CONTROL OF THE
DISTILLATE/GASOLINE RA TIO
OF COAL DERIVED AND OTHER
LIQUIDS
by
JACQUELINE ANNE FAHY
B.Sc. (Hons.)
A Dissertation
Submitted to the School of Chemical Engineering and Industrial Chemistry,
University of New South Wales, in partial fulfilment of the requirements for the
Degree of Doctor of Philosophy.
University of New South Wales
January 1992. CERTIFICATE OF ORIGINALITY
I hereby declare that this submission is my own work and that, to the best of my knowledge and belief, it contains no material previously published or written by another person nor material which to a substantial extent has been accepted for the award of any other degree or diploma of a university or other institute of higher learning, except where due acknowledgement is made in the text. 1
ABSTRACT
Environmental pressures are focusing attention on the removal of aromatics from hydrocarbon feedstocks. The present studies are concerned with the development of a complex catalyst for the hydroalkylation of benzene to cyclohexylbenzene and bicyclohexyl, a high cetane number diesel fuel blendstock.
Studies have focused on nickel and rare earth exchanged 13X zeolite impregnated with platinum, listed in the patent literature as an efficient catalyst.
Considerable problems were found with initial runs in that deactivation and coking occurred. Finally it proved possible to identify conditions such that cyclohexylbenzene was the major product with significant yields of cyclohexane and high molecular weight di- and tri-cyclohexylbenzene isomers also being produced.
Hydrogenation of cyclohexylbenzene to bicyclohexyl was found to be efficient over a supponed nickel catalyst.
The role of the catalyst components was then studied. The catalyst operates as a result of a fine balance of metal and acidic functions. Based on temperature programmed reduction and acidity measurements related to the activity of different catalysts, it was suggested that nickel is the major active hydrogenation catalyst.
Platinum was active for hydrogenation but the main role of the precious metal appeared to be to allow reduction of nickel at low temperatures. The rare eanh salts were necessary to control catalyst acidity and assist in catalyst reduction at low 11 temperatures. The acidity of the zeolite support was responsible for cyclohexylbenzene selectivity.
The effect of operation and pretreatment conditions on the performance of the various catalysts was used to provide support for these suggestions. It proved impossible to totally avoid catalyst deactivation, which precluded detailed study of the kinetics of the reaction.
Hydrogenation of hydroalkylation products was shown to be possible, giving fully saturated bicyclohexyl as the only product. The cetane number performance of the combined hydroalkylation-hydrogenation product indicates its possible
application as a distillate range fuel blendstock.
Hydroalkylation of aromatics contained in coal and petroleum derived fuels
was attempted, but was unsuccessful, apparently due to deactivation resulting from
catalyst poisoning by sulphur. iii ACKNOWLEDGEMENTS
I would like to express my appreciation to the many people who have helped me throughout the course of this project. In particular I would like to thank:
My research supervisor, Professor David Trimm, for his encouragement, guidance and sense of humour, and Professor Mark Wainwright for his assistance with a number of experimental problems.
The National Energy Research Development and Demonstration Council and
BHP for funding the project, and the staff at BHP Melbourne Research Laboratories, especially Noam White and David Cookson for their generosity and enthusiasm.
The technical staff, especially Phillip McAuley, for their efforts in the construction and maintenance of my experimental equipment, and their friendship throughout, and the staff of the School of Chemical Engineering, especially Wendy
Wartho, for their time and patience.
To my fellow postgraduate students, Daniel Thomas, Jennifer Jones and Brett
Moss, for their support and friendship.
And especially to my mother and John Somerville for their special efforts, patience and understanding. iv.
TABLE OF CONTENTS
ABSTRACT
ACKNOWLEDGEMENTS 111. TABLE OF CONTENTS iv. TABLES viii.
LIST OF FIGURES X.
CHAPTER 1. INTRODUCTION 1. 1.1. CRUDE OIL SUPPLY AND DEMAND 1. 1.1.1. Synthetic Fuels 4. 1.1.2. Coal Liquefaction 8. 1.1.3. Characteristics of Coal Derived Liquids 14. 1.2. THE IMPACT OF ENVIRONMENTAL LEGISLATION ON REFINING 17. 1.2.1. Environmental Impact 18.
2. TRANSPORT FUELS 22. 2.1. PETROLEUM BASED TRANSPORT FUELS 22. 2.1.1. Fuel Specifications 23. 2.1.1.1. Gasoline 23. 2.1.1.2. Australian Diesel Fuel 25. 2.1.1.3. Jet Fuel 31. 2.2. COAL DERIVED LIQUIDS AS TRANSPORT FUELS 35. 2.2.1. Suitability of Coal Derived Liquids 35. 2.2.2. Hyd.roprocessing of Coal Derived Liquids 38. 2.3. BACKGROUND TO PRESENT STUDY 43. v. CHAPTER 3. CATALYST DESIGN 53. 3.1. HYDROGENATION OF BENZENE 53. 3.2. HYDROGENATION OF SUBSTITUTED BENZENES 61. 3.3. ALKYLATION 66. 3.3.1. Alkylation of Aromatics 68. 3.4. ZEOLITES 70. 3.4.1. Acidity of Zeolites 75. 3.4.2. Application of Zeolites in Acid Catalysis 79. 3.5. HYDROALKYLATION 81.
4. PROJECT OBJECTIVES 88.
5. EXPERIMENTAL TECHNIQUES 90. 5.1. INTRODUCTION 90. 5.2. MATERIALS 92. 5.2.1. Gases 92. 5.2.2. Chemicals 93. 5.3. CATALYSTS 95. 5.3.1. Cation Exchange of 2.eolites 96. 5.3.2. Impregnation of Catalysts 98. 5.3.3. Calcination 99. 5.3.4. Catalyst Characterisation 99. 5.3.4.1. Bulle Density 100. 5.3.4.2. Total Surface Area 100. 5.3.4.3. Metal Surface Area 101. 5.3.4.4. Elemental Analysis 103. 5.3.4.5. Temperature Programmed Reduction and Oxidation 105. 5.3.4.6. Acidity Measurement 106. 5.3.4.7. Microscopic Analysis 108. vi.
CHAYfER 5. 5.4. HYDROALKYLATION EXPERIMENTS 110. 5.4.1. Apparatus 110. 5.4.2. Procedure 113. 5.5. CHROMATOGRAPlilC ANALYSIS 115.
6. HYDROALKYLATION OF MODEL COMPOUNDS - RESULTS AND DISCUSSION 122. 6.1. HYDROALKYLATION OF BENZENE - RESULTS 123. 6.1.1. Introduction 123. 6.1.2. Initial Testing 124. 6.1.3. Role of Metals 134. 6.1.4. Catalyst Pretreatment 143. 6.1.5. Temperature Programmed Reduction 151. 6.1.6. Effect of Catalyst Support 159. 6.1.7. Catalyst Treatments 164. 6.1.7.1. Steaming as an Alternative to Calcination 164. 6.1.7.2. Chloride Treatment 166. 6.1.7.3. Scxlium Carbonate Treatment 169. 6.1..7.4. Regeneration and Cycles of Use 171. 6.2. HYDROALKYLATION OF BENZENE - DISCUSSION 173. 6.2.1. Initial Testing 173. 6.2.2. Role of Metals 175. 6.2.3. Support Effects 213. 6.2.4. Catalyst Treatments 223. 6.2.4.1. Steaming 223. 6.2.4.2. Chloride Treatment 224. 6.2.4.3. Sodium Carbonate Treatment 225. 6.2.4.4. Regeneration and Cycles of Use 225. vii. CHAPTER 6.3. HYDROALKYLATION OF OTHER MODEL COMPOUNDS 227. 6.3.1. Hydroalkylation of Beni.ene/foluene Mixture 227. 6.3.2. Synthetic Coal Derived Liquid 228.
7. HYDROALKYLATION OF COAL DERIVED AND OTHER LIQUIDS - RESULTS AND DISCUSSION 230. 7.1. HYDROALKYLATION OF COAL DERIVED LIQUIDS 230. 7 .1.1. Feed Characteristics 230. 7 .1.2. Hydroalkylation Trials 231. 7.2. DISCUSSION OF HYDROALKYLA TION OF COAL DERIVED LIQUIDS 233. 7.3. HYDROALKYLATION OF DIESEL FRACTION 235. 7 .3.1. Feed Characteristics 235. 7.3.2. Hydroalkylation Trial 236. 7.4. DISCUSSION OF HYDROALKYLA TION OF DIESEL FRACTION 238.
8. RESULTS AND DISCUSSION - HYDROGENATION AND COMBINED HYDROALKYLATION/HYDROGENATION 240. 8.1. HYDROGENATION OF BIPHENYL 241. 8.2. DUAL CATALYST BED - Combined Hydroalkylation/Hydrogenation 242. 8.3. HYDROGENATION OF CYCLOHEXYLBENZENE 243. 8.4. HYDROGENATION TRIALS AT BHPR-ML 244.
9. CONCLUSIONS AND RECOMMENDATIONS 247. 9.1. CONCLUSIONS 247. 9.2. RECOMMENDATIONS 252. REFERENCES 254. Vlll. TABLES
Table 1.1 Demonstrated Economic Resources and Utilisation Rates of Primary Fuels. 5. Table 1.2 Analytical Data for Gasolines Derived from Wandoan and Yallourn Coal Liquefaction Syncrudes Compared to Petroleum based Gasoline. 16. Table 1.3 Summer _Baseline Gasoline vs. the Clean Air Act Mandate. 19.
Table 2.1 Fuel Boiling Ranges and Primary Requirements. 22. Table 2.2 An Australian Specification for Unleaded Gasoline. 24. Table 2.3 ASTM and Draft Australian Specifications for Diesel Fuels. 26. Table 2.4 Specification for Jet A-1 Aviation Turbine Fuel. 32. Table 2.5 Possible Jet Fuel and Diesel Fuel Blending. Compounds Derivable from Aromatics. 44. Table 2.6 Major Components in Refonnate of Naphtha. 49. Table 2.7 Synthetic Mixtures to Approximate Jet and Diesel Fuels. 49. Table 2.8 Test Results on Synthetic Mixtures. 50. Table 2.9 Test Results on Synthetic Mixtures. 50.
Table 3.1 Structural Parameters of Zeolite Catalysts. 73. Table 3.2 Catalytically Important Properties of Zeolites. 74.
Table 5.1 Gas Specifications. 92. Table 5.2 Chemical Specifications. 93. Table 5.3 Catalyst Specifications. 95. Table 5.4 Gas Chromatography Retention Times and Relative Area Response Factors. 121. ix.
Table 6.1 Hydroalkylation Trials Using Catalyst A (Pt/Ni/RFJzeolite-13X). 127. Table 6.2 Composition of Hydroalkylation Catalysts. 135. Table 6.3 The Effect of Catalyst Metal Content on Benzene Hydroalkylation. 137. Table 6.4 Characteristics of Hydroalkylation Catalysts of Differing Suppons. 160. Table 6.5 Benzene Hydroalkylation with Catalysts of Varying Supports. 161. Table 6.6 Acidity Determinations on Catalysts of Varying Rare Earth Metal Contents (Catalysts I,J,K,A). 218. Table 6.7 Composi~on of Coal Derived Liquid Synthetic Mixture. 228.
Table 7.1 Hydroalkylation of Coal Derived Liquids using Catalyst A. 232. Table 7.2 Diesel Hydroalkylation Trial using Catalyst A. 237.
Table 8.1 Composition of Products BHPR-ML Hydrogenation Trials. 245. X. LIST OF FIGURES Page Figure 1.1 Australian Petroleum Production versus Demand 1988-1998, API Projections. 2. Figure 1.2 Petroleum Products Demand 1989/1999. 3. Figure 1.3 Coal Liquefaction Routes. 9. Figure 2.1 Boiling Point versus Cetane Number for Various Classes of Hydrocarbon. 28. Figure 2.2 Dependence of Diesel Fuel Cetane Number on Aromatic Content. 37. Figure 2.3 Dependence of Jet Fuel Smoke Point on Aromatic Content. 37. Figure 2.4 Simplified Flow Diagram for the Refining of H-Coal Syncrude to Jet and Diesel Fuels. 39. Figure 2.5 Schematic Flow Diagram for the Refining of H-Co~ Syncrude to All Gasoline. 40. Figure 2.6 Syncrude Hydroprocessing. 45.
Figure 3.1 Idealised Projection and Section through a Sodalite Unit of Faujasite. 72.
Figure 3.2 Development of Acidity in ~ Y. 76. Figure 3.3 Dealumination Process. 78. Figure 3.4 Dual Site Mechanism of Benzene Hydroalkylation. 82.
Figure 5.1 Schematic of Hydroalkylation Reactor Studies System. 112.
Figure 6.1 Deactivation testing of Catalyst A (Pt/Ni/RE-13X).
T=17CY'C, P=3.5 MPa, LHSV=l6.56 hr-1, H2:HC=l:1. 125. xi. Figure 6.2 Product Composition for Benzene Hydroalkylation over Catalyst A (Pt/Ni/RE-13X).
1 T=17Q°C, P=3.5 MPa, LHSV=16.56 hr , H2:HC=l:1. 126. Figure 6.3 Influence of Reaction Pressure on Product Yields and Benzene Conversion for Catalyst A (Pt/Ni/RE- l 3X).
T=17Q°C, LHSV=16.56 hr1, H2:HC=l:1. 130.
Figure 6.4 Influence of LHSV on Product Yields and Benzene Conversion for Catalyst A (Pt/Ni/RE-13X).
T=17D°C, P=3.5 MPa, H2:HC=l:1. 131. Figure 6.5 Benzene Conversion and Product Yields as a Function of Nickel Loading. Catalysts E,F,G,A,H. All contain Pt/RE. 17Q°C, 3.5 MPa, LHSV=l6.56 hr1.
benz:H2 1: 1. 138. Figure 6.6 Benzene Conversion and Product Selectivities as a Function of Platinum Content. Catalysts B,A,C,D.
All contain Ni/RE. 17D°C, 3.5 MPa, LHSV=16.56 hr1,
benz:H2 1: 1. 139. Figure 6.7 Benzene Conversion and Product Selectivities as a Function. of Rare Earth Content. Catalysts I,J ,K,A. All contain Pt/Ni. 1700C, 3.5 MPa, LHSV=16.56 hr1,
benz:H2 1: 1. 140. Figure 6.8 Effect of Pretreatment Temperature on Product Yields, Catalyst E (no Nickel). Pretreatment carried out for
15 min. under flowing hydrogen (210 ml.min·1) and 3.5 MPa. Standard Catalyst Test Reaction Conditions. 145. Figure 6.9 Effect of Pretreatment Temperature on Product Yields, Catalyst B (no Platinum). Pretreatment carried out for
15 min. under flowing hydrogen (210 ml.min·1) and 3.5 MPa. Standard Catalyst Test Reaction Conditions. 146. xii. Figure 6.10 Effect of Pretreatment Temperature on Product Yields, Catalyst I (no Rare Earth Metals). Pretreatment carried out for 15 min. under flowing hydrogen
(210 ml.min-1) and 3.5 MPa. Standard Catalyst Test Reaction Conditions. 147. Figure 6.11 Effect of Pretreatment Temperature on Product Yields, Catalyst M (no Platinum/no Rare Earths - only Ni). Pretreatment carried out for 15 min. under flowing
hydrogen (210 ml.min-1) and 3.5 MPa. Standard Catalyst Test Reaction Conditions. 148. Figure 6.12 Temperature Programmed Reduction Profiles (Temperature Derivatives) of Catalysts M (Ni-13X), I (Ni/Pt-13X), and A (Pt/Ni/RE-13X). 152. Figure 6.13 Temperature Programmed Reduction Profiles (Temperature Derivatives) of Rare-earth-13X, Catalysts E (Pt/RE-13X) and B (Ni/RE-l 3X). 152.
Figure 6.14 Product Yields as a Function of Time on Line using Catalyst A after Steaming at 600°C for 2 hours. Standard Reaction Conditions: 17C1'C, 3.5 MPa, LHSV=l6.6 hr-1,
benz:H2=1:1. 165. Figure 6.15 Product Yields as a Function of Time on Stream using Catalyst A after Chloride Addition (17Q°C, 3.5 MPa,
LHSV=16.56 hr-1, benz:H2=1:1). 168. Figure 6.16 Benzene Conversion as a Function of Time on Stream using Catalyst A after treatment with Sodium Carbonate (17Q°C, 3.5 MPa, LHSV=l6.6 hr·1). 170. Figure 6.17 Effect of Catalyst Regeneration on Activity of Catalyst A. Details of Regeneration Procedure given in
Section 6.1.6.4. (17Q°C, 3.5 MPa, LHSV=l6.6 hr1). 172. Xlll.
Figure 6.18 Dual Site Mechanism of Benzene Hydroalkylation. 193. Figure 6. 19 Influence of Catalyst Suppon Total Acidity on Cyclohexylbenzene Selectivity, Catalysts P, Q, I, J, K, N, A and 0. Standard Reaction Conditions: 17D°C, 3.5 MPa,
LHSV=l6.6 hr-1, benz:H2=1:l). 214. 1. CHAPTER 1. INTRODUCTION
The Australian petroleum industry faces a number of emerging pressures over the next decade. The prospect of diminishing crude oil production and reserves, and an increasing demand for petroleum products will increase incentives to exploit our non-petroleum resources. There will also be increasing pressure from government bodies and consumers for these products to meet environmental considerations as well as quality requirements.
1.1. CRUDE OIL SUPPLY AND PRODUCT DEMAND
Australian Institute of Petroleum predictions indicate a rapid decline through the 1990's of Australia's self-sufficiency in crude oil. There is a steadily increasing gap between petroleum products consumption and domestic feedstock production
(see Figure 1.1), which will have to be met by importing increasing quantities of overseas crude oil, or by implementing synthetic fuel technologies. Current predictions (2) indicate a self-sufficiency ratio of 84% in 1990 declining to about
33% within the next ten years. 2.
Thousands or megalilres 50.------,
40 ····································· .....
30 · ...... · ... ·
20 ..... · ·
10 ...... ··········· ... ·
0 .___..___....___....___....___~-~------1988 1989 1990 199119921993 1994 1995 1996 1997 1998 Year
- Products Demand - Feedstocks Output
Figure 1.1 Australian Petroleum Production versus Demand 1988-1998, API
projections ( 1).
Over the same period, demand for middle distillates such as jet fuel and diesel fuel is expected to increase, while that for gasoline will show a marginal decline, see Figure 1.2. This "whitening of the barrel" is expected to present problems to refiners due to the replacement of the sweet, low-sulphur feedstocks
(mainly from Gippsland) with the high-sulphur crude oils from the Middle East (3). 1989
3.
Industrial Diesel 1%
Others 5%
Fuel Oil 6%
1999
... :;:::::::::t}({t/:ffttl/ht::::::::::::::-.. Industrial ... :&i//ii:JltlJiI!Ii:!IllIItIJMI:fi!JIII!I!i:l::1t\.. Diesel .:/\\\/?://:/\/G.a.10.Jlri~f42,%/:/:/://:/\. <1% "''s'S\f!/ff\1~~;\fi 'f Others 5%
Fuel Oil 4%
Figure 1.2 Petroleum Products Demand 1989/1999 (2). 4. It is not the first time that Australia has been faced with a potential shortfall in crude oil supply. The oil embargo of 1973 served to bring the liquid fuel supply problem to the forefront of concern in both Australia and the rest of the world.
Within this framework of events, major research efforts were initiated worldwide to address the liquid fuels supply problem.
1.1.1. Synthetic Fuels
Synthetic fuels from hydrocarbon sources are those that are or can be produced from naturally available carbonaceous materials other than petroleum or natural gas. These materials are found in vast amounts around the world. Like crude oils and natural gases, they are classified as non-renewable energy potential but can share in solving short-term energy requirements. Coal, tar sands and shale oil are amongst the most common sources of synthetic fuels.
These materials are structurally complex and cannot be utilised directly. They must first be changed into simpler energy forms through chemical and physical conversion processes such as pyrolysis, hydrogenation, oxidation and solvent extraction. The end products may be either gaseous, liquid or solid · fuels but, to satisfy market requirements, these new fuels must have similar properties and must perform comparably to conventional hydrocarbon fuels already in use. s. The need for Australia to look toward other resources of fossil fuels is demonstrated in Table 1.1, prepared by the Department of Resources and Energy.
Reserves of brown and black coal can be seen to be extensive, as compared with crude oil reserves, especially in terms of "life-times". There are also significant reserves of natural gas and oil shale.(4)
TABLE 1.1 Demonstrated Economic Resources and Utilisation Rates of Primary
Fuels (4).
Fuel Demonstrated Resources Total "Life" of Economic PJ Production Resources in Quantity Equivalent 1993-94 Years, 1993-94 (Forecast) (Forecast)
Black Coal 31,148 841,000 5,220 161 milltonnes Brown Coal 41,900 398,100 496 803 milltonnes Crude Oil 302,000 11,200 972 12 and Megalitres Condensate LPG 85,000 2,300 62 37 Megalitres Natural Gas 616,000 23,900 1,020 23 millm3 Uranium 474,000 265,400 5,224 51 tonnes U eqmv.
TOTAL 1,542,200 12,994 114 6. To address this scenario a number of synthetic liquid fuel technologies are being researched and developed:-
1) coal liquefaction to produce a range of liquid fuels;
2) natural gas reforming / MTG process;
3) surface or in-situ retorting of oil shale;
and 4) conversion of biomass to ethanol.
1) Coal liquefaction processes can be classified broadly as direct or indirect
The direct processes liquefy solid coal under high pressures, medium temperatures and in a reducing hydrogen atmosphere: these processes include pyrolysis, hydroliquefaction and solvent hydrogenation or extraction. Indirect liquefaction processes involve firstly the gasification of coal to synthesis gas (a mixture of carbon monoxide and hydrogen) followed by chemical reaction to produce a range of liquid fuels. Fischer-Tropsch synthesis is the most widely known indirect liquefaction process used.
2) Production of synthetic liquid fuels from natural gas involves firstly the reforming of natural gas to synthesis gas, followed by chemical reaction to produce methanol. Subsequent reactions with isobutylene, for example, can be used to 7. produce methyl tertiary butyl ether - MTBE or a range of hydrocarbons. The Mobil methanol-to-gasoline (MTG) process is being operated on an industrial scale in New
Zealand at 14,500 barrels per day, as a fixed bed version. In Germany a 100 barrels per day fluid-bed MTG process has been successfully tested and may supersede fixed-bed technology (5).
3) The surface retorting of shales involves the initial crushing of shale and screening to the requirements of the retort. The shale is then heated to separate the oil. The major problems in shale retorting include the development of suitable retorts, the scaling up of these units and the separation of oil from carry-over dust.
Several technologies are under development:- the Union rock-pump retort, capacity
10,000 barrels per day; the Tosco and Petrosix retorts, capacity 800 barrels per day; and the Lurgi, Exxon, Dravo and Paraho retorts which are yet to be scaled up (5).
4) Well-established technology exists for the conversion of starch and sugar feed-stocks to ethanol. Starches are firstly converted to fermentable sugars by acid or enzyme hydrolysis. The sugars are then fermented to ethanol using yeast in a batch or continuous mode (5). 8. 1.1.2. Coal Liquefaction
Due to Australia's extensive coal resources the production of liquid fuels from this material is especially attractive. The known processes for producing liquid fuels from coal can be grouped into four broad categories:
1) pyrolysis
2) solvent extraction
3) catalytic liquefaction and
4) indirect liquefaction.
A schematic flow diagram representing the major processes is shown in Figure 1.3.
In pyrolysis processes (destructive distillation, retorting, and carbonisation), coal is heated to temperatures above 400°C in the absence of air or oxygen to yield heavy oil, light liquids, gases and char. The composition and relative amounts of the products formed are influenced by the coal heat-up rate, the maximum temperature reached, the coal and product residence time, hydrogen partial pressure, the coal particle size and the reactor configuration. Maximum liquid yields result from very rapid heating of small coal particles and short residence times. One of the major difficulties with the pyrolysis process is that a considerable proportion of the coal is converted to coke or char which must be disposed of. Rarely does the proportion of coal converted to tar or oil exceed 20% by weight of the original coal matter expressed on a dry and ash free basis (6). 9.
Liquid Product .. Pyrolysis Oil -- Hydrotreating
Solid or Solvent Solids Liquid Product Extraction Separation
Coal Mining Catalytic and 1-. Hydrogenation Preparation
Liquid Product Catalytic Solids Liquefaction Separation
Liquid Product Synthesis Shift Fischer- Gas ~.co2 ....,... Tropsch f---+ Lt ... Conversion ii Removal Producer Synthesis
Methanol Methanol Synthesis ----.
Figure 1.3 Coal Liquefaction Routes (6). 10. Solvent extraction involves contacting coal at temperatures up to 500°C with a "donor" solvent - one that is capable of transferring relatively loosely bound hydrogen atoms to the coal and maximising the fraction of the coal that enters into solution. Several different extraction processes have been developed, dependent on whether or not extraction is carried out under hydrogen pressure, and if the solvent is hydrogenated in a separate step before being returned to the extraction process (7).
The solvent refined coal (SRC) process adds minimum hydrogen to the coal during extraction, but removes most of the ash and a substantial fraction of the sulphur, to produce a combustible product that is a solid at room temperature.
The two processes which have attracted the most attention for coal liquefaction are indirect liquefaction via Fischer-Tropsch synthesis, and the catalytic liquefaction or hydrogenation/Berguis-Peir process (7).
In the former, the coal is gasified and converted to synthesis gas, this is then converted catalytically to hydrocarbons ranging from light gases to heavy waxes.
Reactors on which the catalyst is fluidised (e.g. Kellogg produce predominantly light gases, whereas reactors containing fixed catalyst beds tend to produce heavier materials. The advantages of the Fischer-Tropsch or indirect liquefaction are that:-
1. almost all coals can be converted to liquid products;
2. the product composition is controllable to a high degree; 11.
3. the "cleanliness" of the product is good i.e., nitrogen and sulphur
compounds are minimal;
4. the technology is commercially available.
The disadvantages are that:-
1. all the coal must be gasified with oxygen and steam, and the gases
must then be purified and recombined to liquid products;
2. a highly complex plant with substantial equipment requirements is
needed and capital cost per unit of production is high;
3. the process has poor thermal efficiency.
Catalytic liquefaction processes are based on the hydrogenation of coal.
Chemically the hydrogenation of coal can be understood by regarding coal as Cfio.8•
Most heavy oils have the approximate formula CHu. By absorbing hydrogen, coal converts to heavy oil. This heavy oil may then be hydrotreated to form light oils, as possible sources of transport fuels. It is presently believed that solvent extraction, discussed above, also proceeds via an hydrogenation mechanism (7).
The hydrogenation process involves mixing of coal with an oil and a donor solvent (solvent or slurrying agent) and the slurry so formed is reacted at pressures between 10-30 MPa and temperatures between 350-500°C for periods up to 4 hours 12.
(generally about 1 hour). Hydrogen is added in most processes, together with a catalyst. Materials from downstream processing may be added or recycled.
The sources of the solvent themselves have been investigated widely (4,5,7).
They may be totally external or may be an oil derived from the liquefaction process itself. External solvents include coal tars from other processes, a residue or fraction from mineral processing, or a similar fraction from shale oil or tar sands oil.
Alternatively a fraction of oil from the liquefaction process may be distilled and recycled. Combinations of external and internal oils are also used and, in some processes, the oil may be hydrotreated to improve its hydrogen donor or solvation properties.
Coals contain significant levels of nitrogen, sulphur and oxygen. Some reduction in the levels of these elements does occur during liquefaction, but heavy oils still require further treatment, as the emission of sulphur and nitrous oxides upon combustion is considered environmentally unacceptable. Oils of this type are also unacceptable as some secondary processing steps cannot tolerate N, S and 0, and the poisoning of catalysts will occur. It is therefore necessary to subject heavy coal-derived oils to further hydroprocessing to reduce the N, S, and O levels to more acceptable levels. 13.
Many studies have been conducted as to the suitability of various coals to coal liquefaction via hydrogenation (8,9,10). The nature of coal liquids produced by hydrogenation is known to be dependent upon coal rank, with very low rank coals such as brown coals and peat producing liquids rich in saturated hydrocarbons.
Generally, however, coal derived liquids are known for their predominantly aromatic composition and the use of coal derived liquids as transport fuels is very dependent upon this. 14.
1.1.3. Characteristics of Coal Derived Liquids
Many studies of coal liquefaction indicate a relationship between the rank of coal, the complexity of the hydrocarbon groups in the coal and the liquefaction products.
An early study of the hydrogenation of lower-rank brown coals (lignite) indicated that they are more reactive, require less hydrogen pressure, and produce smaller polynuclear hydrocarbons than bituminous coals (8). Another study has revealed compositional differences in the pyridine extracts of three coals of different rank and in their liquefaction products produced by the Synthoil process (9). It was found that larger quantities of saturates and higher percentages of lower ring number saturates were observed in the two higher rank coal extracts. However, the ring distributions of the aromatic fractions of the coal extracts indicated slightly higher average ring numbers for the higher ranked coals.
Another study (10), using pyridine extraction of coals of different rank and analysis by proton NMR, indicated more condensed aromatic rings with less substitution and shorter aliphatic chains in the higher ranked coal extracts. Extracts from bituminous coals contained an average of four or five aromatic rings with side chains averaging three or four carbons in length, whereas a lignite extract was found to have an average of one or two aromatic rings with seven-, or eight-carbon, aliphatic side chains. 15. As a result, it is not surprising that coal derived liquids are richer in aromatic species and contain higher concentrations of nitrogen, oxygen and sulphur than petroleum crudes currently processed by Australian refineries. It is these properties that demand attention during the production of transport fuels. Generally, aromatic naphthas make good gasolines. Impending environmental legislation may however change this dramatically and this will be discussed in the next section. Aromatic kerosenes (derived from coal liquids) are considered too "smoky" for commercial jet fuels and aromatic distillates are too low in cetane number for use as specification jet fuels. Table 1.2 gives some typical composition data for gasolines derived from
Wandoan (sub-bituminous, Queensland) and Yalloum (lignite, Latrobe Valley,
Victoria) coals compared to a petroleum based gasoline ( 18). 16.
Table 1.2 Analytical Data for Gasolines Derived from Wandoan and Yallourn Coal Liquefaction Syncrudes (a) compared to Petroleum based Gasoline (18).
Wandoan Coal Yallourn Coal Petroleum based
Density (mg/L) 0.808 - 0.821 0.827 - 0.852 0.757
Hydrogen Content 10.9 - 11.4 9.9 - 11.1 14.1 (wt%) Carbon Aromaticity 54 - 60 55 - 71 55 (atom%) Research Octane 93 - 100 98 - 105 91-93 Number Component Analysis (wt%) n-paraffins 3 - 10 2-7 ) 58.8 iso-paraffins 8 - 15 5 - 12 ) naphthenes 5 - 10 1 - 13 9.2 aromatics 67 - 84 71 - 91 32 Aromatics (wt%) benzene 7 - 9 3 - 11 3 toluene 17 - 26 6 - 18 13 ethyl benzene 5 - 8 7 - 10 5 o,m,p-xylene 14 - 23 14 - 18 18 propylbenzene 1 - 3 3 - 6 ) o,m,p-ethyltoluene 7 - 10 8 - 12 ) 16 trimethylbenzene 5 - 8 5 - 8 ) indan 1 - 2 4-6 ) 17.
1.2. THE IMPACT OF ENVIRONMENTAL LEGISLATION ON
REFINING
Refiners worldwide are being confronted with the prospect of increasingly stringent regulations regarding both refinery emissions (e.g. SOx), and fuel composition and quality. The USA Clean Air Act (CAA) has pioneered legislation in this area and has identified a number of key areas of change with regard to gasoline and diesel fuel quality (11,12,13). Present and proposed U.S. Environmental
Protection Agency rules and the Clean Air Act are calling for specific fuel formulas and fuel specifications to improve air quality in the U.S.A. Alternative fuels may also be introduced to urban areas with especially severe problems.
Important features of the fuels for the 1990s will include mandated oxygen content and lower aromatics content of gasoline, lower gasoline vapour pressure, and perhaps lower olefin content. Alternative-fuel mandates could require the use of methanol (either neat (100%) or mixed with gasoline), liquefied petroleum gases
(LPG's), compressed natural gas (CNG), and gasolines blended with ethanol
(gasohol) in urban areas with the most severe air quality problems.(11)
The CAA has mandated the reduction of aromatics in reformulated gasolines by 1995. It targets total aromatics, with reductions from 32 vol% to 25 vo1%, and benzene, with a maximum benzene level of 1 vol%. Gasoline volatility is also 18. regulated with levels limited to 7.0 - 7.5 psi Rvp (Reid vapour pressure) anticipated in many areas of the United States after 1992.
The US Environmental Protection Agency (EPA) is also considering rules to reformulate highway diesel fuels. EPA wants to limit diesel sulphur content to 0.05 wt% and aromatics content to as low as 10 vol%. The aromatics limit may be replaced by a minimum cetane specification as cetane level may offer a better control over exhaust particulate emissions from diesel vehicles (11).
1.2.1. Environmental Impact
The resurgence of environmental awareness worldwide is presenting a challenge to petroleum refiners. In particular, tougher environmental regulations targeting vehicle emissions will change the nature of the fuels produced by refiners and the processing schemes used to produce these fuels.
Reduced ozone and carbon monoxide levels and reduced airborne toxic compounds are the main goals to be achieved by lower emissions from vehicles.
Motor vehicles are considered to account for almost all carbon monoxide emissions, and approximately 40% of the ozone in urban areas of the United States. They are also a significant source of toxic compound emissions. Specific vehicle emissions that have been targeted for reduction either because they contribute to or cause 19. formation of these toxic emissions include hydrocarbons (evaporative and exhaust), benzene, carbon monoxide, and oxides of nitrogen.
Hydrocarbon emissions are targeted because they combine with NOx in the air and, in the presence of sunlight, promote the formation of ozone (03). Ozone is a lung irritant and causes respiratory problems.
Benzene emissions are targeted because benzene is a known carcinogen. The levels of benzene found in finished gasolines and the CAA requirements are shown in Table 1.3. Benzene is also generated as a result of gasoline combustion in a vehicle's engine.
Table 1.3 Summer baseline gasoline vs. the Clean Air Act mandate (14)
Hydrocarbon type Baseline gasoline CAA mandate
Aromatics, vol % 32.0 25.0 maximum Benzene, vol% 1.53 1.0 maximum Olefins, vol% 9.2 - Saturates, vol% 58.8 - Oxygen content, wt% - 2.0 minimum
Carbon monoxide emissions are listed due to their toxicity. They are especially a problem in cold weather and in regions where there are frequent temperature inversions. NOx is targeted due to its role as reactant in the formation of 20. ozone. NOx levels are affected more by combustion temperatures in an engine than by components in the fuel (10). The direct combination of nitrogen and oxygen becomes significant at temperatures in excess of 1650°C, a value often reached in the combustion chamber.
The most difficult refining scenario to meet reformulated gasoline targets and to maintain quality is considered to be the reduction of aromatics. Ethers are thought to have the potential to replace aromatics, with methyl tertiary butyl ether (MTBE) being the most promising additive. MTBE was introduced as the only non hydrocarbon, refinery pool component to increase octane quality upon the introduction of unleaded gasoline. MTBE has a blending octane ranging between 106 and 110, Rvp between 8 and 10 psi, and photochemical reactivity of 2.6 (12). While
MTBE will continue to increase in volume in the gasoline pool, it is not considered as a one-to-one replacement for aromatics. It would probably accompany other additives, such as oxygenates ETBE (ethyl tertiary butyl ether) and TAME (tertiary amyl methyl ether) (14) and antiknock additive MMT (methylcyclopentadienyl manganese tricarbonyl) (12). Modification to reformer operation and boiling cut point adjustment would also be required.
Diesel fuels have also been targeted for reformulation. Distillate sulphur and aromatics levels have increased during the past two decades due to the influx of heavier crudes, higher conversion in cracking units to produce higher-octane gasoline and the increasing demand for distillates that required blending of cracked stocks 21.
(13). The United States Environmental Protection Agency has proposed to limit highway diesel sulphur content to 0.05 wt% and aromatics content to 20 vol%, or possibly as low as 10 vol%. These rules are expected to come into effect by 1993
(11). A number of different refining strategies are being examined with these restrictions in mind (11,15,16,17). Hydrotreating catalysts are being tested for their hydrodesulphurisation and hydrogenation capabilities. While limits of 0.05 wt% S are attainable using existing hydrotreating systems, some concern over operation with heavier oils is evident (15). Reduction of aromatics in pilot plant studies of diesel range feedstocks has been achieved using hydrotreating catalysts such as Ni
W/ Al203 (15,16) and Ni-Mo/Al20 3 in combination with noble-metal/zeolite catalysts to increase aromatic saturation (16). The importance of hydrogen pressure and liquid hourly space velocity were noted in these studies, but no details of the resulting fuel cetane numbers were given.
From this discussion, it is seen that the aromatic content of gasoline is expected to be reduced in the near future, leading to a glut of aromatic materials. At the same time, fuel quantity and quality will be under pressure. The reaction of refineries to these pressures will be a major test of versatility and efficiency in the
1990s. 22.
CHAPTER 2. TRANSPORT FUELS
In this chapter the characteristics of transport fuels are discussed. Due to their predominance worldwide, the important properties of petroleum derived fuels will be described, followed by discussion of the suitability of coal derived liquids as an alternative source of transport fuels. Finally, the work carried out as a basis of this present study is summarised.
