to commercialize the Gulf-Badger process, and the rights to it were eventually transferred to Shell.

5: Indirect Liquefaction - Where Do We Stand?

171

COOLINQ WATER

c

L

-3

MIXER

[

QAS OUT

t

O A S OUT

Figure 5.24 Schematic diagram of the pilot plant remote-gas-to-dieseljixed-bed process (fromre$ 69).

5.3 Rentech - As noted above, Rentech reports that they have operated during 1982- 1983 a commercial-scale gas-to-liquids facility. This facility utilized a Pueblo, Colorado landfill to produce gas sufficient for a plant with a capacity of about 235 bbl day-' of liquid hydrocarbons. This plant has a capacity that is about the same as the reactor for the Exxon demonstration plant. It was soon learned that the landfill was unable to produce sufficient gas and the operation was discontinued within three months of the attainment of full-scale product ion. Rentech has nine patents to cover their technology. The research topics for all of these patents are either the same or very similar and provide a concise summary of Fischer-Tropsch synthesis that can serve as excellent introductory reading for the novice in this area. The claims are more modest in coverage than the body of the patents. Furthermore, it is sometimes difficult to discern when the cited examples are based on theoretical calculations and when they are based upon data obtained by Rentech. Rentech has operated a 6 inch (0.15 m) diameter, 8 foot (2.54 m) tall slurry reactor with an iron catalyst. They report that they have designed, built and tested two 6 foot (1 -83 m) diameter slurry reactors and have produced catalyst batches of more than 7.5 tons (6810 kg) for use in these reactor^.^' The waxslurry separation, as reported in 1992, was accomplished using an external

I72

Catalysis

cross-flow filter. Using this technique, separation was still a problem. They apparently utilized later in their pilot plant studies an external separator that involved gravity settling to concentrate a catalyst-wax slurry for return to the bottom of the reactor; this approach was apparently very similar in design to the one Mobil Oil utilized in the work they conducted for their DOE contracts. On June 16,1999 an agreement was signed to study, during the next 8 months, the integration of Rentech's Fischer-Tropsch and Texaco's gasification processes. The research will concentrate on the use of petroleum coke, heavy resid and/or coal to generate synthesis gas. Rentech's iron catalyst technology will be utilized. Rentech also signed a memorandum of understanding for BC Projectos, Ltd., a Brazil engineering company, to be Rentech's engineering firm of record. Rentech is also trying to interest owners of existing methanol plants, many operating below capacity or moth-balled because of low methanol prices, to add the Fischer-Tropsch plants to utilize existing gasification capacity. Rentech's 280 bbl day-' plant was transferred to Donyi Polo Petrochemicals and shipped from Pueblo, Colorado to Arunachal Pradesh, India. The plant will process natural gas from the Kunchai field that is currently being flared. The plant is expected to be commissioned in 2000. United Catalysts, Inc. (India) will manufacture the catalyst. Chinese Studies - The Chinese government is supporting several projects aimed at developing technology, including Fischer-Tropsch synthesis, to convert coal to transportation fuels. The Institute of Coal Chemistry, Chinese Academy of Science, located in Taiyuan, is responsible for the large-scale reactor studies. During the past few years these workers have completed a number of studies directed toward understanding of the science and engineering aspects of FTS and with the operation of larger pilot plants. They have utilized fixed-bed tubular (50 x 4500 mm) reactors with a total capacity of 2-5 litres with iron catalysts. A pilot plant of 100 ton year-' (25 x 4000 mm multitube) and a demonstration plant of 2000 ton year- (32 x 7000 mm multi-tube) have been tested. The F-T reactor was used as the first stage with the second stage being a ZSM-5 catalytic process designed to convert the F-T products to gasoline range transportation fuels. They also have built and tested a slurry bubble column reactor (40 x 4500 mm reactor). This reactor has been utilized with an iron-based catalyst and has been employed in conjunction with the ZSM-5 catalytic second stage process. Currently, large fixed-bed reactors are being utilized in studies with ultrafine iron catalysts, magnesium-iron catalysts, and supercritical operation with an iron ~ a t a l y s t . ~ ' 5.4

'

Mobil Oil - During the 1980s, Mobil Oil workers constructed and operated, with DOE funding, a slurry bubble column reactor as the first stage and Mobil's ZSM-5 zeolite catalytic conversion of the F-T products to gasoline range fuels as the second stage.25The slurry bubble column reactor was 5 cm (i.d.) x 7.6 m and in normal operation utilized about 1.5 kg of catalyst. The reactor utilized an external wax filter assembly (Figure 5.25). Thirteen runs were made, the first three with a low-alpha iron catalyst and 5.5

5: Indirect Liquefaction - Where Do We Stand?

