Fluid Catalytic Cracking Catalyst Deactivation - Impact on Selectivity for Complex Feeds

Doctoral thesis for the degree of Doctor of Philosophy

FLUID CATALYTIC CRACKING CATALYST DEACTIVATION - IMPACT ON SELECTIVITY FOR COMPLEX FEEDS

David Anthony Chia B.Sc. (Hons)

School of Chemical Engineering and Industrial Chemistry University of New South Wales July 2002 DECLARATION OF ORIGINALITY

I hereby declare that this submission is my own work and to the be: my knowledge it contains no material previously published or writte1 another person, nor material which to a substantial extent has b accepted for the award of any other degree or diploma at UNSW or other educational institution, except where due acknowledgemer made in the thesis. Any contribution made to the research by oth with whom I have worked at UNSW or elsewhere, is expli acknowledged in the thesis.

I also declare that the intellectual content of this thesis is the produ1 my own work, except to the extent that assistance from others in project's design and conception or in style, presentation and lingu expression is acknowledged.

David Chia ABSTRACT

Fluid catalytic cracking is one of the largest chemical processes in a refinery and as such, slight changes in catalyst activity and product selectivity can result in large swings in profitability. Hence the objective of this study was to increase the efficiency of the refinery by investigating catalyst pre-treatment and deactivation.

The primary focus of the project was on maintaining an industrial perspective and, as such, the investigation did not allow for extensive examination of ideal systems and individual cracking reactions. Rather a fresh approach is adopted, tackling the topic from a more commercial point of view, looking to establish theories on the pre-treatment and deactivation of fluid catalytic cracking catalysts and to investigate what effects this has on the activity, product selectivity and reaction pathways of the catalyst.

The first section of the study investigated the pre-treatment of the industrial catalyst using light catalytic naphtha with the aim of increasing the yield of the valuable diesel blendstream, light cycle oil. Pre treatment of the catalyst with a light hydrocarbon resulted in the decrease of light cycle oil selectivity, combined with an increase of low value products such as dry gas yield (including hydrogen) and coke. Thus it is not an economically viable option to inject a light hydrocarbon into the bottom of the riser with the aim of increasing light cycle oil.

The second section of the study investigated common forms of deactivation affecting catalytic cracking processes, in particular the contamination of the cracking catalyst with nickel, vanadium and coke. These methods of catalyst pre-treatment were also shown to result in

ii decreased conversion of feedstocks, increased formation of hydrogen, and dry gas and a reduction in light cycle oil selectivity.

These observations were explained by the decrease in catalyst acid site strength and site density as a result of the various methods of pre treatment. The reduction in acid site strength and site density was found to affect the reaction pathways of cracking, polymerisation, isomerisation and hydrogen transfer. Further, that only small amounts of pre-treatment were necessary to dramatically reduce the activity of the catalyst and alter the selectivity.

iii ACKNOWLEDGEMENTS

I would like to express my gratitude to all the people who have contributed in any way towards the completion of this thesis. In particular, I would like to thank:

Professor David Trimm for his invaluable advice and direction. I am completely indebted to him for his encouragement, patience, and friendship.

Dr. Laurie Palmer and Caltex Refineries Queensland for their support and sponsorship of the thesis. Without Laurie's faith in me combined with his persuasion of the Caltex management, this project would never have gone ahead.

Dr. John Abbot for his advice and help during the early planning and experimental stages of the thesis.

Stan Robertson for his assistance with building the reactor and ensuring that whenever possible, the down time periods were minimised, without whom the experimental would not have been possible.

My colleagues at the Caltex Refinery and at UNSW, in particular Margarie Hileman, Alison Gregg and Nicholas Burke.

To my friends, in particular James Ellis and Kerri-Ann Lea for their support.

To my Mother, Father and brother Christopher for their never ending encouragement, care and love. Thank you.

iv QUOTE

"I am an empty vessel - not one thought Or look of love to Thee I've ever brought; Yet I may come, and come again to Thee With this the empty sinner's only plea Thou lovest me."