2.1. PETROLEUM BASED TRANSPORT FUELS
The three major classes of petroleum based transport fuels; gasoline, jet fuel and automotive diesel are broadly defined by their different hydrocarbon boiling range fractions. The primary property requirements are given below in Table 2.1.
Table 2.1 Fuel Boiling Ranges and Primary Requirements(18)
Fuel Predominant Boiling Primary Quality Fraction Range Requirements Gasoline Naphtha C5-200°C RON Jet-Fuel Kerosene 200-250°C Smoke Point 20mm Automotive- Distillate 250-300°C Cetane number Diesel 40 23.
2.1.1. Fuel Specifications
2.1.1.1. Gasoline
Historically, automotive gasolines have been classified into two grades -
'premium' and 'regular' on the basis of octane quality. Australian premium grade gasoline (leaded) is presently characterised by a research octane number specification
(RON) of 97 .0 minimum. All new cars sold in Australia since 1986 have catalytic converters to reduce exhaust emissions and these systems are known to be poisoned by lead. The unleaded gasoline specification requires a RON in the range 91 to 93, although based on USA experience this could be increased to approximately 96 as demand for premium unleaded gasoline increases.
ASTM have issued a standard specification for automotive gasoline, ASTM
D439-83. It provides guidance in establishing the required properties to suit various types of engines and climatic conditions. A typical Australian specification for unleaded gasoline is shown in Table 2.2. 24.
Table 2.2 An Australian Specification for Unleaded Gasoline (18)
Method Property Specification ASTM D2699 Octane Number, Research 91.0 to 93.0 ASTM D2700 Octane Number, Motor 82.0 Minimum ASTM D86 Distillation Maximum °C Winter 10% Evaporated, 65 60 30% Evaporated, Report Report 50% Evaporated, 115 110 90% Evaporated, 185 180 Final Boiling Point, 228 228 Residue, Volume % 2 2 ASTM D323 Reid Vapour Pressure,kPa 70.0 80.0 ASTM D2533 or 36:1 V/L, Temperature, specified by a schedule calculated oc ASTM D525 Oxidation Stability, Min. 240 Minimum ASTM D381 Gum, Existent, mg/L 40 Maximum ASTM D130 Corrosion, Copper Strip, 1 Maximum 3 h@ 50°c
IP 30 Doctor Test Negative ASTM D2785 Sulphur Content, % Mass 0.10 Maximum ASTM D3237 Lead Content, mg Pb/L 13 Maximum ASTM D3231 Phosphorus Content, kg/L 1.3 Maximum ASTM D4052 Density @ 15°C, kg/L Report Visual Colour Yellow 25.
The antiknock or octane quality of a gasoline is of prime importance. If it is too low, a high pitched, metallic rapping noise occurs in the engine. In addition to this being audibly annoying, in severe cases it can lead to engine damage. Both the research and motor methods of testing match the knocking characteristics of a test gasoline to standard fuels which are blends of two primary reference fuels: iso octane (RON defined as 100) and n-heptane (RON defined as 0). From this it can be inferred that highly branched paraffins have octane qualities superior to normal paraffins. In addition to branched hydrocarbons, aromatics are generally considered to be the most desirable component class for high octane quality.
2.1.1.2. Australian Diesel Fuel
The properties of a commercial automotive diesel fuel depend on the refining processes used, and on the nature of the crude oil from which it is derived. It is possible, therefore, to produce diesel fuels of widely varying properties. To ensure that adequate engine performance, handling safety and environmental acceptability are obtained it is necessary to specify limits of important fuel properties. The
American Society for Testing and Materials (ASTM) have a standard specification covering three grades of diesel fuels suitable for use in various types of engines.
Table 2.3 summarises the Australian equivalent of this standard. 26.
Table 2.3 ASTM and Draft Australian Specifications for Diesel Fuels (18)
I Property I ASTM 1-D a I ASTM 2-D b I Australian ADF c I Cetane Number 40 (min) 40 (min) 45 (min) Distillation,°C 288 (max) 338 (max) 357 (max) (90% recovered, 282 (min) ASTM D86)
Cloud Point,°C d d d Flash Point,°C 38 (min) 52 (min) 61.5 (min) Oxidation stability, - - 25 (max) mg/L (ASTM 02274)
Copper Corrosion 3 (max) 3 (max) 2 (max) (ASTM Dl30)
Sulphur, wt% 0.5 (max) 0.5 (max) 0.5 (max) Ash, wt% 0.01 (max) 0.01 (max) 0.01 (max) Carbon residue, wt% 0.15 (max) 0.35 (max) 0.2 (max) (on 10% residuum, ASTM 0524)
Water and Sediment, 0.05 (max) 0.05 (max) 0.05 (max) vol% Viscosity, cSt 1.3 (min) 1.9 (min) 1.9 (min) (kinematic at 40°C) 2.4 (max) 4.1 (max) 5.5 (max) Filterability index - - 6 (max) Density, kg m·3 - - 820 (min) 860 (max)
a A volatile distillate fuel oil for engines in service requiring frequent speed and load changes.
b A distillate fuel oil of lower volatility for engines in industrial and heavy mobile service.
c Australian automotive diesel fuel.
d Specification depends on climatic (temperature) conditions. 27.
Cetane Number
Cetane number is a measure of the ease with which a diesel fuel auto-ignites by compression in a diesel engine. It is a quantity which is affected by physical processes (such as spray formation, heating, vaporisation and turbulent mixing) in addition to the chemical ignition process. The cetane number of a particular fuel is determined by an experimental engine test (usually ASTM 0613) where comparison of performance is made against blends of two standards: hexadecane (cetane number
= 100) and 2,2,4,4,6,8,8- heptamethylnonane (cetane number = 15). Thus, cetane number is a performance indicator that is dependent on the chemical composition of the fuel. Hardenberg (19) has examined the role of physical and chemical factors in effecting cetane number in detail. Aromatic hydrocarbons are low in cetane number, normal paraffins have a high cetane number, with isoparaffins and cycloparaffins
(naphthenes) falling in between, see Figure 2.1.
Sulphur, nitrogen and oxygen
It is necessary to ensure that diesel fuels do not contain excessive concentrations of sulphur, nitrogen and oxygen. Sulphur and nitrogen can lead to environmental pollution through exhaust emissions and may also cause corrosion and engine wear if acids are formed during the combustion process. At present, only a sulphur specification is given for diesel fuels. Nitrogen and oxygen levels in petroleum derived crudes are generally very much lower than sulphur levels and are
not usually problematic. If the sulphur level exceeds specification, fuels can be
hydrodesulphurised. 28.
Normal Alkylbenzenes Boiling Temperature C 350r------r-----l--r------~~-
300
Normal Hexadecane
1-Melhylnaphthalene 250
Australian 200 Diesel Fuel - Normal Alkyl Specification Cyclopentane
150 1-Alkenes
100------_J__J___ _L __ _j______J 0 40 80 120 Cetane Number
Figure 2.1 Boiling Point versus Cetane Number for various classes of
Hydrocarbon (19). 29. Boiling Range
Suitable distillation characteristics of a diesel fuel are essential for efficient combustion and good engine performance in high speed engines. Heavy fractions, while high in energy content and giving good fuel economy, may lead to the formation of deposits inside the engine and fuel injector systems. Light fractions, on the other hand, provide easier engine starting and more complete combustion but are generally low in energy content
Cloud Point and Flash Point
When a diesel fuel is cooled, there is a point at which the higher molecular weight hydrocarbons (i.e. the waxes) begin to separate and to appear to cloud or haze the fuel. This temperature is the cloud point, and is an indication of the low temperature useability of the diesel fuel. Cloud point specifications depend on geographical location and season: the most stringent specification in Australia is -
4°C (max). The flash point of a diesel fuel is not directly related to engine performance. It is a legal requirement concerning the safety precautions set down for fuel handling and storage.
Density and Viscosity
The density of a diesel fuel is an important parameter as it relates to the energy content of the fuel. Higher density fuels have a greater energy content per unit volume. Satisfactory viscosity is necessary to ensure the pumpability of the fuel.
Atomisation of the fuel at the injectors is also affected. 30. Storage (oxidation) and Thermal Stability
Refiners generally require their diesel fuel to have sufficient stability for storage of at least one year. Australian automotive diesel fuels are presently very stable because:
(i) much of the fuel is hydrotreated (desulphurised);
(ii) the fuels may be stabilised by additives; and
(iii) much of the diesel is derived from sweet indigenous crudes.
The accelerated stability test ASTM D2274 is generally included in automotive diesel specifications. Storage instability results from complex interactions between molecular oxygen, organocompounds of nitrogen, sulphur and oxygen, and the more reactive hydrocarbons species such as olefins and aromatics (18).
Thermal stability refers to the diesel fuels resistance to formation of deposits at the elevated temperature existing in fuel injectors and engine cylinders. Present limits include limits on ash and carbon residue to ensure engine operability. 31. 2.1.1.3. Jet Fuel
There are two basic types of jet fuel in general use: the kerosene type and the wide-cut gasoline type. The former is a modified illuminating kerosene: the latter a wider boiling-range which includes some gasoline fractions. The basic civil jet fuel specification used in the USA is ASTM D1655 giving the requirements for three types of fuels:
(i) Jet A, a nominal -40°C freeze-point kerosene;
(ii) Jet A-1, a nominal -50°C freeze-point kerosene; an
(iii) Jet B, a wide-cut gasoline grade fuel.
Jet A is used by domestic carriers within the United States and Jet A-1 is the standard grade jet fuel used for international flights. Within Australia, only kerosene type aviation turbine fuel (Jet A-1) is used.
Jet fuels consist almost entirely of pure hydrocarbon compounds. As they are produced from straight-run distillate, they contain virtually no olefins. Olefins are limited by fuel specifications (see Table 2.4) as are aromatics. Aromatics are limited because they do not bum as cleanly as paraffinic hydrocarbons and can cause smokiness and carbon deposition in the engine. Aromatics are also said to degrade elastomers in the fuel system and to increase the luminosity of the combustion flame, which is said to affect the combustion chamber life (18). 32.
Table 2.4 Specification for Jet A-1 Aviation Turbine Fuel (18)
Property Test Specification
Appearance Visually clear and bright and free from sediment, suspended matter and undissolved water at temperature of delivery.
4 (max) Colour IP l7B Composition
Total Acidity D3242 0.015 mg KOH/g (max) Aromatics Dl319 20% vol (max) Olefin Content D1319 5% vol (max) Total Sulphur D2785 0.30% mass (max) Mercaptan Sulphur or IP 104 0.003% mass (max) Doctors Test IP 30 Negative Volatility
Distillation: D86 Initial boiling point Not limited, to be reported 50% and 90% recovery points 200°c (max) 20% volume recovery 300°C (max) Final boiling point 1.5% vol (max) Residue 1.5% vol (max) Loss 775 kg/L (min) Density at 15°C D1298 830 kg/L (min) 38°C (min) Flash Point (Abel),°C IP 170 Fluidity
Freezing Point,°C IP 16 -47 (max) Viscosity at -20°C, cSt D445 8 (max) Corrosion
Corrosion, Silver Strip 4 h @ 50°c IP 227 1 (max) Copper Corrosion, Bomb 2 h@ 100°c D130 1 (max) 33.
Stability D1840
Thermal Stability (JFfOT) D3241 Filter P, mmHg Pre-heater deposit 25 (max) Less than code 3, no 'Peacock' colour deposits. Contaminants
Existent gum, mg/L D381 70 (max) Copper content,ug/kg IP 255 150 (max) Water Separation Index, modified D2550 85 (min) Water Reaction D1094 Interface Rating 16 (max) Separation Rating 2 (max) Combustion
Aniline Gravity Point D661 and 4800 (min) or D1298 Specific Energy D2382 42.8 MJ/kg (min) Smoke Point, mm Dl322 19 (min) plus Naphthalene D1840 3% vol (min)
Volatility
The volatility requirements for jet fuel are controlled by flash point, density and a small number of distillation points. These distillation points are chosen to give a properly balanced fuel. The distillation end-point is chosen to exclude heavy material which would give poor vaporisation and thus affect engine performance. 34.
Fluidity
As jet aircraft fly at high altitudes and experience low temperatures, it is necessary to ensure that waxes (solidified hydrocarbons) do not separate from the fuel and block lines, filters, nozzles etc. Thus a maximum freezing point is specified.
Also, the fuel viscosity at low temperature must be such that adequate fuel flow and pressure are maintained under all operating conditions.
Combustion Quality
The range of a jet aircraft depends on the calorific values of the fuel, predetermined by its heat of combustion and density. Typically, paraffinic fuels have
a higher gravimetric calorific value than those of naphthenic fuels, but the latter
have superior calorific value on a volume basis. The relative importance of volume
or weight is dependent on aircraft design and the flight pattern. As well as having
high heat of combustion, it is necessary that the fuel bums with a minimum of
smoke formation, carbon deposition and flame retardation.
Thermal Stability
During flight the fuel is subjected to considerable heat input from kinetic
heating of the airframe and from the use of bulk fuel as a coolant for engine oil,
hydraulic and air-conditioning equipment. Thus the fuel must perform satisfactorily
without the formation of deposits which may effect engine operability. 35.
2.2. COAL DERIVED LIQUIDS AS TRANSPORT FUELS
2.2.1. Suitability of Coal Derived Liquids
The suitability of coal derived naphtha as a source of high octane gasoline has been mentioned. These gasolines could also be very useful as blendstocks, with the capability of compensating for the inferior octane quality of other (e.g. petroleum derived) blendstocks. Special attention is required concerning the relatively high concentrations of heteroatom compounds (particularly oxygen and nitrogen) as these compounds have undesirable effects on fuel stability, may poison naphtha reforming catalysts and are converted to atmospheric contaminants during combustion.
Similarly, their high aromatic content may in future prove problematic.
Coal derived distillates are considered too high in aromatic compounds for use as specification diesel fuels. Hardenburg discusses the effect of aromatics on diesel fuel performance in his study " Thoughts on an ideal diesel fuel from coal"
(19). The content of aromatics in conventional diesel fuel is about 15%: this includes monocyclic to tricyclic and, in some cases, even tetracyclic, aromatics. These are found to be highly undesirable, as they give rise to unacceptable exhaust emissions and decrease the cetane number of the fuel considerably. Cetane number decreases with increasing aromatic content (see Figure 2.2), particularly if the aromatics are 36. polycyclic. As ignition delay increases with decreased cetane number, Hardenburg
(19) believes this to be the major cause of the exhaust problems with highly aromatic diesel fuels. High concentrations of heteroatoms may also present significant problems with storage stability and undesirable emissions.
The aromatic content of coal derived kerosenes also present problems for their use as jet fuels. Figure 2.3 shows the effect of aromatic content on smoke point. Current jet fuel specifications require a smoke point of 19 and it is considered that if the smoke point requirement is met, all other specification properties will be acceptable.
Thus, for coal derived liquids (specifically the distillate and kerosene fractions) to be used as diesel and jet fuels, extensive hydroprocessing is required to reduce the aromatic and heteroatom contents. 37.
Celane Number 50...------~
45
40
35
30
25
20
15
10 L___ .,__ __ ...,__ __ ---'------'------'------'---.,__--.....__-~ 0 10 20 30 40 50 60 70 80 90 Aromatics, Volume % Figure 2.2 Dependence of Diesel Fuel Cetane Number on Aromatic Content (19).
Jet Fuel Smoke Point 28 ~------,
24
20
16
12
8
4
OL__ ___----1... _____,_ ____ .,__ ___ ~----~---~ 0 10 20 30 40 50 60 Aromatics, Volume % Figure 2.3 Dependence of Jet Fuel Smoke Point on Aromatic Content (19). 38. 2.2.2. Hydroprocessing of Coal Derived Liquids
An extensive study of the upgrading of coal hydrogenation liquids into specification grade diesel and jet fuels has been made by Sullivan et al. (21).
In this study Sullivan subjects three coal hydrogenated syncrudes to various modes of refining process: hydrotreating, hydrocracking, and reforming. All erodes
studied are characterised by low hydrogen content and high concentrations of oxygen
and nitrogen impurities, when compared to petroleum crudes.
Three refining routes (to give differing product slates) were evaluated:
maximum jet fuel plus some gasoline; all gasoline; and maximum heating fuel plus
some gasoline. Only the first two modes are relevant here, namely the Jet-Fuel Mode
illustrated in Figure 2.4, and the All-Gasoline Mode illustrated in Figure 2.5. The
syncrudes were hydrotreated in fixed bed, down-flow, pilot plants. High and low
pressure product separators, recycle hydrogen and extensive control and monitoring
facilities were used.
Proprietary Chevron catalysts were used in different pilot plant simulations of
the syncrude hydrotreater: ICR 106 and ICR 113. The ICR 106 catalyst contained
nickel, tungsten, silica and alumina; the ICR 113 catalyst contained nickel,
molybdenum, silica and alumina. 39.
Gas i Light Naphtha
Gas ~ Motor Gasolin, i i ; C .2 H-Coal _., Jet Fuel Mode ~ ~ : Reforming Syncrude -.; Heavy Naphtha Reformate Hydrotreating i5 h
: Hydrogen Kerosene Jet Fuel
Diesel Fuel
- Gas Oil Refinay Fuel
Figure 2.4 Simplified Flow Diagram for the Refining of H-Coal Syncrude to Jet
and Diesel Fuels (21 ). 40.
Gas f I I Light N:iphtha
Ii
'I .2 i Hydrogen Gas ] '' •• -;; Motor G 0 All Gasoline Heavy orBTX H-Coal Naphtha Producti, Mode I Syncrude - Refonning - Hydrotreating 0 ! ! ; ·~ i ' Hydrogen I I I i Recycle Hydrocracking • Hydrogen
Figure 2.S Schematic Flow Diagram for the Refining of H-Coal Syncrude to All
Gasoline (21 ). 41.
For the Jet-Fuel Mode, the syncrude hydrotreater was operated at high severity and the by-product naphtha was fed to a second stage catalytic reformer to produce a BTX (benzene, xylene and toluene) rich liquid which was blended with the light naphtha to produce gasoline. The distillation column also had take off points for kerosene and gas oil, which may be suitable for jet and refinery fuels respectively, and may be blended to produce diesel fuel. The severe hydrotreater conditions used were: 2300 psi hydrogen partial pressure; 0.5 LHSV; between 750°F and 800°F average catalyst temperature; and 8000 SCF/bbl recycle gas rate.
In All-Gasoline Mode the syncrude was subjected to a less severe hydrotreatment to cleanse the oil and stabilise the reactive components. Distillation of the product was carried out together with a recycle of heavies from the distillation column. Light gases were removed from the column, and heavy naphtha was drawn off for reforming to a BTX rich liquid. The light naphtha and BTX liquid were then blended to produce gasoline. Hydrotreatment conditions used in this gasoline mode were: 2300 psi hydrogen partial pressure; 1.0 LHSV; 750°F; and catalyst ICR 106. 42.
Results of this study have shown that hydrotreating and, when necessary, hydrocracking are suitable routes for convening coal derived liquids to transpon fuels, specifically gasoline, jet and diesel fuels. The yield of transponation fuel was found to be almost equivalent to the original volume of syncrude and the heteroatoms are ultimately removed as hydrogen sulphide and ammonia. It was found that cetane number was the limiting specification for diesel fuel and smoke point was the limiting specification for jet fuel. When these specifications were met all other specifications were met, but the reverse was not found to be so. In order to meet specifications for jet and diesel fuels the aromatic content of the fuels had to be reduced dramatically, to less than 4% LV (21). 43. 2.3. BACKGROUND TO PRESENT STUDY
Diesel and jet fuels produced from petroleum erodes can contain up to 20% aromatics and still satisfy cetane number and smoke point requirements respectively.
This is due to the high proportion of linear paraffins in the petroleum erodes, such that high levels of aromatics can be tolerated.
Coal derived liquids, on the other hand, have a very low paraffin content, rarely greater than 10%. This, coupled with a high naphthene content, makes them unsuitable as diesel and jet fuels without extensive hydrotreatment. Upon hydrotreatment aromatics are converted to naphthenes. The cetane number and smoke point of these compounds is also inferior to linear paraffins and this restricts their use as diesel and jet fuels.
Coal derived liquids can be hydrotreated to produce specification jet and diesel fuels, if the aromatic content is reduced to less than 4% on a liquid volume basis (21).
The ideality of a process which could convert aromatic compounds from coal derived liquids to linear paraffins during syncrude hydroprocessing can therefore be seen. 44.
Rather then reducing the aromatic content of coal derived liquids to produce specification grade jet and diesel fuels, a study by White (22) has shown that these fuels can be obtained by blending selected aromatics and compounds derivable from
aromatics. Table 2.5 gives a listing of possible blending compounds and their relevant properties. The majority of compounds listed are present in coal derived
liquids, though not necessarily in large quantities. It is considered that, if access is
available to the compounds listed in Table 2.5 (particularly bicyclohexyl and
cyclohexylbenzene), they may be blended to produce specification grade jet and
diesel fuels (22).
Table 2.5 Possible Jet Fuel and Diesel Fuel Blending Compounds Derivable From Aromatics (22).
No Compound Bpt •c Freezing Octane Cetane Cloud Pnt Smoke Pnt Density Point ·c No. No °C °C g/ml
I npropylcycloheune 157 .95 17.8 45 50 45 0.794
II Hydridanc 162 -22 23 0.862
ill nButylcyclohcxanc 181 -75 <10 50 54 50 0.799
IV Dccalin (cis) 196 -42 32 40 35 22 0.897 (!rans) 187 -30 35 0.870
V Tctralin 207 -35 96 15 <-20 6 0.969
VI Naphthalene 218 80 <<-20 1.015
VII Bicyclohcxyl 238 4 53 48 30 0.891 . <-10 0.950 vm Cyc!ohcxylbcnzcnc 235 7 0.S66 IX Biphcnyl 256 70 21
.. ..
.i. .i.
v, v,
Motor Motor
Gasoline Gasoline
4 4
BTX BTX
H H
i---
r r
Gas Gas
l l
Reformer Reformer
H2 H2
2 2
H
r r
Reactor Reactor
Hydroalkylalion Hydroalkylalion
Naphtha Naphtha
Light Light
i... i...
Naphtha Naphtha
n-PCH n-PCH
n-BCH n-BCH
DEC DEC
T T
C, C,
a, a,
~HIN ~HIN
H2 H2
Reactor Reactor
Hydrogenation Hydrogenation
I I
I I 1------
e e
:, :, C C
-
~ ~
i i
C C
C C
'.;j '.;j
Gas Gas
·.i ·.i ~8 ~8
o o
:; :;
:3 :3
.2 .2
-
IN IN
Kerosene Kerosene
Distillate Distillate HIN HIN
• •
I--
~ ~
BP BP
BCH BCH
CB CB
Column Column Distillation Distillation
.. ..
• •
• •
• •
Hydrocracker Hydrocracker
Recycle Recycle
Hydrolreater Hydrolreater
H.i-~ H.i-~
H2 H2
.. ..
Syncrude Syncrude
N N
N N
::, ::,
:n :n
V, V,
(') (')
-._, -._,
(b (b
a a
,,--. ,,--.
0 0 '"' '"'
0. 0.
(JQ (JQ
:c :c
(b (b
0. 0.
'"' '"'
i:: i::
(') (')
::, ::,
Cl) Cl)
"'O "'O
0', 0', '< '< N N
.., ..,
ti, ti,
"rl "rl C C
'< '< ~ ~ 46. Such processing may be represented as in Figure 2.6. The codes used in
Figure 2.6 have the following meanings:
1>EC .:= dec.ane.. HIN = hydridane
IN = indane
n-PCH= n-propylcyclohexane
n-BCH= n-buty lcyclohexane
BCH = bicyclohexyl
CB = eye lo hex y lbenzene
BP = biphenyl
BTX = benzene, toluene and xylene
* = blending components for jet and diesel fuels
Syncrude is hydrotreated to reduce sulphur, nitrogen and oxygen levels so that poisoning of catalysts in subsequent treatments is avoided. A catalyst comprising of oxides of nickel and/or cobalt together with tungsten and/or molybdenum on an alumina support may be used. The catalyst is sulphided prior to use and reacted under all-gasoline mode or under low severity conditions.
Distillation of the product is then carried out, with kerosene and distillate fractions being taken off to be used as blendstock for diesel and jet fuels. Bottoms from this distillation, some hydrotreated product and a recycle stream from the hydrocracker, then enter the main distillation column. Off takes of n-propylcyclohexane, n butylcyclohexane, indane, hydridane and decalin are possible, with the light naphtha 47. fraction (180 - 190°C) sent to the reformer. Non-distilled or heavy products from the main distillation column and recycle hydrocarbons (essentially indane and hydridane) are combined and treated in the hydrocracker to increase the yields of hydridane and substituted cyclohexanes. The reformer treats heavy naphtha from the main distillation column to form a BTX rich liquid fraction: this stream can then be combined with the light naphtha from the main distillation to produce gasoline blendstock. The reformer catalyst is typically Pt, or Pt/Re, on an alumina support.
The BTX fraction could, alternatively, be converted to non-fused double ring compounds such as bicyclohexyl or cyclohexylbenzene, in a hydroalkylation reactor, using a suitable catalyst and conditions. Light fractions from the hydroalkylation reactor, which would include unconverted BTX and naphthenes, could be sent to the reformer for recovery of hydrogen and BTX. The ratios of bicyclohexyl (YID to cyclohexylbenzene (V1ll) could be adjusted in the hydrogenation reactor using a suitable hydrogenation catalyst.
Thus White (22) has shown that all compounds (I to IX) are accessible from coal derived syncrudes using this detailed process scheme, in addition to straight run kerosene and distillate. 48.
To illustrate the operability of the process, a sample of anthracene oil
(boiling range 250 - 350°C) was used to represent the syncrude. The oil was hydrogenated using a Co-Mo/Al20 3 catalyst at conditions of:- 1.2 LHSV; reaction temperature = 240°C; pressure = 24 MPa; and hydrogen to liquid feed rate of 1500
L H2 STP/L liquid feed. A naphtha fraction was distilled off and catalytically reformed using an 0.3% Pt and 0.6% Cl on alumina catalyst. The reformate was characterised by gas liquid chromatography, see Table 2.6. A yield of >70% of components suitable for jet and/or diesel fuel usage was obtained from the naphtha.(22) 49.
Table 2.6 Major Components in Reformate of Naphtha (22)
Compound Weight% Compound Weight% Naphthenes Aromatics
Most predominant Benzene 5.76 naphthene, Hydridane Toluene 6.28 1.03 Ethylbenzene 14.69 Xylenes 5.25 n-propylbenzene 18.45 Ethyltoluene 14.01 Indan 17.77 n-butylbenzene 1.80
Total: 84.01
Table 2.7 Synthetic Mixtures to Approximate Jet and Diesel Fuels (22).
Component Percentage by Volume
Kl K2 K3 K4 Dl 02 D3
n-propylcyclohexane 25.1 27 28 - 9.9 - 12 n-bu ty lcyclohexane 24.9 29 42 60 14.7 5 13 Decalin 19.9 22 13 20 19.7 5 23 Tetralin 20.2 12 12 5 5.0 - 6 Cyclohexylbenzene - - -- 4.9 -- Bicyclohexyl 9.9 11 5 15 34.7 90 40 Biphenyl -- - - 11.1 - 6 50.
Table 2.8 Test Results on Synthetic Mixtures (22)
Test Standard Unit "Jet Fuel" Kl "Diesel Fuel" Dl Spee. observed spec. observed
Density 04052-81 gm/L 0.830 0.864 0.889 20"C
Smoke Point IP 57/55 mm 25min 17 na na Flash Point 03243 or 056 °C 38min 42 38min 60 Cloud Point "C 1 max -10 Freezing IP16n3 °C -50°Cmax -30 +3 max 5 Point Aniline Point 0611-77 "C 28.3 Kinematic 0445-79 cSt l.9-4.l l.81 Viscosity at40 °C
Table 2.9 Test Results on Synthetic Mixtures (22)
Test "Jet Fuel", Observed Diesel Fuel observ ..
Standards, units and K2 K3 K4 02 03 specification as in Table 2.9. Density - - - - - Smoke Point 24 24 - -- Flash Point 49 - - 80 60 Freezing Point -45 -40 -80 -20 -15 (crystals) Freezing Point -25 -30 -30 0 -10 (clear) Cetane number na na na 45+ 50.5 Kinematic na na na 2.9 l.9 Viscosity 51.
Synthetic mixtures were prepared from various components listed in Table
2.5: both kerosene simulations (designated K) and distillate simulations (designated
D) were made (see Table 2.7 for compositions). The first two mixtures, Kl and Dl, were subjected to various petroleum industry standard tests to gauge their suitability for use as jet and diesel fuels respectively (see Table 2.8 for results). Considering the arbitrary nature of the mixtures, the results were considered very promising. The cetane number of Dl was estimated at 20. Engine testing of the mixture however, gave a value of 43, which was considered surprising given the high aromatic content of the mixture. Modifications to the mixtures were carried out to improve the freezing characteristics of both fuels and the smoke point deficiencies of the jet mixture. Test results of these samples is given in Table 2.9. It can be seen that near
specification jet and diesel fuels were prepared by blending rather unconventional mono-substituted cyclohexanes and two-ring nonfused cycloalkanes.(22)
Most processes referred to and required for Whites' (22) approach to
syncrude upgrading involve off-the-shelf technology, but two important exceptions
exist:- the hydroalkylation of aromatics and subsequent hydrogenation. Two key
reactions have been identified which require · catalyst/reaction development and
optimisation:- 52. (i) hydroalkylation of benzene to produce
cyclohexylbenzene;
and (ii) hydrogenation of cyclohexylbenzene to produce
bicyclohexyl.
These reactions are also considered to be of interest to refiners faced with the task of removing benzene and other aromatics from the gasoline pool, and ultimately with the task of utilising them, presumably as petrochemical feedstocks. A number of alternatives exist for removal of benzene such as, for example, solvent extraction, hydrogenation, or removal of benzene precursors (by boiling cut adjustment) from reformer feedstocks. Potentially, hydroalkylation of aromatics provides a viable option for upgrading gasoline stocks to meet the proposed limits on aromatics content. Furthermore, given the likelihood of a glut of aromatics for use as petroleum feedstocks, this technology provides a timely option for increasing distillate fuel production to meet expected future demand. 53.
CHAPTER 3. CATALYST DESIGN
It has been shown in the preceding chapters that, in order to use coal derived liquids as diesel and jet fuels, extensive hydroprocessing is required: firstly, to reduce the levels of heteroatoms such as sulphur, nitrogen and oxygen; and secondly,
to reduce the levels of aromatics. A possible route for aromatic reduction has been
identified by White (22) wherein hydroalkylation of benzene to produce
cyclohexylbenzene followed by further hydrogenation to bicyclohexyl are the key
steps. It is therefore necessary to consider some aspects of these reactions and the
catalysts that promote them.
3.1. HYDROGENATION OF BENZENE
The hydrogenation of benzene on metals is widely used and studied as a test
and model reaction in heterogeneous catalysis (35). There is, however, a great deal
of controversy regarding the reaction mechanism. Basic considerations such as the
nature of the active sites, the respective role of benzene and hydrogen adsorption,
and the possible changes in mechanism over a wide temperature range remain
unresolved.
Since the pioneering work of Sabatier (23), nickel has been known to be a
very effective hydrogenation catalyst. The hydrogenation of benzene has been 54. investigated over a range of Ni-Si02 catalysts with regard to the influence of surface structure (24,25,26,27 ,28). The rate of reaction per unit metallic area for nickel catalysts is generally considered to be structure insensitive (24,25). However, studies by Coenen et al.(26) have found small nickel particles (size smaller than 1.2 nm) to be less active than large particles by a factor of 5. Reports of a drop in activity which exceeded the fall of specific surface area when nickel catalysts were thermally sintered have also been made (27).
Martin and Dalmon (28) examined the structure sensitivity of benzene hydrogenation over nickel catalysts, both supported and unsupported, at high and low temperatures. They found that, in agreement with the trends observed by Coenen et al.(26), the reaction is structure-sensitive only for particles less than 6nm and at low temperatures. As particle size increases the activity first increases, reaches a maximum, then remains constant or decreases.
It was also reported that the structure-sensitivity of benzene hydrogenation is effected by the temperature at which the investigations are carried out, with very little structure-sensitivity being noted under higher temperature conditions (572 K)
(28).
Martin and Dalmon (28) further examined the role of preparation conditions and their effect on activity. They found, in agreement with other studies using similar conditions (24,25,26,29,30), that when samples were reduced at moderate 55. temperatures (< 750 K) variation of activity with particle diameter is not very pronounced. However, if high temperature reduction is carried out, the large particle diameters obtained show a sharp decrease in activity as the diameter increases due to the poorly dispersed samples having strong metal-support interaction. This result was in agreement with the findings of Shelwood et al. (27).
Investigations of the nature and amount of adsorbed species and the occurrence of competitive or non-competitive adsorption of benzene and hydrogen onto the catalytic surface have also yielded conflicting theories.
Van Meerten et al (31,32) deduced from gravimetric and magnetic experiments that hydrogen and benzene adsorbs on different sites on a nickel surface, the benzene coverage being independent of the H2 pressure. On the basis of this non-competitive adsorption and the surface heterogeneity, they calculated the respective abundance of benzene and hydrogenated species on the active part of the surface from kinetic data. A large part of the nickel surface was considered to be occupied by cyclohexane precursors. Similar conclusions were obtained by Franco and Phillips (33).
Kehoe and Butt (34), believed that the zero-order behaviour with respect to
benzene indicated that the surface was saturated with benzene, with the attack of
gaseous hydrogen molecules on the n-bonded benzene complex being rate
determining. 56.
Using an isotopic transient studies on nickel catalysts, Mirodatos (35) reported that quantitative amounts of material adsorbed on the catalysts could be directly measured in the course of the benzene hydrogenation. In the low temperature range (< 120°C), the surface occupancy was described as follows:
(i) the whole surface of the catalyst is covered with a layer of reversibly
adsorbed benzene which appears to be little dependent on the chemical nature
of the surface. This adlayer decreases very rapidly as temperature is
increased;
(ii) the main species specifically chemisorbed onto the metallic surface is
hydrogen;
(iii) the concentration of irreversibly adsorbed benzene, leading through a
unidirectional step to gaseous cyclohexane, is very low, occupying less than
l % of the exposed metallic surface. This concentration of reactive adspecies
or reactive sites does not depend on the nickel dispersion and is similar for
supported and unsupported nickel, thus indicating the structure-insensitive
nature of the reaction. (35)
Combining this transient approach with a kinetic study carried out in the steady-state regime Mirodatos, Dalmon, and Martin (36) investigated the mechanism of the hydrogenation of benzene to cyclohexane. Experiments were carried out in the 57. low temperature range ( <400 K) as at higher temperatures, side reactions (such as benzene hydrogenolysis and cyclohexene formation) occur. The proposed rate limiting step in benzene hydrogenation on nickel at low temperature was suggested to be a two-body process between (i) an ensemble of about 4 Ni atoms occupied by a chemisorbed molecule of benzene (this benzene being supplied by the loosely adsorbed reservoir that coats the entire surface of the catalyst, with the chemisorption occurring as soon as suitable sites appear on the Ni surface covered mainly with hydrogen) and (ii) a molecular species of hydrogen, probably coming from the gas phase with its concentration related to PH2 (36).
Studying Pt/y-Al20 3 catalysts, Basset et al. (37) confirmed that benzene hydrogenation is structure-insensitive, and found that the activity per platinum site
does not depend on the complex used for impregnation (H2PtC16 or Pt(N0i)i(NH3}i},
Pt dispersion (12%-67%), or the chlorine content of the catalyst. The authors concluded that benzene and hydrogen are adsorbed on two adjacent platinum sites.
However, the dilution of the catalyst by y-Al20 3 was not taken into account and the possibility of hydrogen spillover not considered.
The effect of hydrogen spillover on benzene hydrogenation has been examined by a number of workers (37,38,39,40,41,42), with Antonucci et al. (41) observing an increase in specific activity by a factor of 4 when the degree of 58. dilution of Pt/y-Al20 3 with y-Al20 3 was 50/1 (volume ratio). These results indicated the hydrogenation reaction occurred not only on the Pt surface, but also between spillover hydrogen and benzene adsorbed on y-Al20 3 and between spillover hydrogen and benzene from the gas phase (41 ).
This was further examined by Ceckiewicz and Delmon (42) using 0.6 wt%Pt/y-Al20 3 mechanically mixed with y-Al20 3• Taking into consideration the possibility of interference from reactant impurities and y-alumina contamination, they found a definite synergistic effect from the mixed catalyst with conversions always higher than the individual mixture components taken alone over the temperature range 70°C-150°C and up to a dilution ratio of 1:20.
Support effects have also been noted with palladium catalysts
(43,44,45,46,47). It has been found that the specific activity of Pd catalysts can be enhanced if acidic supports such as Si02-Al20 3, Ti02 or zeolites are used, although these supports by themselves do not catalyse the reaction. One explanation for this support effect has been that a metal-support interaction exists which changes the chemical properties of the dispersed metal (45). Another possibility is that this enhanced activity could be due to the creation of additional active sites at the metal support interface (44,45). Chou and Vannice (46) have examined the kinetic behaviour of supported and unsupported Pd catalysts for this reaction. Finding that
Pd/Si02-Al 20 3 and Pd/fi02 catalysts when reduced at low temperatures (448K) were 59. much more active than Pd/Si02 and Pd/Al20 3 catalysts, they postulated the direct involvement of the support in the reaction process. Peri (48) has shown that benzene adsorbs on acid sites existing on oxide surfaces, and it was considered that adsorption was also possible on the metal-support interface region. Such adsorbed gas would readily react with hydrogen activated by the metal (45), thus increasing the surface concentration of adsorbed benzene available for reaction.