173

rJ

-06-

V-6

v-t

B Ib

To Wax Recefver

Figure 5.25 Schematic of the external wax$lter assembly (from ref 25).

switching to the high-alpha mode during the third run and the runs thereafter. During the course of these runs, many operational problems were encountered. The diameter of the reactor is at, or perhaps below, the diameter now considered to be free of significant wall-effects which impact the measured kinetics. In spite of any operational problems, these studies provide the most detailed open-literature data for slurry phase operations and have been the basis of all, or nearly all, of the economic and process studies conducted by DOE contractors, such as Bechtel and Mitre. In most instances, the F-T data used in these evaluations are from the last Mobil run using a high-alpha iron catalyst. A summary of these results is given by Fox and Tam.72 'Mobil gave detailed product breakdowns which exhibit a characteristic break in the Schulz-Flory relationship at a carbon number between 20 and 25. In addition, methane production is several times what it would be if the correlation line for carbon numbers 2 to 20 were simply extrapolated back to methane . . . . The a values are constant over the specified range of carbon numbers and the summed weight fractions for each range are adjusted such that the total yield adds up to 1.0. There is a sharp break in the slope of the product distribution curve at C20-C25.Vapor-liquid equilibrium predicts that the C22 component is split almost evenly between vapor and liquid phases. Since the liquid phase remains in the reactor much longer than the vapor phase, it is believed that the longer liquid residence time is responsible for the sharp break in slope, at about C22,to a higher value of a.'

Catalysis

174

I

I

I 1

I I

I I

I

0.1

-- 7

:::56

0 . 01

w f / n 0.0 101

9.49

1

0.00 101 0.000 01 0.0000 10 1

0

10

I

1

I

1

I

I

I

I

20

30

40

50

60

Carbon Number, n

Figure 5.26 Design Schulz-Flory plots based on Mobil Oil data (from ref: 72).

Fox and Tam conclude that for Mobil's F-T catalyst system and reactor configuration, setting the operating temperature fixes the entire carbon number distribution (Figure 5.26). The Mobil product distribution data stand in sharp contrast to the ones reported by Kolbel and Ralek for the operation of the Rheinpruessen-Koppers plant. Whereas very low methane and wax yields were reported for the German plant, Mobil obtained a higher methane make than expected from the ASF distribution. With an iron catalyst the water-gas shift reaction can make an important contribution to the overall consumption of CO. The Mobil operation should be viewed to be representative of processing where the WGS is approaching its equilibrium value. Thus, Mobil has reported WGS ratios that are typically 25 whereas the value at equilibrium is expected to be in the range of 60.

6

Process and Economic Evaluations

The following will present typical reports rather than attempt to be comprehensive. G r e g ~ summarizes r~~ the economic advantage of the F-T synthesis that is utilized within a petroleum refinery. He concludes that the F-T distillate fuels have exceptional properties but that the F-T naphtha is difficult to reform into high octane gasoline. When F-T is viewed as a supplement rather than an alternative to petroleum, the refiner would likely utilize the lighter F-T products in petrochemical applications rather than fuel. The lighter F-T products have been utilized as gasoline fuel in South Africa but until now they have been able to add lead to boost octane number and therefore do not have to depend on naphtha reforming alone to boost octane to acceptable levels. As more and more F-T plants come on-stream, the ability to utilize the lighter fractions as petrochemical feedstocks will become plagued with overcapacity.

5: Indirect Liquefaction - Where Do We Stand?

175

Most of the process and economic assessments carried out by the US DOE have been based on coal as the feedstock. Only recently have these DOE studies been based upon natural gas.74 In the past few years, the DOE effort has concentrated upon producing F-T products within an electric power generating plant or as a plant that co-produces power. Lange and Tijm75have considered many schemes for converting methane to liquid hydrocarbon fuels. They conclude that F-T transportation fuels are not competitive at oil prices less than U S 2 0 per bbl. These authors state that while natural gas offers the advantage of a lower cost feedstock, compared to petroleum, the higher capital cost associated with GTL offsets the feedstock cost and makes GTL unattractive at prices less than about $20 per bbl. They indicate that the analyses of numerous fuel manufacturing plants show that the capital cost correlates with the energy loss as well as with the heat and momentum transfer duty of the process and process segments. These authors conclude, perhaps not surprisingly, that when the various methane conversion schemes are compared, the Shell SMDS process is the cheapest scheme proposed at that time (1996). The authors found that a direct scale-up to a 50000 bbl day-' plant from that of a 10000 bbl day-' plant results in a reduction of capital cost sufficient to lower the cost by about $5 per bbl. However, formidable difficulties are faced in increasing the size of plant components by this five-fold amount. G r a d a ~ s i ,based ~ ~ on his selection of recent literature, concluded that a Fischer-Tropsch plant can generate close to $6 billion in cash flow over the lifetime of the plant, and would have product sales revenue of $200-400 million per year. However, the initial investment will be at risk for at least six years before revenues have returned capital cost, taxes and operating expenses.

7

Potential Commercial Operations

Currently, the interest in monetizing remote natural gas is very high. In some instances this interest is driven by environmental considerations. Methane is at least 30 times as potent for developing the 'greenhouse effect' as carbon dioxide is. Thus, there is the expectation that very severe restrictions will be developed throughout the world that will limit the release of natural gas during recovery of petroleum crude. Even today, some regions require measures such as reinjection of natural gas that is obtained in association with crude. Limitations on carbon dioxide release will also place restrictions on flaring of associated natural gas. In addition, producing countries are beginning to place charges on the total carbon recovered, and not just the petroleum crude. Syngas generation, even with a natural gas feedstock, comprises about 50% of the cost of a Fischer-Tropsch operation. There has been a renewed interest in adapting the rapidly advancing membrane technology to the production of syngas at a much reduced cost compared to partial oxidation or steam reforming. Recently it was announced that Phillips Petroleum would join a syngas alliance77made up of Sasol, Amoco, British Petroleum, Praxair and