Mary Shekleton, 1827-83

V TABLE OF CONTENTS

DECLARATION OF ORIGINALITY ...... i

ABSTRACT ...... ii

ACKNOWLEDGEMENTS ...... iv

QUOTE ...... V

TABLE OF CONTENTS ...... vi

LIST OF TABLES ...... xii

LIST OF FIGURES ...... xv

CHAPTER

1 INTRODUCTION ...... 1

2 LITERATURE REVIEW ...... 3

2.1 CATALYST OVERVIEW ...... 3

2.2 REFINING AND CATALYSIS ...... 5

2.2.1 FUEL QUALITY ...... 14

2.3 THE CHEMISTRY OF CATALYTIC CRACKING ...... 16

2.3.1 THERMAL CRACKING ...... 17

2.3.1.1 Thermal Cracking Reactions ...... 18

2.3.2 CATALYTIC CRACKING ...... 19

2.3.2.1 FCC Catalyst Development ...... 19

2.3.2.2 Active Sites ...... 21

2.3.2.3 Mechanism of Catalytic Cracking Reactions ...... 22

2.3.2.4 Coke Formation and Catalyst Deactivation ...... 35

Vi 2.3.2.5 Summary of Catalytic Cracking Reactions ...... 37

2.4 CRACKING CATALYST STRUCTURE AND PROPERTIES .... 38

2.4.1 ZEOLITES ...... 40

2.4.2 MATRIX ...... 42

2.4.3 CRACKING CATALYST PROPERTIES ...... 43

2.4.3.1 Activity and Selectivity ...... 43

2.4.3.2 Stability ...... 44

2.4.3.3 Accessibility ...... 45

2.4.3.4 Acidity ...... 46

2.4.3.5 Framework Aluminium ...... 46

2.5 DEACTIVATION ...... 49

2.5.1 THERMAL DEACTIVATION ...... 52

2.5.2 POISONING ...... 52

2.5.3 FOULING ...... 54

2.5.4 COKING ...... 55

2.5.4.1 Coke Formation ...... 57

2.5.4.2 Coke Deactivation ...... 58

2.5.5 METAL CONTAMINANTS ...... 64

3 PROJECT OBJECTIVES ...... 69

4 EXPERIMENTAL ...... 71

4.1 MATERIALS ...... 71

4.1.1 GASES ...... 71

4.1.2 CHEMICALS ...... 73

4.2 ACTIVITY AND SELECTIVITY TESTING ...... 74

4.2.1 REACTOR AND FURNACE SYSTEM ...... 74

vii 4.2.2 PRODUCT COLLECTION SYSTEM ...... 79

4.2.3 CATALYST ...... 80

4.2.4 INDUSTRIAL FEEDSTOCK ...... 85

4.2.5 PRE-TREATMENT OF CATALYSTS ...... 87

4.2.5.1 In-Situ Coking ...... 87

4.2.5.2 In-Situ Ammonia Pre-Treatment ...... 87

4.2.5.3 Ex-Situ Ammonia Pre-Treatment...... 88

4.2.5.4 In-Situ Heat Pre-Treatment...... 88

4.2.5.5 Metals Impregnation ...... 89

4.2.5.6 Metals Impregnation on a Sand/Catalyst Split System 90

4.2.5. 7 Ex-Situ Pre-Coking ...... 90

4.2.5.8 Summary of Catalyst Pre-Treatment Methods ...... 91

4.2.6 PRODUCT ANALYSIS ...... 93

4.2.6.1 Gas ...... 93

4.2.6.2 Liquid ...... 94

4.2.6.3 Solid ...... 95

4.2.7 CATALYST CHARACTERISATION ...... 96

4.2.7.1 Temperature Programmed Desorption ...... 96

4.2. 7 .2 SEM / EDAX Machine and Method ...... 96

4.2.8 STATISTICAL SIGNIFICANCE OF RESULTS ...... 98

5 HYDROCARBON ON CATALYST PRE-TREATMENT ...... 99

5.1 INTRODUCTION ...... 99