Bronsted acid sites are commonly found on Si02-Al20 3 and zeolites and can also exist on titania which contains impurities (49). The presence of chloride and the reduction temperature used have significant effects on the activity, with catalysts prepared from PdC12 showing higher activity when the catalysts were reduced only at low temperatures (448K). Activity enhancement has also been reported by addition of chloride by impregnation with HCl (45). It is known that halogens can polarise
surface hydroxyl groups on oxides and increase their acidity (50), and hydrogen chloride generated during the reduction in catalysts prepared from PdC12 can adsorb
on supports such as Ti02 and produce strong acidic sites (51). The observed activity
enhancement is attributable to the chloride which can generate strong Bronsted acid
sites. It has been shown that chlorination of Al20 3 at lower temperatures produces
Bronsted acid sites, while Lewis acid sites develop at higher temperatures (52). The
effects of chloride on the benzene hydrogenation activities of Rh/A1 20 3, Rh/Si02,
Ir/Si02, and Ir/Si02 were investigated by Marques da Cruz et al (53) and Tiep et al.
(54). These authors considered that no effect of chlorine was observed. However,
under conditions of low reduction temperature (448K) and with Si02-Al20 3 or Ti02 60. used as supports, results comparable to Chou and Vannice (46) were obtained.
Kinetic data obtained at higher temperatures showed evidence of deactivation for unsupported Pd samples and, although these samples were found to be completely regenerated in pure hydrogen, the process is relatively slow. Chou and
Vannice (46) considered the inhibition to be due to carbonaceous species generated by the dehydrogenation of adsorbed benzene. They proposed that the benzene adsorbed on acidic sites on the oxide surface near Pd particles in the adlineation region also reacts with hydrogen but that the sites are less susceptible to deactivation. Consequently steady-state specific activities are greater over catalysts containing these sites. 61. 3.2. HYDROGENATION OF SUBSTITUTED BENZENES
Kinetic analysis of the competitive hydrogenation of benzene and toluene has revealed that the ratio KT/B of the adsorption coefficients of toluene and benzene can be used to indicate changes in the electronic structure of supported metals (57 ,58). In this way structure insensitive reactions such as hydrogenations can be used to reveal relationships between the catalytic properties and the surface structure of the metals.
Toluene has been found to be more strongly adsorbed on metallic surfaces than benzene (57 ,58,59), due to the substituent effect of the methyl group which makes toluene more of an electron donor than benzene. Toluene molecules are therefore expected to form stronger Il-bonds with the metal (57). Large KT/B ratios have been associated with the electron-deficient character of the metal, with KT/B increasing upon increasing acidity of the support or upon the presence of electron acceptor
adsorbates, such as H2S. KT/B decreases in the presence of electron donor adsorbates,
such as NH3, or with reduction of support acidity e.g., by neutralisation by NaOH
(57).
Divergent results have been reported for the effect of alkyl-substitution on the
rate of hydrogenation of benzene derivatives. Since all work has been carried out in
the liquid phase, some discrepancies may be attributable to varying degrees of
diffusion limitation. Reaction rates for hydrogenation of mono-substituted n
alkylbenzenes are found to decrease with length of side-chain for the first two
carbon atoms but a further increase has little effect (60). The number of substituted 62. alkyl-groups is also found to effect the rates of hydrogenation, via the proposed relation (61):
2 rD = r- 0
Here r0 is the rate of benzene hydrogenation and r0 is the rate for a benzene derivative containing n substituents. It is also considered that not only the number of substituents but also their arrangement determines the reactivity of the molecule.
This has been shown for the xylenes, where the para-isomer is always the most reactive presumably because it is the least strongly adsorbed (62), and by trimethylbenzenes. Chain length has also been found to have an effect, with decreased rates of reaction possibly being attributable to (i) a lower concentration in the adsorbed layer due to steric effects; and (ii) an increase in the strength of adsorption of the reactant through the electron-releasing effects of the alkyl groups
(61).
Generally, hydrogenation of polyalkylbenzenes leads to complex mixtures of isomers and only the products from the isomeric xylenes have been estimated with
any accuracy. All three xylenes appear to yield products having predominantly the
cis-configuration on hydrogenation at room temperature. With increasing temperature
the proportion of trans-isomers are found to increase for the o- and p- xylenes.
Above 150°C the trans-isomers are found to be the major products. 63.
As interest in the catalytic hydroprocessing of heavy synthetic and petroleum based crudes in the petroleum industry increases, the search for catalysts active for hydrogenation of aromatics, hydrodesulphurisation (HDS), hydrodenitrogenation
(HON), and hydrodeoxygenation (HDO) of the high molecular weight compounds found in these materials becomes of great importance. Catalysts studied include
Co0-Mo0/y-Al20 3 (63,64), Ni0-Mo0/y-Al20 3 (65), and Ni-Wfy-Al20 3 (66).
Hydrogenations carried out using these catalysts are typically the slowest of the above listed reactions. Model compounds such as biphenyl and naphthalene are commonly studied.
As reported, kinetic data for the hydrogenation of aromatic compounds has generally been limited to benzene, with phenol and toluene being studied to a lesser extent Extensive reaction network studies have been required for the catalysts referenced above. Sapre and Gates (63,64), have carried out batch reactor experiments on benzene, biphenyl, naphthalene and 2-phenylnaphthalene hydrogenation using a sulphided Co0-Mo0/Al20 3 catalyst. In the experiments carried out with benzene as the reactant, cyclohexane was the only primary product observed. In the biphenyl study, cyclohexylbenzene was found to be the primary product, and this was subsequently convened to bicyclohexyl. Three unknown hydrocarbon products were also obtained. These were considered to be isomerisation products from cyclohexylbenzene namely, cyclopentylbenzene with methyl substituents on the five-membered ring. 64.
This was supported by the fact that the distribution of these products were in a ratio of roughly 3:3: 1 molar proportions, since cyclohexylbenzene would be expected to give only the three isomers shown above, and the formation of the third would be expected to be slower due to steric hindrance.
When naphthalene was hydrogenated, 1,2,3,4-tetrahydronaphthalene (tetralin) was the primary product, which was further hydrogenated to form cis- and trans decahydronaphthalene (decalin), trans-decalin being the predominant isomer.
Hydrogenation of 2-phenylnaphthalene gave two primary products, 2-phenyltetralin and 6-phenyltetralin, each of which was further hydrogenated to give phenyl substituted decalins. All experiments were carried out at 325°C and 75 atm (63).
Sapre and Gates (64) determined a model for biphenyl hydrogenation
catalysed by sulphided Co0-Mo0/Al20 3 using differential data obtained form a flow microreactor. The recommended equation for the forward reaction was considered to
be of the Langmuir-Hinshelwood form, expressing inhibition by biphenyl, H2S, and
H2 (64).
Studies on the hydrogenation of a series of substituted benzenes (including biphenyl and cyclohexylbenzene) and a series of ortho- and para-substituted phenols
using a Ni0-Mo0/y-Al20 3 catalyst at 340°C and 70 bar H2 were carried out by
Aubert et al.(65). They found that the presence of saturated groups such as ethyl or 65. bulkier groups such as cyclohexyl and cyclohexylrnethyl, did not significantly effect the rate of hydrogenation of the benzene ring, thus excluding the inteivention of steric factors in the processes of TI-adsorption onto the catalyst. When a second group capable of hydrogenation (such as phenyl) was directly bound to the benzene ring a slight increase in the rate of hydrogenation was obseived. Phenols were found to be more rapidly hydrogenated than their corresponding hydrocarbons, with the presence of ortho- and para- substituents leading to a lower reactivity when compared to phenol alone. As both ortho- and para- phenols were less reactive, the effect was considered to be due to electronic effects via delocalisation of TI-electrons due to the electron-donating ability of OH groups by resonance (65). 66. 3.3. ALKYLATION
Alkylation reactions, such as the alkylation of paraffins with olefins, have been studied since 1935. The reaction is believed to be ionic and can be understood in terms of carbonium ion chemistry. There has been some debate over whether the reaction occurs in the catalyst phase, the hydrocarbon phase, or at the interface.
For the alkylation of paraffins with olefins the reaction sequence is considered to follow a number of steps:
1. The initial formation of ions by addition of a proton to olefins,
ionisation of an alkyl halide, or by other means of producing ions.
2. Generation of ions from isoparaffin by a hydrogen transfer reaction.
3. Addition of ions to olefins forming higher molecular weight ions.
4. Saturation of the heavier ions by hydrogen transfer reactions with
isoparaffins present, before or after isomerisation.
5. Continuation of the reaction chain by means of ions produced in the
final hydrogen transfer step. 67.
These primary reactions have secondary reactions which occur consecutively and simultaneously:
1. Condensation of the original isoparaffin with itself, "self-alkylation".
2. Reaction of the original isoparaffin with produced isoparaffins to yield
isoparaffins of intermediate carbon content.
3. Disproportionation of produced isoparaffins to yield isoparaffins of
higher and lower molecular weight.
4. Polymerisation of the initial olefin followed by saturation of the
polymer through hydrogen transfer. (67)
Catalyst properties considered to be important are (1) the ability to add a proton to an olefin, and (2) the ability to abstract a hydride ion from a paraffin. The balance of these properties determines the properties of the catalyst and the course of the alkylation reaction (67).
Two types of acid-acting catalysts are active for alkylation: (a) the Friedel
Crafts type, which include A1Cl3, AlBr3, ZrC14, and BF3 in conjunction with small
amounts of the corresponding hydrogen halides; and (b) protonic acids, of which
sulphuric acid and liquid hydrogen fluoride are the principal acids used. 68.
3.3.1. Alkylation of Aromatics
The alkylation of aromatic hydrocarbons with olefins is of great industrial importance, with processes including styrene production from ethylbenzene, synthesis of acetone and phenol from isopropylbenzene, and manufacture of sodium alkylbenzenesulphonates for detergents.
The catalysts used for the ring alkylation of aromatic hydrocarbons consist of three categories of acids: (a) protonic acids, (b) Friedel-Crafts catalysts, and (c) oxide catalysts. The activity of the protonic acid decreases in the following order:
HF > H2S04 > H3P04 > CiH5S03H. Phosphoric acid or its modification, silicophosphoric acid (also known as "solid phosphoric acid") is used commercially for the reaction of propene with benzene to form isopropylbenzene (68). Sulphuric acid does not catalyse the ethylation of benzene. It is however effective for alkylation of benzene with higher alkenes. Hydrogen fluoride is the most efficient catalyst for the alkylation of butenes and higher alkenes with benzene (68).
The general sequence of catalytic activity of Friedel-Crafts catalysts is
considered to be: AlBr3 > A1Cl3 > GaBr3 > GaC13 > FeC13, SbCl5 > ZrC14 > BF3 >
ZnC12 > BiCl3• Completely anhydrous metal halides are inactive as catalysts for the
alkylation of aromatics and a co-catalyst is required. Addition of HCl or HBr, alkyl
halide, or small amounts of alcohol or water activates metal halides (68). 69.
Many acidic oxides and mixed oxides have been used as alkylation catalysts, notably silica-alumina. Typically, alkylations of aromatic hydrocarbons (69,70,71), phenols (72), and heterocyclics (73) have been reponed with a variety of alkylating agents using this acidic catalyst. The reactions have generally been suggested to proceed via carbonium ion mechanisms (74,75).
Early studies by Venuto et al. (76) revealed the possibility of the alkylation of aromatics using highly acidic faujasites or zeolites. The catalysts were prepared by cation exchanging the synthetic sodium X or Y aluminosilicates with 5% rare
eanh solution (REC1 3.6H20) until a sodium level of 0.99 wt% was obtained. The catalysts were then calcined and reduced. Benzene, phenol and thiophene were found
to be alkylated with a wide variety of olefins ({;-C16), alcohols, haloalkanes, ethers and paraffins (76). The benefits of these oxidic catalysts, such as their high acidity, easy separation from products, absence of corrosive propenies and regenerability, has led to numerous developments in this field (77,78,79,80,81). 70. 3.4. ZEOLITES
Zeolites as synthesised or formed in nature are crystalline, hydrated aluminosilicates of Group I or Il elements. Structurally they comprise a framework based on an infinitely extending three-dimensional network of Si04 and Al04 tetrahedra linked together with common oxygen atoms. The isomorphic substitution of Si by Al gives rise to a net negative charge compensated by cations. Zeolites can be represented by the formula
M21nO.Al20 3.xSi02.yH20 where M is the compensating metal cation with valency n.
Structurally, the formula of zeolites can be represented by the unit cell composition:
Mll/D[(AlOi)I (SiOi)y].zH20 where the term in brackets is the crystallographic unit cell.
The most important zeolites in catalysis are, in decreasing order: synthetic faujasite (X and Y), mordenite, zeolites T, A and L and clinoptilolite (82).
Zeolite X and Y are synthetic aluminosilicates having a framework structure essentially the same as that of naturally occurring mineral faujasite. The faujasite
structure has been described by numerous authors. According to Breck (83,84), three 71. types of cation sites were distinguishable. Type I sites are located at the centres of the hexagonal prisms connecting the sodalite units and number 16 per unit cell.
These sites are accessible only to ions and molecules which can pass through the 2.4
A entrances to the hexagonal prisms. Type II sites are located on the six-membered rings constituting the unjoined hexagonal faces of the truncated octahedra. A unit cell contains 32 Type II sites. Type III sites are located in the supercages and number 48 per unit cell. The occupancy of the various sites depends upon the particular cation and the silicon-aluminium ratio. Zeolite X, with its higher concentration of framework aluminium atoms, contains a larger number of occupied
Type III sites as compared with zeolite Y.
With the availability of more precise x-ray diffraction data it became possible to define additional cation sites (85,86). The majority of workers have adopted the notation developed by Smith (86). Smith distinguishes between Type I and Type I' sites, the former being within the hexagonal prisms and the latter being mirror imaged in the sodalite cages. Similarly Type II sites are categorised as being situated just within the sodalite cages and adjacent to the unshared six-membered rings (Type
II'), and being slightly displaced from the plane of the unshared six-membered rings into the supercage (Type II); or as being further displaced into the supercages and adjacent to the unshared six-membered rings (Type II*), see Figure 3.1. 72.
l'
.I' 3 -. :r 3 bridging 2 /,,. 4· ring o· • I' • n' free . .u 6-nng 4 •a' br1dgin~_ ·.r· 4-ring 2 4 ·a free 6-ring ·a•
Figure 3.1 Idealised Projection and Section through a Sodalite Unit of Faujasite
(86). 73.
The factor which distinguishes between zeolite X and zeolite Y is the silicon- to-aluminium ratio. This ratio falls within the ranges 1 ~ Si/Al ~ 1.5 and 1.5 ~ Si/Al
~ 3, respectively, for zeolite X and zeolite Y. Table 3.1 shows the characteristics of the various classes of zeolites. The type of cation and its location are the principal factors responsible for the catalytic activity of zeolites.
Table 3.1 Structural parameters of zeolite catalysts (82).
Unit Cell composition Symmetry Type of porosity, Type accessible through windows of size (nm)
A Na12(Al02)1i(Si02)12 cubic cages (0.42)
X Na8iAl02)8iSi02)106 cubic supercages (0.75)
y Na5iAl02) 5iSi02) 136 cubic supercages (0.75)
Z(eolon) Na8_iAl02)8,iSi02)39.3 orthorhombic pores (0.67 x 0.7)
L Kg(Al02)g(Si02)21 hexagonal pores (0.7) erionite Nag( Al 0 2)/Si 0 2)27 hexagonal pores (0.64)
offreite Ks(Al02)s(Si02)13 hexagonal pores (0.64)
clinoptilolite NaiAl02) 6(Si02) 3o monoclinic pores (0.52 x 0.36)
A variety of active catalysts can be prepared by replacing the sodium cations
in zeolite X or Y frameworks with various multivalent cations through standard ion
exchange procedures. The only limitation to the nature of the cations that can be
introduced by ion exchange is the acid stability of the zeolite. Generally, the zeolite
is more stable in acidic solutions when the Si:Al ratio is higher. When the exchange
solution is strongly acidic, protons are also exchanged, thus creating strong acid sites 74. in the zeolite structure. The acid stabilities of various zeolites are shown in Table
3.2.
Table 3.2 Catalytically important properties of zeolites (82).
Zeolite Reaction in Temperature to Largest molecule 0.1 N HCl which thermally adsorbed (kinetic stable (K) diameter, nm)
A decomposes 973 ethylene (0.39)
X decomposes 1045 ({;F5) 3N (0.80) y decomposes 1066 z stable 1273 benzene (0.62)
L stable 1273 (C4F9) 3N (0.81) erionite stable 1023 n-paraffins (0.43) offretite stable 1238 cyclohexane (0.6) clinoptiloli te stable 1273 oxygen (0.35) 75.
3.4.1. Acidity of Zeolites
Both Bronsted and Lewis acid sites are found in zeolites. The former are protons attached to lattice oxygen atoms, while the latter can be charge compensating cations or at cation positions. Protons can be introduced into the
structure through ion exchange, thermal decomposition of the NH4-exchanged fonn, hydrolysis of water of hydration of cations or reduction of cations to a lower valence state (82).
Development of acidic OH groups via ion-exchange with polyvalent cations such as Ca2+, Mg2+, CeJ+, or Lal+ can be expressed as follows:
[Ca(OHJ]2+ -+ [Ca(OH)t + H+
Water molecules coordinated to polyvalent cations are dissociated by heating,
to give the following local structure.
[Ca(OH))+ H 0 0 0 0 6 0 0 / '-s/I ' Al / ' Si / ' Si / Al/ ' Si / ' /'./ './ ',,/ ',,/ ',,/'
Acidic sites are also formed by the reduction of transition metal cations. 76.
Hydrogen Y catalysts can be prepared by treatment with dilute mineral acids or by ammonium ion exchange followed by a calcination step in which ammonia and water are evolved at 600-650 K, and 770-820 K, respectively. Similar infra-red adsorption patterns are observed in the hydroxyl region for catalysts prepared by the two methods, indicating that the active sites are the same in each case (88,89). The transformation of NH4 Y can be schematically represented as in Figure 3.2:
0 0 0 0 0 0 0
/ 'si/ 'AI/ " Si / " Si / " Al / '-s./I ' / ' / ' / ' / ' / " / ' i-NH, H H 0 6 0 0 6 0 0 /'-/ /'-./'-./ /'-./'-. Si Al Si Si Al Si (HY) / " / " / ' / ' / ' / ' i-H,O o o o 0 6 0 o / " Si e Al / " Si / Si " / " Al / '-s/1 ' / '-./ ' / ' / ' / ' / " ( I ) ( n )
Figure 3.2 Development of Acidity in NH4 Y.
The transformations are confirmed by the changes in intensity of OH bands in the infra-red spectrum. Ward (90) showed that these bands (3540 and 3643 cm·1) increase with the calcination temperature up to 673 K, are constant at 673 - 773 K and decrease above 773 K. The band at 3740 cm· 1, attributed to OH groups of the amorphous part of the zeolite, behaves differently. By studying the adsorption of pyridine on Bronsted and Lewis acid sites, Ward (91) was able to monitor changes in intensity of characteristic infra-red bands at 1540 and 1420 cm·1, respectively. 77.
This was attributed to the change of acid site populations with the temperature of calcination, and evidence was obtained which proved the OH groups of HY zeolites are acidic, see Figure 3.2. The acidic character of HY can be expressed by the following equilibrium.
H+ H 0 0 0 0 0 0 0 0 / '\. / / '\. / /,/'\./'\./ Si Al Si Si Al Si / '\. / '\. / '\. /'\,./'\./'\.
Lewis acid sites are seen to develop above 770K. They were correlated to tri coordinated aluminium in the zeolite structure (I in Figure 3.2) formed by dehydroxylation of HY at higher temperatures. Thus two Bronsted acid sites are converted into one Lewis acid site. It is widely considered that the local structure (I)
in these zeolites is not stable and aluminium ions can be easily dislodged and exist
in the pores as cationic species such as (AlOt or the polymeric form, shown below
(92,93,94).
(AIQ)+ 0 0 0 0 0 0 0 / '-s( '-s( 's( "s( '\.AI/ "s(' / '\./ '\./ '/ '/ '\./'
Catalytic activity is largely related to the number of Bronsted acid sites rather
than Lewis acid sites. Enhancement of catalytic activity has been reported by
dealumination of zeolites (95,96), where the effect of a decrease in numbers of
active centres by loss of acidic OH groups is surpassed by an increase in acid
strength due to the changed Si/ Al ratio. Dealumination can be achieved by treatment
with acids, steam, EDTA or SiCl4 • The process is represented in Figure 3.3. 78.
Si Si I I 0 0 H 0 H I I I Si-0-Al O-Si-0 + 3H20 - Si-OH HO-Si + Al(OH), I I H 0 0 0 I I Si Si ( m )
0 H 0 0 0 I I I I I 0-Al 0-Si-0 + Al(OH), - O-Al-0-Si-0 + Al(OH)i+ + H20 I I I I 0 0 0 0
Figure 3.3 Dealumination Process (87).
The increase in acid strength of OH groups via dealumination is caused by interaction with aluminium species dislodged from the framework and left in the cavities, which are thought to be represented by
H (AJO+)P I,/ 0
""/"-Al Si / / ""/"' 79.
3.4.2. Applications of Zeolites in Acid Catalysis
The acidity of synthetic zeolites has attracted considerable attention in the context of their use as catalysts in alkylation reactions. They have the advantages of high acidity, ease of separation of the catalysts from reaction products, regenerability, absence of corrosive substances such as halogens and volatile acids, and lack the environmental disposal problems of more traditional catalysts, such as aluminium chloride.
One example of the use of zeolites industrially, involves the production of p xylene via alkylation of toluene with methanol. P-xylene is used in the manufacture of polyester fibres. The equilibrium amount of the para isomer in xylenes is approximately 24% of the total, and separation of the isomers is not easy, due to the closeness of their boiling points. The ability to alkylate toluene directly to p-xylene is of great industrial importance.
Yashima et al (77) investigated the distribution of xylene isomers produced
by alkylation of toluene with methanol using a variety of cation exchanged zeolites.
A relatively high selectivity towards the desired para-isomer was obtained. This was
greatly improved upon by using modified ZSM-5 zeolites, with selectivities up to
97% being reported (78,79). 80.
Another example of an industrially important alkylation reaction is m production of ethylbenzene, for use in styrene manufacture. The Mobil/Badger ethylbenzene process employs ZSM-5 as a catalyst due to its low rate of deactivation and low yields of non-selective byproducts (80,81). Alkylation is carried out in the gas phase at approximately 400°C and 2.0 MPa, with a molar benzene/ethylene ratio of 6 to 7. Conversion of ethylene is 100%.
Bi-functional catalysts that combine the acidic characteristics of a zeolite support and the hydrogenative ability of cation exchanged and impregnated metal species have been prepared and used (102,107-125). One example of development of these catalysts has been for hydroalkylation reactions and these are discussed in detail in the next section. 81. 3.5. HYDROALKYLATION
In 1934 Truffault (98) reported the isolation of cyclohexylbenzene and dicyclohexylbenzene from the hydrogenation of benzene over a nickel catalyst in the presence of P20 5• He concluded from these studies that cyclohexene and cyclohexadiene were intermediates in benzene hydrogenation. Interest in the hydroalkylation of benzene arose in the 1960's with a series of patents claiming that cyclohexylbenzene was obtained from the hydrogenation of benzene over certain supported metal catalysts (99,100,101). Catalysts used were a Group VIII metal on silica-alumina (101), a hydrogenation metal on a cracking catalyst (99), and a combination of a hydrogenation metal and heteropoly acid on a variety of supports
(100).
In 1968 Slaugh and Leonard (102,103) reported the production of cyclohexylbenzene (CHB) in high yield over a variety of supported transition metal catalysts. Both batch (autoclave) and flow experiments were carried out in these
studies and a variety of catalyst compositions were tested:- nickel/tungsten;
palladium; and platinum. Supports studied included zeolites and alumina.
The poisoning effects of both sulphur and water were noted in this work, as
were the beneficial effect of halogen (fluorine) addition. Highest yields of CHB were
reported using Ni/W/Al20 3, Ni/W/Al20 3-Si02, and Ni/zeolite, though operability was
also evident with Pd/Al20 3-Si02 and Pt/Al20 3-Si02• No Pt or Pd/Ni/Si02-Al20 3 82.
catalysts were tested. Selectivities of up to 79% CHB with 20% conversion of
benzene were reported. Noting the importance of both acidity and hydrogenation
activity a mechanistic scheme was proposed in which cyclohexene was featured as
the intennediate, see Figure 3.4. 0 Metal Site I H2
Acid Site Metal Site 0 ~ Q ...... 0-0 AddStte. MetalSite /2 / H2 Acid Site
+
+k~- Tetramers
Figure 3.4 Dual site Mechanism of Benzene Hydroalkylation (102). 83.
Supporting evidence for this mechanism was obtained by carrying out a 14C- tracer study of hydrogenating equal molar amounts of benzene and cyclohexene-14C over a Ni/F/W/zeolite catalyst. It was found that (1) only 0.1 % of the cyclohexene-
14C was converted to benzene; (2) ~ 74% of the cyclohexylbenzene resulted from reaction of cyclohexene-14C with benzene; and (3) 90% of the bicyclohexyl resulted from two molecules of cyclohexene-14C (102). This last point is especially interesting, indicating that the bicyclohexyl byproduct is not formed via hydrogenation of cyclohexylbenzene, but rather by alkylation of cyclohexene by cyclohexene.
Louvar and Francoy (104) investigated the hydroalkylation of both benzene and toluene. For the high and low pressure studies of benzene, physical mixes of a series of alkylation catalysts and hydrogenation catalysts were prepared and placed in an autoclave. Both Si02-Al20 3 and alumina supports, impregnated with nickel, palladium, platinum and arsenic, were tested in conjunction with acid catalysts
A1Cl3, A1Br3, P20 5 and BF3• Low conversion to cyclohexylbenzene were found
(<8%) and a major production of cyclohexane was reported, even in the presence of fluorides. The hydroalkylation of toluene was more successful, even under the low pressure conditions used (60psig). With Ni/Si02-Al20 3 and Pd/Si02-Al20 3 catalysts the major products formed were 1-(p-toly)- and 1-(m-toly)-1-methylcyclohexane. The authors considered the formation of hydroalkylation products from toluene to reflect a more stable tertiary carbonium ion intermediate, as compared to the more difficult to form cyclohexyl cation in the case of benzene. This was taken as evidence in 84. support of a carbonium ion mechanism. However, different catalyst systems were used in the respective studies, so no true comparison can be made.
A later study by Yamazaki et al.(105) also studied nickel supported on silica alumina as favoured by Slaugh and Leonard (102). By studying physical mixes of
Ni/Si02 with Si02-Al20 3 it was found that 42wt% Al 20 3 gives the optimal cyclohexylbenzene selectivity for benzene hydroalkylation. These results were correlated to the various acid strength ranges of silica-alumina with different compositions measured according to Benesi's method. Other products formed were cyclohexane, m- and p-dicyclohexylbenzene, and traces of methylcyclopentane and phenylmethylcyclopentane.
Experiments with a 5% Ni supported on silica-Al20i42%) catalyst gave results comparable with Slaugh et al (102) using nickel on alumina containing both tungsten and fluorine. Using optimal conditions for these batch experiments
(T=200°C and P=26 to 18 kg/cm2> a selectivity for CHB of 85.7% at X=9.3% was obtained.
Reaction of methylbenzenes were also examined under the same conditions as benzene hydroalkylation. Generally, the reaction time was found to increase with methyl substitution i.e. benzene < toluene < p-xylene < m-xylene < 1,2,4- trimethylbenzene. Selectivity to hydroalkylated products was 83-86%, except in the 85. case of p-xylene, where lower selectivity may have resulted due to steric effects.
When mixtures of benzene and methylbenzenes (1:1 molar ratio) were hydroalkylated, the main products were found to be cyclohexylmethylbenzenes. The ratio of cyclohexylmethylbenzenes to cyclohexylbenzene was found to increase with increase in the basicity of the methylbenzene. This was considered to support the formation of cyclohexylmethylbenzenes by electrophilic attack of cyclohexyl cation on methylbenzene. Yamazaki et al (105) concluded that the cyclohexyl cation may be formed from cyclohexene as an intermediate and gave further support to the mechanism of Slaugh and Leonard (102).
Using a series of catalysts made up of various supported noble metals and fused salts, Kamiyama et al. (106) also investigated benzene hydroalkylation. This
study found a maxima in CHB yield of 54.7% at 140°C after 5 hours using a
combination of Pd-Al20 3 and NaCl-AlC13 fused salts. Platinum, rhodium, ruthenium,
iridium and salts other than NaCl gave poor yields. The role of temperature in this
catalyst system was also crucial, with higher temperatures resulting in increasing
cyclohexane production.
Hydroalkylation of a mixture of benzene and toluene (1: 1 molar ratio)
resulted in the formation of six products, the major products being 4,4' -
dimethylcyclohexylbenzene (yield 19.8% ), cyclohexylmethylbenzene (yield 11.0% ),
methyldicyclohexyl (3.2%) and 4,4'-dicyclohexyl (3.3%). This product distribution 86. was significantly different from that reported by Yamazaki et al.(105), especially as no cyclohexane was detected.
The most extensive investigations of benzene hydroalkylation have been presented in the patent literature of the last twenty years, notably those of Crone and
Suggitt et al. (107-116) and Murtha et al.(117-125). Crone and Suggitt concentrated mainly on Ni and Co/W/zeolite catalysts (107-116), although some attention was given to rhenium (107), and mention was made of noble metals (111).
Discussion of increased activity and selectivity by steam treatment (atmospheric pressure, 1100°F, 0.5-8 hrs.) of the catalyst (rather than calcination) was made, with increases of conversion from 6.6 to 15.6% and of selectivity from 36% to 59% being observed for a Ni/W/zeolite catalyst (108). Catalysts of the same composition as mentioned above with rare earth metals added in a cation exchange step were also tested (110): again slight increases in selectivity were observed, although this approach was not developed. Consideration of the use of recycle was made in the later patents (114,115,116), as was catalyst regeneration in a non-oxidative atmosphere (113).
The most detailed investigation of catalysts for CHB production was given by
Murtha and co-workers in a series of patents from 1978 to 1982 (117-125). These studies employed calcined, acidic, nickel and rare-earth treated crystalline zeolites
(generally 13X), with Group VIII noble metals (0.01 -1.0 wt%, usually 0.1 wt%) 87. impregnated. Metals used included Rh,Pd,Pt,Re and Ir. Generally, results of 10% activity and 60-80% selectivity for CHB were obtained. Improvements in selectivity by treatment with halide ions (specifically organic chlorides and bromides (118)) were shown, as were selectivity increases due to small additions of water (25-100 ppm) in the feed or resulting from catalyst regeneration in air (124). Reaction conditions employed were P=500psi, LHSV=5-25, T=l50-200°C; calcination temperatures were variable with maximums of 524°C and 325°C being used (120). 88. CHAPTER 4. PROJECT OBJECTIVES
This study was undertaken to investigate the hydroalk.ylation of benzene using a multi-metallic zeolite supported catalyst. Catalyst and reaction optimisation, including elucidation of the roles of metals and support in this complex catalyst, was necessary. Testing of the optimal catalyst for hydroalkylation of coal derived liquids and petroleum derived fuels to investigate their possible conversion to jet and diesel fuel components was also carried out.
The aims of this work include:
(i) Development and optimisation of the hydroalkylation reaction conditions, in
reactor trials using benzene;
(ii) Study of the hydroalkylation catalyst composition, in order to identify the
optimum catalyst formulation and elucidate the roles of specific species in the
catalyst;
(iii) Study of the hydroalkylation catalyst acidity, including the relationship
between metal species and support acidity.
(iv) Examination of the effect of catalyst treatments on cyclohexylbenzene yield,
including regeneration after cycles of use. 89. (v) Investigation of the hydroalkylation of coal derived liquids using the
optimised catalyst and conditions.
(vi) Investigation of the hydroalkylation of a petroleum derived diesel fuel using
the optimised catalyst and conditions.
(vii) Investigation of the hydrogenation of model compounds.
(viii) Investigation of the hydrogenation of hydroalkylation reaction products, to
yield a distillate fuel blendstock. 90. CHAPTER 5. EXPERIMENTAL TECHNIQUES
5.1. INTRODUCTION
Experimental techniques employed in this study can be divided into four main categories: preparation of catalysts; characterisation of catalysts; design, building and operation of reaction studies equipment; and characterisation of reactor feed and product streams.
For the investigation of the hydroalkylation of benzene, it was decided to examine initially the previous work of Murtha et al. (117-124) and specifically to
study the metal exchanged and impregnated zeolite bifunctional catalysts. A variety
of catalysts were prepared, and the techniques employed and equipment used in their
preparation are discussed in this chapter, as are their resultant compositions and
characterisation.
A high pressure reaction studies system was designed and built to carry out
catalyst testing and reaction monitoring at high temperatures. Initially problems were
encountered in operation and extensive modifications were necessary: these will be
discussed in following sections. Due to safety concerns regarding the use of
relatively large volumes of benzene and other carcinogenic chemicals during these
studies, the reactor system was built to be wholly contained within a fume cupboard. 91. Analyses of reactant and product streams were predominantly carried out using on-line gas-liquid chromatography. Due to the complexity of some analyses and the difficulty of some separations, modifications and optimisation was also necessary. These will also be outlined in this chapter. 92.
5.2. MATERIALS
5.2.1. Gases
All gases used in this study were supplied by the Commonwealth Industrial
Gases Ltd., Australia. Table 5.1 lists the gas purity and area of application.
Table 5.1 Gas Specifications
GAS SPECIFICATION USE
Hydrogen Ultra High Purity Catalyst Reduction (99.999%) Reactant Industrial Dry (99.5%) Catalyst Reduction Reactant Gas Chromatography Nitrogen High Purity (99.99%) Diluent G.C. Carrier Gas Thermogravi metric analysis Air Industrial (<25 ppm Gas Chromatography water) Catalyst Preparation and Calcination Catalyst Regeneration Thermogravimetric analysis 5% Hydrogen in 4.8 ± 0.1 mol% Thermogravimetric Nitrogen Hydrogen, Volumetric analysis Standard Nitrogen in 30.2 ± 0.2 mol% Single-Point BET Helium Nitrogen, Volumetric Surface Area Standard Measurements Helium Ultra High Purity Metal Surface Area (99.999%) Measurements Carbon Monoxide Chemically Pure (99.5%) Metal Surface Area in Helium Measurements 93.
5.2.2. Chemicals
Table 5.2 lists the source, purity and application of all chemicals used in this study.
Table 5.2 Chemical Specifications
I CHEMICAL I SPECIFICATION I USE I Benzene Ajax (99.999%) Reactant Identification Toluene Rhone-Poulenc Reactant (95%) Identification Biphenyl Ajax (99.5%) Reactant Identification Xylene Ajax (99.99%) Reactant Identification
Methylcyclopentane Riedel De Haen Reactant (99%) Identification Hexane Ajax (95%) Reactant n-Pentane Ajax (95%) .. n-Heptane Ajax (99.5%) .. ..
Octane Ajax (99.999%) " Cyclohexane BDH (99.5%) Reactant Diluent Identification n-Decane Sigma (99%) Diluent Cyclohexylbenzene Tokyo Kasei (>99%) Identification Bicyclohexyl Tokyo Kasei (>99%) "
1,4- Aldrich (99%) " dicyclohex ylbenzene
Zeolite 13X Davison, Grace Catalyst Preparation 94.
Zeolite SA Davison.Grace II
Zeolite - Zeolon The PQ Corporation II Zeolite - Y W.R. Grace " (Octacat RN·)
Nickel chloride Ajax (98%) II
Ammonium chloride Ajax (99.5%) "
Cerous nitrate Ajax II
Hexachloroplatinic Strem II acid hexahydrate
Rare earth chloride Catoleum <·> II solution
Carbon Ajax (95%) Catalyst Treatment tetrachloride
Sodium carbonate Ajax (99.9%) "
Triethylamine Ajax (99.9%) Acidity Measurements
<"> Rare earth c··> - Lap3 45% chloride solution Ce02 6% contains - 48 wt% cone. HCI Pr60,, 6% 30 wt% Lanthanum Ndp3 18% concentrate <.. > Cl 15% max. 22 wt% water Sr0 1.5% " CaO 0.5% II
Fe20 3 0.35%" II Na20 0.6% 95. 5.3. CATALYSTS
The catalysts used in this study are listed in Table 5.3, they were prepared using the techniques of cation exchange, impregnation and calcination.
Table 5.3 Catalyst Specifications
Description
Catalyst Support Approx Cation Exchanged Approx Impregnated Designation Metals (wt%) Metals (wt%)
Ni REM Pt Ni
A 13X 5.0 10.0 0.10 0.00
B 13X 5.0 10.0 0.00 0.00
C 13X 5.0 10.0 0.15 0.00
D 13X 5.0 10.0 030 0.00
E 13X 0.00 10.0 0.10 0.00
F 13X 1.0 10.0 0.10 0.00
G 13X 2.50 10.0 0.10 0.00
H 13X 10.0 10.0 0.10 0.00
I 13X 5.0 0.00 0.10 0.00
J 13X 5.0 2.50 0.10 0.00
K 13X 5.0 5.0 0.10 0.00
L 13X 5.0 10.0 0.15 0.15
M 13X 5.0 0.0 0.0 0.0
N 5A 5.0 10.0 0.10 0.00
0 y 5.0 10.0 0.10 0.00
p Zeolon 5.0 10.0 0.10 0.00
Q Si02-AJi03 5.0 10.0 0.1 0.00
R "(·Al203 0.00 0.00 0.00 10.0
("l Rare Earths 96. 5.3.1. Cation exchange of zeolites
As discussed in the preceding chapter, cation exchange is commonly used to introduce catalytically active metal species into a zeolite matrix, via exchange with inactive matrix ions such as sodium.