176

Catalysis

Statoil. Each of the members will contribute technical expertise, R&D and funding to the project. The original alliance was announced7*to 'expeditiously and aggressively move this [membrane] technology toward commercialization'. The original alliance companies announced that they have enlisted support from universities and government laboratories and would consider other companies or groups. DOE has announced an eight year project to develop technology to convert natural gas to liquid for transport through conventional pipelines.79 Air Products' proposal was chosen to develop ceramic membranes that could sharply reduce the cost of generating synthesis gas. In the first two years, members will develop membrane materials, catalysts and other key components. During the next three years the technology will be scaled up to a 12 Mcfd (thousand cubic feet per day) experimental unit. During the final two year phase the technology will be scaled to 15 MMcfd (million cubic feet per day) precommercial unit at Air Products' industrial gas complex at LaPorte, Texas - the location of the DOE slurry phase F-T reactor. In addition to Air Products, the team consists of Babcock & Wilcox, Cerametec, Eltron Research, ARCO, Argonne National Lab., Pacific Northwest National Lab., Penn State University and the University of Pennsylvania. Several of the companies currently looking to develop commercial FischerTropsch plants are not the well-known petroleum companies. Reema International Corporation, located in Denver, Colorado, signed on December 17, 1997 a Memorandum of Understanding with the National Gas Company of Trinidad and Tobago Limited for a project that would convert gas to liquid transportation fuels. Reema expects to finance, build and operate a plant costing about US275 million that will convert about 100 million ft3 day-' of natural gas into about 10000 bbl day-' of high quality transportation fuel. It is projected that the plant design and construction will require about three years. Exxon announced feasibility studies with Qatar's state-owned oil company to build a commercial plant at the 50 000 to 100 000 bbl day-' scale. The final decision on building this plant appears to have been delayed from the initial date expected for the decision.80Apparently the Exxon affiliate wants Qatar to sell the natural gas for less than they now sell it to petrochemical producers, which is believed to be about US$O.SO/Mcf (thousand cubic feet). Exxon expects to spend several million dollars in the next year to review an Alaskaspecific GTL application of its AGC-21 process. The potential Alaskan plant would be at the 50000 bbl day-' scale. This plant would have to compete for natural gas feedstock with liquified natural gas (LNG) production. An advantage of the GTL process is that the product could be transported through the existing oil pipeline and in this way extend the life of the pipeline as crude production declines. Sasol has been about as active as Syntroleum in forming joint ventures. Sasol together with Qatar General Petroleum Corp. (QGPC) and Phillips Petroleum have agreed to assess the possibility of a joint venture that would lead to a F-T plant with a 20000 bbl day-' facility.8' In April, 1998 Chevron

5: Indirect Liquefaction - Where Do We Stand?

177

1oa

80

2

. I

>

. I

60

c,

0

a

.-a?> 4-

I

40

a?

U

20

0

Gascat**

SiO, Supported Catalyst

Ti4-Suppor ted

Figure 5.27 Activity of cobalt-based catalysts. Reaction conditions; 220 "C, pressure = 450 psig. (** = Energy International catalyst) (from ref: 84).

and Sasol announced that they had reached agreement to pool their resources to begin design and engineering for construction of a 20 000 bbl day- GTL products plant in NigeriaqS2Chevron views this to include the possibility of expanding this new technology to other worldwide applications. Statoil and Sasol have formed an alliance for the conversion of natural gas to synthetic crude oil and liquid fuels by Fischer-Tropsch technology. Sasol's technology for the conversion of gas to fuels and Statoil's offshore and floating production technology for oil and gas are mutually complementary and are a basis for the alliance. Statoil prefers to cooperate rather than continue the development of their own Fischer-Tropsch technology, which had been underway for several years. The two companies will cooperate exclusively in developing floating and offshore applications and on a case-by-case basis for other applications. Wexford P.A. Syncrude Ltd. has, together with Bechtel Corp., made a design/economic assessment of a once-through natural gas Fischer-Tropsch plant with co-production of power.74 Presumably, the detailed design and scale-up technology that serves as the basis of this study is proprietary at this time, and belongs to Syncrude. The company has patented a catalyst that is highly selective for production of liquid hydrocarbons that contains cobalt and molybdenum or molybdenum and zirconium on an inorganic support .83 Energy International, a subsidiary of Williams Field Services, Inc., Tulsa, Oklahoma, announced their GasCat F-T process that is claimed to achieve higher productivity and improved catalytic performance (Figure 5.27) compared to existing Fischer-Tropsch technique^.^^ The process utilizes an

'