To prepare the catalysts referenced above, cation exchange of various commercially available zeolites and silica-alumina was carried out with an aqueous solution containing some or all of the following:- mixed rare earth chloride solution; cerous nitrate; nickel chloride hexahydrate; ammonium chloride, all dissolved in approximately 2 litres of distilled water. As the mixed rare earth chloride solution used in this study was low in ceria when compared with that used by Murtha et al.(118), in most instances cerous nitrate was added such that a solution of approximately 23 wt% lanthanum and 43 wt% cerium was used, the remainder being made up of other rare earth chlorides. A LHSV of 0.25 was used in all preparations and the temperature used was 100°C.
Generally, a fixed bed of 100 grams of zeolite or support was placed in a
Vycor tube (approximately 45mm O.D, length=l.2 m) with teflon plugs at either end and suitable Swagelok fittings attached to allow passage of the exchange and waste solutions. This configuration was placed in the constant temperature zone of an electrically heated furnace with a type K thermocouple placed in the centre of the zeolite bed to monitor temperature. Temperature was controlled using a dual set- 97. point Ero Electronic furnace controller capable of temperature programming: a digital read-out of the catalyst bed temperature was available. The volumetric flow of cation exchange solution used was in the range 0.1-0.5 mVmin. As no pump was available which could supply the required volume at these low flow-rates consistently, a pressurised flow apparatus was constructed. This consisted of a stainless steel bulb (volume approximately 1 L) into which the exchange solution could be loaded. A pressurised air line and a Fairchild back pressure regulator was attached to the inlet such that a constant pressure was maintained above the solution.
A needle valve was used to regulate the flow of solution from the vessel outlet, and a Dobbie pressure gauge (0 - 250 kPa) mounted to monitor this flow. Using this equipment and the conditions noted above, the cation exchange process was generally continued for approximately 48 hours.
In the case of zeolite Y (particle size = 40 µm), cation exchange was carried out batch-wise using a slurry configuration, due to a large pressure drop across the zeolite bed when the flow technique was attempted.
The cation exchanged zeolite was then washed repeatedly in distilled water (6
X 175 ml) to remove any excess ions, and dried under vacuum overnight. 98.
5.3.2. Impregnation of catalysts
All catalysts containing platinum were prepared using impregnation, with hexachloroplatinic acid hexahydrate as the platinum precursor. Where noted, addition of nickel was also performed using this technique, nickel chloride being used as a precursor.
The amount of (H2PtC16).6H20 to yield the desired weight percent loading of platinum was calculated and dissolved in 25 ml of distilled water. The solution was added to the cation exchanged zeolite (generally about 30 grams) to form a slurry, which was gently stirred for approximately 15 minutes to ensure complete penetration of the solution into the pores. The slurry was then heated to approximately 363K (whilst still stirring) to slowly and evenly remove the water by evaporation. This treatment was followed by drying under vacuum overnight. 99.
5.3.3. Calcination
Calcination of all catalysts tested was carried out using the cation exchange equipment described previously except that, in this case, the vessel was left empty and the back-pressure regulator and needle valve were used to regulate gas flow.
After loading into the Vycor tube the catalysts were heated slowly to 200°C over a period of 4 hours. After this the temperature was increased at 1°C/min to 520°C and held at this value for a total calcination time of 8 hours. The catalysts were then allowed to cool in air to ambient temperature. Flow ,mes u~ v..e.re apf)rox,rntdeJ'J tco"' 1/Mi" •
5.3.4. Catalyst Characterisation
All of the catalysts used in this study were tested for bulk density and total surface area, and their composition was checked by atomic absorption spectroscopy.
Some catalysts were also characterised using temperature programmed reduction
(TPR), temperature programmed oxidation (TPO), differential scanning calorimetry
(DSC), scanning electron microscopy (SEM), transmission electron microscopy
(TEM), x-ray fluorescence (XRF), and x-ray diffraction (XRD). 100. 5.3.4.1. Bulle Density
Bulle density measurements enable the determination of residence times for reactions over solid catalysts. A sample of catalyst (1-2cm3) was dried in a pre weighed glass tube, and cooled and weighed. The volume of catalyst was then determined by displacement of water using a 5 cm3 graduated measuring cylinder.
5.3.4.2. Total Surface Area
The total surface area of all catalysts were measured using a single-point
BET flow technique. The apparatus and procedure used are fully described by
Thomas (126). A calibrated mixture of nitrogen in helium (nominally 30% nitrogen) was used as the carrier gas. The adsorption and desorption of nitrogen from this gas was used to determine the total BET surface area of the sample, as described by
Nelson and Eggertsen ( 127). 101. 5.3.4.3. Metal Surface Area
Measurement of metallic surface area was carried out for the hydroalkylation catalysts using a carbon monoxide pulse method. A catalyst sample (100mg) was
placed in a Vycor micro-reactor and held in position using plugs of Kaowool. The
sample was then reduced in flowing hydrogen at 500°C for 2 hours. The catalyst was
then cooled to room temperature in oxygen free helium which was subsequently directed via a 6-port sampling valve to a gas chromatograph fitted with a thermal conductivity detector.
To determine the amount of CO that could be chemisorbed on the surface,
pulses of a gas mixture containing 7 .5% CO in high purity helium were sent over
the reduced catalyst at room temperature by changing the sample valve position.
After a few pulses, the integrated peaks of carbon monoxide detected by the gas
chromatograph detector were small, as the majority of CO in the pulse was
chemisorbed. The peak area was then observed to increase and become constant and
the metal was then assumed to be fully saturated by CO. The total amount of CO
chemisorbed could then be calculated by difference, knowing the number of pulses
injected over the catalyst As the stoichiometry of CO adsorption upon different
metals is not known, no estimate of absolute metal surface area was possible. 102.
Unfortunately, due to difficulties with diffusion of the carbon monoxide into the zeolite pores and channels, very broad peaks were obtained for the zeolite supported catalysts. This led to problems in estimating the amount of carbon monoxide chemisorbed as the chromatographic peaks were not integrated correctly.
Subsequently these results could not be reproduced, and this technique was not pursued. 103. 5.3.4.4. Elemental Analysis
Analysis of all catalyst samples for nickel, platinum and sodium content was carried out using atomic absorption spectroscopy (AAS) on a Varian-Techtron AAS
Atomic Absorption Spectrophotometer. Catalyst samples were ground to a particle size of < 100 ,um to ensure homogeneity and to increase surface area so as to assist in acid digestion. A small amount (generally 0.1-0.2 grams) was then weighed into a
Teflon beaker and wet with 2 ml distilled water. The sample was then digested in 10
ml HN03 and 10 ml HF and evaporated to dryness. The residue was dissolved in 10 ml HN03 and 10 ml H20 with boric acid solution added to complex any remaining fluoride ions. If on addition of HN03 the residue was not found to dissolve, the digestion with HF was repeated. The resulting solutions were then diluted to give a concentration within the recommended range for the spectrophotometer.
A series of standard solutions were prepared from BDH Atomic Absorption
Standard solutions (1000 ppm) in the required range of concentrations for each species being investigated. A blank solution was also prepared with equivalent amounts of acid added in order to avoid background and matrix effects.
Analysis of a small number of catalyst samples was also carried out using a
Labtam V .25 Ion Coupled Arc Plasma Spectrometer. Sample preparation was as described for AAS. Standard solutions were prepared containing the same volume of acids as the test solutions to minimise matrix effects, as this more sensitive 104. technique cannot tolerate these interferences. Analysis for platinum, nickel, sodium, cerium and lanthanum were made, and a good correlation was found with results from atomic absorption spectroscopy.
The Si0/Al20 3 ratio for catalysts on differing supports was obtained using
X-ray fluorescence. Analysis for other oxides such as Fei03 was also performed using this technique. The catalyst sample (0.5 grams) was diluted to 10 grams with high purity quartz, then ground and pressed into a 40 mm boric acid-backed disc.
The disc was analysed using a Siemens SRS-300 Sequential Spectrometer with a rhodium end-window tube.
X-ray diffraction measurements were carried out on the zeolite support 13X and catalyst A (the fresh catalyst, the used catalyst and the regenerated catalyst) to determine the effect of these procedures on the degree of crystallinity and on the major oxides present. The samples were ground to a fine homogeneous powder
(particle size Rigaku Geiger Flex x-ray diffractometer, with 25 kV voltage, 30 mA current and 0.02 degree step width. 105. 5.3.4.5. Temperature Programmed Reduction and Oxidation Thermal analysis of some catalyst samples was carried out using a Dupont 951 Thermo-Gravimetric Analyser (fGA). Weight changes in the catalyst samples were monitored under various conditions of drying, oxidation and reduction. In some cases, sequences of the above treatments were carried out in an attempt to parallel changes to the catalysts during use. The gas flow to the sample was regulated using a rotameter and a three way valve, such that switching between gases was instantaneous. Between 10 and 50 mg of catalyst was placed in a platinum pan and loaded into the TGA furnace. The sample was usually dried in nitrogen (20 mVmin) by heating to 300°C at 10°C/min, followed by cooling to room temperature. Temperature programmed reduction (TPR) was then carried out using a gas stream of 5% H2 in N2 (15 mVmin) and heating to 800°C at 1D°C/min. The sample was then cooled to room temperature in nitrogen. Temperature Programmed Oxidation (TPO) was carried out in essentially the same manner, with air used as the oxidising gas at a flow rate of 20 mVmin at temperatures of up to 500°C attained by heating at 10°C/min. TPO was performed in sequence with TPR to ensure complete oxidation of samples for subsequent analysis. 106. 5.3.4.6. Acidity Measurement As noted previously, the acidity of zeolites is a very important and interesting feature catalytically. It is known to be influenced by the silica-alumina ratio, the degree of cation exchange, the cation species used in the cation exchange, the acidity of the exchange solution and the calcination temperature used. There are a number of methods employed to measure and characterise acidity: these include:- amine titration using indicators (128,129); gaseous base adsorption followed by temperature programmed desorption (128,130,131,132); differential scanning calorimetry (133,134,135,136); infra-red transmittance spectroscopy of surface hydroxyl groups (137,138,139); and potentiometric acid-base titration (140). Methods employing base adsorption have gained the most recognition, with temperature programmed desorption of gaseous bases such as ammonia and pyridine being widely reported (128,130,131,132). As decomposition of ammonia can occur under the conditions used, these methods necessitate the use of a mass spectrometer to measure the composition of the effluent stream dS less rigorous detection methods can lead to erroneous results~ Differential Scanning Calorimetry (DSC) has been successfully applied to acidity measurements of zeolites, with studies including those by Aboul-Gheit et al. 107. (133,134,135) and Auroux et al. (136). With this technique, the size of the exotherm associated with desorption of a base from the catalyst is measured and is assumed to be directly related to the acid strength. Although this technique possesses the general characteristics of the TPD method whereby the desorption of a base from stronger sites takes place at higher temperatures than from weaker sites, it has the advantage of excluding effects produced by the solid adsorbent, differentiating between physical and chemical adsorption of the base and providing quantitative acid strength distributions. In this study, the procedure developed by Aboul-Gheit et al.(133) using adsorption of triethylamine (IBA) was adopted. All samples were prepared by grinding and sieving to a particle size of < 1oqi.m, and were heated to 350 °C for 2 h before being cooled to 800C. Half of each sample was then removed and placed in a desiccator: subsequently to be used as reference samples. The remaining sample was thoroughly mixed with excess triethylamine while still hot and left to soak overnight. The sample was then filtered and dried at 60°C for 1 hour prior to DSC measurements. For analysis, the sample cell (an aluminium pan) was loaded with 10.2mg of IBA-presorbed sample, and the reference cell with 10.0mg of IBA-free sample. Aboul-Gheit et al. (133) report that this ratio of sample to reference weights gives the optimum results: differences in weight up to 0.3mg were however considered tolerable. 108. All DSC measurements were carried out at atmospheric pressure with a 15 cm3/min flow of nitrogen used as a purge gas for all analyses. A Dupont Thermal Analyser with Differential Scanning Calorimeter was used under the following conditions: initial temperature, 5Q°C; rate 2C>°Cmin·1; final temperature, 600°C; mass, 10.2 mg. It was found that not placing a lid on the aluminium pan made matching of sample weights easier and gave superior results. 5.3.4.7. Microscopic Analysis Some catalysts were examined using a Cambridge Instruments Stereoscan 360 Scanning Electron Microscope. Samples were mounted on to stainless steel discs and a conductive coating of Ag/Pd applied prior to analysis. The images obtained were magnifications X 3,500, with particles of approximately 3,cro visible on the catalyst surface with good resolution. Attempts were made to identify the elemental composition of specific particles using energy dispersive x-ray microanalysis, KEVEX/EDAX. Even though signals corresponding to platinum, nickel, cerium and lanthanum were obtained, there were often interferences with large silicon and aluminium peaks and the resolution was not sufficient to enable assignment of chemical composition to individual particles as desired. 109. Attempts were then made to use Transmission Electron Microscopy (IBM) to examine thin slices of the catalyst surface under higher magnifications. Sample preparation using whole catalyst particles was first carried out by soaking the particles in acetone and then in a 1:1 mixture of acetone:Spurr's resin, followed by embedding the samples in 100% Spurr's resin. The samples were then sectioned to a width of approximately 100nm using a Reichert Jung Ultramicrotome Ultra Cut E with glass knives. Unfortunately as the resin had not penetrated into the specimen sufficiently, the samples were found to break. Preparation of fragmented samples were also unsuccessful, and TEM could not be pursued. 110. 5.4. HYDROALKYLATION EXPERIMENTS 5.4.1. Apparatus Testing of catalysts and reaction studies were carried out using a reactor system designed to accommodate a wide range of reaction conditions: LHSV = 5 - 25 hr"1 Temperature = 140 - 200°c Pressure = 1.4 -6.9 MPa H2 flow rate = 0.2 - 1.0 mole/mole feed The apparatus used for all experiments is shown in Figure 5.1. Gas flows, both reactant and diluent, were supplied using Brooks 5850E thermal mass flow controllers. Hydrogen flow could be controlled between 0-400 cm3.min·1, while the diluent or nitrogen line operated in the range 0-20cm3.min·1, Swagelok one-way valves were placed in all gas lines to protect the mass flow controllers from back flushing with any liquid reactants. Liquid reactant at accurately controlled flow rates (0.0 to 9.9 cm3.min"1) was supplied using a Waters HPLC dual piston pump model 501 with a pressure limit of 6000 psi. 111. The reactor (3/8" O.D stainless steel of length 35 cm) was heated by a small electrically heated furnace. This was mounted next to a heated bypass box and connected via heated lines to the subsequent pressure let down and sampling points. Stainless steel ¾" tubing was used for all system lines, and only stainless steel needle valves and other fittings were included in the system. Some problems were experienced in the initial stages with product condensation in sample lines, and extensive modifications were needed. The modifications included: the use of a nitrogen stream (up to 4000 cm3/min) preheated to 2500C which served as a diluent for the product stream; heating of all transfer lines using mineral insulated heating wires controlled by a Variac voltage controller; installation of a heated reactor bypass system to enable direct G.C analysis of the feed to the reactor (benzene, H2, NJ (this also provided a mixing point for benzene and hydrogen prior to the reactor inlet); flushing lines prior to each run to remove any high molecular weight material which may have accumulated in the reactor system; and using higher gas chromatograph oven temperatures to prevent condensation of sample in the sample loop. These adjustments were made in order to improve the accuracy and reproducability of catalyst activity measurements. ~ O'Q C., !'I) !.II - ELECTRICAL FURNACE HYDROGEN MFC >I5D?)DDD'O REACTOR Cl) (") ::r 0 3 ~ ::t. (") NITROGEN MFG 0-. PREHEATER II I"".. ,.. It 2 2 2 2 2'.,I I I :c '< HEATER BOX GAS CHROMATOGRAPH a0. FD ~ ;i,;- '< ....~ ------~ o· :::s :,;:, LIQUID FEED 0 ~ CHILLED ,,----- (") .....ISOLATION VALVE .... =.J KOV (GMl 0.., - METERING VALVE /BAct. 'P Catalysts prepared from zeolite 13X were used as supplied (particle size 1.70-2.12mm), as were zeolite 5A catalysts (1.70-2.12mm), zeolon catalyst (300- 500pn), and zeolite Y catalyst <«,m). The Si01 -Al20 3 catalysts were ground in a mortar and pestle, and screened to produce the desired size fraction (300-50C)'m). All experiments used 2g of catalyst loaded into the reactor and held in position with glass wool plugs. Glass beads placed above the catalyst served as a preheater. The reactor was then placed in the constant temperature zone of the electrical furnace and the reactor temperature monitored by a 1/16" type K thermocouple placed in the catalyst bed through a Swagelok "T" union. The catalyst was generally reduced at 170°C and 3.5 MPa in flowing hydrogen at 200 cm3min· 1 for 15 min prior to contacting with the benzene or other liquid feed. Where other reduction conditions have been used, these are stated specifically. Following pre-treatment, the reactor was cooled to approximately 20"C below the desired reaction temperature so as to minimise the increase in catalyst temperature due to the exothermic nature of the reaction. In all cases the liquid feed and hydrogen were introduced into the reactor in a down-flow manner. Reactor pressure was maintained such that all reactants except hydrogen were present in the liquid phase and thus the reactor was operated as a trickle bed. Pressure was maintained using a Whitey needle valve heated to approximately 114. 200°C, and monitored by a Dobbie Instruments pressure gauge (0-10 MPa) on the inlet hydrogen line. The pressure drop across the catalyst bed was always negligible and was checked prior to each run with the outlet needle valve fully open. Bypass runs were carried out after any change of conditions to monitor the concentration of feed being delivered. Caution was necessary when switching to bypass and returning again to the reactor as the seals in three Whitey on-off valves used in this procedure were found to leak after pericxls of use due to the temperature and environment to which they were subjected, i.e., exposed at approximately 215°C to benzene and other solvents. Checks were therefore carried out for each run to ensure no internal system leaks were occurring between the bypass and reactor lines. This was done by passing hydrogen to the desired valve and checking for leakage with a Quantum Instruments Gas Leak Probe hydrogen detector. Checks were also made for external leaks after pressurising the system before each experimental run. Gas samples were taken at approximately 25 min intervals using a 6-port Valeo sampling valve with 0.1ml volume sample loop. Liquid samples were also taken periodically from an ice water cooled knock-out pot. Total gas flows were monitored during all experimental runs using a gas meter at the system outlet. A carbon trap was placed before the gas meter to ensure no organic reactants or products were vented to atmosphere. 115. 5.5. CHROMATOGRAPHIC ANALYSIS The reaction products for all experiments were analysed by on-line gas chromatography using a 6-port Valeo sampling valve with a 0.1 cm3 sample loop. The products were also periodically collected from an ice water cooled knock-out pot for later off-line analysis. The chromatographic analysis was performed using a Varian 3300 temperature-programmable gas chromatograph equipped with a flame ionisation detector. On occasion a Gow Mac gas chromatograph equipped with Thermal Conductivity detector was also used. Peak areas for both gas chromatographs were determined by a Hewlett Packard 3390A integrator/recorder. High purity nitrogen was used as the carrier gas for most studies with an Oxy-trap installed in the line to protect the column from any oxygen present. High purity hydrogen was used as the carrier gas for analysis with the TCD gas chromatograph: again ·an Oxy-trap was used to ensure gas purity. The retention times and relative molar response factors for the various components were determined by injecting gravimetrically prepared liquid standards. Confirmation was also made by on-line analysis of a prepared mixture fed to the gas chromatography via the bypass configuration. Where identification of product species was not possible by comparison of retention time with a known standard, 116. estimation of boiling · point and comparison with tabulated boiling point data was carried out, and gas chromatography - mass spectroscopy was used. An AEl MS 12 Mass Spectrometer with 70 ev, 8000 V accelerating voltage and ion source at 2200C was used, coupled with a Shimadzu GC6 AMP gas chromatograph with an all glass ~IVAl'f (!10~) straight split and OV-101 column. Spectra acquired were processed by a VG 2000 " data system. Gas Chromatography-Mass Spectrometry was also used to assist in analysis of more complex product and feed streams. It was anticipated that chromatographic analysis for this study would be difficult. Considerable time and effort was spent determining the optimum column type and chromatographic conditions for the analysis of the product species formed. A survey of the literature revealed that analysis of coal derived materials was almost exclusively undertaken using high resolution capillary gas chromatography (141,142,143,144,145). A review by Charlesworth (141) of the analysis of coal derived materials by gas chromatography identified fused silica capillary columns with stationary phases OV-101 and SE-54, employing temperature programming and detection by FID, as possible separation columns. The Varian 3300 was not equipped for use with capillary columns and the desired separations were initially investigated with a Altech 6' x Ya" stainless steel column packed with OV-101 (mesh range = 80-100, 3 wt% on W-HP support). Using temperature programming, separation of bicyclohexyl and cyclohexylbenzene 117. was achieved. Benzene and cyclohexane were found to elute together, irrespective of column temperature or carrier flow rate. The need to employ capillary chromatography was evident, and the Varian 3300 was adapted for this purpose. As on-line gas sampling ·was used with this chromatograph, conventional adaption kits were not suitable for this modification. After the experimental procedures used for flow switching in high resolution gas chromatography were examined (146,147,148,149,150), the following changes were carried out:- (1) a make-up gas line (1/16" stainless steel) was installed off a "T" union on the main carrier gas line. A needle valve was installed to regulate the gas flow; (2) the detector was fitted with a length of glass-lined stainless steel tubing, through which the capillary column could pass, to reduce dead volume. The make-up gas line was joined to this tubing to further reduce dead-volume effects at the detector and to increase the gas flow over the detector; (3) a "T" piece was installed from the outlet of the sampling valve to accommodate the capillary column and the line to the splitter valve and vent. The spitter valve consisted of a needle valve and coil of tubing to compensate for any pressure drop due to splitter usage; 118. (4) as all lines and most fittings added were 1/16", graphite ferrules were used to seal both the detector and injector ends of the capillary column. A 15 m x 0.54 mm Alltech Econopac SE-54 (film thickness = 1.2,m) megabore capillary column was installed. Although all separations were achieved using this column, temperature programming from ambient temperature was necessary to separate benzene and cyclohexane. This presented a problem, as condensation in the sample loop was anticipated under such conditions. It was decided to install a second gas chromatograph in series with the Varian to analyse for benzene/cyclohexane. A Gow-Mac gas chromatograph with thermal conductivity detector was used with a Graphpac GB 60/80 packed column installed. This column was found to separate benzene and cyclohexane at 1500C. Another V alco gas sampling valve was installed to enable sample injection and a knockout pot placed before the valve to trap heavier compounds. Unfortunately during experimental runs the packed column was found to become overloaded with heavy compounds not trapped in the knockout pot. Although attempts were made to alleviate this by installation of a second liquid N2 cooled knockout pot, this could not be avoided and resulted in crystallisation of some products and complete blocking of lines. Thus, analysis of benzene/cyclohexane was still not possible on-line. 119. Similarly, the capillary column was found to be overloading during initial experimental runs. This was considered unusual given the large volume of nitrogen acting as diluent to the product stream. Investigation revealed that the splitter vent line, installed initially to aid the separation, was leaking condensed product into the column. This line was capped and no further overloading was observed. Analysis for cyclohexane/benz.ene was performed off-line on another gas chromatograph using a BP 20 megabore capillary column programmed from 60 - 190 °C at 5°Cmin·1, ·while the SE-54 capillary column was used to analyse for cyclohexylbenz.ene and heavier products on-line using a temperature program 130 - 190°C at 8°Cmin·1• This procedure proved unworkable due to the unavailability of the second gas chromatograph and due to variances in the overall mass balances thought to be due to condensation of high boiling materials in the on-line sample loop. A column was therefore sought which could deliver the required separations at an even higher temperature. A Carbowax megabore capillary column (15m, 0.54mm I.D., I.~ film thickness) was installed into the Varian 3300 for analysis of on-line gas samples. Separation of benz.ene/cyclohexane was achieved using this column at 190°C with low carrier gas velocity. Other desired separations were also delivered within a reasonable time by temperature programming. 120. Conditions used were: column temperature 190°C (3 min) -> 200°C at 5°Cmin·1 carrier gas flow 3 ml.min·1 detector temperature 2500C injector temperature 25D°C The retention times and relative area response factors are summarised in Table 5.4. 121. Table S.4 Gas Chromatography retention times and relative area response factors Compound Retention Time Relative Area (mins.) Response Factors cyclohexane 0.96 1.01 benzene 1.01 1.00 methylcyclopentane 1.33 0.71 isomers 1.48 - of 1.75 - methylcyclopentylbenzene 1.91 - bicyclohexyl 2.29 0.59 cyclohexylbenzene 2.47 0.49 biphenyl 2.73 0.43 p-dicyclohexylbenzene 14.68 0.04 m-dicyclohexylbenzene 16.40 0.04 122. CHAPTER 6. RESULTS AND DISCUSSION HYDROALKYLATION OF MODEL COMPOUNDS The first and major stages of the present study have been focused on the hydroalkylation of model compounds such as benzene and toluene and on mixtures of compounds present in coal derived liquids. An investigation of the hydroalkylation catalyst composition has been carried out in reactor trials using benzene, as benzene is considered to be the component of major interest in this study and to be representative of the aromatic nature of coal derived liquids and other petroleum based fuels. Given the complexity of the hydroalkylation reaction, coupled with interferences from side reactions such as hydrogenation, isomerisation and alkylation, benzene is also a sensible choice with which to monitor changes to the catalyst system and reaction pathway. A synthetic mixture of major compounds present in coal derived liquids has been prepared from model compounds and tested for the hydroalkylation reaction. 123. 6.1. HYDROALKYLA TION OF BENZENE - RESULTS 6.1.1. Introduction Murtha et al. (117-125) reported that a catalyst comprising of a nickel and rare-earth treated zeolite support impregnated with platinum was capable of hydroalkylating benzene to cyclohexylbenzene with a conversion of benzene of up to 10 wt% and selectivity to cyclohexylbenzene of 74 wt%. In the present study the veracity of this claim was first investigated, using the catalyst proposed by Murtha et al. (118). The effects of process variables such as temperature, pressure, LHSV, and hydrogen and benzene partial pressures were examined, in order to optimise reaction conditions for maximum yield of cyclohexylbenzene. Due to the complex composition of the hydroalkylation catalyst, an extensive study of the role of the various metals and their loadings was undertaken to determine the active metal sites on the catalyst and the influence of variables such as reduction temperature on catalyst behaviour, as was an investigation of the effect of varying the catalyst support and acidity. A number of catalyst treatments, including regeneration after cycles of use, were also carried out, and these are discussed in this chapter. 124. 6.1.2. Initial testing Experiments using an empty reactor confirmed that the reactor system, including preheater and reactor, exhibited no reactivity toward the benzene feed under typical conditions of use. The initial catalyst chosen for subsequent study (catalyst A), was a nickel (5.0 wt%) and rare-earth (10.0 wt%) cation exchanged zeolite (13X), impregnated with platinum (0.1 wt%). Catalyst activity and catalyst deactivation as a function of time on line was examined (see Figure 6.1). On bringing the fresh catalyst on line, a period of instability (ea 120 min.) was observed, after which the conversion settled down to about 17 wt% after 8 hours under the stated conditions. A trend of deactivation is seen to be present, but deactivation was not marked, with the catalyst activity falling by less than 10% over an 8-hour period. As a result, the catalyst was used for longer times on line. The scatter of data points in this figure is caused by a degree of pulsing of products during sampling over short intervals. 125. Conversion of Benzene (wt%) 35 .------~ 30 ... 25 - . • 20 - • . . . . • • • 15 ... . 10 ... 5 >- 0 I ' I ' 0 100 200 300 400 500 Time on line (min) Figure 6.1 Deactivation testing of catalyst A (Pt/Ni/RE-13X). T=1700C, 1 P=3.5 MPa, LHSV=l6.56 hr , H2:HC=l:l. The product distribution of major components as a function of time on line for this catalyst (A) is shown in Figure 6.2. Cyclohexane was observed to be the major product initially, although production of the desired product, cyclohexylbenzene, increased to approximately 75 wt% as time on line increased. Production of dicyclohexylbenzene isomers, via further alkylation of cyclohexylbenzene, was also evident with the p-isomer predominating over the m isomer. Only traces of the o-isomer were produced due to steric effects. Higher 126. boiling tricyclohexylbenzene isomers were also produced in small quantities, these being only detected in off-line analyses of liquid reaction products. Similarly, traces of the isomerisation product, methylcyclopentane, were also detected in off-line analysis, as were three products of unknown composition with boiling points intermediate to benzene and bicyclohexyl. The fully hydrogenated product, bicyclohexyl, was also observed during the initial stages of the experiment in trace quantities. Selectivity (wt%) 100 ,------~ BO 60 40 20 100 200 300 400 500 Time on line (min) -- Cyclohexane -t- CHB --- m-HMW ~ p-HMW Figure 6.2 Product composition for benzene hydroalkylation over catalyst A (Pt/Ni/RE-13X). T=170°C, P=3.5 MPa, LHSV=16.56 hr1, H2:HC=l:l. 127. Catalyst A was also used to examine the effects of process variables on the hydroalkylation of benzene. The results are presented in Table 6.1 in tenns of the weight per cent conversion of benzene, and product selectivity for cyclohexylbenzene (CHB), cyclohexane (CH), and high molecular weight products (HMW) (predominantly dicyclohexylbenzenes), again as weight percentages of total products. Table 6.1 Hydroalkylation trials using Catalyst A (Pt/Ni/RFJzeolite-13X) Temp. Press. LHSV ffi:benz Benz.ene CHB CH HMW ("C) (MPa) (hr"I) Molar Conversion Selectivity Selectivity Selectivity Ratio (wt%) (wt%) (wt%) (wt%) 146 3.S 16.6 1:1 26.8 76.7 19.4 3.8 170 3.S 16.6 1:1 25.9 77.4 1S.5 6.1 19S 3.S 16.6 1:1 25.9 77.2 1S.l 6.3 170 2.8 16.6 1:1 17.3 76.3 18.3 4.3 170 4.2 16.6 1:1 3S.O 60.S 16.1 18.7 170 3.S 8.3 1:1 S0.7 78.2 12.0 7.1 170 3.S 33.2 1:1 11.2 84.4 10.1 2.7 170 3.S 16.6 O.S:l 40.8 88.4 9.7 1.8 170 3.S 16.6 2:1 12.8 79.2 16.2 4.6 170 3.S 16.6° 1:0.S 19.2 76.9 12.2 10.6 170 3.S 16.6° l:0.7S 9.2 79.9 13.S 6.6 * A decane diluent was employed to keep the total LHSV constant. 128. The effect of increasing reaction temperature from 146 to 195°C was not marked (Table 6.1), with very little change in benzene conversion (26.8-25.9 wt%) or product selectivities (OIB 76.7-77.2 wt%) being observed. The dependence of reaction rate on reaction temperature was examined in terms of the Arrhenius equation In k = In A - E/RT (6.1) where k is the rate constant, A is a constant, E the activation energy, R the universal gas constant and T the absolute temperature (K). Due to the number and complexity of the reactions taking place, no correlation was possible, and no estimation of the activation energy could be made. Indeed, were a value of activation energy to be determined from this study, no literature values are available for comparison. Comparison with thermodynamic estimates would be questionable due to the extensive numerical estimates needed to generate fundamental thermodynamic constants for the compounds produced by the reaction. No Gibbs free energy data or heat of formation data are tabulated in the literature for these compounds. Heats of vaporisation and heat capacities for cyclohexylbenzene (and for the higher molecular weight compounds) have also not been reported. 129. The effect of increasing reaction pressure from 2.8 to 4.2 MPa (at 17fJ'C and 16.56 ht1) is shown in Figure 6.3, with product selectivities shown as weight percent yields for clarity. The benzene conversions given represent their summation. The observed increase in benzene conversion is expected as an increase in total reactor pressure corresponds to an increase in hydrogen partial pressure, in that the composition and flow rate of the feedstock were not changed on increasing pressure. Liquid hourly space velocity was found to have a major effect on product yields (see Figure 6.4), with benzene conversion increasing from 11.2 wt% at 33.2 ht1 to 50.7 wt% at 8.3 hr·1• Yields of cyclohexane and higher molecular weight compounds steadily increased with decreasing space velocity. These effects may indicate a change in the flow regime of the reactor, with an accompanying change in the interaction between gas and liquid in the system, and particularly in the extent of liquid hold-up. 130. 40------Product Yield (wt%) 35 30 25 20 15 10 0-,:______.....______, 2.8 3.5 4.2 Reaction Pressure (MPa) CHB ~ - cyclohenne · 6.· hich IA products ~ Benzene Convenion Figure 6.3 Influence of reaction pressure on product yields and benzene conversion for catalyst A (Pt/Ni/RE-13X). T=17D°C, LHSV=l6.56 1 hr , H2:HC=l:l. 131. Yield (wt%) 60.------,Product 50 ...... · ... · ...... · · ... · · · · ... · · ... · · ... · · · ····· ...... 30 ...... 20 ...... 10 ------~--- ···· ···················· ·······-·······-~--~-~- ~-~.-:-::.:::--..~.:-:-:- .. ~ 0 8.3 16.6 24.9 33.2 LHSV (hr-1) CHB ~ - cyclobexane · ·l!.· · hi1h MW products -¼- Benzene Conversion Figure 6.4 Influence of LHSV on product yields and benzene conversion for catalyst A (Pt/Ni/RE-13X). T=1700C, P=3.5 MPa, H2:HC=l:1. 132. The final experimental parameter to be considered was the hydrogen to benzene ratio. Firstly, the hydrogen flow rate was varied while maintaining a constant benzene liquid hourly space velocity and an overall pressure of 3.5 MPa. Higher benzene conversions were observed at the lower hydrogen:benzene ratio and this was accompanied by a slight increase in CHB selectivity. Secondly, variation of the benzene partial pressure at constant hydrogen flow and constant liquid hourly space velocity was examined. These experiments were carried out by passing benzene over the catalyst with an inert liquid diluent (decane) in the required proportions, thus keeping the overall LHSV constant Higher hydrogen to benzene ratios were observed to result in a decrease in benzene conversion, with selectivities to cyclohexylbenzene remaining essentially unaltered. Due to relatively high conversions reported in the hydrogen and benzene partial pressure experiments discussed above, this data was not used for kinetic modelling of the hydroalkylation reaction (55). As it was not known if the moderate catalyst deactivation observed during the hydroalkylation of benzene would be intensified with the use of coal derived and petroleum derived liquids, no experimental design was carried out to collect data relevant for kinetic modelling of the hydroalkylation reaction at this stage. This phase of the program concentrated on optimising the reaction conditions. 133. From this initial investigation, conditions for subsequent reactor studies were chosen. The standard conditions chosen were :- reaction temperature 170°C; reaction 1 pressure 3.5 MPa; liquid hourly space velocity 16.56 hr ; and liquid reactant : H2 feed rates of 1:1 on a molar basis. Where necessary, these conditions were adjusted to investigate particular parameters (see below). 134. 6.1.3. The Role of Metals The hydroalkylation catalyst tested thus far (catalyst A) consisted of a nickel and rare earth chloride exchanged zeolite 13X support impregnated with platinum. This catalyst represents quite a complex system, with the action of the different metal species on the hydroalkylation reaction not known. The influences of the metals on each other, the catalyst support, the catalyst acidity and the role of pretreatment conditions on performance was also unknown. The first point to be studied was the role of the various metals in the hydroalkylation catalyst A series of catalysts of varying metal content were prepared using the techniques described previously. Their compositions, as determined by Atomic Absorption Spectroscopy and Ion Coupled Arc Plasma Spectroscopy are given in Table 6.2. Also shown are their total BET surface area as found by single point measurement and the Si02:Al20 3 molar ratio for the various catalyst supports, determined using X-ray Auorescence Spectroscopy. 135. Table 6.2 Composition of Hydroalkylation Catalysts. Metal Content (wt%) Total BET Si~AJiO, Catalyst Surface Area mol• Designation Nickel Platinum Rare Sodium (m2g·•) ratio Earths A 53 0.1 9.1 0.6 389.S 2.86 B 4.9 0.0 8.5 0.6 362.5 not detemmed C 4.8 0.14 9.1 0.3 274.9 " D 5.2 0.18 9.0 1.0 349.4 " E 0.0 0.10 16.4 0.6 391.5 " F 1.7 0.09 9.5 4.9 303.5 " G 2.9 0.09 9.4 1.6 348.3 " H 6.9 0.13 9.2 0.9 214.9 " I 5.9 0.10 0.0 1.5 294.1 " J 5.7 0.10 2.4 1.1 220.7 " K 5.6 0.16 4.7 0.7 301.4 " L 6.1 0.14 93 0.6 292.8 " M 7.7 0.02 0.1 1.5 267.3 " When a comparison is made between the desired and resultant catalyst compositions (see Tables 5.3 and 6.2), discrepancies are evident Most notable of these are for high nickel loadings and high platinum loadings, with the desired concentration being reached in neither case. The nickel concentration achieved was 6.9 wt%, and the platinum concentration was 0.18 wt%, compared to the desired values of 10 wt% and 0.3 wt% respectively. The sodium levels of the catalyst appear to be influenced by the rare-earth and nickel concentrations in the ion exchange solution, and were found to vary between 0.1 and 4.9 wt%. 136. All catalysts prepared were tested for their hydroalkylation ability in reactor studies using benzene. In all cases 2 grams of catalyst was loaded into the reactor and reduced at 1700C for 15 mins. under flowing hydrogen at 3.5 MPa. Standard reactor conditions, described previously, were employed to examine the catalyst behaviour. The results obtained are shown in Table 6.3. In all cases a period of catalyst "break in" was observed. The reaction products were monitored for approximately 5 hours to ensure no marked catalyst deactivation was occurring. 137. Table 6.3 The Effect of Catalyst Metal Content on Benzene Hydroalkylation1 Catalyst Benzene CHB CH HMW Designation Conversion Selectivity Selectivity Selectivity (wt%) (wt%) (wt%) (wt%) A 20.1 75.4 19.2 4.6 B 1.2 0.0 100.0 0.0 C 22.5 75.2 13.1 10.9 D 23.7 93.2 1.6 2.0 E 3.8 31.1 65.5 0.0 F 2.3 37.8 62.0 0.0 G 23.8 53.6 41.7 3.8 H 19.8 58.3 24.2 16.7 I 5.5 1.5 94.8 0.0 J 2.5 69.2 21.2 9.5 K 15.4 63.6 32.4 3.0 L 19.5 81.1 12.3 5.07 M 1.3 0.0 100.0 0.0 1. Reaction Conditions: T=170°C, P=3.5 MPa, LHSV=16.56 hr·1, H2 flowrate=210 ml.min·1, Benzene:H2=1:l. The effects of varying the metal loadings can be seen in Figures 6.5 - 6. 7. The top graph of each shows the effect on overall benzene conversion (as weight percent), while the lower figure gives the product selectivities (also as weight percent) for the three major products: cyclohexane (CH); cyclohexylbenzene (CHB); and high molecular weight (HMW) di- and tri-cyclohexylbenzenes. 138. 25....------~Benzene Conversion wl'1: 20 · - - ... · ... · 15 10 5 0 .___...... __.....___ _..__ __..__...... L. _ __,_ _ ___,J 0 2 3 4 5 8 7 Nickel content wt% 80,------Selectivity wt:r. 110 40 20 0 2 3 4 5 8 7 Nickel content wt% -+- CH -¼- CIIB -ir IIICW Figure 6.5 Benzene conversion and Product Yields as a function of Nickel loading. Catalysts E. F, G. A. H. All contain Pt/RE. 17D°C, 3.5 1 MPa. LHSV=l6.56 hr· • benz:H2 1:1. 139. Benzene Conversion •~ 25,------, 20 .. · 15 0 .______,1. ______..__ ___ ....L.. __ _.J 0 0.05 O.l 0.15 0.2 Platinum content wt% Selectivity wt~ 120 ,----..:,._------. 0 0.06 0.1 0.16 0.2 Platinum content wt% -t-- CH -.- CHB -a- l!Wll' Figure 6.6 Benzene Conversion and Product Selectivities as a function of Platinum Content. Catalysts B, A, C, D. All contain Ni/RE. 17CY'f"., 3.5 MPa, LHSV=l6.56 hr"1. benz:H 2 1:1. 140. Benzene Convenion wt7. 25~------~ 20 15 10 5 0 2 4 8 8 10 12 Rare earth content wt% Selectivity wt7. 100 ~------~ 0 2 4 II 8 10 12 Rare earth content wt% --T- CU ~ CUD ""tr" IJl,flr Figure 6.7 Benzene Conversion and Product Selectivities as a function of Rare-Earth Content. Catalysts I, J, K, A. All contain Pt/Ni. 170"C. 3.5 MPa, LHSV=16.56 hr 1, benz:H: 1:1. 141. The catalysts of varying nickel content (Figure 6.5) show a dramatic increase in activity between 1. 7 and 2.9 wt% Ni, accompanied by an increase in cyclohcxylbenzene selectivity. While the benzene conversion is found to approach a level of approximately 20 wt% at higher nickel loadings, the selectivity towards cyclohexylbenzene increases steadily up to a value of 75.4 wt% at 5.3 wt% Ni Production of high molecular weight compounds also increases, with a significant quantity (16.7 wt%) being produced for the highest nickel catalyst. The catalysts containing no nickel and a very low loading of nickel (l.7wt%) were found to exhibit very low activity (3.8 and 2.3 wt% benzene conversion, respectively) and cyclohexane was the major product formed. The data in Figure 6.6 indicates the marked effect of platinum loading on catalyst activity. There is a rapid increase in benzene conversion to a level of 23.7 wt% at 0.18 wt% Pt, with an accompanying increase in cyclohexylbenzene selectivity (93.2 wt%). In the absence of platinum, only cyclohexane is produced, and at very low levels (1.2 wt%). Variation in the concentration of rare-earth chlorides used in the cation exchange of these catalysts resulted in a series of catalysts of significantly different rare earth content. The effect of this compositional difference is shown in Figure 6.7, where a steady increase in catalyst activity is observed with increasing rare earth loading. Cyclohexylbenzene selectivity was found to increase with increasing rare earth content, resulting in significantly higher yields of the desired product. A 142. maxima in higher molecular weight compounds was observed at 2.4 wt% rare carths oxides, the level of these products reducing as the rare earth metal content increased further. In the absence of rare-earths metals, activity is very low (1.S wt%) and cyclohexane is the only product 143. 6.1.4. Catalyst Pretreatment It has been shown that benzene conversions up to SO. 1 wt% and cyclohexylbenzene selectivities of as much as 88.4 wt% can be achieved with catalyst A, i.e. a catalyst comprising 0.1 wt% Pt on a 5 wt% Ni and 10 wt% rare earth cation exchanged 13X zeolite support. This was possible following a rather mild catalyst pretreatment employing conditions of 170"C and 3.5 MPa for 15 minutes under flowing hydrogen. When catalysts deficient in one or more of the metallic components (such as catalyst E (no nickel), catalyst B (no platinum), catalyst I (no rare earths) and catalyst M (no platinum/no rare earths)) were pretreated under the same conditions, they were found to be virtually inactive. To investigate the role of pretreatment conditions on these metal deficient catalysts a number of pretreatment temperatures were used, followed by reaction 1 under the standard conditions of 170"C, 3.5 MPa, 16.56 hr and benzene:H2 = 1:1. The effects of increased reduction temperature can be seen in Figures 6.8 to 6.10 for these catalysts, with results represented as individual product and total yields. In all cases (no nickel (Figure 6.8), no platinum (Figure 6.9), no rare earths (Figure 6.10) and no nickeVno rare earths (Figure 6.11)) increased pretreatment temperature results in a dramatic increase in catalyst activity, with total yields approaching those observed for catalyst A (with its prereduction at l 70°C) or higher when reduction temperatures of 400°C or higher were used. The rate of pretreatment 144. temperature enhancement of catalyst activities is, however, not the same for all the catalysts examined. In the case of the "no rare-earth" catalyst (see Figure 6.10) activity was found to be equivalent to catalyst A after reduction at 300°C (19.9 wt%) and was increased to approximately 45 wt% when a temperature of 400°C was used. With this catalyst cyclohexane was the major product. 145. Product Yield (wt%) 20 .------Pt/RE/zeolite 13X (no Ni) 15 · ·········· .... · ... -- --- · · · 10 ...... · · - 5 0 .___ __...______.__ ___._ __ __._ ____., ____,1 150 200 250 300 350 400 450 Pretreatment Temperature (C) Total Yield -+- Cyclobexane _. Cyclohe:1ylbenzene Figure 6.8 Effect of Pretreatment Temperatme on ·Product Yields, Catalyst E (no Nickel). Pretreatment carried out for 15 min. under flowing hydrogen (210 ml.min. 1) and 3.5 MPa. Standard Catalyst Test Reaction Conditions. 146. Product Yield (wt.%) 20 ------Ni/RE/zeolite 13X (no Pt) 15 ... · · · · ... · · ..... · ... · · ...... · · · · 10 · ... 5 ..... 0 L:Z::==:::J::======.--====="===±--J 150 200 250 300 350 400 450 Pretreatment Temperature (C) Total Yield -+- Cyclohexane -¼- C1clohnylbenzene Figure 6.9 Effect of Pretreatment Temperature on Product Yields, Catalyst B (no Platinum). Pretreatment carried out for 15 min. under flo"'ing hydrogen (210 ml.min" 1) and 3.5 MPa. Standard Catalyst Test Reaction Conditions. 147. 50r-Product_____ Yield ....;..._....;_ (wt%) ______ Pt/Ni/zeolite 13X (no RE) 40 ······································································ 30 20 0 Lti--=:::t::::====::I::===l..-:=:::t:===:::§!____j 150 200 250 300 350 400 450 Pretreatment Temperature (C) Total Yield ~ Cyclohexane ""*- Cyclobnylbenzene ~ HIIY Figure 6.10 Effect of Pretreatment Temperature on Product Yields, Catalyst I (no Rare Earth Metals). Pretreatment carried out for 15 min. under flowing hydrogen (210 ml.min- 1) and 3.5 MPa. Standard Catalyst Test Reaction Conditions. 148. Product Yield (wt%) 20.....------, Ni/zeolite 13X (no Pt/no RE) 15 ···································································· ······················ 0 l-....!a:-...... 1..--e,o::i:::;;.__ .L.- _ __.__ ___._ ___._ _ ___. 150 200 250 300 350 400 450 500 Pretreatment Temperature (C) Total Yield -+- Cyclobe:a:ane --- Cyclohe:iylbenzene ~ HKW Figure 6.11 Effect of Pretreatment Temperature on Product Yields, Catalyst M (no Platinum/no Rare Earths - only Ni). Pretreatment carried out for 15 min. under flowing hydrogen (210 ml.min"1) and 3.5 MPa. Standard Catalyst Test Reaction Conditions. 149. The effect of prereduction temperature on product selectivities for these metal deficient catalysts is not straight forward. For catalyst E (no nickel) the increase in CHB selectivity is from 65.5 wt% to 83.3 wt% over the prereduction temperature range 17<1'C to 400°C, with cyclohexane being the only other product detected. For catalyst B (no platinum) cyclohexylbenzene virtually replaces cyclohexane as the major product under more harsh reduction conditions, with selectivities for cyclohexylbenzene of 96.6 wt% and 93.6 wt% being found. No high molecular weight products were observed with catalyst B, even after high temperature reduction. Some interesting results were found for catalyst I (no rare earths). Cyclohexane is the major product after reduction at l 70°C with a selectivity of 94.8 wt% and only small amounts of cyclohexylbenzene detected. Increasing the reduction temperature to 300°C results in a large increase in catalyst activity (19.9 wt%), although cyclohexane remains the major product (83.0 wt%) compared to cyclohexylbenzene (9.4 wt%). After reduction at 400°C, the benzene conversion more than doubles (44.4 wt%) and, although there is a significant increase in cyclohexylbenzene selectivity (31.4 wt%), cyclohexane still predominates, accompanied by the production of high molecular weight compounds for the first time with this catalyst. 150. For the single metal catalyst (catalyst M, no platinum/no rare earths) cyclohexane is again replaced as the sole product after harsh pretreatment However, significant production of cyclohcxanc (selectivity 35 wt%) is still observed after pretreatment at 475°C, accompanied by heavy compounds (17.2 wt%) and cyclohexylbenzenc (57 .8 wt%). For all cases of catalysts being prereduced at higher temperatures the catalysts exhibited a rapid decline in activity with time on-stream. Qualitatively. this activity decline appears to increase in rate with increasing pretreatment temperature. For example, for catalyst E (no nickel) complete deactivation was observed after 122 min. when reduced at 400°C compared with 198 min for the catalyst reduced at 300°c. 151. 6. 1.5. Temperature Programmed Reduction The observation of the dramatic effect of pretreatment temperature on catalyst activity prompted a further investigation of the properties of these catalysts using temperature progranuned reduction (TPR). Weight losses during reduction of the metal oxides present on the catalyst with hydrogen were measured using thermogravimetric analysis (TGA). A stream of 5% hydrogen in nitrogen (12 - 15 mVmin.) was passed over a 40 - 70 mg sample in the TOA furnace (lC>°C/min. to 800°C), and the per cent weight change recorded as a function of temperature. Depending on the reducibility of the components in the catalyst surface, one or more maxima (in a derivative curve) were obtained from the reduction of sample weight at characteristic temperatures. Five catalysts (A,B,E,I and M) were examined in this fashion, and the thermograms obtained are summarised in Figure 6.12 and 6.13. Catalyst A (Pt/Ni/rare earths) exhibited a maximum at 192°C and a very low broad maxima at 583°C. Catalysts M, I, E and B, on the other hand, exhibited a number of broad maxima - for catalyst M (Ni) a peak at 637°C with a shoulder at 475°C and a plateau from 150-400°C: for catalyst I (Pt/Ni) a series of unresolved peaks from 200- 4000C, with maxima at 223°C, 360°C with a shoulder at 475°C, accompanied by a low area broad peak at 592°C: for catalyst E (Pt/RE) unresolved peaks at l 7C>°C, 210°C and 430°C and resolved peaks at 58C>°C and 760°C: for catalyst B {Ni/rare 152. Arbitrary Units 110..------, .....·'· . 100 . . '" : ,,, ·. .. 90 ··.. ..~~ ...... ··•····...... 80'----.....___ ...... _ ___._ __ ~ ______._ ____,.__ _ __. 0 100 200 300 400 500 600 700 800 Temperature (C) ···•--- Catalyst A - Catalyst I - Catalyst M Figure 6.12 Temperature Programmed Reduction Profiles (Temperature Derivatives) of Catalysts M (Ni-13X), I {Ni/Pt-13X) and A {Pt/Ni/RE-I~. Arbitrary Units 110~------ 100 0 100 200 300 400 500 600 700 800 Temperature (C) - Rare Earth/13X · •·· Catalyst E - Catalyst B Figure 6.13 Temperature Programmed Reduction Profiles (Temperature Derivatives) of Rare-earth-13X, Catalysts E (Pt/RE-13X) and B (Ni/RE-13X). 153. earths) at 194°C and 602°C with a shoulder at 517°C. A rare earth-13X sample was also examined. The the~grarn of this sample (Figure 6.13) resulted in peaks at /I 186°C, 61 OOC and a significant weight loss approaching 800°C. Assignment of major peaks of the TPR derivative curves to various reducible metal cations was earned out where possible with reference to several literature sources and internal comparisons within this study. Hurst et al (183) examined the reduction of Ni2+ cations in NaX zeolites using TPR and identified two peaks, at 527°C and 727°C, attributed to Ni2+-+ Ni0 in supercages and to Ni2+ on sodalite sites. The observance of significant changes in TPR profiles with varying sample weights was made in the Hurst et al. study, with a lower sample weight giving three peaks for nickel reduction, at 277°C, 530°C and 777°C. Hurst et al. also acknowledged the difficulty in comparison of TPR peak positions from different authors, due to their sensitivity to experimental conditions. The reduction of Ni2+ ions in the pores of zeolites was also examined by Mahoney et al.(184) with a single peak identified at 577°C for Ni2+ reduction. Platinum reduction profiles were found to be dependent upon the platinum salts used in catalyst preparation, with H2[PtCIJ having a peak maxima at 127°C (183). Calcination temperature was also found to effect the reduction temperature of pt2• ions with broad peaks at 120°C and 220°C observed for Pt/HNaY catalyst calcined at 360°C and 550°C respectively (165). 154. Harrison et al.(180) examining the effect of cerium on platinum in auto- exhaust catalysts, found that the profile of ceria consisted of a peak at 500°C corresponding to surface capping oxygen, and a peak at 800°C corresponding to bulk oxygen. In combination with platinum, peak maxima at 2300C, 480°C and 890°C were obtained, indicating a reduction in the temperature required for reduction of surface species in the presence of platinum. TPR studies by Yao et al.(182) on unsupported Ce02, alumina-supported Ce02, and noble metal doped ceria, give some insight into the complexity of definitive peak assignment and interpretation for these systems. They found unsupported ceria gave two peaks, one of moderate size at 470- 4900C (attributed to surface capping oxygen anions) and a large peak at 770-790°C considered to be due to bulk oxygen anions, in agreement with Harrsion et al.(180). Addition of 3% Pt resulted in a low temperature peak at l 10°C (attributed to the reduction of platinum oxide), a series of small peaks at 2200C, 320°C and 4500C (for which no definitive assignment was considered possible, although they were related to interaction between Pt and capping oxygen in ceria), and a broad peak with a maxima at 740°C (considered to be the unaffected bulk oxygen atoms). The results for the alumina-supported ceria samples were more complex, with a small peak in the low temperature range (100°C) found to diminish with Ce02 loading (attributed to a small amount of removable oxygen anions on the bare amorphous Al20 3 surface), and a group of unresolved peaks from 300-900°C, with maximas of approximately 53C>°C, 650°C and 850°C (considered to be due to reduction of surface capping oxygens, bulk oxygen anions and shared oxygen anions 155. in the interface between the bulk ceria and alumina, respectively) (182). Upon addition of platinum to alumina-supported ceria a multiplicity of peaks were reported between 100-35C>°C with maxirnas at lOO'C and 170°C. Peaks below 150°C were considered to be due to reduction of Pt oxides, while those between 150°C and 350°C could not be designated unambiguously. The high temperature regions of the thennogram, with peaks at 650°C and 88C>°C were largely unaltered by platinum addition. Temperature Programmed Reduction was also carried out by Chien et al. (174) to examine the effect of cerium addition on NiX zeolites. Without cerium, TPR of a 5 wt% Ni/X sample revealed a large maxima at 37CY'C, with a shoulder from 400-550°C, including a small peak at 410°C. Addition of cerium resulted in the diminution of the high temperature peak shoulder, and was considered to be representative of the improved reducibility of the catalyst No high temperature peaks (cf. 800°C) were reported for cerium in this study. From this data, the peak maxima obtained in TPR profiles of the present study were allocated as follows. Catalyst M {Ni/zeolite) gave a broad plateau from 200-400°C and two peaks at 520°C and 616°C (Figure 6.12). In agreement with Hurst et al.(183), the latter two peaks were assigned to Ni2+ reduction in supercage (Sn) and sodalite (Si) positions since the sample was calcined at 524°C, this was not unexpected. The weight loss recorded at this temperature range (2.8 wt%) correlated to reduction of all nickel present in the sample. 156. Addition of platinum changes the TPR considerably (see TPR for catalyst L Figure 6.12). The low temperature plateau is replaced by a series of unresolved peaks with maximas at 223°C, 377°C and 474°C, with a broad peak at 592°C. Due to the very small loading (0.1 wt%) no peak due to reduction of platinum can be expected. However the effects of addition on reduction of nickel can be seen. Significant lowering of the temperature for nickel reduction in both matrix positions was found upon platinum addition (from 500°C and 637°C to 474°C and S92°C). The total weight loss registered (ea. Swt%) corresponded to complete reduction of nickel over the temperature profile carried out. The unresolved peaks at < 400°C cannot be attributed directly to any specific metal, although they indicate the lowering of the nickel reduction temperature by association with platinum To examine the reduction profile of rare earth cations on a zeolite support, temperature programmed reduction was carried out on a l 3X support cation exchanged with the standard rare earth chloride solution. The TPR profile for this lar~~ sample consisted of a small peak at 186°C, a peak at 610°C and a very,,. weight loss near to 80S°C (Figure 6.13). The weight losses obseived with this sample were found to be slightly higher than expected (8 wt% rather than 4 wt%), although complications arising from calculating the weight losses for such a complex mixture of rare earth oxides may be contributing. The calculations were based on the reduction of Ce02 and ~03 in approximately equal proportions, according to the equations: 157. Ce02 + 2H2 ~ ee2• + 2H20 + Je- No weight changes for higher oxides or other rare earth oxides present were predicted. The high temperature peak (805°C) is in agreement with the profile of cerium reduction discussed by Harrison et al.(180) upon reduction of bulk oxygen. The 610°C peak is intermediate to the 500°C assigned by Harrison et al. for surface capping oxygen, and the 65Q°C peak considered by Yao et al.(182) to be due to bulk oxygen anions for alumina supported ceria samples. The latter assignment would negate the 805°C assignment to bulk oxygen and infer a correlation to shared oxygen anions between the bulk and the support as postulated by Yao et al. The TPR profile of catalyst E (Pt/RE-13X) (Figure 6.13), shows the effect of platinum addition on reduction of rare earth metals. The unresolved peaks at 17Q°C, 210°C and 430°C, are in general agreement to the observations of Yao et al. (182). Although they were not specifically identified, they were considered to be modifications due to the influence of platinum on reduction of surface capping oxygen. The high temperature peaks of bulk cerium reduction were found at slightly lower temperatures in the Pt/RE sample (580°C and 760°C cf. 610°C and 805°C), contrary to the observations of Yao et al.(182). This indicates the addition of platinum results in rare earth oxide reduction at a slightly lower temperature. The weight loss with this sample was again twice that to be expected from reduction of all platinum and rare earth metals present. 158. From the TPR profile of catalyst B {Ni/RE-13X) peaks at l 93°C and 602°C with a shoulder at S l 7°C indicate a shift in nickel reduction to slightly lower temperatures with rare-earth incorporation. The weight loss accompanying the 193°C peak corresponds to reduction of all nickel in the catalyst, while that of the high temperature peaks is slightly over the amount expected for reduction of all rare earth metals. This behaviour indicates the lowering of the nickel reduction temperature upon incorporation of rare earth oxides, and represents a probable hydrogen spillover effect While this trend is in agreement with Chien et al.(174) the peak maxima temperatures of the two studies are very different In the case of catalyst A (Pt/Ni/RE-13X) (Figure 6.12) a maxima at 193°C was observed accompanied by a very broad low area peak at S83°C. These results show the combined benefits on nickel reduction of the presence of platinum and rare earth cations. The peak at 192°C correlates with that observed for catalyst B {Ni/RE- 13X). It is also considered to be related to the 223°C peak observed for catalyst I {Ni/Pt-13X, Figure 6.12), with addition of rare earths resulting in increased metal reduction at this lower temperature. For this sample the weight loss corresponding to the 192°C maxima reveals complete reduction of all nickel and 75% of the rare earth metals at this temperature, the reduction of the remaining 25% rare earth metals taking place at approximately 583°C. 159. 6.1.6. Effect of Catalyst Support The studies discussed thus far have focused on the use of zeolite 13X as the catalyst support. In an investigation of the effects of the nature of the catalyst support on benzene hydroalkylation, Pt/Ni/RE catalysts were prepared on four additional supports, namely Si02-~03 and zeolites SA, Y and zeolon 900-Na. In all cases the loadings of metals were 0.1 wt% platinum, 5.0 wt% nickel and 10 wt% rare-earth oxides, and the standard techniques of cation exchange and impregnation were performed, except in the case of catalyst O where a batch wise cation exchange was carried out due to excessive pressure drop through the fixed bed configuration. The support acidity of these catalysts was examined using acidity measurements designed to give quantitative measurements about the number of acid sites, as well as a qualitative indication of the strengths of those sites (see experimental section for details). The results obtained are shown in Table 6.4. Also included are their BET surface areas for comparison. 160. Table 6.4 Characteristics of hyd.roalkylation catalysts of differing supports Catalyst A Q 0 N p Designation Catalyst Support zeolite 13X Si02-A120 3 zeolite Y zeolite 5A Zeolon ~ Na Peak Range - 250 - 450 300 - 450 200 - 400 250 - 400 - Lewis Acid Sites ("C) Peak Maxima - 338,388 broad 375 320 337 - Lewis Acid Sites ("C) Peak Range - 450 - 550 450 - SOO 400 - SOO 400 - SOO 380 - SOO Bronsted Acid Sites ("C) Peak Maxima - 509 482 450 431 420 Bronsted Acid Sites ("C) Lewis Acid 18562 165 2681 9906 - concentration (arbitrary units) Bronsted Acid 17937 487.5 38350 3150 302.3 concentration (arbitrary units) Total Acid 36499 652.5 41031 13056 302.3 concentration (arbitrary units) Total BET 389.46 70.85 205.46 81.27 286.41 Surface area (m2/g) The heat of desorption of triethylamine (IBA) from the various acid sites of the catalyst is measured as a function of catalyst temperature during acidity determination (133,134,135,136). Thus, the temperature of desorption indicates the strength and type of acid sites, while the amount of heat evolved (due to the quantity of TEA desorbed) can be used to evaluate the relative concentrations of different 161. groups of acid sites. The two values give an idea of the number of acid sites of a particular strength and an overall idea of the acid concentration. The measurement can only be treated as comparative, however, due to the surface heterogeneity of the solid swface. It can be seen from these measurements that the catalysts prepared have very different acid characteristics, with catalyst P (zeolon support) showing very little acidity, catalyst O (zeolite Y support) having very high Bronsted acidity, and catalyst A (zeolite 13X support) exhibiting high Lewis and Bronsted acidity in almost equal proportions. The results from the benzene hydroalkylation experiments with these catalysts, conducted under a standard set of conditions, are presented in Table 6.5. The Si02:Al20 3 molar ratios given in Table 6.5 were determined by X-ray fluorescence and show the wide range of values covered by these catalysts. Table 6.S Benzene Hydroalkylation with catalysts of varying supports Catalyst Support Benzene CH CHB Si0/Al20 3 Designation Conversion Selectivity Selectivity Molar Ratio (wt%) (wt%) (wt%) A zeolite 13X 20.1 19.2 75.3 2.86 N zeolite 5A 26.6 97.5 1.7 2.17 0 zeolite Y deactivated 5.8 90.6 3.51 p Zeolon deactivated 1.3 93.2 11.19 Q Si02-A120 3 inactive - - 4.58 1 Reaction conditions: 170"C, 3.5 MPa. LHSV=l6.56 hr , 1:1 H2:benzene 162. Although the activity of catalyst N, supported on zeolite SA, is comparable to that of catalyst A supported on 13X, the aperture size of the zeolite SA (0.42nm) is considered too small to allow access of the desired intennediates and products. Since benzene has a kinetic diameter of 0.62nm, entry to the zeolite cages is unlikely, reaction to cyclohexane must be occurring on the catalyst surface. The other catalysts tested were found to be either inactive (catalyst Q with Si~-Al20 3 support) or to deactivate rapidly (catalyst O exhibiting total deactivation within 110 minutes and catalyst P within 95 minutes). In both cases selectivity to cyclohexylbenzene was very high. The high acidity of catalyst O (supported on zeolite Y) not only results in very high cyclohexylbenzene selectivity, it is also considered to be the cause of the observed rapid deactivation of this catalyst, as production of high molecular weight compounds would be favoured under such highly acidic conditions (82). As no high molecular weight compounds were evident in the products obtained it is considered that they were retained by the catalyst It was interesting to note that the spent catalyst was not significantly blackened or "coked" after use:- thus the inference is that the deactivation was internal. 163. The failure of the silica-alumina supported catalyst (catalyst Q) to produce the hydroalkylation product, cyclohexylbenzene, is explained due to its low acidity (both Lewis and Bronsted) especially when compared with catalyst A (13X supported). However, the lack of any cyclohexane in the product stream via benune hydrogenation is somewhat surprising. Deactivation of catalyst P (supported by zeolon-900Na) appears to be via a 0 different mechanism, as here the Bronsted acidity was the lowest of all catalysts tested and no Lewis acidity was detected. Examination of table 5.2, reveals benzene as the largest molecule adsorbed by zeolon molecular sieves. As a result, deactivation would appear to be due to channel blockage by reaction products or intermediates. 164. 6.1. 7. Catalyst Treatments A number of catalyst treatments were carried out in an attempt to improve catalyst performance. The treatments used were: steaming the catalyst; treatment with a halide containing compound; treatment with sodium carbonate and regeneration/cycles of use. 6.1.7.1 Steaming as an Alternative to Calcination In an attempt to increase the activity of the hydroalkylation catalyst, steaming the cation-exchanged and impregnated zeolite (catalyst A) was carried out as an alternative to calcination. Crone and Arkell (108) noted an improvement in activity of a Ni/W/Si02-Al20 3 bifunctional catalyst when steaming at 600°C was used as an alternative to calcination. Steaming is considered to increase carbonium ion activity by generation of extra-framework aluminium ions (151,152). Steaming of the hydroalkylation catalyst (catalyst A) was carried out in situ. Distilled water was passed over the catalyst (2g) in the reactor at 0.8 ml/min. The catalyst was then heated to 600°C for two hours, cooled, dried and pre-reduced, under the usual conditions, prior to the introduction of benzene. The results obtained are shown in Figure 6.14. Catalyst activity is significantly reduced compared to that of the calcined catalyst (26% ), with a marked decline in activity over the duration of the experiment. It is interesting to note the production of the fully hydrogenated product, bicyclohexyl, as opposed to cyclohexylbenzene. 165. 20 .------, 15 1 !,, 'i benzene conversion > 10 ..u ' ,,::, e ~ 5 0 L..b::::::2::;;:o=LL;;.;;;40~.:.:L..,L_~80~.;;.L...... L.;..:..-..:80...;.;.;;.;.J..._.t:.;_,;.;.10=-o....;.:.J---1-1":'.20=-----..J Time on Stream (mlns) [] bicyclohexyl D cyclohexane Figure 6.14 Product Yields as a Function of Time on Line using Catalyst A after Steaming at 600°C for 2 hours. Standard Reaction Conditions: l 7C1'C, 3.5 MPa, LHSV=l6.6 hr1, benz:H2=1:l. 166. 6.1.7.2. Chloride Treatment Halide addition is common with reforming catalysts, where it is believed to enhance surface acidity (153,154). Murtha (125) noted an improvement in cyclohexylbenzene selectivity when the catalyst was treated with a halide-containing compound, the halide being added either prior to reaction, or simultaneously with the feedstock. The amount of halide added is critical to catalyst activity in that the catalyst will be rapidly poisoned by excess halide. Too little halide, on the other hand, will not improve selectivity. To investigate the effect of halide treatment on catalyst activity, the hydroalkylation catalyst (catalyst A), pre-reduced in the usual fashion, was treated under reaction conditions with benzene to which carbon tetrachloride had been added. The amount of halide present corresponded to 8 mg per gram of catalyst, at the high end of the range considered by Murtha to be optimal, viz. 0.5 - 10 mg/g (125). After two hours of processing this feed, fresh benzene was fed to the reactor for the remainder of the trial. The results obtained are presented as a bar chart in Figure 6.15. The higher activity achieved initially with the doped feed is probably due to catalyst break in, as the conversion rapidly drops to around 12% with a corresponding decrease in cyclohexylbenzene selectivity (from 95% to 85%). Upon changing the feed to pure benzene, a dramatic improvement in catalyst performance is observed. Benzene conversion ultimately increases to approximately 30%, with cyclohexylbenzene selectivity at 94 wt%. Selectivities for cyclohexane and high 167. molecular weight products are 3.7 and 1.9 wt% respectively. When compared with data for the untreated catalyst (benzene conversion 25.9 wt%, and cyclohexylben1.C11c, cyclohcxanc and di- and tricyclohexylbenzenes sclectivitics of 77 .4, 15.S and 6.1 wt% respectively), this represents a significant improvement Addition of chlorine to the zeolite matrix would be expected to affect the acid sites in the surface and to increase their concentration (82). Reactions favoured by high acidity (e.g.,alkylation reactions) would be promoted, leading to a greater production of alkylated products. The data in Figure 6.15 suppon this, with cyclohexylbenzene accounting for 96% of the product stream, compared to 83% for the untreated catalyst The marked improvement in catalyst activity upon changing to a pure benzene feed is considered to be due to the effect of chlorine redistributing the active metal species into a more dispersed state (153,154), an effect with is not nullified without exposing the catalyst to an oxidising atmosphere or high temperature. This is supponed by the fact that product sclectivities remain largely unaltered. Clearly the acidity of this catalyst plays a very imponant role in the hydroalkylation reaction, and treatment with chlorine can give a substantial improvement to the acidic function of the catalyst, at the same time as improving the metal distribution, leading to increased activity. 168. 40r------ - 30 !,, 'i cNcl1de-free ;:: 20 benZllne 13 ,,:, l ~ A. 10 30 15 110 140 175 210 240 270 300 Time on Stream (mlns) [\] cyclohexylbenzene D high MW products • cyclohexane Figure 6.15 Product Yields as a Function of Time on Stream using Catalyst A after Chloride Addition (17Q°C, 3.5 MPa, LHSV=l6.56 hr· 1, benz:~=1:1). 169. 6.1.7.3. Sodium Carbonate Treatment Further evidence of the importance of catalyst acidity was provided in a hydroalkylation trial in which the catalysts acid sites were "poisoned" by treatment with N~C03 (3 wt% in distilled water). The catalyst was heated in the carbonate solution for 2 hours and dried under vacuum overnight Pre-reduction and reaction conditions used for the benzene hydroalkylation experiment were standard. The results are presented in Figure 6.16. The carbonate treatment was found effectively to neutralise all acid sites on the catalyst, with cyclohexane being the only product of reaction. The treatment was also found to have a detrimental effect on catalyst activity, with only low benzene conversions resulting and deactivation evident Acidity determination on this catalyst using desorption of triethylamine revealed all Bronsted acid sites to be completely neutralised and less than 1% of Lewis sites active when compared with the untreated catalyst A. Measurement of the BET surface area of the treated catalyst revealed a significant reduction in surface area, from 389.46 m2g·1 to 231.22 m2g·1• 170. Benzene Conversion (wt%) 10 ------, 8 6 4 2 30 Time on Stream (mins) Figure 6.16 Benzene Conversion as a Function of Time on Stream using Catalyst A after Treatment with Sodium Carbonate (l 7D°C, 3.5 MPa, LHSV=16.6 hr· 1). 171. 6.1.7.4. Regeneration and Cycles of Use Regeneration of catalysts is commonly practised in the petroleum industry, particularly in situations where catalyst deactivation is the result of coking. Aside from the economic benefits of recycling catalysts, it is possible to regain most of the activity of deactivated catalysts by a relatively simple treatment. Here the effect of regeneration on catalyst activity and selectivity was investigated for hydroalkylation catalyst A. Regeneration was carried out by heating the catalyst to 500°C in an atmosphere which contained increasing amounts of oxygen (5% to 20% ). The catalyst was held at 500°C for three hours. After cooling, the catalyst was reduced at 170°C and 3.5 MPa in flowing hydrogen for 15 minutes (i.e., standard pretreatment conditions) prior to any experiments being carried out. To investigate the catalyst behaviour after cycles of use, the regenerated catalyst was repeatedly deactivated by running for extended periods on line and regenerated under the conditions stated. The results obtained are shown in Figure 6.17. Both catalyst activity and cyclohexylbenzene selectivity were observed to increase after the first two regenerations. Thereafter a progressive decline in activity is observed. This observation is consistent with that of Murtha (125) who noted an improvement in the activity of this catalyst after one regeneration. 172. Conversion/Selectivity (wt%) 100 ..------ BO 60 40 20 0 ...______.______.______. ______.______.. 0 1 2 3 4 5 Number of Regenerations ___._ Benzene Conversion --- CHB Selectivity Figure 6.17 Effect of Catalyst Regeneration on Activity of Catalyst A. Details of regeneration procedure given in section 6.1.6.4. (17CY'C, 3.5 MPa, LHSV=l6.6 hr·1). 173. 6.2 HYDROALKYLATION OF BENZENE - DISCUSSION 6.2.1. Initial testing Experiments carried out in this study have confirmed the ability of metal exchanged and impregnated zeolites to hydroalkylate benzene to cyclohexylbcn7.Cne with high selectivity and reasonable yield. As stated previously, a nickel (5.3 wt%) and rare earth chloride (9.1 wt%) exchanged zeolite 13X support impregnated with 0.1 wt% platinum was used initially. This catalyst was found to be active, and the product distribution obtained closely resembled that of the previous work of Murtha et al (118) using the same catalyst under similar conditions. Benzene conversion was found to be considerably higher than previously published results (20.1 wt% cf. 9.0 wt%), with the desired product, cyclohexylbenzene being the major product formed (75.4 wt% cf. 74.4 wt%). This was accompanied by a slight increase in production of cyclohexane (19.2 wt% cf. 12.2 wt%) while heavy molecular weight compounds, m- and p-dicyclohexylbenzene, were considerably less (4.6 wt% cf. 10.4 wt%). The yields of cyclohexylbenzene observed are amongst the highest reported, especially given the very mild conditions used. Confirmation of the product spectrum obtained was carried out using gas chromatography-mass spectrometry, as were attempts to identify three unknown products. These products, with boiling points intermediate to benzene and bicyclohexyl, were not referenced in any available registry of mass spectral data. 174. From the spectrum obtained they were considered to have a molecular weight of 160. This molecular weight corresponds to the possible structure of methylcyclopentylbenzene, whose three isomers were postulated as products from biphenyl hydrogenation using a Co0-Mo0Jy-Al20 3 by Sapre and Gates (63). In support for their assignment of these structures was the observation of a 3:3: 1 molar ratio corresponding to the meta-, ortho- and para- isomers, which, due to steric hindrance, would be expected for isomerisation products derived from cyclohexylbenzene. This ratio was also found in this present study. These products, and the production of traces of methylcyclopentane, would confirm the small degree of isomerisation activity of the hydroalkylation catalyst due to its acidic function. The initial investigation of the hydroalkylation of benzene (using catalyst A) has optimised the reaction conditions and revealed the complexity of the reaction products (especially as a function of time) involving not only hydroalkylation, but hydrogenation, isomerisation and alkylation to yield high molecular weight compounds. 