178

Catalysis

alumina-supported cobalt catalyst that shows significantly higher activities, among several advantages which include longer catalyst life, higher degree of regenerability, and lower cost than titania-supported catalyst. The process results from more than 20 years of effort by the Pittsburgh-based engineers and scientists, first as the Alternate Energy Development Department of Gulf Oil Corp. in the 1970s, and later as Energy International. The parent, Williams International Co., is exploring opportunities to licence the Gas-Cat catalyst and to co-invest in a grass-roots plant using the process. The liquid hydrocarbon output from a GasCat plant is claimed to be competitive with crude oil at a price of $16-17 per bbl. Energy International has recently completed a three year study, funded at approximately $3 million, on Fischer-Tropsch cobalt catalyst technology for the US Department of Energy. The DOE has also funded a study by Energy International Corp. on a 25000 bbl day-’ floating gas-to-liquids plant featuring Fischer-Tropsch synthesis. Howe-Baker Engineers, Inc., an international company specializing in the design and construction of plants to generate syngas and hydrogen, has announced that they are expanding their syngas operations to include FischerTropsch synthesis for the production of transportation fuels. They anticipate offering small plants suitable for use at the 1000 to 10000 range as well as the larger scale plants. Conoco/du Pont has initiated a large effort to develop Fischer-Tropsch technology for the conversion of natural gas to transportation fuels. To date, few details have been made public. Today, Rentech offers two approaches to the use of their Fischer-Tropsch technology. In the first, the 250 bbl day-’ plant constructed in Colorado has now been shipped to be located in the Kumachi gas field in India.85 Various dates have been circulated for the startup operation of this Indian plant, and discussions for two other plants are underway. Rentech is also negotiating with the Texaco Group, Inc. to establish a business relationship to accelerate the development and licensing of Rentech’s Process technology towards commercializing the technology on a worldwide basis.86 In the second approach, Rentech proposed the gasification of the 15000 bbl day- of heavy resid that a typical 100000 bbl day-’ refinery must dispose of today. A portion of the synthesis gas derived from the heavy resid could meet the refinery’s estimated 50 MW electrical requirement, with the remainder being used for Rentech Fischer-Tropsch technology to produce about 5500 bbl day-’ of liquid hydrocarbon products. As indicated above, an agreement has been reached whereby Rentech and Texaco will develop this technology. Carbon Resources is a Cyprus company that has announced that Automated Transfer Systems Corporation will provide capitalization to establish the commercial design parameters of its ‘SYNGEN’ process and the proprietary Fischer-Tropsch process of the N.D. Zelinsky Institute of Organic Chemistry of the Russian Academy of Sciences. Carbon claims strong prospects for plants from Algerian, Egyptian and Nigerian concerns. A feature of this operation is the conversion of natural gas via high-energy plasma using technology patented by Prof. Albin Czernichowski, University of Orleans

’

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(France). The syngas is generated by passing the natural gas, which may contain high contents of COZ, at a moderate temperature through a plasma arc generated by electricity. This process is claimed to provide substantial saving in capital costs and operating costs over conventional syngas formation. The syngas is generated at low pressure and this could be a disadvantage, except for Fischer-Tropsch processes that operate at a lower pressure. The second stage converts the syngas to hydrocarbons using a proprietary catalyst developed at Novocherkassk Plant of Synthetic Products located in Russia. Prof. A. Lapidus, N.D. Zelinsky Institute of Organic Chemistry, Moscow, is providing direction for Dr V.S. Boudtsov at the Novocherkassk plant. Prof. Lapidus’s published work on Fischer-Tropsch synthesis has emphasized cobalt catalysts that are operated at atmospheric pressure. The proprietary catalyst is claimed to have ‘chain-limiting’ properties and has been reported to have been confirmed in their lab in Orleans, France. The recent history of this company is complicated. The company was acquired by Synergy Technologies Corporation, a Colorado corporation. An agreement was signed in June, 1999 so that Texas T Petroleum Ltd. would acquire 50% of Carbon Resources, Ltd. Two of three early entrance co-production plants (EECP) selected for funding by US DOE will feature Fischer-Tropsch synthesis. The US DOE will fund these projects at the $6-8 million range during about five years, and the industries must provide up to 50% cost sharing. These projects are designed to be the first step toward developing advanced technology modules that would ultimately be integrated into an ultra-high efficiency, near pollution-free energy concept labeled ‘Vision 21’. One of these projects will be conducted by Waste Management and Processors, Inc. (WMPI), located in Frackville, Pennsylvania. The other will be managed by Texaco Natural Gas, Inc., of Houston, Texas. WMPI will evaluate the concept of using coal residue as a feedstock. They plan to utilize a Texaco gasification process and Sasol technology for the Fischer-Tropsch synthesis. In their study, Texaco will use Rentech’s Fischer-Tropsch technology to produce high-quality transportation fuels and electricity from coal and petroleum coke. The 5000 barrel day-’ plant considered for Pennsylvania has won a $47 million tax break from the State General Assembly. Plant developers can take up to 15% of monies that would normally be paid on state taxes to fund the cost of the project. The cost per barrel of daily capacity would be $62 400 and would be decreased to $53000 per barrel with the tax break. Even the lower cost appears to be about double the values quoted by most potential developers. The Texaco project fits into the ‘bottomless refinery’ concept that Texaco is advancing. The refiner should be able to generate their power need, sell excess power, produce a cleaner slate of fuels, run a heavier crude and maximize the value of refinery bottoms if the Texaco project reaches the commercial stage. At this time, there is significant activity and optimism concerning the introduction of a number of commercial operations using Fischer-Tropsch technology to convert natural gas to transportation fuels. An oil analyst for Morgan Stanley Dean Witter, Doug Terreson, anticipates that by 2005 refinery output worldwide will be up from the current 76 million bbl day-’ to 90