175. 6.2.2. Role of metals Three metal species were present on the zeolite catalyst used to produce cyclohexylbenzene from benzene in the initial hydroalkylation study discussed above:- platinum, nickel and a mixture of rare earth chlorides. The forthcoming sections consider the catalyst composition and examine the effects on the hydroalkylation reaction, in order to illuminate the following points:- the role of each metal in promoting hydroalkylation; whether there is a single active metal or if they all contribute; the types of interactions that occur between the metal species; the types of interaction that occur between the metals and the acidic zeolite support. Previous investigators have claimed varying degrees of success in the hydroalkylation of benzene, and often the catalysts employed have involved a single metal. Slaugh and Leonard (102) were amongst the first to examine the hydroalkylation of benzene and, included in their studies of a variety of mixed metal catalysts, was a report of a 5% nickel-zeolite (MS-A-3 = 75%Si02-25%Al20 3) catalyst having selectivity to cyclohexylbenzene of 64 wt%, to cyclohexane of 15.9 wt% and to C18 hydrocarbons of 12.9 wt%, at a benzene conversion of 16 wt%. The highest yields of cyclohexylbenzene reported by these investigators was with a NiFJ(NH4) 2W40 1,/zeolite catalyst The experiments were carried out in an autoclave reaction vessel, at a temperature of 200°C and pressure of 800 psi. Catalyst pretreatment consisted of calcination to 500-55<>°C with catalyst reduction in hydrogen at 300°C for 20 minutes prior to use. 176. Based on the results from studies using these catalysts and a number of other single and multi-metallic catalyst compositions, Slaugh and Leonard (102) proposed a mechanism for benzene hydroalkylation in which cyclohexene features as the key reaction intennediate (Figure 3.4). They considered that cyclohexene is fonned on a metal hydrogenation site (M) and is either further hydrogenated to cyclohexane or desorbs and migrates to an acid site where alkylation with benzene occurs to yield cyclohexylbenzene. Further alkylation of cyclohexylbenzene on an acid site was believed to yield trimers, m- and p- dicyclohexylbenzene, and tetramers. In the present study this product spectrum was confirmed (using catalyst A), coupled with isomerisation side reactions which would occur on the catalyst acid sites, producing methylcyclopentane (from cyclohexane isomerisation) and methycyclopentylbenzene (from cyclohexylbenzene isomerisation), albeit in very low yields. Yamazaki et al.(105) also reported good results with a 5%Ni on Si02-Al20 3 catalyst which had been dried at 120°C for 12 hours and calcined at 500°C for 2 hours prior to reduction in flowing hydrogen at 350-500°C for 3.5 hours. These reactor trials using benzene and methylbenzenes (toluene, m- and p-xylene, tri- and tetramethylbenzene) were carried out batch wise at a temperature of 200°C, and pressure of 56~12 kg/cm2 for periods of 1-5 hours. Yamazaki et al.(105) considered their studies on benzene and methylbenzene hydroalkylation to be supportive of the mechanism of Slaugh and Leonard (102) discussed above. 177. It is interesting to note that investigations using modified catalysts deficient in one or more metals were generally found to be inactive or to produce only cyclohexane under the conditions stated in the present study. Catalyst B, with nickel and rare earth metals but no platinum, was found to produce only cyclohexanc at very low yield (1.2 wt% conversion) (Table 5.2). Catalyst E, containing only platinum and rare earth metals (no nickel), exhibited very low activity (3.8 wt% benzene conversion) and the major product was cyclohcxane (Table 5.2). Catalyst L with no rare earth metals (platinum and nickel), showed slightly higher activity (benzene conversion 5.5 wt%) but cyclohexane was the major product with only a small amount of cyclohexylbenzene being detected (1.5 wt% of total products) (Table 5.2). A single metal catalyst comprising only nickel on zeolite 13X suppon (catalyst M) was also found to exhibit low activity (1.3 wt%) with cyclohexane the only product detected (Table 5.2). All of the above results were obtained from catalysts treated under "standard" conditions i.e., calcination carried out in air to a temperature of 524°C followed by reduction in flowing hydrogen (220 mVmin.) at 170°C, 3.5 MPa for 15 minutes prior to benzene introduction. The reactor conditions involved 2 grams of catalyst, T = 170°C, P = 3.5 MPa, liquid hourly space velocity of 16.56 hr1 and a benzene:"2 molar ratio of 1:1. While it is accepted that, in order to promote hydroalkylation, a catalyst must contain both hydrogenation or metal sites and acid sites (102,104,105,106), the roles of specific metals in multi-metallic catalysts is not clear. Especially in the case of catalysts supported by zeolites, questions arise regarding the fonn and accessibility 178. of individual metal species, their inter-relationships, and their interaction with the support and inherent 7.COlite acidity. Zeolites are easily cation exchanged with reducible transition metal ions. Depending on the particular 7.COlite structure, the cation exchange capacity (CEC - a measure of the maximum degree of exchange possible), and the number and nature of the cation sites are all variable parameters. In all 7.COlites except ZSM-5, transition-metal ions show a high affinity for the zeolite. For degrees of exchange not exceeding 50% of the CEC, all ions are fixed in the matrix (155). In the case of zeolites X, Y, and mordenite, exchange is incomplete due to the existence of sites that are hidden from direct exchange, thus placing a limit on the total metal loading possible by exchange (155). As mentioned in section 6.1.3, the maximum attainable levels of nickel and platinum loadings were less than the desired quantities. This is likely to be due to the cation exchange capacity of the zeolite supports being used and, in the case of nickel and rare earths metals, may reflect a degree of competition for overall exchange sites. After cation exchange and during degassing, metal ions distribute themselves over different sites. The site populations are considered (156) to be determined by the following: 179. (1) Energetic and coordinative differences e.g., in the hexagonal prism of a faujasite, octahedral coordination to lattice oxygens is possible, whereas in a six-membercd ring, only a one sided trigonal coordination is possible. (2) Competition for a site with other cations which can give catalysts of distinctly different properties due to consecutive or competitive ion exchange. As the energy differences amongst sites for a given structure are dependent upon the chemical composition, the two features are difficult to rationalise (156). It is during catalyst dehydration that the hydrated transition metal ions take up their initial positions in the zeolite matrix. In order to gain a high level of dispersion, it is generally considered that the rate of temperature increase should be very slow, so as to avoid the production of occluded oxide clusters or collapse of the zeolite structure. In order to activate the metal ion precursors thus deposited, reduction is necessary. This may be by using molecular hydrogen, atomic hydrogen, ammonia decomposition or auto-reduction employing the zeolite framework itself (155). The reducibility of transition metal ions in zeolites is determined by the following: (1) the structure and chemical composition of the zeolite matrix, (2) the nature and amount of co-cation, (3) the site locations in the structure, 180. (4) the presence of oxidising sites such as surface hydroxyl groups, (5) the presence of residual water, and (6) heat liberated by the exothermic reduction. Thus, reproduction of reduced metal catalysts may not always be possible, as a dynamic equilibrium exists such that as certain metal ions are reduced at certain sites there will be rearrangement of cations to ensure thermodynamic equilibrium of the system. In this study the effects of platinum, nickel and a mixture of rare-earth chlorides (predominantly lanthanum and cerium) have been examined in relation to their ability to promote the hydroalkylation of benzene when an acidic support, generally zeolite l 3X, is used. Before the results obtained in this work can be discussed fully it is necessary to briefly examine the available literature concerning the state and location of these metal species in zeolite systems. 181. Platinum Extensive studies have been carried out by Gallezot and co-workers (157,158,159) of platinum aggregates or particles associated with X and Y zeolites, in order to investigate the location, particle size distribution and structure of the particles after various preparation and pretreatment conditions and modes of use. Gallezot et al (157) considered three states of platinum dispersion: (1) atomically dispersed platinum i.e., isolated monoatomic Pt<0> atoms encaged in socialite units or dispersed in an amorphous matrix when the framework has been destroyed at high temperatures. (2) platinum agglomerates, i.e., particles of several Pt atoms which can fit into the supercages. (3) platinum crystallites, i.e., metal crystals of a size generally greater than 15 A. Using X-ray diffraction, small angle X-ray scattering (SAXS) and chemisorption, Gallezot et al (157) found that, after pretreatment at 300°C for 15 hours, pf+ ions mainly occupy the supercages of PtY zeolites. After reduction at 300°C for 15 hours platinum remains inside the porous network forming 6-13 A agglomerates, which fit into the supercages. These atoms are all exposed to gas and 182. should chemisorb hydrogen with H/Pt= 1 as reported by Dalla Betta and Boudan (160). Growing and crystal organisation of agglomerates into 25-30 A crystallites occurs between 800 to 900°C, with hydrogen chemisorption undergoing a simultaneous decrease from a H/Pt ratio of 1 to 0.15. Complete breakdown of the zeolite structure occurs after hydrogen treatment at 900°C. Crystal structure analysis shows that platinum-Y zeolites pretreated at 600°C for 15 hours contain almost all J>t2+ ions occupying Sr sites in the sodalite cages. Upon reduction at 300°C, most of the Pt atoms are found to occupy random positions throughout 2.5 A spheres concentric with the sodalite cages. These reduced atoms have no electrostatic requirements to occupy cation sites any longer and randomly distribute along the cage walls. However, 3 out of the 10 Pt'0> initially present in sodalite cages were found on Sr sites, probably due to a slight interaction with the aluminosilicate framework via a partial electron transfer from the platinum atoms to electron acceptor sites of the zeolite. Lewis sites are known to be formed during the 600°C pretreatment by dehydroxylation of silanol groups. Before reduction, J>t2+ ions are considered to have the highest dispersion since electrostatic requirements prevent the overcrowding of sodalite cages by multivalent cations. After reduction at 300°C, Pt atoms of 2.8 A diameter remain in their sodalite cages and are unable to pass through the 2.2 A apertures. Heating above 300°C causes Pt atoms to be progressively removed without any structural damage. When 183. the zeolite is evacuated at increasing temperature, Pt atoms progressively migrate out of the sodalite cages to form crystallites (20 - 30 A.) and agglomerates. The Pt atoms leave the sodalite cages as soon as they have enough energy to pass through the 2.2 A. barrier to the large pore system. The electron deficient nature of Pt aggregates in zeolites is widely accepted (161,162,163,164), with early evidence from CO and NO adsorption being sited and XPS and EXAFS giving more direct evidence for the differences from the bulk metal behaviour. While it is believed that the aggregates of platinum in supercages of faujasites are electron deficient, it is not clear as to what extent this is the result of intrinsic size effects or of electron transfer from metal to support. The influence of co-exchanged ions on platinum supported Na Y and HNa Y has been examined by Tzou et al. ( 165). They found the size and location of Pt particles in the zeolites to be strongly dependent on the distribution of Pt ions between the socialite cages and supercages, which is controlled by the calcination temperature. In agreement with Gallezot et al. (158,159). Tzou et al. (165) found that the presence of co-exchanged multivalent cations, such as La1•, Fe2+. or Ca1•, can effectively block sodalite cages and hexagonal prisms, thus forcing ft2+ ions to stay in the supercages and preventing formation of large particles on the external zeolite surface even at high calcination temperatures. 184. Nickel The reducibility of Ni2+ ions in zeolites by molecular hydrogen is known to require high temperatures. Under these conditions, metal aggregates will sinter very rapidly and agglomerate at the external surfaces of the crystals or in defects in the bulk of the zeolite crystals. All treatments or modifications which enhance the Ni reducibility will therefore decrease the degree of metal sintering and provide superior dispersions. Such possibilities include the following: (1) replacement of molecular hydrogen with atomic hydrogen; (2) preparation of hydrolysed Ni species; (3) incorporation of co-cations such as Ca2•, Ce2+, Mg2+, Ba2• and Sr•; (4) incorporation of Pt or Pd aggregates. (155) Reduction with atomic hydrogen enhances the reducibility of all transition metal ions in a zeolite (provided the reduction is not kinetically controlled) (166). Ex-situ generation of H atoms has been achieved using a microwave discharge in hydrogen, with lifetimes of atoms being sufficient to enhance reduction of the metal ions (167). In situ generation of H atoms is possible by the use of occluded Pt or Pd aggregates on which molecular hydrogen is activated and then spills over on to the metal ion to be reduced (61). It has been demonstrated that complete reduction of Ni2+ in zeolites is possible at lower temperatures than with molecular hydrogen. Less sintering of the metal occurs and a superior dispersion of Ni is obtained (168,169,170,171). 185. Using X-ray Diffraction, Delafosse and co-workers (168,169,170,171) have investigated the position, size and state of nickel ions in reduced uolite catalysts under varying pretreatment and reduction conditions, support characteristics, and in the presence of differing nature and concentration of co-cations. Examination of Ni2+ reduction and Ni0 particle siz.e in exchanged sodium X and Y uolites has been carried out using X-ray Diffraction, NMR, and temperature programmed reduction and oxidation. Many factors were found to be involved in the nickel cation reducibility. The initial locations of the Ni2+ cations were found to be distributed over the S1, Su, Sr and SIi' sites, with the more easily reduced species located at Su positions. The concentration of cations in this position was found to be dependent on the Si02:A120 3 ratio, the degree of exchange, the pretreatment conditions and the nature of any other cations present ( 170). Reducibility was also found to be dependent on the composition and redox properties of the zeolite framework, with Ni2• more reducible in z.eolite A than X or Y. Higher concentrations of zeolite framework protons, such as sodium, were also found, according to Riekert (172), to have an adverse effect on nickel reducibility in zeolite supports. The redox equilibrium (see equation 6.2) was displaced toward the left by high proton concentrations. Z-0 'Ni 2+ + H2 .., 2ZOH+ + Ni0 (6.2) z-d 186. The presence of other cations was found drastically to modify the Ni2+ reduction (168,170,171), with platinum or palladium and rare earth cations (such as Ce3+ and Lal+) being particularly effective. Platinum is considered to favour higher nickel dispersion by either increasing the reducibility through hydrogen activation/spillover, or via a stabilising effect on the nickel particles themselves. Jeanjean et al. (168) considering this scenario with nickel and platinum exchanged Na-X zeolites found that, after activation at 460°C under vacuum, 12 Ni2+ out of 17 present in the unit cell were occupying S1 sites (hexagonal prisms) and 3 Ni2+ were located at Su positions (supercages). Static hydrogen reduction for 16 hours gave no reduction enhancement for a NiNaPt-X zeolite as compared to a NiNa-X zeolite. Reduction of Ni1+ cations in hexagonal prisms (S1 sites) was found to occur at high temperatures (515°C) thereby favouring agglomeration. Only 5 Ni1+ were left on S1 sites. When reduction at a moderate temperature (350°C) was pcrformed, only Ni2• cations from Su sites were reduced, and particles located in the supercages were found to be smaller than those formed in Ni-X zeolites containing no platinum. This stabilising effect was especially evident when reduction under flowing hydrogen (at atmospheric pressure for 16 hours) was carried out The presence of platinum greatly increased the degree of reduction and a homogeneous dispersion of 25A. particles was reported. The authors considered this stabilising effect to result from small platinum clusters which provide seeds around which migrating nickel atoms could agglomerate (168). 187. Examining the effect of palladium on the reduction rate of Ni2+ ions in X zeolites, Guilleux et al. (171) considered a spill-over mechanism to be operating whereby metallic palladium atoms dissociatively adsorb hydrogen molecules to form hydrogen atoms which migrate toward the Ni2• ions and facilitate their reduction. This mechanism was only considered to operate for catalysts reduced at tempcramrcs less than 350°C. Above this temperature, Pd' atoms were considered to migrate toward the external surface and were no longer available to activate the reduction of N1•2+ • Briend Faure et al. (170) also report the beneficial effect of platinum and palladium addition to NaX zeolites, with homogeneous distributions of Ni0 atoms of particle size 25-30A being reported, regardless of the reduction temperature. In the absence of platinum or palladium, samples reduced at less than 2500C contained very small particles occluded in the zeolite framework and larger agglomerates of 60- 1ooA on the outside surface. Reduction at higher than 300°C was reported to produce very large nickel particles of size > 150A. 188. Rare Earth Metals The modifying effect of multivalent rare earth cations, such as La)+ and Ce)+, on nickel reduction for zeolite catalysts has been examined by several investigators (169,170,173,174), most notably in the context of faujasite systems. In mixed Ni2+-La)+ cationic forms of NaX, the Ni2+ ions are considered preferentially to fill the hexagonal prisms under completely anhydrous conditions (173). During reduction the hexagonal prisms are progressively emptied and La)+ ions occupy the Sr sites in the sodalite cages, as, for electrostatic reasons, both types of sites cannot be occupied at the same time (173). Comparison of the activation energies of reduction with NiNaX reveals a decreased reducibility of the lanthanum containing sample proposed to result from the formation of a mixed cation fonn of the type La)+-O-Ni2+ with extra-framework oxygen in supercages during dehydration. Cerium, on the other hand, is noted for its enhancement of nickel dispersion and reduction (174,175,176,177,178), and improvement of catalytic reactivity in zeolite systems. Ce)+ cations are considered to act as "chemical anchors" and catalyst promoters (174,175). Numerous studies have shown that one of the main factors determining the metallic dispersion in catalytic systems is the strength of interaction between the metal and the support. These interactions are thought to arise from electron transfer 189. of metallic particles to the electron acceptor sites of the support (158). It has also been observed that the presence of a modifying element which can interact strongly with the support and with the atomic metal species (or particles) decreases their mobility and favours a high degree of dispersion (179). Guilleux and co-workers (176,177) have examined the effect of Cc3+ exchanged in NiX zeolites on location and reducibility of Ni2+ ions and on the stabilisation of a highly dispersed metallic nickel species. Using X-ray Diffraction and e.p.r spectroscopy they characterised catalyst samples of varying compositions dehydrated at 500°C and 35D°C for 16 hours, to show that the Ce3+ ions not located at Sr sites were found at S1 positions. The temperature of pretreatment determines the amount of residual water molecules and therefore the cation positions. After dehydration at 3500C, oxygen were located in the extra-framework position, and were attributed to water molecules or hydroxyl groups. CeJ+ was at Sr sites and coordinated with three framework oxygen atoms and one to three water molecules occupying S0• sites. In samples with no residual water, multivalent cations such as Ce3+ and Ni2+ are considered to occupy Sr sites bonded to extra-framework oxygen and framework oxygen in tetrahedral or octahedral coordination. The CeJ+ concentration at S1 sites increases with the temperature of dehydration, with partially dehydrated samples having CeJ+ ions located at Sr sites. In the absence of water, CeJ+ ions and Ni1+ ions are found in octahedral coordination at S1 sites. Thus, during dehydration, there is competition between CeJ+ and Ni1+ to 190. occupy S1 sites. After dehydration, S1 sites are preferentially occupied by Cel+ ions leaving Ni2+ ions essentially out of the hexagonal prisms (176). The reducing properties of these catalysts have been noted to depend on the Ce3+/Ni,._ ratio, with the presence of Ce3+ ions (which are strongly electron-donating and have a high polarising power) facilitating electron transfer from the reducing centres of the zeolite to an electron-accepting molecule. The presence of Ni2+, which is electron accepting, leads to a decrease in the electron accepting or Lewis acid power of these centres. In support of this I..a3+, which does not possess electron donor properties, is not found to favour nickel reduction and leads to a poorly dispersed state (173). Djemel et al.(176) concluded that the degree of nickel reduction was dependent on several closely related factors: (a) the pretreatment conditions which determine the cation location; (b) the redox properties of the support; and (c) the extent of Na+ exchange. They considered the stabilising effect observed with Ce3+ addition to be related to the Ce3+/N12+ ratio which influences the initial position of the Ni2+ ions and the redox properties of the system. The presence of Ce3+ ions, with their high polarising field, can thus act as anchoring points, preventing the migration of nickel particles toward the outside surface of the zeolite. 191. Reports of enhanced catalytic activity for carbon monoxide (174,175) and benzene (169) hydrogenation with Ni-Ce/X catalysts have also been published, with a shift in product selectivity to higher molecular weight compounds evident in the first case. This was attributed to better reducibility of nickel, more accessible sites for carbon monoxide adsorption and a large surface concentration of hydrogen. This infers that cerium cations act not only as chemical anchoring sites but also as catalyst promoters. The promoting effect of ceria with respect to platinum has also been examined by several investigators, in view of the application in three-way auto exhaust catalysts (180,181,182). Addition of Ce is believed to promote the water-gas shift reaction, to provide storage of oxygen in the varying environment of the converter, to prevent thermally induced sintering of the alumina support, and to increase the dispersion of platinum and rhodium (180,181,182). The role of ceria in combination with platinum group metals was examined by Harrison et al. (180) using temperature programmed reduction/oxidation. Examination of ceria reduction with hydrogen revealed two peaks, a low temperature peak at 500°C assigned to the reduction of a surface species (postulated to be surface capping oxygen anions attached to a surface Ce4+ ion (28)) and a higher temperature peak (>800'C) corresponding to the reduction of bulk oxygen and the formation of lower oxides of cerium. Addition of a platinum group metal (Pt, Pd, Rh) altered the profile substantially, with the temperature of the surface oxygen reduction being 192. significantly lowered in all cases. Rhodium was found to be particularly effective in this process when comparison of the relative sizes of the surface and bulk reduction peaks was made, with the heat of adsorption of water considered to be the determining factor (it is much higher for rhodium than for platinum or palladium). Harrison et al. concluded that hydrogen dissociatively adsorbed on the platinum group metal reacts with an oxide ion in the ceria surface to form a hydroxyl group. As this hydroxyl group would be more strongly adsorbed by rhodium than by platinum or palladium, rhodium would catalyse the formation of water. This is then free to leave the metal surface. This creates an oxygen vacancy in the ceria lattice at a site near a noble metal crystallite (180). It can be seen from this overview of previous work that the behaviour, interactions and variation of mixed metal zeolite supported catalysts can be extremely complex, with the identification of specific effects complicated by the number of variables in catalyst preparation, pretreabnent and conditions of use. Examination of results are further complicated by the accepted multi-site mechanism for benzene hydroalkylation which involves both metal and acidic support sites, and by the occurrence of side reactions. 193. Hydroalkylation Catalyst It is useful to consider the role of various components of the catalyst in tmns of a postulated reaction mechanism As discussed earlier, the hydroalkylation of benzene can be expressed in terms of the reactions: 0 Metal Site/ H2 Acid Site 0 Metal Site... 0 ... . 0-0 MetalSite/H2 / , Acid Site + iAddS Tetramers Figure 6.18 Dual site Mechanism of Benzene Hydroalkylation (102). 194. The activity and selectivity of a hydroalkylation catalyst is seen to depend on a balance between the hydrogenation function and the acid function. Hydrogenation is catalysed by metals (61) and by some oxides (193). Isomerisation may also occur over both metals and oxides (194). Initial studies were focused on the nature and the role of the metal species in the catalyst. To understand the behaviour of the various species in the hydroalkylation catalysts it is necessary to examine the conditions under which they are treated, especially during the processes of calcination and reduction. As mentioned previously, the four metal deficient catalysts tested, B (no Pt), E (no Ni), I (no Rare Earths) and M (no Pt, no Rare Earths), when calcined at 524°C and subsequently reduced at 17Q°C for 15 minutes under flowing hydrogen, were found to have very low activity. Cyclohexane was the major product in all cases. In contrast to this, catalyst A (the complete metal catalyst) was found to produce cyclohexylbenzene in considerable yield under the same conditions. It therefore appears that an interaction or series of interactions are operating in this "tri-metal" catalyst. During the preparation of these catalysts, metal ions are deposited into the zeolite matrix, with the initial positions of the various species determined primarily by the other co-cations present, and the calcination temperature. From the work discussed above, it is anticipated that, following calcination at 524°C, all Pr+ ions 195. would be present at Sr sites in the sodalite cages. The majority of Ni2• ions would be initially located at S1, S0 , Sr and Sir sites following cation exchange but would migrate to S1 (hexagonal prisms) sites and to S0 (supercage) positions on calcination at S24°C. Rare earth cations such as Ce!+ would be competing for S1 and Sr sites. depending on the degree of dehydration, with the concentration of Cc!+ at S1 sites increasing with the temperature of calcination. La!+ would be expected to occupy supercage positions with any residual Na• present located at S0 sites near hexagonal windows. Of course. the positions of the various cations are not mutually independent, and important interactions can take place, especially during catalyst reduction. An equilibrium can be expected and it is not possible to predict exact cation positions. The degree of reduction of the metals is also known to be dependent upon a number of factors, most notably the zeolite matrix and hydroxyl concentration, the nature and concentration of co-cations present and the initial cation positions. It is known that, following high temperature calcination, pf• ions in sodalite cages reduced at 300°C or less remain in the sodalite cages. The Pt0 atoms are 2.SA. in diameter and are blocked by the 2.2A apenures. At >300°C, they have sufficient energy to pass through the zeolite apertures and migrate to form agglomerates and crystallites of 20-30A. in the supercages and on the external surface. The process is reasonably slow (ea. hours) (157,158,159) and the presence of 1..a2+ can retard this 196. relocation by blocking sodalite cages and hexagonal prisms, forcing J>t2• to remain in the supercages. The temperature of reduction of nickel is lowered considerably by the presence of platinum, with Nil+ ions at Su sites forming a 25A dispersion after reduction at 3500C (168). At temperatures greater than 3500C, very large nickel particles (> 150A) are formed. The number of Nil+ ions at Su sites is considered to be dependent on the presence of Ce3+, with Ce3+ and Ni2+ competing for S1 sites (176). 2 As Ni + at Sn (supercages) sites are reduced at a lower temperature than those at S1 sites (183), addition of CeJ+ would be expected to assist low temperature nickel reduction. For the hydroalkylation catalyst system, it can be postulated that the following scenario may be occurring on calcination at 524°C followed by reduction at 1700C for 15 minutes (standard conditions). Platinum cations are reduced at Sr 0 sites to Pt • These assist in the reduction of Ni2+ at this low temperature by dissociative hydrogen spillover and form a well dispersed metal distribution. The presence of Ce3+ (which preferentially fill S1 sites) forces Ni2+ into the more easily reduced Sn positions and promotes their low temperature reduction. In addition the Ce3+ ions hinder the mobility of the well dispersed nickel atoms by acting as an anchoring point for the nickel atoms. The beneficial effects of Ce3+ can however be hampered by high concentrations of residual Na+ or excessive La2• loadings. 197. Attempts were made to confirm these suggestions from temperature programmed reduction studies (Figures 6.12 and 6.13). Ni/zeolite (catalyst M) showed little low temperature reduction, with hydrogen uptake becoming important at over ea. 4500C. Peaks at 500°C and 637°C were assigned to Ni in the supercage (Sn) and sodalite (Si) positions, and this agrees with the suggested catalyst structure. Addition of rare earth was found slightly to change the temperature of reduction (517°C and 603°C respectively) with the amount of Ni associated with S1 sites (hydrogen taken up at ea. 600 - 630 C) being markedly reduced (Figure 6.13). This supports the proposal that Ce3+ and Ni2+ compete for S1 sites, and that Cel+ displaces Ni2+. Some reduction of rare earth metals was found in the TPR profile of RE-13X (Figure 6.13) at the rather low temperature of 186°C. Upon addition of platinum (catalyst E, Figure 6.13) this peak was also present, shifted to the lower temperature of 170°C. For Ni/RE-13X (catalyst B, Figure 6.13) the peak at 193°C is therefore attributed, at least in part, to the low temperature reduction of rare earth metals. As the weight loss for this peak is larger than the weight loss for reduction of rare earth in this temperature range (ea. 186°C, Figure 6.13) the size of this peak also indicates the low temperature reduction of nickel upon rare earth addition. The most probable 198. explanation of this low temperature peak rests with the possibility of reduction of rare earths and of nickel, enhanced by spillover of hydrogen from reduced rare earth metal. It is seen from figure 6.13 that rare earth oxides on the 7.COlite can also be reduced, with peak maxima occwrlng at l 86°C and at 61 ere. There is the possibility of a further peak at higher temperatures, which could not be reached using the DuPont Thermal Analyser due to difficulties with the furnace thermocouple. There is little doubt that the peak observed at ea. 603°C for Ni/RE (catalyst B) originates, at least in part, from reduction of rare earth oxides. As a result, Ni remaining in the S1 position (ea. 517°C) is seen to be small. The effect of addition of Pt to the system is obvious in all cases. In the case of the rare earths (catalyst E (Pt/RE-13X) compared with Rare Earth- 13X, Figure 6.13), it is seen that the temperature of reduction of rare earth oxides is reduced. The high temperature peak at 6100C is displaced to 580°C while the low temperature peak is displaced from 186°C to 17Q°C. Unresolved peaks at 210°C and 43Q°C were also found. None of these peaks can be attributed to Pt since the metal was present in quantities too small for weight losses due to their reduction to be significant They do show, however, the marked effect of platinum addition on rare earth reduction, and the low temperature peaks are similar to peaks reported by Yao et al. (182) and attributed to interaction of Pt with surface capping oxygen in ceria. 199. Addition of Pt to Ni-13X (catalyst I) had a significant effect The temperature of the high temperature reduction maxima was reduced, to 474°C and 592°C respectively, indicating that Pt was assisting in the reduction of Ni The broad plateau observed with catalyst M (Ni- 13X, Figure 6.12) was altered upon Pt incorporation with the observation of poorly resolved peaks at 223°C and 367°C (Figure 6.12). Although these peaks were not due to Pt alone, they show the increased weight loss due to low temperature reduction of nickel in the presence of platinum. Previous studies (195) of Ni and Ni/5%Pt on alumina show that reduction of nickel oxide occurs at 38D°C and 300°C, respectively. In the Pt/Ni-13X system, which contains only 0.1 wt% Pt, reduction at less than ea. 250°C is unlikely. As a result, the weight loss at ea. 223°C may not be entirely due to reduction of nickel oxide. This supports the role of rare earth metals in promoting low temperature reduction of nickel in this catalyst system. Turning to the hydroalkylation catalyst A (Pt/Ni/RE- 13X, Figure 6.12), it is seen that the amount of Ni remaining in the S1 site is small (peak at ea. 583°C). Reduction is concentrated at low temperatures. Again comparison of Ni/Pt-13X with and without rare earth (Figure 6.13) reveals clearly the role of the rare earth m displacing Ni from S1 sites (loss of peak area at ea. 583°C). 200. The low temperature peak at 192°C is in agreement with that observed with catalyst B (Ni/RE-13X, Figure 6.13), and represents a significantly higher weight loss. It also correlates with displacement of the 223°C peak of Ni/Pt- l 3X (catalyst I, Figure 6.12) to a lower reduction temperature upon addition of rare earth oxides. This peak represents a considerable increase in the weight loss in this low temperature region by the combined actions of nickel, platinum and rare earths. Some efforts to confirm these suggestions were made. Examination of this series of hydroalkylation catalysts for crystal structure determination using X-ray diffraction was not possible. It was hoped that microscopy (both scanning electron microscopy (SEM) and transmission electron microscopy (IBM)) would give valuable visual representations of the catalysts inner and external surfaces but, unfortunately, sample preparation and instrument limitations proved problematic. However, the postulated cation behaviour discussed above with respect to temperature programmed reduction was examined further in respect to the effects of metals on catalyst activity and selectivity for the hydroalkylation reaction, and the influence of catalyst reduction temperature. Due to their inactivity, all metal deficient catalysts (B, E, I, M) were investigated under increasingly harsh reduction conditions to improve reducibility and (possibly) activity. The single metal catalyst, M, which consisted of nickel supported on zeolite 13X was inactive under mild pretreatment conditions, producing 201. only small quantities of cyclohexane (yield 1.3 wt%) (Figure 6.11). This is to be expected as the nickel cations would not be reduced under these low temperature conditions, as is shown in the TPR profile of this catalyst (catalyst M, Figure 6.12), with reduction of nickel requiring temperatures >400°C. Increasing the ·reduction temperature to 240°C resulted in a slight increase in benzene conversion (5 wt%) but, after reduction at 475°C, conversion approximating that of catalyst A was obtained (18.5wt% cf. 20.1 wt%) with cyclohexylbenzene as the major product (selectivity 57.8 wt%) accompanied by significant quantities of cyclohexane (35.0 wt%) and dicyclohexylbenrene isomers (17.2 wt%). The literature indicates that large nickel particles would be formed under high temperature reduction conditions (168,169,170,171) and such particles are known to coke easily ( 196). This is supported by the production of significant quantities of dicyclohexylbenzene, indicating the occurrence of further alkylation of the hydroalkylation. Further polymerisation leading to coke formation would not be unlikely and the deactivation observed with this catalyst was not unexpected. Inclusion of rare earth oxides to the catalyst (catalyst B: Ni/RFJzeolite 13X) led to virtually no activity being observed after reduction at low temperatures and the system produced only small quantities of cyclohexane (1.2 wt% yield) (Figure 6.9) Some assistance from cerium ions in reducing and stabilising the nickel ions had occurred (as expected from Figure 6.13) and increasing the reduction temperature would be expected to increase the reduction of Ni. A significant 202. increase in benzene conversion (17. 