Catalysis

180

million bbl day- * and that gas-to-liquids plants could be contributing two to three million bbl day- .86 Based on the current research activity in the US and other countries, there is reasonable expectation that Fischer-Tropsch synthesis will assume a significant role in producing transportation fuels. This is made more likely by the projections that by about 2000 the world production of petroleum crude will peak and then decline from that point onward.87 The pessimist will take the view that they have heard this before - the US geared up in the 1950s and again in the 1970s to use synfuels to replace a predicted shortage of petroleum. The optimist will take the view that the date predicted for US petroleum production to peak (1960s) and then decline was correct, and that supply factors will dictate that Fischer-Tropsch synthesis commercialization will accelerate. The optimistic view must be correct at some date, and there are many reasons to believe that it will be within the next few years. According to a recent announcement, Shell International believes that it has gained much experience with Fischer-Tropsch synthesis during the operation of the Bintulu Middle Distillate Synthesis plant and are now pursuing additional opportunities form commercial projects. A recent coverage in the Remote Gas Strategies newsletter indicates that there is mention of a 50 000 bbl day- plant in Bangladesh. In this article, Jack Jacometti, Manager, Technology and Commercial Support for Shell International Gas Ltd., London states that, ‘We are looking at a variety of opportunities worldwide.’ He continues, ‘I don’t believe introducing new technology is a problem’, implying that Shell will move forward from the fixed bed reactor technology that serves for the Bintulu plant, and it is likely that the new technology will include a slurry phase reactor. 8

Summary of Current Status

There are two contrasting viewpoints of the current state of Fischer-Tropsch technology. On the one hand, there is the viewpoint of the investor and venture capitalist who are looking for situations where there is expected to be rapid growth during the short to mid-term period. On the other hand, there is the viewpoint of the scientific and engineering communities. It is not surprising that the two viewpoints may not agree and, in many instances, may be completely contradictory. The investor outlook currently appears to be very positive. The recent report entitled ‘Fischer-Tropsch Technology: Gas-to-Liquids, Solids-to-Solids, Liquids-to-Liquids’, by Howard, Weil, Labouisse, Fredrichs, Inc., a New Orleans-based investment firm, is representative of the optimistic outlook.88 For the short term, this report indicates that gasification will be utilized in downstream processing as an approach to handle heavy bottoms. It has been reported that Texaco is planning to increase refinery throughput by 40000 bbl day-’ using heavier curde at a 150000 bbl day-’ plant and using a deasphaltizer unit and a gasifier to handle the bottoms. The addition of a FT unit would increase the revenue from the plant significantly, and presumably this would be the short-term goal of the Rentech-Texaco joint effort. The

5: Indirect Liquefaction - Where Do We Stand?

181

report indicates that Exxon has a license agreement with Texaco to use the IGCC process at its refineries and chemical plants worldwide. For the short-term, the report indicates that the requirements to reduce CO2, such as required by the Kyoto Protocol, could make a dramatic expansion of Fischer-Tropsch technology. This could even reach to the gasification of coal and producing both electricity and hydrocarbon products. For the longer term, the report indicates that the gas-to-liquid process provided by the Fischer-Tropsch technology will permit today’s major refining companies to prevent the petroleum-producing countries from acquiring significant fractions of refinery operations. Most of these investment-type organizations indicate that there are five to six companies that are in the forefront of a larger group of potential developers of Fischer-Tropsch technology. For example, Tower89recently cited six such companies: Sasol, Shell, Exxon, Rentech, Syntroleum and Conoco. Nearly all of these organizations include Rentech and Syntroleum in the same grouping as Sasol and Shell and, because of the relative size of the companies, are impressed by the potential for growth by Rentech and Syntroleum. Clearly, Rentech and Syntroleum have been much more effective in publicizing their processes to the public than either Shell or Sasol has done to date. For instance, one entry to a web chat-site devoted to Syntroleum activities indicated that ‘he had heard that some company in South Africa had some activity that was similar to Syntroleum’. From the technical perspective, we would have a much different view of the ranking of current and potential company’s commercial Fischer-Tropsch technology. Sasol has operated three types of reactors at all scales from the small laboratory pilot plant to large commercial scales, and are therefore the most advanced in commercialization. They have operated fixed-bed reactors to produce high wax products for about 40 years, and currently produce more than 5000 bbl day-’ of products with these reactors. In spite of numerous operational problems, the Sasol operators were successful in developing ‘chain limiting’ technology and currently produce about 90 000 bbl day- of products using circulating, and now stationary, fluid-bed reactors. These operators developed their technology to utilize slurry bubble column reactors, scaling from laboratory size to a 3 m diameter large pilot plant and then to a 5 m diameter commercial reactor with an output of 1500 bbl day-’. No other company has reported such extensive operations. Shell has commercialized the fixed-bed reactor at an even larger scale than the Sasol operations, and expects to have a production of about 15000 bbl day-’ when their Bintula plant is back in operation. However, it appears that Shell’s largest bubble column reactor operation has been limited to short runs at the US DOE facility at LaPorte, Texas, and details of this operation have not been made public, To date, it does not appear that Shell has operated the fluid-bed reactors at a large scale. Exxon and Rentech have operated bubble column reactors of 200-300 bbl day- scale and have reported reliable operation at this scale. Exxon’s research and development effort has been reported to have cost more than $300 million