7 wt%) was observed after reduction at 400°C, with cyclohexylbenzene as the major product (selectivity 93.55 wt%) (Figure 6.9). This represents a significant improvement in cyclohexylbenzene selectivity on reduction of metal. However, the catalyst was also found to deactivate rapidly, probably due to the size of nickel particles formed at these higher reduction temperatures and to their coking. Thus, the role of the rare earth oxide in this catalyst appears to involve a small decrease in the reduction temperature for nickel and a significant increase in selectivity in the hydroalkylation reaction. The role of Pt was then examined using a Pt/RE- 13X catalyst (catalyst E, Figure 6.8). Platinum should be reduced at low temperatures without difficulty. However this catalyst exhibited low activity (3.8 wt%) with cyclohexylbenzene as the major product (selectivity 65.5 wt%). An increase in reduction temperature led to increased benzene conversion (19.2 wt%) with cyclohexylbenzene as the major product (83.5 wt%), and a CHB:CH ratio of 5:1. Deactivation was rapid. The increase in benzene conversion observed in these results is somewhat surprising, given that it would be expected that platinum would be reduced after treatment at low temperature. Thus both platinum and nickel give reasonable conversion of benzene to cyclohexylbenzene once efficient reduction of the metals has been achieved. 203. Combinations of Pt and Ni on zeolite 13X were then examined (catalyst I: Ni/Pt-13X, Figure 6.10). No rare earths were added to this catalyst. It was found that, following low temperature reduction, the catalyst exhibited low activity with cyclohexane as the major product (yield 5.3 wt%) (Figure 6.10). A very large increase in catalyst activity (44.4 wt%) was observed upon increasing the reduction temperature to 400°C, with cyclohexane remaining the major product (65.8 wt%) and a CHB:CH ratio of 0.5:1 being produced. Rapid catalyst deactivation was also observed. The benzene conversion with this catalyst (44.4 wt%) was approximately twice that observed over the Ni/zeolite catalyst (18.5 wt%). The selectivity ratio CHB:CH was 0.5 in the Ni/Pt-zeolite catalyst and 1.6 over the Ni-zeolite catalyst This result led to the suggestion that Pt could also be active for beni.ene conversion. A Pt/RE-13X catalyst (E) was found to give benzene conversions of ea. 19 wt% but with a much higher selectivity (CHB:CH = 5). Similarly, the Ni/RE-13X catalyst gave a benzene conversion of 17.7 wt% and selectivity of CHB:CH = 14.5. Thus it would seem that both Pt and Ni are active in the conversion of benzene but that rare earth oxides have a major role to play in dictating selectivity. 204. The role of rare earths and zeolites in these catalysts is also clear from selectivity studies involving metals supponed on other materials. Numerous investigations of the hydrogenation of benzene over nickel supponed catalysts have been carried out, most notably with Ni/Si~ and Ni/Al20 3 (24,25,26,27,28,36). In all cases, under widely varying catalyst reduction and reaction conditions, cyclohexane was the only product reported and no higher molecular weight species were produced. Similarly for platinum supponed catalysts (37,38,39,40,41,42) activity towards hydrogenation has been reponed with the absence of any side reactions. The catalyst most commonly studied in this case was Pt/y-Al20 3• These findings highlight the observations of this present study, that the presence of rare earth oxides and an acidic suppon (zeolite 13X) are responsible for cyclohexylbenzene selectivity via hydroalkylation. Comparisons with the complete catalyst (catalyst A, Figure 6.12) also help to elucidate the role of the various components. This catalyst was found to exhibit considerable activity (benzene conversion 20.1 wt%) after low temperature reduction at 17Q°C for 15 minutes and reaction under standard conditions (Table 6.3) and to deactivate very slowly (Figure 6.1). Examination of the temperature reduction profile for this catalyst (catalyst A, Figure 6.12) reveals reduction at a considerably lower temperature than for the other catalysts tested. The bulk of the nickel is seen to be reduced in the low temperature region (ea. 192°C) with only a small amount of Ni remaining at the high reduction 205. temperature of S83°C (associated with reduction of nickel at S1 sites). The activity of this catalyst upon moderate temperature reduction and its temperature programmed reduction profile support the postulates that nickel reduction is favoured at lower temperatures by the presence of platinum and rare earth oxides, by action of spillover of hydrogen from reduced well dispersed platinum centres formed at this low temperature and by displacement of Ni2+ into more easily reduced S0 positions, respectively. No activity/selectivity studies were carried out over this catalyst after reduction at higher temperatures, as it was considered that any benefit in reduction would be negated by the onset of catalyst deactivation due to the displacement and agglomeration of platinum species and the formation of very large nickel particles. The absence of marked deactivation with this catalyst was also attributed to the stabilising influence of rare earth oxides, which are known to act as anchoring sites for reduced nickel particles in z.eolite systems (174,175) via electron transfer from metallic particles to the electron acceptor sites of the support (158). The production of cyclohexylbenz.ene with high selectivity (75.4 wt%) over this catalyst indicates the balance of metallic hydrogenation and support acid functions in this system. The presence of these two catalytic functions is further confirmed by the side reactions taking place: hydrogenation of benz.ene to produce cyclohexane (selectivity 19 .2 wt%) (over the metal function); further alkylation of 206. cyclohexylbenzene to produce dicyclohexylbenzene isomers (4.6wt%) (probably over the acid function); · traces of isomcrisation products (methylcyclopentane and methylcyclopcntylbenzenes) (probably due to either acid or metal functions); and traces of bicyclohexyl, due to further hydrogenation of cyclohexylbenzene. From these comparisons, it is clearly seen that the support and the rare earth oxides that form part of the support have an important role to play. The observations support the benzene hydroalkylation reaction mechanism proposed by Slaugh and Leonard (102) (Figure 3.4) and indicates that not all metal sites on the hydroalkylation catalyst are equivalent. Platinum metal and rare earth cations or oxides have specific roles in providing activity and aspects of selectivity for cyclohexylbenzene production. The effect of rare earths on the dimerisation selectivity of the reaction also implies interaction between these oxides-metals and the zeolite support. Attention was therefore directed to the reactions of Figure 6.18 which involve an acidic function, such as dimerisation, alkylation and isomerisation, in order to elucidate the respective influences of rare earth oxides and the catalyst support (see section 6.2.3). This last point will be discussed funher in section 6.2.3, which deals with the hydroalkylation catalyst support effects. 207. Using standard reduction (17 The nickel series showed a large increase in catalyst activity between l. 7 and 2.9 wt% Ni (Figure 6.5). This may have been caused by approximately 0-2 wt% being accepted into the ion exchange sites of the zeolite matrix, leaving loadings above this amount available as easily accessible active sites. Cation exchange capacity (CEC) for NaX zeolites is referenced as 6.25 metal equivalent/gram (191) with the %CEC for Ni2+ reported as being in the range 43-72% and strongly dependent on the pH of the exchange solution, the temperature of exchange and the presence of co-cations (192). In the present study, cation exchange of the zeolite supports with nickel was carried out in combination with the mixture of rare earth cations and it is difficult to predict the amount of nickel present in exchange sites of the zeolite matrix for the various loadings. It would appear that not only must metal reduction be promoted (this catalyst also contained platinum and rare earth chlorides) to achieve activity, but the nickel loading must also be sufficient. Once sufficient nickel was present (between 1. 7 wt% and 2 wt%), benzene conversion was found to reach a maxima at approximately 24 wt% and to level off to 20 wt% for higher nickel loadings. 208. The influence of nickel loading on product selectivities observed is somewhat surprising (Figure 6.5, lower graph). It appears that, as the nickel content increases, the cyclohexylbenz.ene selectivity increases at the expense of cyclohexanc production. This may be caused by approximately 2 wt% Ni2• being present in more difficult to reduce sites which are not in the optimum positions for interaction with the rare eanh oxides and acidic sites of the support. Incorporation of higher loadings of nickel, leads to other sites being occupied with production of cyclohexylbenzene increased and accompanied by production of high molecular weight species formed by funher alkylation of cyclohexylbenzene. A maxima in cyclohexylbenzene production was found at 5.3 wt% nickel. Funher addition of nickel led to a significant increase in high molecular weight products at the expense of cyclohexylbenzene. From the reaction network for hydroalkylation (Figure 6.18) this increase in higher molecular weight compounds would be expected to result from increased acidic function. This is not expected upon nickel addition, however, the addition of Ni2+ to previously unfilled sites may lead to accessibility of reaction intermediates to more acidic sites on zeolite support. Metal salts have also been reported to have activity for dimerisation reactions (61). The effect of increased platinum loading on benzene conversion is dramatic (Figure 6.6). Without platinum present the catalyst is seen to be virtually inactive under the moderate pretreatment conditions employed. Addition of 0.1 wt% Pt results in a benzene conversion of 20 wt% being achieved. Changes caused by 0-1 wt% Pt are uncertain. Whether there is a "jump" in activity upon Pt addition which 209. stays relatively constant upon addition of more platinum, or whether the increase is steady as depicted on the diagram is uncertain. What is clear however, is the importance of platinum in promoting catalyst activity. This observation agrees with the postulate of well dispersed platinum metal promoting reduction of nickel at this low temperatuIC. The slight increase in benzene conversion observed with further increases in platinum loading (0.1-0.18 wt%) may result from further improvement to dispersion and reducibility of Ni2+ or to the activity of the reduced platinum. The observed effect on selectivity for this series of catalysts supports the former postulate. Once platinum is present in the catalyst, cyclohexane production decreases dramatically, and cyclohexylbenzene is the major product (selectivity approximately 80 wt%). Production of high molecular weight compounds was also observed. Again it is not clear what is occurring between O and 0.1 wt% Pt. However, the major role of platinum in promoting hydroalkylation is confirmed. Similar behaviour is noted for reaction product selectivities upon rare earth incorporation (Figure 6. 7). Without rare earths, cyclohexane is the only product observed. Upon addition of 2.5 wt% of rare earth metals, cyclohexane is replaced by cyclohexylbenzene as the major product (approximately 70 wt%) and high molecular weight products are fonned. Further addition of rare eanhs did not alter the product distribution greatly, although a slight maxima in cyclohexane was found at 5.2 wt% rare earths. These results clearly show the influence of rare earth species on hydroalkylation activity with this catalyst. 210. The effect of rare earth loading on benz.ene conversion (Figure 6. 7, top graph) is seen to be relatively constant As total reduction of rare eanh oxides would not be expected at the low temperature at which these catalysts were pretreated (sec Figure 6.13), the increased activity is thought to be due to the role of rare eanhs in enhancing the reduction of nickel at this temperature (and in combination with platinum) via displacement of Ni2• from difficult to reduce S1 sites to more easily reduced S0 sites (see Figure 6.13). The action of rare eanh oxides in providing anchoring sites for nickel would also lead to an improved distribution and higher activity. It has been found that, to achieve activity and selectivity toward cyclohexylbenzene, the role of each type of metal is very important, and two major influences due to metals are occurring:- Firstly, it appears that, in the absence of platinum on the catalyst, activation via molecular hydrogen reduction requires higher temperatures. Catalyst activity, once achieved, is short lived due to rapid catalyst deactivation of large nickel particles produced under these conditions. This behaviour suggests that Pt is required 2 to assist Ni + reduction probably by a dissociative H2 spillover mechanism, and that nickel is the major active metal species in the hydroalkylation reaction. Temperature programmed reduction profiles of these catalysts support this scenario (Figure 6.12). 211. Secondly, when comparison of the product selectivities of catalyst M (Ni/zeolite) with catalyst B (Ni/RFJzeolite) arc made after activation at high temperatures, a marked increase in cyclohexylbcnzene selectivity is observed upon incorporation of rare eanh metals into the catalyst. The ratio of CHB:OI for catalyst M is 1.6, while that for catalyst B is 14.5. Catalyst A (Pt/Ni/RFJzeolite) under standard reduction and reaction testing conditions results in a CHB:OI ratio of 4:1. Clearly, the role of the rare earth metals is not only to stabilise the active nickel particles, but also to optimise the cyclohexylbenz.ene selectivity. The importance of the platinum-nickel relationship was verified by carrying out the hydroalkylation of benzene with a physical mix of two catalysts, one deficient in nickel (catalyst E, Pt,RFJzeolite), and one deficient in platinum (catalyst B, Ni,RE/zeolite). Thus all metal species were present in the reactor but the nickel and platinum were no longer on the same surface. This combination was found to be inactive, giving further suppon for the presence of a dissociative hydrogen spillover mechanism operating from platinum to nickel on the hydroalkylation catalyst. Examination of benzene hydroalkylation results of some previous investigators also reveal the importance of balancing metal functions and the pretreatment conditions. Louvar and Francoy (104), in their attempts to hydroalkylate benzene with a series of Group VIII metals in the presence of acids (SiOJAl 20 3, A1Cl3, BF3), found low benzene conversions with cyclohexane as the major product. 212. Following the results of the present study it can be seen that, as single metals were employed in the work of Louvar and Francoy (104), the reduction conditions used were generally inappropriate to give the dispersions of reduced metals necessary. 213. 6.2.3. Support effects Examination of the hydroalkylation network (Figure 6.18) reveals that not only arc metal functions necessary for cyclohexylbenzcne production, as has been discussed in the previous section, but that the important step of dimerisation or alkylation requires acid sites. It is therefore necessary to examine the role of the acidic catalyst support in the hydroalkylation catalyst. It is necessary not only to consider the support in the context of the acidity requirement of the catalyst, but also in terms of the geometric consideration of zeolites. Pore size and aperture diameters must be considered. The development of zeolite acidity is known to be a function of a number of catalyst preparation variables, including cation exchange with polyvalent cations, reduction of transition metal cations and the calcination temperature, which effects the development of Lewis acid sites. In all cases of catalyst preparation carried out in this study, cation exchange of the various catalyst supports was carried out under the same conditions of metal loading (0.1 wt% Pt, 5 wt% Ni and 10 wt% rare earth chlorides) and exchange media, with calcination at 524°C used to maximise the population of Bronsted acid sites (90,91). The activity and selectivity to benzene hydroalkylation of catalysts prepared with various supports is shown in Table 6.5, while their acid characterisation results are detailed in Table 6.4. 214. The acidity of the catalyst suppon, by providing an alkylation site on the catalyst surface, is of prime imponancc in the production of cyclohexylbenzenc (101,102,105,106). This influence can be seen in Figure 6.19 where cyclohexylbenzcne selcctivities (from Table 6.5) are shown as a function of the total acid concentration (Table 6.4) for the variously supported catalysts, and for catalysts of varying rare eanh content (Tables 6.3 and 6.6, respectively). 100.::.==--=~---=:...... :..-...:..._CHB Selectivity (wt%) ______-, 80 .... · · · ··············· · · · · ··········· ...... · ...... ····· ...... );!: ;)(········· 60 .. ;. · · · ..... 40 ...... · .. . 20 o~--====::::t:::==tL---.L_---___l.-----'------' 0 10 20 30 40 50 Total Acid Concentration (XlOOO) -e- Various Supports · -~·· RE Loading-13X Figure 6.19 Influence of Catalyst Support Total Acidity on Cyclohexylbenzene Selectivity, Catalysts P, Q, I, J, K, N, A and 0. Standard Reaction Conditions: 17 Catalyst N supponed by zeolitc 5A was found to have higher hydrogenation activity than catalyst A (13X), although it exhibited reasonable acidity. It can be noted however, that this catalyst did have weaker acid sites than 13X and this, coupled with the lower acid concentrations may have caused this behaviour. As hydrogenation activity is enhanced by electron transfer from metals to suppons, this may also have been occurring in this catalysl As expected cyclohexylbenzene selectivity was observed to increase dramatically with total acid concentration for hydroalkylation catalysts on differing supports. Catalyst P (z.eolon support) however, exhibited very low acidity and high CHB selectivity. This catalyst was also found to have low activity and to deactivate very rapidly. Zeolon is a z.eolite of orthorhombic symmetry with pore size of 0.67x0.7 nm (Table 3.1 ). Due to its kinetic diameter of 0.62 nm, benzene is the largest molecule adsorbed by z.eolon. It is therefore considered that this catalyst would be expected to exhibit low catalyst activity (due to slow mass transfer of benzene into the cavities and to the active metal sites) and to deactivate very rapidly. Even a small amount of coke formation would slightly block zeolite pores and reduce the access of benzene. The high cyclohexylbenzene selectivity found with this catalyst is attributed to the limit of detection of the gas chromatographic analysis carried out, as, this catalyst exhibited very low activity (cf. 5wt%) prior to its complete deactivation. Thus the quantity of cyclohexylbenzene produced was very small, and was probably formed due to reaction with surface acid sites prior to deactivation. 216. Catalyst deactivation was also observed with catalyst Y as a consequence of its very high support acidity. In this case, initial catalyst activity was very high (cf. 45 wt%) prior to deactivation. It has been found (175) that the acidity of the suppon can generate strong mctal-suppon interactions leading to electron depletion of the metal and, consequently, to coverage during the reduction process by strongly adsorbed hydrogen species. Losses in reactivity due to the absence of free metallic surface sites can then occur. This scenario may provide an explanation for the rapid deactivation of the zeolite Y supported catalyst Perhaps more likely however is the possibility of acid catalysed polymerisation leading to fouling of the catalyst (67). The catalyst supponed by zeolite 13X (catalyst A) exhibited moderately high acidity, and good selectivity for cyclohexylbenzene. This catalyst contained 0.1 wt% Pt, 5 wt% Ni and 9 wt% rare eanh metals as chlorides. Although catalysts of differing metal loadings and catalysts deficient in metal species were generally not tested for their acid characteristics, examination of selectivity results indicate the crucial influence of the acidic 13X suppon is unchanging. Upon high temperature reduction of the metal deficient catalysts, selectivities to cyclohexylbenzene were found to be high for Pt/RE-13X and Ni/RE-13X (83.5 and 93.5 wt%, respectively) and significant for Pt/Ni-13X and Ni-13X (65.8 and 57.8 wt%, respectively)(see Figures 6.8-6.11). Thus it can be seen that suppon acidity plays a major role in dimerisation activity via alkylation and that one role of the rare eanh may be to control acidity in the catalyst. The inactivity of the Si02-Al20 3 supported catalyst is attributed to the blocking of the support pores by metal species during catalyst preparation. This is supponed by the very low total surface area found for this catalyst. 217. Examination of Figure 6.6, showing hydroalkylation selcctivities as a function of rare canh metal loading, reveals, very little cyclohexylbcnttne is produced in the absence of rare earth cations and cyclohexane is the major product. Upon introduction of 2. 7 wt'1, rare eanh metals, production of cyclohexylbeD7.Cne (69 .2 wt%) is found. This behaviour indica1es that production of cyclohexylbenzcne is influenced by the presence of rarc canh metals. Rare eanh cations were exchanged into the 1.COlite catalyst as chloride ions and acidity enhancement due to the presence of increased Cl"' was considered possible. Determination of the acidity of catalysts with varying rare earth content (catalysts I, J, K) was carried out to examine this point The results are presented in Table 6.6. The cyclohexylben7.ene selcctivities of these catalysts (fable 6.3) as a function of their acidity (fable 6.6) are included in Figure 6.19. Acid site strengths for these catalysts, as reflected by the temperature of desorption of triethylamine, are generally in the same range. The concentration of sites however, varies dramatically with rare eanh content, with catalyst I (no rare eanh metals) having very low total acid concentration, especially in comparison to catalyst A (9.1 wt% rare eanh). The behaviour of the intermediate rare eanh loaded catalysts does not give a linear increase in acid concentration with rare earth loading. These acidity detenninations show that the hydroalkylation catalyst acidity is not a result of in....-reased chloride content due to cation exchange with rare earth chlorides, as, in that case, a linear relationship would be expected. Neither is it a function purely of t..'le 7.COlite suppon as, if this were the case, the variation in rare earth metal cont~:-:: would have little effect. 218. Table 6.6 Acidity Determinations on Catalysts of Varying Rare Earth Metal Contents (Catalysts I, J, K, A). Catalyst I J K A Designation Rare Earth Metal 0.0 2.7 4.7 9.1 content (wt%) Catalyst Support zeolite 13X zeolite 13X zeolite 13X zeolite 13X Peak Range - 200 - 400 200 - 400 225 - 400 250 - 450 Lewis Acid Sites ("C) Peak Maxima - 368 343 248,350 338,388 Lewis Acid Sites ("C) Peak Range - 400 - 500 400 - 550 400 - 500 450 - 550 Bronsted Acid Sites ("C) Peak Maxima - 460 498 461 509 Bronsted Acid Sites ("C) Lewis Acid 486 1325 1139 18562 concentration (arbitrary units) Bronsted Acid 920 380 1080 17937 concentration (arbitrary units) Total Acid 1406 1705 2219 36499 concentration (arbitrary units) Total BET Surface 294.1 220.7 301.4 389.5 Area (m2/g) We have shown that the acidity and pore size of the zeolite is important in the hydroalkylation of benzene. It is unclear however, what influence beyond acidity control the rare earth oxides on the support fulfil with respect to dimerisation activity. In the last section, a relationship between rare earth oxides present in the support and cyclohexylbenzene selectivity was postulated for the fully reduced catalysts. This is now .examined in more detail. 219. The stability of metals encaged or occluded in a zeolite matrix is known to depend on a number of factors, including the interaction of metal atoms with anchoring sites on the su~ and the geometry of the porous network. Anchoring sites can be either Bronsted or Lewis acid sites or high electrostatic field cations (such as Cc)+ discussed previously), with an induced electron transfer or polarisation of the metal atoms increasing the strength of association with the zeolite framework (175). This last point is especially interesting in the context of catalyst stabilisation. The catalysts tested for hydroalkylation of benzene have often exhibited rapid catalyst deactivation. In some cases this was attributable to catalyst suppon characteristics such as high acidity, which led to polymerisation reactions, and fouling. For catalysts reduced at high temperatures, catalyst deactivation was considered to be caused by sintering of metal species which led to rapid coking. The complete hydroalkylation catalyst (catalyst A, Pt/Ni/RE- 13X), while exhibiting superior stability to the above mentioned catalysts, was also found to deactivate very slowly. The findings discussed above regarding stabilisation of catalysts via anchoring of metal sites by high electrostatic field cations (for example CeJ+ and La2•) could be a feasible explanation for the improved stability of the hydroalkylation catalyst. The action of these cations was not sufficient to retard all deactivation of this catalyst. As a consequence, kinetic investigation of this reaction scheme was not possible. 221. Summary Investigation of the hydroalkylation of benzene has shown the many different reaction pathways that are possible with this system, and the importance of understanding the operation of the catalyst used in order to maximise its efficiency for production of cyclohexylbenzene. Examination of the role of metals and suppon in the hydroalkylation catalyst has revealed a complex series of interactions and influences determining catalyst behaviour. The type of metals, metal loadings, pretreatment conditions, support used and reaction conditions were all found to play some pan in this system. The metal present on the catalyst (principally nickel) was found to be responsible for catalyst activity. In order for reaction to occur, reduction with molecular hydrogen was required, and to achieve this at low temperatures the presence of platinum was necessary, such that hydrogen spillover from platinum to nickel could occur. This low temperature reduction was considered to result in highly active dispersed nickel. In the absence of platinum, higher reduction temperatures were required and these led to rapid catalyst deactivation due to coking of large nickel particles produced under these conditions. The presence of rare earth metals was also found to assist in lowering the temperature of reduction for nickel, and one role of ceria in enhancing reduction appears to involve the displacement of Ni2+ from S1 sites to the more easily reduced S0 sites. 222. The action of the acidic zeolite suppon in promoting the alkylation reaction was clearly seen, with pore si7.c effects of the suppon also effecting catalyst activity and stability. Supports with very high acidity were found to deactivate rapidly, probably as a consequence of polymerisation reactions and fouling. Rare earth oxides present on the support were found to effect catalyst selectivity, with dimerisation enhanced by their presence. This was considered to be due to their role in modification of zeolite acidity, and possibly by promotion of the mobility of reaction intermediates (such as cyclohexene) enabling their migration to acid sites for alkylation. Stabilisation of highly dispersed nickel was also attributable to rare earth species, with the observation of improve catalyst life upon its incorporation and in the absence of high reduction temperatures. As catalyst deactivation was still observed with the optimised hydroalkylation catalyst, no determination of reaction kinetics was possible. Clearly, further improvement in catalyst stability is necessary. 223. 6.2.4. Catalyst Treatments Following the detailed investigation of the hydroalkylation mechanism and catalyst functions, a series of catalyst treatments were carried out to attempt to improve the catalyst performance (steaming and addition of chloride), to clarify the effect of acidity on the catalyst performance (sodium carbonate treatment and chloride addition), and due to the catalyst deactivation observed, to investigate catalyst regeneration after cycles of use. 6.2.4.1. Steaming As an alternative to calcination, steaming of catalyst A was carried out in-situ at 600"C in an attempt to increase the carbonium ion activity of the catalyst by generation of extra-framework aluminium ions (151,152). The activity was found to be less than for the calcined catalyst A, with cyclohexane the major product and significant catalyst deactivation (Figure 6.14). This behaviour can be explained by the observations made by Jacobs et al.(181) in their examinations of steamed NiY zeolites. It was found that the amount of hydrogen chemisorbed on the nickel phase of steamed catalysts decreased drastically compared with the non-steamed sample, indicating a high degree of metal sintering had occurred in the steamed samples. This supports the rapid deactivation noted earlier upon high temperature catalyst reduction, and the postulate that this was caused by sintering of nickel. 224. 6.2.4.2. Chloride Treatment Addition of chloride in the form of carbon tetrachloride to the benzene feed in a hydroalkylation experiment resulted in some interesting results, sec Figure 6.15. After an initial period of high catalyst activity (considered to be a feature of catalyst "break-in") and extremely good selectivity to cyclohexylbenzene, there was a dramatic loss of activity, to a value less than with an undopcd benzene feed. Upon switching to a chloride free benzene feed during the run, the benzene activity and cyclohexylbenzene selectivity were restored to the extremely high values first encountered. Halide addition is known to increase acidity (153,154) and this is reflected in the high cyclohexylbenzene selectivities as expected. The effect on catalyst activity is attributed to the poisoning of active sites by undesirable choride compounds. Upon the introduction of a chloride free benzene feed, these are removed from the catalyst and the beneficial effects of chloride addition are evident Volatilisation and redispersion of metals is considered to be favoured by Cl containing compounds, with CCl4 used as a feedstock additive to regenerate the activity of hydrocracking Pd/zeolite catalysts (182). 225. 6.2.4.3. Sodium Carbonate Treattnent As discussed above the role of zcolite acid sites in the hydroalkylation catalyst is very important In an attempt to modify the catalyst activity after preparation, catalyst A was "poisoned" with a dilute N3.iC03 solution. This catalyst treattnent was found to effectively remove all acidic functions of the catalyst, with cyclohexane the only product recovered. It also led to rapid catalyst deactivation. The low catalyst activity and deactivation are attributed to the poisoning of active sites by sodium and possible changes in electronic properties of the catalyst via its interaction with metal species. 6.2.4.4. Regeneration and Cycles of Use Regeneration of deactivated catalyst A was carried out by heating in a stream containing an increasing concentration of oxygen (5% """ air) to a temperature of 500°C for 3 hours. The catalyst was then reduced in flowing hydrogen at 170"C for 15 minutes prior to any funher experimentation. This procedure was repeated to examine the effect of cycles of use on the catalyst, see Figure 6.17. It can be seen that there is a significant improvement in benzene conversion after the first two regenerations, accompanied by an increase in cyclohexylbenzene selectivity. Funher use and regeneration procedures lead to a levelling of catalyst 226. activity to a value approximating the fresh hydroalkylation catalyst (cf 20. l wt%). The effect of further cycles of use and regeneration on cyclohexylbenz.cne selectivity is severe, with a drop from 92wt% following the second regeneration to 47wt% after the founh, with cyclohexane production increasing. As all regenerations were carried out in-situ, no samples of catalyst were available for characterisation during these cycle studies. It does appear however, that regeneration under the conditions stated leads to an improved catalyst state which favours cyclohexylbenz.ene production and is significantly more active. This may be via redistribution of metal species to a preferred configuration, or by altering of the acid function of the support Subsequent treatments (#3 and #4) result in a loss of selectivity, indicating that the ideal configuration has been removed, almost certainly as a result of the thermal deactivation and sintering. Control of temperature on a local scale during repeated oxidation cycles being almost impossible. 227. 6.3. HYDROALKYLATION OF OTIIBR MODEL COMPOUNDS 6.3.1. Hydroalkylation of Benz.enc/Toluene Mixture To examine the competitive hydroalkylation of aromatic compounds, a 1: 1 wt% mixture of benzene and toluene was passed over catalyst A under the standard reaction conditions. Although the conversion was uncharacteristically low (3.5wt% ), hydroalkylation of both benzene and toluene was achieved with a product selectivity of 20.7wt% cyclohexylbenzene, 42.3wt% methylcyclohexylbenzene and 8.0wt% methylcyclohexyltoluene. Small quantities of cyclohexane and methylcyclohexane were also detected. The product distribution obtained indicates the preferential hydroalkylation of toluene to form methylcyclohexylbenzene and cyclohexylmethylbenzene, in agreement with Yamakazi et al. (105). This is considered to be due to the substituent effect of the toluene methyl group, with its electron donor capacity causing toluene to be more strongly adsorbed on the metal surface than benzene, and can be related to an electron-deficient character of the metal (57,58,59). 228. 6.3.2. Synthetic Coal Derived Liquid Prior to evaluating the hydroalkylation catalyst with a coal derived liquid, preliminary studies utilising a synthetic mixture of model compounds was undertaken. The synthetic mixture comprised a blend of hydrocarbons intended to simulate the composition of a coal derived gasoline product (DR072) supplied by BHP Melbourne Rese.arch Laboratories. The composition of the mixture is given in Table 6.7. Only the eight major components, identified from the true coal derived fraction using gas chromatography - mass spectrometry, were included due to the complexity of the true coal derived liquid. Table 6.7 Composition of Coal Derived Liquid Synthetic Mixture Component wt% Benzene 30.8 Toluene 27.2 Pentane 7.2 Hexane 10.6 Methylcyclopentane 13.5 Heptane 5.2 Octane 1.1 Xylenes 4.4 Using catalyst A (Pt/Ni/RFJzeolite) at 170"C, 3.S:MPa and 16.6 hr- 1, a conversion of 18% (based on the mass of feed) was achieved. The experiment was conducted over a period of 3 hours and no significant catalyst deactivation was 229. observed. The principal products were cyclohexylbenz.ene, methylcyclohexylbenz.ene, methylcyclohexyltoluene and dicyclohcxylbenzene, with selectivities of 25.9, 42.2, 6.24, and 2.5 wt% respectively. With traces of cyclohexanc and methylcyclohexarie~~ resulting from hydrogenation of aromatics in the feed. The product distribution obtained is essentially the same as for the benzene/toluene hydroalkylation experiment discussed above, indicating that the presence of other hydrocarbons typical of a coal derived liquid do not interfere with the hydroalkylation r~action. The conversion of 18% on a feed basis converts to approximately 29% on an aromatic basis, which is comparable to studies on benzene hydroalkylation. The results from this synthetic coal derived liquid hydroalkylation therefore indicate that hydroalkylation is a prospective upgrading route for coal derived liquids. 230. CHAPTER 7 RESULTS AND DISCUSSION COAL DERIVED AND OTHER LIQUIDS 7.1. HYDROALKYLATION OF COAL DERIVED LIQUIDS One of the objectives of this study was to examine hydroalkylation as a possible route for upgrading the largely aromatic coal-derived liquids to specification jet and diesel fuels. Following the very promising results obtained from the model compound simulated "coal-derived" liquid trial discussed in the last chapter, true coal-derived liquids were tested. Of specific interest in this phase of the work was the prospect of catalyst deactivation, which had been noted in the model compound studies but had not been severe. 7 .1.1. Feed Characteristics Two coal derived liquids were provided by B.H.P. Research - Melbourne Laboratories (BHPR-ML) for hydroalkylation studies. The samples were gasolines derived from a Yalloum coal syncrude. The syncrude being the product of a multiple-cycle coal hydroliquefaction trial on the BHPR-ML continuous coal liquefaction reactor (187). The syncrude product from this process was a wide boiling range oil, with a high aromatics and heteroatom content. The syncrude was 231. hydrotreated over a Ni/Mo/A120 3 catalyst under conditions which ensured complete hydrogenation of aromatics and removal of heteroatomic species. After distillation, the hydrotreated oil yielded various fractions, including a heavy naphtha (70-l 70°C). The naphtha was fed to a reformer charged with a Pt/Re/Al 20 3 reforming catalyst, where it was converted to a gasoline of high octane number (RON ea. 100), this being the sample supplied for hydroalkylation studies. Two samples were provided, Yallourn Reformate #1 a blend of RF025/3/l and RF025/4/l, and Yallourn Reformate #2 a blend of RF025/5/l, RF025/6/l, RF025n/1 and RF025n/2.. Analysis of these coal derived liquid feeds was carried out using gas chromatography - mass spectrometry (G.C.-Mass Spee.) at both the University of N.S.W., School of Chemistry, and at BHPR-ML. The principle aromatic components of these coal derived liquid feeds are shown in Table 7 .1. It can be seen that the composition of the coal-derived gasolines are very similar, with 13 compounds (all aromatic) accounting for 80% or more of the sample volume. 7 .1.2. Hydroalkylation Trials Two hydroalkylation trials on coal-derived liquids were performed, using hydroalkylation catalyst A, the 0.1 wt% Pt/5 wt% Ni/10 wt% RE on 13X catalyst. In each case the catalyst was reduced in flowing hydrogen at 17D°C and 3.5MPa for 15 minutes prior to feed introduction. The first trial was on Yallourn Reformate # 1 using standard conditions (170°C, 3.5MPa,LHSV=l6.6hr· 232. 1,hydrogen:feed=approximately 1:1) for a period of 105 minutes, while in the second trial (on Yallourn Refonnate #2) the liquid hourly space velocity was reduced to 8.3hr-1, with a concomitant doubling of the H2:feed ratio to 2: 1. The second trial ran for a period of 205 minutes. Samples were taken at regular intervals during both trials. The analyses of these product samples was carried out at BHPR-ML, and the results obtained are presented in Table 7 .1. Table 7.1 Hydroalkylation of Coal-Derived Liquids Using Catalyst A. Trial #1 Trial #2 Component (l 70"C,3.5MPa.16.6hr 1) (l 70"C,3.5MPa.8.3hr1) Fee.d Prod. Prod. Fee.d Prod. Prod. 15rnin 75rnin 30rnin 200rnin (wt%) (wt%) (wt%) (wt%) (wt%) (wt%) benzene 9.3 8.3 6.8 12.4 8.4 5.7 toluene 20.4 39.7 23.8 19.6 21.4 18.2 ethylbenzene 8.0 7.1 9.2 6.8 8.6 8.4 m,p-xylene 13.3 12.0 15.7 12.2 16.4 15.6 o-xylene 5.1 4.6 5.6 4.8 5.7 6.1 n-propylbenzene 3.5 2.7 3.9 2.8 3.3 3.8 m,p-ethyltoluene 7.1 5.6 8.1 5.9 7.5 8.3 J ,3.5-trimethylbenzene 1.5 1.3 1.7 1.4 2.0 2.0 o-ethyltoluene 1.9 1.4 2.1 1.5 1.8 2.1 J ,2,4-trimethylbenzene 4.5 3.5 5.1 3.8 4.5 5.4 indan 4.6 3.4 4.8 4.0 4.3 5.3 Hydrogen Content 10.37 10.07 10.16 11.00 10.40 10.32 (wt%) 233. 