’

’

182

Catalysis

and that they have operated integrated gasification and Fischer-Tropsch units for about two years. On the other hand, Rentech’s operation at this scale has been limited to about three months. In addition to a longer operating experience, Exxon has vast experience in the design, startup and operation of a wide range of proven and of revolutionary process units. Other potential Fischer-Tropsch developers are, based on public information, much less advanced than the above companies. It is always hazardous to make such an assessment since some companies are effective in scaling to a significant size before any information reaches the public. For example, there are rumors that BP-Amoco is near to a commercial operation; however, neither company has a patent position that would indicate any activity beyond the laboratory or small pilot plant level. Amoco, for example, has had a very active Fischer-Tropsch research activity for more than ten years and this included the operation of several large-scale pilot plants. While these potential developers may have adequate, or even superior, technology, it remains to be advanced to a scale that most developers would consider adequate to scale to the size needed for a 10 000 bbl day- plant. Acknowledgement

This work was supported by US DOE contract number DE-AC22-94PC94055 and the Commonwealth of Kentucky. References 1 2

3 4 5

6 7 8 9 10 11 12 13 14 15

16 17 18

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Anonymous, Oil & Gas J., Dec. 30,1996. Syntroleum news release, 12/12/97. Syntroleum news release, 03/10/97. Houston Business Journal, 03131/97. Syntroleum news release, 04/10/97. Anonymous, International Journal for Measurement and Control, Jan., 1998. Syntroleum news release, 02/24/98. Syntroleum news release, 08/04/97. Syntroleum news release, 02/05/98. Syntroleum news release, February 6,1997. G.L. Beer (Inventor), Syntroleum Corp (Applicant), PCT/US97/19722, filed 30 October 1997. R.A. Dictor and A.T. Bell, Ind. Eng. Chem., Proc. Des. Dev, 1983,22 678. T.P. Kobylinski, US Pat. 4,088,67 1, May 9,1978. T.P. Kobylinski, C.L. Kibby, R.B. Pannell and E.L. Eddy, US Pat. 4,605,680, August 12,1986. H. Beuther, C.L. Kibby, T.P. Kobylinski and R.B. Pannell, US Pat. 4,413,064, November 1, 1983. H. Beuther, C.L. Kibby, T.P.Kobylinski and R.B. Pannell, US Pat. 4,399,234, August 16,1983. A.H. Singleton and S. Reiger, Proceedings of the 62nd Annual Convention - Gas Processors Association, 1983,pp. 162-165. C.B. Benham and M.S. Bohn, ‘A decade of research and development in FischerTropsch applications’, Pittsburgh Coal Conference, October 15,1992. Dr. J-L. Zhou, seminar at CAER, February 1998. J.M. Fox, I11 and S.S. Tam, Top. Catal., 1995,2,285. J.H. Gregor, Catal. Lett., 1990,7,317. G.N. Choi, S.J. Kramer, S.S. Tam, J.M. Fox, N.L. Carr and G.R. Wilson, 1997 Coal Liquefaction & Solid Fuels Contractors Review Conference, Sept. 3-4, 1997,Pittsburgh, PA. J.-P. Lange and P.J.A. Tijm, Chem. Eng. Sci., 1996,51, 2379. M.J. Gradassi, in ‘Natural Gas Conversion V’, A. Parmaliana et al., eds., Elsevier, Stud. Surf: Sci, Cat., 1998,119,35-44. Anonymous, C&E News, May 1 1, 1998,p. 10. Anonymous, C&E News, July 28,1997,p. 22. Anonymous, O&G J.,June 23,1997,p. 20. L.K. Aalund, O&C J, Special Issue, April 27,pp. 33-42,1998. Anonymous, The Oil Daily, Sept. 23,1997. Chevron News Release, April 15,1998. G.R. Wilson and N.L. Carr, US Pat. 5,639,798, June 17,1997. A.H. Singleton, O&G J., Aug. 4,1997,pp. 68-72. Anonymous, Oil & Gas J., 1997,95,36. Anonymous, Remote Gas Strategies, December 1997. C. Campbel and J. Laherrere, Sci. Am., March 1998. Gas-to-Liquids News, 1999,vol. 11, no. 1. B. Tower, ‘Strategic Implications of FT Technology on the Oil and Gas Industry’, Proc. of Gas-to-Liquids Processing ’99,San Antonio, TX,March 17- 19, 1999.

1997.

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75 76 77 78 79 80 81 82 83 84 85

86 87 88

89

6 Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts BY AKIFUMI UENO 1

Introduction

Steam reforming of methane into syngas (CO/H2) is the first step in the production of methanol and formaldehyde. Steam reforming is endothermic and requires significant energy, so that the direct conversion of methane into oxygenates in a single step is preferred (see Table 6.1). Despite decades of work, only a few papers have reported the yields of oxygenates higher than 4% based on methane.'-3 This is due to the inverse relationship between methane conversion and the selectivity to oxygenates; a high conversion corresponds to a low selectivity to oxygenates. Since formation of CO and C 0 2 during methane partial oxidation reduces the selectivity of oxygenates, suppression of CO and C02 formation, as well as an increase in methane conversion, is necessary to achieve high yields of oxygenates. CO is formed by the oxidation of methanol and formaldehyde, and C02 is produced mainly by the direct oxidation of methane.4 The oxidation of methanol and formaldehyde might be suppressed by a reactor design in which the residence time of the oxygenates on catalysts will be ~ptimized.~ The direct oxidation of methane into C02 might be controlled by the catalyst design. Thus, the key technologies for the direct conversion of methane into oxygenates (methanol and formaldehyde) are in reactor and catalyst design. Most of the work done before 1985 is reviewed by Pitchai and Klier,6 where Mo03-based catalysts were reported to be one of the most promising catalysts for partial oxidation of methane. Recent work published before 1990 is summarized in a review by Brown and P a r k y n ~elucidating ,~ the interaction between methane and catalyst surfaces in order to clarify the active sites for the selective production of oxygenates. In this decade, considerable work has been published on the formation of silicomolybdic acids (SMAs) by the solid state reactions between M o o 3 and Si02 in MoO3/Si02 catalysts in the presence of water vapor. It is still ambiguous, however, whether or not silicomolybdic acid plays an important role in the partial oxidation of methane to oxygenates. This paper will review the recent work concerning Mo03/Si02 catalysts for direct conversion of methane into oxygenates, and then focus on the generation, thermal stability, and the catalytic activity of SMA dispersed on Si02 surfaces for the direct conversion of methane. Catalysis, Volume 15 0The Royal Society of Chemistry, 2000 185

186

Catalysis

Table 6.1 List of thermodynamicparameters for some reactions of interest AH298 (kcal mol- *)

Reactions of interest

AG298 (kcal mol-

CH4 + H20 -+ CO + 3H2 H2 + C02 + H20 + CO 2H2 + CO + CH30H CH30H + 11202 HCHO + H20

49.1 9.8 -22.2 - 34.9

33.9 6.8 -6.5 - 39.7

CH4 + 1/202 -+ CH30H CH4 + 0 2 + HCHO + H20

- 30.7

- 65.6

- 27.1 - 66.8

-+

2

I)

Partial Oxidation of Methane on MoOJSiOz and Alkali Metaldoped MoOJSi02

The effect of molybdenum content in Mo03/Si02 catalysts on the partial oxidation of methane to formaldehyde was studied by Banares et al. using two kinds of Si02 supports; one with the surface area 86 m2 g-' and the other with 200 m2 g-1.8 The loading of Moo3 was converted into the number of Mo atoms exposed on Si02 unit surface (1 nm2), assuming complete dispersion of Moo3. (This assumption is suspicious because Mo ions coagulate to form Moo3 crystallites before complete dispersion.) The catalysts were used to carry out the partial oxidation of methane at 863 K using gases composed of CH4/ 0 2 (1 : 1 molar ratio) mixture. Figure 6.1 shows that the methane conversion and the yield of formaldehyde increased as the number of exposed Mo atoms increased up to approximately 1 Mo nmA2 for both of the Si02 supports, although the maximum with the lower surface area Si02 support is more pronounced. The methane conversion and the formaldehyde yield decreased as the number of Mo nm-2 was higher than 1.5, suggesting that the structure of active species generated on Si02 supports strongly depends upon the surface concentration of Mo ions. According to the paper by Deltcheff et al.,9 Moo3 crystallites are the main species on the catalyst surface at high Mo loading, and at low loading several kinds of molybdenum-oxo-species are formed, which are mainly responsible for catalytic reactions exhibited on MoO3/SiO2 catalysts. Banares et al.* further studied the effects of oxidants upon the selective oxidation of methane using 0 2 and N20. They concluded that pathways for the selective oxidation to HCHO and for the complete oxidation to C 0 2 were a redox cycle. The interaction between CH4 and lattice 02-ions in the Moo3/ Si02 catalyst is as shown in Scheme 1. 0 2 was concluded to be more effective lattice 02CH4(9)

-1

HCHO

adsorbed 0--

CO

lattice 02-

co2

-1

Scheme 1

gas phase oxidation

O2 or shift reaction

6: Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts

k,

3

0 2

5

O0 A

-

187

b

0

A

pQ,

NurnberofMoatomhm'

Number of Mo atoms/nm'

Figure 6.1 Conversion of methane and yield of formaldehyde on MoO&3iOz catalysts with (a) 86 m2g-' S O 2 support, and (b) 200 m2g-' Si02 support; 0means methane conversion and A means yield of formaldehyde