7.2. DISCUSSION OF HYDROALKYLATION OF COAL DERIVED LIQUIDS The results in table 7 .1 indicate that no significant conversion of aromatics to higher molecular weight products has taken place in either hydroalkylation trial. Indeed, the hydrogen content data show a drop from feed to products in both trials. Only traces of cyclohexylbenzene and methylcyclohexylbenzene were found in the earliest liquid product samples, no cyclohexane or methylcyclohexane being detected. These results indicate that significant catalyst deactivation has occurred, possibly as a result of catalyst poisoning from small amounts of sulphur in the coal derived liquid feedstocks, or as a result of more severe catalyst coking than had been found for the model compound studies. At the end of trial #1 the feed was switched to benzene to flush out the reactor and system lines. As no cyclohexylbenzene was evident in the exit liquids at this stage, this supports the assertion that the catalyst had deactivated as a result of contact with the coal-derived liquid feed. Catalyst deactivation due to poisoning by the coal-derived feed is also supported by the rapid decline observed in reaction exotherm upon feed introduction, with increases of 35°C and 5°C for trials #1 and #2 respectively falling to the base catalyst temperatures within 5 minutes on-line. 234. Despite the inactivity/deactivation, there are clearly some compositional differences between the feedstock and products from both trials. There is no obvious explanation for the observed increase in toluene concentration in trial #1, particularly as the concentration ~turns to a level close to that of the feedstock for the final product (75 mins.). This effect is not reproduced in trial #2. Benzene concentrations show a decline in both trials, though no significant production of cyclohexylbenzene was observed. As benzene is the most volatile component in the samples, it is possible that some evaporative losses occurred over the period between the trials and sample analysis. It is clear that the results obtained from the coal-derived liquid hydroalkylation trials are not consistent with the extensive data collected on model systems with the same catalyst and under similar reaction conditions. A more definitive explanation of the observed effects requires further experimentation, with particular attention to the nature and extent of catalyst deactivation. 235. 7.3. HYDROALKYLATION OF DIESEL FRACTION Given the likelihood of the introduction of stringent requirements regarding the aromatics content of both gasoline and distillate, it was of interest in this study to assess the suitability of hydroalkylation as a means of catalytically treating diesel fuel. Refiners are already feeling the pressure of legislation in the USA to reduce diesel aromatics cont~nt, and are faced with high capital costs to upgrade current refining processes to achieve the lower limits demanded (15,17). Hydrotreaters, for example, will have to operate at higher pressures in order to hydrogenate the aromatic components of distillate. Hydroalkylation offers an interesting alternative, although it must be appreciated that hydrotreannent of the product will still be required after the hydroalkylation step. 7 .3.1. Feed Characteristics A sample of lfght recycle oil was supplied by Shell and was examined using gas chromatography-mass spectrometry to identify the major components. As the mass spectral analysis indicated a predominance of straight-chain alkanes in the range C13 - C21, the sample was doped with naphthalene (2.3 wt%) which served both to increase the aromatic content, as well as to act as an internal aromatic marker, readily distinguishable in the product gas chromatographs. 236. 7.3.2. Hydroalkylation Trial Hydroalkylation of the petroleum diesel sample was undertaken using Catalyst A (Pt/Ni/RFJl 3X), which had been reduced with flowing hydrogen at l 7f1'C ,3.SMPa for IS minutes prior to feed introduction. The standard reactor conditions were chosen for the trial: l 7f1'C, 3.SMPa, liquid feed rate=0.8ml.min·1, hydrogen feed rate=210ml.min·1• Reaction products from the trial were collected at frequent intervals and analysed by gas chromatography. Unlike the trials employing model compounds (and to a lesser extent coal liquids), it was not possible in the case of the diesel hydroalkylation study to trace the fate of specific chemical components in the fuel, other than perhaps naphthalene. Bulle properties of the feedstock and products were monitored in order to ascertain whether any significant hydroalkylation had taken place. Hydrogen content was determined for all samples, and analysis by 13C Nuclear Magnetic Resonance Spectroscopy to give carbon aromaticity and simulated distillations were carried out on selected samples. These measurements were all carried out at BHPR-1\11.. The results obtained are given in Table 7 .2. 237. Table 7.2 Diesel Hydroalkylation Trial Using Catalyst A. Hydrogen Carbon Simulated Distillation (wt%) Sample Content Aromaticity (wt%) (%) <196° 196- 235- >317°C C 235°c 311°c Diesel Fuel 13.38 16.2 2.5 5.0 60.0 32.5 Diesel+2.3% Naphthalene 13.28 19.0 3.6 7.4 59.2 29.7 Product 1 13.26 Product 2 1.3.14 Product 3 13.21 Product 4 12.98 18.4 2.5 6.2 59.0 32.3 Product 5 13.06 Product 6 13.01 Product 7 13.12 Product 8 13.13 2.3 6.0 58.5 33.2 238. 7 .4. DISCUSSION OF HYDROALKYLATION OF DIESEL FRACTION The hydroalkylation reaction is a hydrogen-consuming process, resulting in products with a higher hydrogen concentration than the feed (c.f. cyclohexylbell7.Cne and benzene). The data in Table 7.2 indicate, however, that the hydrogen content of the products is lower (12.98-13.21 wt%) than the doped feedstock (13.28 wt%). The same behaviour that was noted for the coal-derived liquid trials discussed previously. 13C NMR data, on the other hand, indicate a slight decrease in aromatic carbon content, which would be expected upon hydroalkylation of aromatic species. The decrease is, however, very close to the precision of the instrument, and not conclusive evidence for hydroalkylation. Hydroalk:ylation would also be expected to be accompanied by a shift in the boiling range distribution of the products when compared to that of the products, this shift being to a higher boiling range. The simulated distillation data in Table 7 .2 indicate a marginal shift in boiling distribution to higher boiling point although, as with the NMR data, the changes are considered too small to provide supporting evidence for hydroalkylation. 239. Thus the data from this trial are considered inconclusive with regard to whether hydroalkylation of aromatic species in the diesel fuel has occurred. A number of questions arise, however, and warrant further attention. If no hydroalkylation has taken place, can this be attributed to the small, but significant, sulphur content of the diesel fuel? Are the catalyst and process conditions, which are considered optimum for benzene hydroalkylation, optimum for the diesel fuel? Is the catalyst deactivating in the presence of diesel? Regrettably, these questions could not be answered in detail without detailed analysis and specific doping of diesel fuels. 240. CHAPTER 8 RESULTS AND DISCUSSION HYDROOENATION AND COMBINED HYDROALKYLATIONIHYDROGENATION Hydroalkylation is the first step in the process of converting aromatics to distillate fuel. Hydrogenation to cyclic aliphatics (naphthenes) is also necessary before a distillate fuel of acceptable quality can be produced (22). One of the minor objectives of this research was to investigate the hydrogenation of products from hydroalkylation trials. This has been achieved in two phases. In the first phase, hydrogenation trials were performed using biphenyl, as insufficient cyclohexylbenzene had been collected from hydroalkylation experiments and the high cost prohibited purchasing the necessary quantities. In these preliminary studies, experiments using a dual catalyst bed (hydroalkylation/hydrogenation) were carried out so that the direct hydrogenation of hydroalkylation products could be investigated. In the second phase of the hydrogenation studies, samples of cyclohexylbenzene provided by BHP and produced in the hydroalkylation reaction were hydrogenated. 241. 8.1. HYDROOENATION OF BIPHENYL Biphenyl was. chosen for the initial investigation of the hydrogenation reaction as it is structurally equivalent to cyclohexylben:zene and was considered a more difficult compound to hydrogenate. A 3 wt% biphenyl in cyclohexanc solution was prepared as the liquid feed. The catalyst chosen for these studies was 10 wt% Ni on y-Al20 3 (particle siz.e 300-SOO;m), as nickel is known to be an effective hydrogenation catalyst for aromatic compounds. The catalyst (2 grams) was reduced in flowing hydrogen at 500°C for 3 hours at atmospheric pressure prior to the introduction of the feed. The reaction conditions used are given below: Reaction temperature = 138°C Reaction pressure = 3.5 MPa LHSV = 6 hr"1 Liquid flowrate = 0.4 ml.min·1 1 H2 flowrate = 240-440 ml.min· Under the conditions employed, the hydrogenation of biphenyl was successful. Conversion of biphenyl was virtually quantitative (100 wt% and 99.8 wt%), with bicyclohexyl selectivities of 27 wt% and 38 wt% found for the two trials carried out with hydrogen flows of 240 ml.min·1 and 440 ml.min·1 respectively. In each case cyclohexylbenzene was found to be the major product. 242. 8.2. DUAL CATALYST BED - Combined Hydroalkylation/Hydrogenation As a result of the preliminary trial using biphenyl, an attempt was made to process benzene directly into biphenyl, using a dual catalyst bed. Catalyst E (Pt/RE- 13X) was chosen as the hydroalkylation catalyst as it gave good yields of cyclohexylbenzene w~en prereduced under the harsh condition required for reduction of the hydrogenation catalyst One gram of each catalyst was placed in the reactor, separated by glass wool. Glass beads were placed above and below the catalyst bed. The catalysts were then reduced at 500°C for 3 hours prior to benzene introduction. Conversion of benzene at 200°C, 3.5 MPa and LHSV=7 hr·1 was extremely low (1.1 wt%), with only trace amounts of bicyclohexyl and cyclohexylbenzene detected (selectivities being 23.1 and 76.9 wt% respectively) in the product stream. Subsequent hydroalkylation studies on catalyst E have revealed that following high temperature reductions, such as was carried out in the dual bed system, rapid catalyst deactivation is observed. This is considered to be due to rapid coking of the large metal agglomerates formed due to sintering of metal sites at these high temperatures (see section 6.2.2). As all hydroalkylation catalysts prereduced under high temperatures were found to deactivate rapidly (see section 6.1.4), the dual catalyst bed approach to hydroalkylation/hydrogenation was not pursued, as reduction conditions for the two types of catalysts were incompatible. 243. 8.3. HYDROGENATION OF CYCLOHEXYLBENZENE The 10 wt% Ni/'f-Al20 3 was again used in hydrogenation trials with cyclohexylbenune. The feed used was a distilled (to remove benune and cyclohexane) product from BHPR-ML hydroalkylation trials using a catalyst equivalent to catalyst A. Small quantities (ea. 20%) of di- and tri cyclohexylbenzenes were also present in this higher boiling sample. Given the volume of material required for a long-term trial and the limited sample available, the cyclohexylbenzene/higher molecular weight feed sample was diluted (1: 1 volume basis) with cyclohexane prior to processing. The catalyst (2 grams) was reduced at 500°C for 3 hrs at atmospheric pressure in flowing hydrogen and then cooled to 13Q°C prior to the introduction of the liquid feed. Reaction conditions used were 137°C, 3.5 MPa and LHSV=6 hr1• Conversion of cyclohexylbenzene was found to be very low (5 wt%), with bicyclohexyl being the only product detected. This may have been the result of catalyst deactivation by polyaromatic compounds in the feed, although subsequent hydrogenation studies on cyclohexylbenzene in a stirred autoclave at BHPR-ML indicate that the severity of the process conditions (170"C, 3.5 MPa, LHSV=6 hr1) may have been too low to ensure complete conversion to bicyclohexyl. To show the operability of the hydrogenation of hydroalkylation products the results from studies at BHPR-ML are reproduced below (188), with the permission of the authors. 244. 8.4. HYDROOENA TION TRIALS AT BHPR-ML In any commercial application of hydroalkylation technology, distillation after the hydroalkylation stage would be used to generate material for hydrogenation (distillate production), as well as to generate a stream for recycling back to the hydroalkylation reactor. Thus, prior to the hydrogenation trials, products from selected hydroalkylation trials were blended and distilled to remove cyclohexane and benzene. The resultant two samples comprised cyclohexylbenzene (ea. 70 wt%), bicyclohexyl (ea. 10 wt%) and di- and tri-cyclohexylbenzenes (ea. 20 wt%). Hydrotreatment was expected to yield a product consisting mainly of bicyclohexyl, with smaller amounts of higher molecular weight saturated, cyclic species. The above feedstocks were hydrotreated in a stirred autoclave using a commercial Ni/Al 20 3 catalyst (Mallinckrodt E235TR, 43.1 % Ni). The catalyst was reduced under flowing hydrogen for 18 hours at 300°C and 0.5 MPa. For the hydrogenation experiment, the reactor temperature was decreased to 1300C and the liquid (ea. 500 ml) introduced into the reaction vessel. The pressure was increased to 3.5 MPa, while the hydrogen flow was maintained at 100 ml.min·1, and introduced into the base of the reactor. The hydrogenation continued for 72 hours, when analysis of the liquid product revealed complete conversion of cyclohexylbenzene to bicyclohexyl. Given the slow rate of hydrogen delivery to the reactor, the duration of this trial was not surprising. On an industrial scale however, such a reaction would 245. be carried out in a conventional (fixed-bed) hydrotreater, where throughput would be significantly increased. Products from the two hydrogenation trials performed at BHPR-ML were analysed by gas chromatography (see Table 8.1). They contained predominantly bicyclohexyl, although significant amounts of higher molecular weight components were also detected. These are presumably isomers of tricyclohexyl, and boil in the range (nominal) 300-340°C. Irrespective of their exact composition, these high molecular weight components still report to the diesel boiling range and, as such, are valued products. The same can be said of the lower range component (nom. 210- 2200C), which may be fully hydrogenated methylcyclopentylcyclohexane. Table 8.1 Composition of Products from BHPR-ML Hydrogenation Trials. Hydrotreated Product Analysis Feedstock Source (wt%) HA003-8 HA013-14 unidentified (nom.range 200-220"C) 5.2 3.3 bicyclohexyl 73.7 63.9 cyclohexylbenzene 0.3 0.1 unidentified (nom.range 300-340"C) 20.8 32.7 Cetane Number 51 54 246. Table 8.1 also gives Cetane numbers for the hydrotreated products. These assignments were carried out at BHPR-ML using a small diesel test throttling engine. The test method used is based on IP41/60 (189), and relies on the determination of misfire pressure in the engine after throttling. This pressure correlates closely with cetane number and thus, after generation of a calibration curve (using secondary diesel standards), the cetane number of an unknown can be determined with a reasonable degree of accuracy ( 2 CN) ( 190). Cetane numbers measured for the hydrotreated products were 51 and 54 respectively (cf. Australian Standard of 45), confirming their potential as distillate fuel blendstocks. 247. CHAPTER 9. CONCLUSIONS AND RECOMMENDATIONS 9.1. CONCLUSIONS The hyclroalkylation of benz.ene over a nickel and rare earth treated zeolite supported catalyst impregnated with platinum produces cyclohexylbenz.ene as the major product with significant yields of the minor products, cyclohexane and high molecular weight di- and tri-cyclohexylbenzene isomers. Production of traces of isomerisation products derived from benz.ene and cyclohexylbenzene indicates the complexity of the reaction scheme taking place. The reaction proceeds over the catalyst at moderate temperatures (146-195°C) and pressures (2.8-4.5 MPa), with good selectivity to cyclohexylbenzene (60.5-88.4 wt%) achieved at reasonable conversions (9.2-50.7 wt%). Increasing reaction temperature has little effect on the yields of cyclohexylbenzene obtained. Lowering of the benzene liquid hourly space velocity and hydrogen partial pressure were found to increase production significantly. Moderate catalyst deactivation was observed under all test conditions with this catalyst and, as a consequence, no kinetic modelling of the hydroalkylation reaction could be carried out. 248. Catalyst performance was found to be critically dependent on the presence of two metals (i.e., platinum, nickel) and the rare earth salts, with low catalyst activity and production of cyclohexane as the major product found with metal deficient catalysts. Increased metal loadings were found to improve cyclohexylbenzene yields. The role of catalyst pretreatment was found to be crucial, especially for modified hydroalkylation catalysts deficient in one or more metals. Examination of these catalysts under increasingly high prereduction temperatures (up to 475°C) resulted in benzene conversion comparable or exceeding that of the "complete" multi-metallic hydroalkylation catalyst reduced at 170°C. Cyclohexylbenzene selectivity was also found to increase. In all cases of high temperature reduction, catalyst deactivation was found to be rapid. The observed effects of prereduction temperature were correlated with catalyst temperature reduction profiles. This behaviour was interpreted in terms of the roles, interactions and positions of the various metal species present within the zeolite framework. It is considered that nickel is the primary active metal species, with the presence of platinum in a highly dispersed state assisting in low temperature nickel reduction via a dissociative hydrogen spill-over mechanism. The role of the rare earth metal salt of cerium, appears to be three-fold. Firstly, the salt assists in the low temperature reduction of nickel by occupying S1 sites in the zeolite framework, and forcing Ni2+ into the more easily reduced S11 positions. Secondly, ceria acts as a "chemical anchor" for nickel atoms thereby retarding their mobility and possible sintering. 249. Thirdly, ceria promotes dimerisation via electron density enrichment of the nickel particles, resulting in reaction intermediates (probably cyclohexene) being less strongly adsorbed on the metal surface. The intermediates are then able to migrate to an acidic site of the zeolite support for alkylation. The function of the catalyst support is also important. Its acidity was found to be directly responsible for cyclohexylbenzene selectivity, via alkylation. Zeolite geometry was also found to effect the catalyst performance, with catalyst deactivation occurring for small aperture sizes. Production of cyclohexylbenzene was found to be promoted by a fine balance between the metal and acid functions of the catalyst, with the hydroalkylation catalyst operating via an extremely complicated combination of metal and support effects. This coupled with the complexity of the hydroalkylation network (involving numerous possible side reactions and reaction sequences) made the system very difficult to study. Catalyst treatments were found to have variable effects on benzene conversion and cyclohexylbenzene selectivity. Addition of carbon tetrachloride to the benzene feed was found to result in a significant improvement in activity and selectivity, probably as a result of redispersion of the active metal species and increases in zeolite acidity. Regeneration of the spent catalyst after cycles of use was found to improve catalyst performance (for the first and second regeneration) and to 250. decrease performance for subsequent regeneration. This was considered due to the modification of the catalyst metal distributions or acidity to some optimum configuration, which, upon further regeneration, was lost. Hydroalkylation of a benzene/toluene mixture produced cyclohexylbenzene, methylcyclohexylbenzene and methylcyclohexyltoluene, with traces of cyclohexane and methylcyclohexane. The relative yields of products indicated preferential hydroalkylation of toluene. This was considered to result from preferential adsorption on the metal as a result of the electron donor capacity of toluene. A mixture of model compounds designed to simulate a coal derived gasoline product was also successfully hydroalkylated. Hydroalkylation of coal derived liquids over the optimised hydroalkylation catalyst was unsuccessful under the conditions used, with no conversion of aromatic compounds in these liquids to higher molecular weight products and no increase in the hydrogen content of the product stream observed. It is considered that rapid catalyst deactivation has occurred, due to poisoning of the catalyst by sulphur compounds present in the coal derived liquids. 251. Hydroalkylation of a petroleum derived diesel fraction with the same catalyst was also unsuccessful, with no increase in the hydrogen content of the product stream observed. Again catalyst deactivation due to the presence of sulphur may have occurred. Hydrogenation of the hydroalkylation reaction products was also carried out. Initially, due to the high cost and inavailability of sufficient cyclohexylbenzene, the hydrogenation of biphenyl was performed (biphenyl being a more difficult compound to hydrogenate and structurally similar to cyclohexylbenzene). This was carried out over a Ni/y-Al20 3 catalyst at 138°C and 3.5 MPa with cyclohexylbenzene and bicyclohexyl (approximate yields 70wt% and 30wt% respectively) produced. Attempts to hydrogenate products from the hydroalkylation reaction with the same catalyst and under the same conditions resulted in very low conversions. Subsequent studies at BHPR-ML (where complete conversion to bicyclohexyl was achieved) have revealed the above stated hydrogenation conditions were not of the required severity. Cetane number determinations at BHPR-ML on the hydrogenated hydroalkylation product confirms their potential as distillate fuel blendstocks. The combined hydroalkylation and hydrogenation of aromatic compounds, such as those present in coal derived liquids and petroleum derived fuels, is therefore considered a possible route for distillate range fuel blendstocks. 252. 9.2. RECOMMENDATIONS The results obtained in this work suggest several interesting areas for future work, including: (i) Improvement of hydroalkylation catalyst stability. Catalysts tend to deactivate with time and the cause and minimisation of deactivation is required. (ii) Investigation of catalyst resistance to sulphur poisoning is also necessary, especially in view of the possible application of the catalyst in upgrading petroleum derived and coal derived fuels. (iii) Further investigation of toluene and xylene hydroalkylation, in order to investigate the-possible use of the reactions with BTX streams. (iv) Studies of reaction kinetics and reaction modelling. These could not be performed due to catalyst deactivation. (v) Physical examination of the catalyst surface (both external and internal) is required to confirm the proposed location of metals. (vi) Investigation of mass and heat transfer, given that this a gas-liquid-solid reaction. 253. (vii) Investigation of different reactor types and configurations would be useful in examination of reaction kinetics and catalyst stability, especially in view of mass and heat transfer problems. The use of recycle could also be addressed. (viii) Given the emerging restrictions of aromatics in petroleum derived gasoline, testing of the reactions with this feedstock would be timely. 254. REFERENCES 1. Anon, "Oil and Australia 1989", Petroleum Gazette, 2, 11-22 (1989). 2. Anon, "Worsening Supply/Demand Gap During the Next Decade", Petroleum Gazette, 3, 2-16 (1989). 3. Anon, "Downstream industry facing numerous and far-reaching changes", Petroleum Gazette, 1, 10-14 (1991). 4. Anon, "Synthetic Liquid Fuels", OECDIIEA (International Energy Agency), Paris, 1985. 5. Anon, "Energy 2000 - Synfuels", Dept. Resources and Energy, Discussion Paper No. 7 (1986). 6. Nowaki, P., Chem. Tech. Rev., 131, 5 (1979). 7. Matar, S., "Synfuels", Pennwell, Oklanoma (1982). 8. Wu, W.R.K., Storch, H.H., Bu. Mines Bull., 633, 195 (1968). 9. Azcel, T., Williams, R.B., DOE!MERC-8007-1 (Pan 1), 237pp (1976). 10. Takeya, G., Pure and Appl. Chem., SO, 1099-1115 (1978). 11. Corbett, R.A., Oil Gas J., June 18, 33-57 (1990). 12. Unzelman, G.H., Oil Gas J., April 9, 43-48 (1990). 13. Unzelman, G.H., Oil Gas J., April 23, 91-93 (1990). 14. Unzelman, G.H., Oil Gas J., April 15, 44-48 (1991). 15. Yoes, J.R., Asim, M.Y., Oil Gas J., May 11, 54-58 (1987). 16. Suchanek, A.J., Oil Gas J., May 7, 109-119 (1990). 17. Nash, R.M., Oil Gas J., May 29, 47-56 (1989). 18. Palmer, L.D., EOGR No. 625, NERDDPEG/871625, Sept.,(1986). 255. 19. Hardenburg, H.O., Sth. African Mech. Eng., 30 (1980). 20. Sefer, N.R., Moses, C.A., SAE Tech. Paper Series, Oct. (1981). 21. Sullivan, R.F.~ O'Rear, D.J., DOE Contract No. AC22-76ET 10532 Sept. (1981). 22. White, N., US Patent 4,645,585 (1987). 23. Sabatier, P., "Catalysis in Organic Chemistry", D. Van Nostrand Co., Inc., New York (1922). 24. Nikolajenko, V., Bosacek, V., Danes, V.L., J. Cata/., 2, 127 (1963). 25. Aben, P.C., Platteuw, J.C., Stouthamer, B., Ree. Trav. Chim., 89, 499 (1970). 26. Coenen, J.W., Van Meerten, R.Z.C., Rijnten, H.Y., "Proc. 5th lnt. Congr. Catalysis", Elseveir, Palm Beach, 1, 45 (1973). 27. Shelwood, P.W., Adler, S., Phillips, T.R., J. Amer. Chem. Soc., 76, 2231 (1954). 28. Martin, G.A., Dalmon, J.A., J. Cata/., 75, 233 (1982). 29. Dixon, G.M., Singh, R., Trans. Faraday Soc., 65, 1128 (1969). 30. Taylor, W.F., Staffin, H.R., Trans. Faraday Soc., 63, 2309 (1967). 31. Van Meerten, R.Z.C., Coenen, J.W.E., J. Cata/., 46, 1 (1977). 32. Van Meerten, R.Z.C., Coenen, J.W.E., J. Cata/., 46, 13 (1977). 33. Franco, H.A., Phillips, M.J., J. Cata/., 63, 346 (1980). 34. Kehoe, J.P.G., Butt, J.B., J. Appl. Chem. Biotechnol. 22, 23 (1972). 35. Mirodatos, C., J. Phys. Chem., 90, 481 (1986). 36. Mirodatos, C., Dalmon, J.A., Martin, G.A., J. Cata/., 105, 405 (1987). 256. 37. Bassett, J.M., Dalmai-Imelik, G., Primet, M., Murin, R., J. Cata/., 37, 22 (1975). 38. Sander, K.M., J. Cata/., 20, 106 (1971). 39. Vannice, M.A., Neikam, W.C., J. Cata/., 23, 401 (1971). 40. Teicher, S.J., Maz.abrand, A.R., Pajonk, G., Gardes, G.E.E., Hoang-Van, C., J. Colloid Interface Sci., 58, 88 (1977). 41. Antonucci, P., Van Truong, N., Giordano, N., Maggiore, R., J. Cata/., 75, 140 (1982). 42. Ceckiewicz, S., Delmon, B., J. Cata/., 108, 294 (1987). 43. Figueras, F., Gomez, R., Primet, M., Adv. Chem. Ser., 121, 480 (1973). 44. Primet, M., Garbowski, E., Mathieu, M.V., Imelik, B., J. Chem. Soc. Faraday I, 76, 1953 (1980). 45. Vannice, M.A., Chou, P., "Proc. 8th Intl. Cong. Cata/." Dechema, Berlin, S, 99 (1984). 46. Chou, P., Vannice, M.A., J. Cata/., 107, 129 (1987). 47. Chou, P., Vannice, M.A., J. Cata/., 107, 140 (1987). 48. Peri, J.B., J. Phys. Chem., 70, 1482, 2970, 3168 (1966). 49. Busca, G., Saussey, H., Saur, 0., Lavalle9, J.C., Lorenzelli, V., Appl. Cata/., 14, 245 (1985). 50. Chapman, I.D., Hair, M.L., "Proc. 3rd Intl. Cong. Cata/.", Elsevier, Amsterdam, 2, 1091 (1964). 51. Siriwardane, R.V., Wightman, J.P., J. Colloid Interface Sci., 94, 502 (1983). 257. 52. Marczewski, M., Derewinski, M., Malinowski, S., Canad. J. Chem. Eng., 61, 93 (1983). 53. Marques da Cruz, G., Djega-Mariadassou, G., Bugli, G., Appl. Cata/., 17, 205 (1985). 54. Tiep, L.V., Bureau-Tardy, M., Bugli, G., Djega-Mariadassou, G., Che, M., Bond, G.C., J. Cata/., 99, 449 (1986). 55. Levenspiel, 0., "Chemical Reaction Engineering", Wiley & Sons, New York, (1972). 56. Boudart, M., McConica, C.M., J. Cata/., 117, 33 (1989). 57. Tri, T.M., Massardier, J., Gallezot, P., Imelik, B., "Metal-Support and Metal Additive Effects in Catalysis", Elsevier, Stud. Surf Sci. Cata/., 11, 141 (1982). 58. Szymanksi, R., Charcosset, H., Gallezot, P., Massardier, J., Tournayan, L., J. Cata/., 97, 366 (1986). 59. Minot, C., Gallezot, P., J. Cata/., 123, 341 (1990). 60. Smith, H.A., "Catalysis" (Emmett, P.H., ed.), Reinhold, New York, 5, 174 (1957). 61. Bond, G.C., "Catalysis by Metals", Academic Press, London (1962). 62. Smith, R.A., Rader, C.P., Proc. 2nd Intl. Cong. Cata/., Elsevier, Paris, 1, 1213 (1961). 63. Sapre, A.V., Gates, B.C., Ind. Eng. Chem. Process Des. Dev., 20, 68 (1981). 64. Sapre, A.V., Gates, B.C., Ind. Eng. Chem. Process Des. Dev., 21, 86 (1982). 65. Aubert, C., Durand, R., J. Cata/., 112, 12 (1988). 258. 66. Wilson, M.F., Fischer, I.P., J. Catal., 95, 155 (1985). 67. Emmett P.H., "Catalysis" Vol. VI, Reinhold, New York (1958). 68. Pines, H., "The Chemistry of Catalytic Hydrocarbon Conversions", Academic Press, New York (1981). 69. O'Kelly, A.A., Kellet, I., Ind. Eng. Chem., 39, 154 (1947). 70. Dolgov, B.N., Cherkasov, A.S., J. Gen. Chem. USSR., (English Transl.) 24, 825 (1954). 71. Cullinane, N.M., Chard, S.I., J. Am. Chem. Soc., 804 (1948). 72. Ipatieff, V.N., Orlov, M., Petrov, A., Zh. Russ. Fiz. Khi. Obshch., 59, 45 (1927). 73. Kutz, W.M., Corson, B.B., J. Am. Chem. Soc., 68, 1477 (1946). 74. Patinkin, S.H.; Friedman, B.S., in "Catalysis" (Emmett, P.H., ed.), ~ei,nhold, New York, 2, 31 (1958). 75. Schriesheim, A., in "Catalysis" (Emmett, P.H., ed.), . Qei·nhold, New York, 2, 524 (1958). 76. Venuto, P.B., Hamilton, P.S., J. Catal., 13, 385 (1969). 77. Yashima, H., Amhad, H., Yamasaki, K., Katsuta, M., Hara, N., J. Catal., 16, 273 (1970). 78. Chen, N.Y., Kaeding, F.G., Dwyer, F.G., J. Am. Chem. Soc., 101, 6784 (1979). 79. Kaeding, F.G., Chu, C., Young, L.B., Weinstein, B., Butter, S.A., J. Catal., 67, 159 (1981). 259. 80. Dwyer, F.G., "Catalysis of Organic Reactions", Marcel Dekker, New York (1981). 81. Venuto, P.B., Hamilton, L.A., Ind. Eng. Chem. Prod. Res. Dev., 6, 190 (1967). 82. Jacobs, P.A., "Carboniogenic Activity of 2-eolites" , Elsevier, New York (1977). 83. Breck, D.W., J. Chem. Educ., 48, 678 (1964). 84. Breck, D.W., "2.eolite Molecular Sieves, Structure, Chemistry and Use", Wiley, New York (1974). 85. Olson, D.H., Kokotailo, G.T., Charnell, J.F., J. Colloid Interface Sci., 28, 305 (1968). 86. Smith, J.V., Adv. Chem. Ser., 101, 171 (1971). 87. Tanabe, K., Misono, M., Ono, Y., Hattori, H., "New Solid Acids and Bases", Stud. Surf Sci. Cata/., 51 Kodansha, Tokyo (1989). 88. Ward, J.W., J.·Phys. Chem., 73, 2086 (1969). 89. Bolton, A.P., Lanewala, M.A., J. Cata/., 18, 154 (1970). 90. Ward, J.W., J. Cata/., 9, 225 (1967). 91. Ward, J.W., J. Cata/., 9, 396 (1967); 16, 386 (1970). 92. Breck, C.W., Skeels, G.W., Proc. 6th Intern. Congr. Catalysis, The Chemical Society, London, 645, (1977). 93. Frank, W.H., Skeels, G.W., Proc. 5th Intern. Congr. Zeolites, Heydon, Naples, 334, (1980). 94. Jacobs, P.A., Beyer, H.K., J. Phys. Chem., 83, 1174 (1979). 260. 95. Eberly, P.E., Kimberlin, C.N., Ind. Eng. Chem. Prod. Res. Dev., 9, 335 (1970). 96. Koradia, P.B., K.iovsky, J.R., Asirn, M.Y., J. Catal., 66, 290 (1980). 97. Kerr, G.T., Adv. Chem. Ser., 121, 219 (1979). 98. Truffault, R., Bull. Soc. Chim., 1, (5), 391 (1934). 99. Stamicarbon, N.V., French Patent 1,326,722 (1963). 100. Logemann, J.D., US Patent 6,153,678 (1964). 101. Louvar, J.J., Francoy, A., US Patent 3,317,611 (1967). 102. Slaugh, L.H., Leonard, J.A., J. Cata/., 13, 385 (1969). 103. Slaugh, L.H., Leonard, J.A., US Patent 3,412,165 (1968). 104. Louvar, J.J., Francoy, A., J. Cata/., 16, 62 (1970). 105. Yamazaki, Y., Masuda, A., Bull. of lap. Pet. Chem. Ins., 18, 25 (1976). 106. Kamiyama, T., Enomoto, S., Inoue, M., Chem. Lett. Chem Soc. Jpn., 261 (1979). 107. Suggitt, R.M., Arkell, A., Crone, J.M., US Patent 3,760,018 (1973). 108. Crone, J.M., Arkell, A., US Patent 3,760,019 (1973). 109. Suggitt, R.M., Crone, J.M., Arkell, A., US Patent 3,839,477 (1974). 110. Suggitt, R.M., Crone, J.M., Fishkill, A.A., US Patent 3,784,618 (1974). 111. Suggitt, R.M., Crone, J.M., Fishkill, A.A., Arkell, A., US Patent 3,784,617 (1974). 112. Suggitt, R.M., Crone, J.M., US Patent 3,864,421 (1975). 113. Arkell, A., Suggitt, R.M., Fishkill, A.A., US Patent 3,760,017 (1973). 114. Suggitt, R.M., Crone, J.M., US Patent 3,869,523 (1975). 261. 115. Arkell, A., Crone, J.M., US Patent 3,957,687 (1976). 116. Suggitt, R.M., ·Crone, J.M., US Patent 3,926,842 (1975). 117. Murtha, T.P., Zuech, E.A., US Patent 4,094,918 (1978) 118. Murtha, T.P., Zuech, E.A., Jones, W.A., US Patent 4,093,671 (1978) 119. Murtha, T.P., Jones, W.A., Zuech, E.A., Johnson, M.M., US Patent 4,122,125 (1978). 120. Murtha, T.P., Zuech, E.A., US Patent 4,118,434 (1978). 121. Murtha, T.P., US Patent 4,152,362 (1979). 122. Murtha, T.P., US Patent 4,219,689 (1980). 123. Murtha, T.P., Zuech, E.A., US Patent 4,268,699 (1981). 124. Murtha, T.P., Zuech, E.A., US Patent 4,329,531 (1982). 125. Murtha, T.P., Zuech, E.A., US Patent 4,094,920 (1978) 126. Thomas, D.L., "Studies of the Hyclrogenolysis of Dimethylsuccinate and Diethyloxalate Over Copper Based Catalysts", Ph.D. University of New South Wales (1990). 127. Nelson, F.M., Eggertsen F.T., Anal. Chem., 30, 1387 (1958). 128. K. Tanabe, "Solid Acids and Bases", Kodansha, Tokyo (1970). 129. K. Tanabe, "Catalysis: Science and Technology", Springer-Verlag, Berlin, Vol. 2, Chapt. 5 (1981). 130. Sato, H., Ishii, N., Hirose, S., Nakamura, S., "Proc. 7th Intern. Zeolite Con/", p.755, Kodansha, Tokyo (1986). 131. Wang, G., Itoh, H., Hattori, H., Tanabe, K., J. Chem. Soc., Faraday Trans. 1, 79, 1373 (1983). 262. 132. Hashimoto, K., Masuda, T., Mori, T., "Proc. 7th Intern. Zeolite Conf ", p.503, Kcxlansha, Tokyo (1986). 133. Aboul-Gheit, A.K., Al-Haijaji, M.A., Summan, A.M., Thermochimica Acta, 118, 9 (1987). 134. Aboul-Gheit, A.K., Al-Haijaji, M.A., Summan, A.M., Abdel-Hamid, S.M., Thermochimica Acta, 126, 397 (1988). 135. Aboul-Ghiet, A.K., Al-Haijaji, M.A., Analytical Letters, 20, (4), 553 (1987). 136. Auroux, A., Bolis, V., Wierzchowski, P., Gravelle, P.C., Vedrine, J.C., J. Chem. Soc. Faraday Trans. I, 75, 2544 (1979). 137. Peri, J.B., J. Phys. Chem., 69, 220 (1965). 138. Peri, J.B., J. Phys. Chem., 69, 211 (1965). 139. Kutsov, L.M., Borovkov, V. Yu., Kazansky, V.B., J. Cata/., 72, 149 (1981). 140. Kazansky, V.B., Borovkov, V. Yu., Kutsov, L.M., "Proc. 8th Intern. Congr. Catal.", Verlag Chemie, Weinheim, 3, 3 (1984). 141. Charlesworth, J.M., Fuel Processing Technol., 16, 99 (1987). 142. Hayes, P.C., Pitzer, E.W., J. Chromatography, 253, 179 (1982). 143. Gerasimenko, V.A., Kirilenko, A.V., Nabivach, V.M., J. Chromatography, 208, 9 (1981). 144. Matisova, E., Krupcik, J., Cellar, P., Garaj, J., J. Chromatography, 303, 151 (1984). 145. Bertsch, W., Anderson, E., Holzer, G., J. Chromatography, 126, 213 (1976). 146. Deans, D.R., J. Chromatography, 203, 19 (1981). 147. Bertsch, W., J. High Resol. Chrom. & Chrom. Commun., Aug. 85 (1978). 263. 148. Gordon, B.M., Rix, C.R., Borgerding, M.F., J. Chromat. Sci., 23, 1 (1985). 149. Jennings, W., J. Chromat. Sci., 22, 129 (1984). 150. Miller, R.J., Stearns, S.D., Freeman, R.R., J. High Resol. Chrom. & Chrom. Commun., Feb., 55 (1979). 151. Beyerlein, R.A., McVicker, G.B., Yacullo, L.N., Ziemiak, J.J., A.C.S. Div. Petr. Chem., New York, 31, 190 (1986). 152. Beyerlein, R.A., McVicker, G.B., Yacullo, L.N., Ziemiak, J.J., J. Phys. Chem., 92, 1967 (1988). 153. Sinfelt, J.H., "Catalysis - Science and Technology", Springer-Verlag, New York, pp 258-300 (1981). 154. McPherson, L.J., Olive, M.F., "Modem Petroleum Technology", Springer Verlag, New York, pp 395-478 (1984). 155. Jacobs, P.A., Stud. Surf Sci. Cata/. (Met. Clusters Catal.), 29, 357 (1986). 156. Mortier, W.J., "Compilation of Extra Framework Sites in 2.eolites", Butterworth, New York (1982). 157. Gallezot, P., Alcaron-Diaz, A., Dalmon, J.A., Renouprez, A.J., Imelik, B., J. Catal., 39, 334 (1975). 158. Gallezot, P., Catal. Rev.-Sci. Eng., 20, 121 (1979). 159. Gallezot, P., Bergeret, G., "Catalyst Deactivation", Chemical Industries, 30, 263 (1987). 160. Dalla Betta, R.A., Boudart, M., "Catalysis", Elsevier, Amsterdam, 1, 1329 (1973). 161. Primet, M., Fouilloux, P., lmelik, B., J. Cata/., 61, 553 (1980). 264. 162. Dalla Bena, R.A., Boudart, M., Gallezot, P., Weber, R.S., J. Cata/., 69, 514 (1981). 163. Vedrine, J.C., Dufaux, M., Naccache, C., Imelik, B., J. Chem. Soc. Faraday /, 74, 440 (1978). 164. Figueras, F., Gomez, R., Primet, R., Advan. Chem. Ser., 121, 480 (1973). 165. Tzou, M.S., Teo, B.K., Sachtler, W.M.H., J. Cata/., 113, 220 (1988). 166. Huang, Y.Y., Anderson, J.R., J. Cata/., 40, 143 (1975). 167. Oliver, D., Richard, M., Bonnerot, L. Che, M., "Growth and Properties of Metal Clusters", Elsevier, Amsterdam, p.165 (1980). 168. Jeanjean, J., Delafosse, D., Gallezot, P., J. Phys. Chem., 83 (21), 2761 (1979). 169. Delafosse, D., "Catalysis by Zeolites", B. Imelik et al. (Eds.), Elsevier, Amsterdam, p.235 ( 1980). 170. Briend-Faure, M., Guilluex, M.F., Jeanjean, J., Delafosse, D., Proc. Sym. Zeolite Szeged. Hungary, p.99 (1978). 171. Guilleux, M.F., Kermarec, M., Delafosse, D., J.C.S. Chem. Comm., p.102 (1977). 172. Rieckert, L., Ber. Bunsenges Phys. Chem., 73, 331 (1969). 173. Briend-Faure, B., Jeanjean, J., Delafosse, D., Gallezot, P., J. Phys. Chem., 84, 875 (1980). 174. Chien, S.-H., Chiang, W.-L., Appl. Cata/., 961, 45 (1990). 175. Sauvion, G.Y., Tempere, J.F., Guilleux, M.F., Djeda-Mariadassou, G., Delafosse, D., J. Chem. Soc., Faraday Trans. I., 81, 1357 (1985). 265. 176. Djemel, S., Guilleux, M.F., Jeanjean, J., Tempere, J.F., Delafosse, D., J. Chem. Soc., Faraday Trans. 1., 78,835 (1982). 177. Jeanjean, J., Djemel, S., Guilleux, M.F., Delafosse, D., J. Phys. Chem., 85, 4145 (1981). 178. Sauvion, G.N., Djemel, S., Tempere, J.F., Guilleux, M.F., Delafosse, D., "Catalysis by Zeolites", B. Imelik et al. (Eds.), p.245 (1980). 179. Yennakov, Y.I., Proc. 7th Intern. Congr. Cata/., Kodansha, Tokyo, 66, 121 (1980). 180. Harrison, B., Diwell, A.F., Hallett, C., Plat. Metal Rev., 32 (2), 73 (1988). 181. Chojnacki, T., Krause, K., Schmidt, L.D., J. Cata/., 128, 161 (1991). 182. Yao, H.C., Yu Yao, Y.F., J. Cata/., 86, 254 (1984). 183. Hurst, N.W., Gentry, S.J., Jones, A., Cata/. Rev.-Sci. Eng., 24 (2), 233 (1982). 184. Mahoney, F., Rudham, R., Summers, J.V., J. Chem. Soc., Faraday Trans. 1., 75, 314 (1979). 185. Jacobs, P.A., Nijs, H., Verdonck, J., J. Chem. Soc. Faraday Trans. 1, 15, 1196 (1979). 186. Martin, H.Z., US Patent 3,318,802 (1967). 187. Smith, B.E., Cookson, D.J., "Continuous Coal Hydrogenation Processes and Products", NERDDC EOGR, Dec. (1989). 188. Cookson, DJ., Percival, G.G., White, N., Fahy, J.A., "Control of Distillate/Gasoline Ratio for Coal-Derived and Other Liquids", ERDC EOGR Proj. No. 1400, Aug. (1991). 266. 189. Anon, "Ignition Quality of Diesel Fuels, Method B-Throttling", Inst. Petroleum Std., IP41/60. 190. Johnston, R.R.M., Christie, E.A., "Proc. Aust. Inst. Energy Nat. Conf.", 55, 671 (1985). 191. Maes, A., Cremers, A., Advan. Chem. Ser., 121, 231 (1973). 192. Schoonheydt, R.A., vanJlamme, L.J., Jacobs, P.A., Uytterhoeven, J.B., J. Cata/., 43, 292 (1976). 193. Thomas, C.L., "Catalytic Processes and Proven Catalysts", Academic Press, New York ( 1970). 194. Trimm, D.L., "Design of Industrial Catalysts", Chemical Engineering Monographs II, Elsevier, Amsterdam (1980). 195. ol Harque, I., "Mono and Bi-metallic Catalysts for Steam Reforming", Ph.D., University of New South Wales (1990). 196. "Catalyst Handbook", Scientific Text, Woolfe, London (1970).