than N20 in reoxidizing the catalyst; whereas adsorbed 0- reacts nonselectively to produce CO. Aoki et al. prepared MoOJSi02 catalysts by two different methods: by a conventional impregnation with low Moo3 dispersions and by a soYgel method with high They tested these catalysts in the direct conversion of methane at 873 K in an excess amount of water vapor, and reported relatively high yields of oxygenates (ca. 4% yield of methanol and formaldehyde) on the highly dispersed catalysts, as given in Table 6.2. They concluded that the highly dispersed Moo3 over Si02 support might generate some complex active species, such as silicomolybdic acid, for the selective oxidation of methane in an excess amount of water vapor. Spencer et al.I2T1 studied the relationship between methane conversion and selectivity to formaldehyde during the partial oxidation of methane over Si02supported Moo3 and V2O5 catalysts at various temperatures. Figure 6.2 shows that the dependence of conversiordselectivity on temperature is weak on the MoO3/SiO2 catalyst, but strong on V205/Si02. This suggests that the mechanism of methane oxidation on Mo03/Si02is somewhat different from that on V205/Si02. Faraldos el al. l4 also compared the catalytic activities of Moo3/ Si02 with those of V205/Si02for selective partial oxidation of methane. The methane conversion on V205/Si02 catalyst was much higher than that on Mo03/Si02, but the selectivity to formaldehyde was higher on the MoO3/SiO2 catalysts at any methane conversion, as exhibited in Figure 6.3. The complementary selectivity trends between HCHO and CO in Figure 6.3 strongly suggest that HCHO was further oxidized into CO on both catalysts. The selectivity to C02 shows, however, a different trend; on the Moo3 catalyst C 0 2 was formed even at the very low conversion of methane, but no C02 was detected on V205 at the very low conversion and a further increase in the methane conversion resulted in an increase in C 0 2 formation. This indicates that HCHO and C02 are the primary products on the Moo3 catalyst, whereas on V2O5 catalyst C 0 2 is formed by further oxidation of CO. Most of the recent

1.o

2.9 8.2 12.0

0.58 1.2 4.0 6.6

conv. (010)

(K)

773 823 873 923 773 823 873 923

Methane

Temp.

I,

73 52 35 11

1

32 13 12

4

0 13 8 11

1

2 1

CHjOH 25 35 33 18 0 10 17 23

co

Selectivity (YO)

62

30 37

14

41 51 54 81

co2

0.20 0.17 0.52 0.07 0.86 1.7 3.8 1.8

HCHO -I-CHjOH

Yield (YO)

respectively, and the amount of catalyst employed was 1.5 g in all experimental runs.

HCHO

Flow rates of CH4.02 and water vapor were 1.8,0.2 and 2.0 L h-

Sol/gel

Impreg.

Catalyst

Table 6.2 Activity and selectivity of Mo031Si02 catalysts prepared by impregnation and sollgel methods for partial oxidation of methane in an excess amount of water vapor

1

873k 840k

921k

3

4

CH, conversion (%)

2

5

6

7

Q

i

i

3

4

s

CH, conversion (96)

6

7

9

5b h

g

Figure 6.2 Selectivitylconversion relationship for formation of HCHO during oxidation of methane with oxygen at different temperatures; (a) 3 Mo03ISi02 catalyst, and (b) V2O,lSiO, catalyst

0

0

A

8

* 896k

Catalysis

190 100

b

A

80 -

80 0

n

E

8 60 P :s

-

c

*L 1

5 6'0

h

0

I

$40

0

20 0

0

0

1

2

3

CH,conversion (%)

4

0

1

2

3

4

5

G

CH, conversion (%)

Figure 6.3 Selectivity to HCHO (O), CO (A) and CO, (0) vs. CH4 conversionfor (a) 0.8M0 and (b) 0.08 V: ( x M o and y V mean the number of Mo and V atoms on 1 nm2 of SiO,, surface)

work concerning the methane partial oxidation over Mo03/Si02 catalysts has accepted that HCHO and C 0 2 are the primary products of the reaction. Banares et a l l 5 studied the effects of additional alkali metal cations upon the structure and catalytic activity of molybdenum oxides in a 2.4% Moo3/ Si02 catalyst. New bands were observed in Raman spectra at 890 and 830 cm-' when alkali metal cations such as Na, K and Cs were doped into the catalyst. Since these alkali metal cations have higher affinity for surface molybdenum oxide species, it is concluded that the alkali-molybdates such as Na2M0207, K2M02O7 and Cs2M0207 were formed on silica surface. These alkali-molibdates were found to be thermally stable and not affected by hydratioddehydration treatments. The formation of alkali-molybdates on Si02 decreased the number of molybdenum oxide species (probably, isolated molybdate), hence the catalytic activity of Mo03/SiO2 for partial oxidation of methane decreased as the amount of doped Na ions increased, as shown in Figure 6.4(a). The number of remaining isolated molybdenum oxide species was estimated by monitoring changes in the intensity of Raman band at 986 cm- l, assigned to Mo=O band of the isolated molybdenum oxide species, with the amount of doped Na ions. The intensity was normalized to the intensity of Raman band of the S O 2 support at ca. 490 cm-', which could be considered as constant. Thus, the results given in Figure 6.4(b) show the turnover frequency (TOF) on the isolated molybdenum oxide species for HCHO production on Na-doped Mo03/Si02 catalysts. The TOF in Figure 6.4(b) is fairly constant, except for the sample with higher Na doping. This means that the poisoning mechanism involves the interaction of each Na ion with Mo=O bond in the isolated molybdenum oxide species, but not with Mo-0-Mo lattice bond in polymolybdate crystallites which was proposed by Spencer et a l l 6 Erdohelyi et a1.17 investigated the partial oxidation of methane on supported potassium molybdate catalysts in a fixed-bed continuous-flow reactor at ca. 860-923 K. The structures of the potassium molybdates deposited by impregnation were found to depend strongly upon the pH value

6: Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts

191

30

1

b A

h

r

cy

‘cn20

30 Y P