5.2 PRE-INJECTION OF LIGHT CATALYTIC NAPHTHA ...... 103

5.2.1 LIGHT CATALYTIC NAPHTHA CHARACTERISATION .. 104

5.2.2 INDUSTRIAL FEEDSTOCK CHARACTERISATION ...... 105

viii 5.2.3 STANDARD AND PRELIMINARY REACTIONS ...... 108

5.2.3.1 Coke Selectivity on Hydrocarbon Pre-Treated Catalyst ...... 111

5.2.3.2 Hydrogen and Dry Gas Selectivity on Hydrocarbon Pre- Treated Catalysts ...... 113

5.2.3.3 Light Cycle Oil Selectivity on Hydrocarbon Pre-Treated Catalysts ...... 115

5.2.3.4 Paraffin to Olefin Ratio Shifts ...... 119

5.3 SUMMARY ...... 121

6 HEAT PRE-TREATMENT OF CRACKING CATALYST ...... 124

6.1 INTRODUCTION ...... 124

6.2 ACTIVITY LOSS ...... 125

6.3 CRACKING OF LIGHT CATALYTIC NAPHTHA ...... 129

6.3.1 HYDROGEN AND DRY GAS SELECTIVITY ...... 131

6.3.2 OLEFIN SELECTIVITY ...... 132

6.4 SUMMARY ...... 137

7 CRACKING OF PURE COMPOUNDS ...... 140

7.1 INTRODUCTION ...... 140

7.2 CRACKING OF DECANE AND HEXADECANE ...... 140

7.2.1 SHIFTS IN PRODUCT SELECTIVITY ...... 142

7.3 TRI-METHYL PENTANE CRACKING ...... 147

7.4 SUMMARY ...... 151

8 AMMONIA PRE-TREATMENT OF CATALYST ...... 153

8.1 INTRODUCTION ...... 153

8.2 PRELIMINARY EXPERIMENTS ...... 154

ix 8.3 TEMPERATURE PROGRAMMED DESORPTION ANALYSIS155

8.4 REDUCED CRUDE CRACKING ...... 160

8.4.1 REACTION PRODUCT DISTILLATION ...... 166

8.4.2 PRODUCT SELECTIVITY SHIFTS OF EX-SITU AMMONIA PRE-TREATED CATALYST ...... 167

8.4.3 PURE HYDROCARBON CRACKING ON AMMONIA PRE- TREATED CATALYSTS ...... 174

8.5 SUMMARY ...... 175

9 CRACKING OF SQUALANE ...... 177

9.1 INTRODUCTION ...... 177

9.2 SUITABILITY OF SQUALANE AS A REPRESENTATIVE TEST MOLECULE ...... 179

9.2.1 CRACKING OF SQUALANE ON AN AMMONIA/ HEAT PRE- TREATED CATALYST ...... 181

9.2.2 PRODUCT SELECTIVITY SHIFTS ...... 182

9.2.3 SHIFTS IN EXTENT OF ISOMERISATION REACTIONS 188

9.2.4 SHIFTS IN EXTENT OF HYDROGEN TRANSFER REACTIONS ...... 191

9.3 SUMMARY ...... 200

10 METALS AND COKE DEACTIVATION OF CATALYST ...... 201

10.1 INTRODUCTION ...... 201

10.2 CATALYST CHARACTERISATION BY SEM / XRD ...... 203

10.2.1 SCANNING ELECTRON MICROGRAPH$ ...... 203

10.2.2 ELEMENTAL ANALYSIS ...... 208

10.2.3 XRD ANALYSIS ...... 211

10.2.4 SURFACE AREA MEASUREMENTS ...... 213

X 10.3 NICKEL PRE-TREATED CATALYSTS ...... 214

10.3.1 HYDROGEN AND METHANE SELECTIVITY ...... 216

10.3.2 LIGHT OLEFIN SELECTIVITY ...... 219

10.3.3 LIGHT CYCLE OIL SELECTIVITY ...... 223

10.4 SAND/ EQUILIBRIUM CATALYST SPLIT SYSTEM TESTING ...... 225

10.5 VANADIUM PRE-TREATED CATALYSTS ...... 232

10.6 PRE-COKING OF CRACKING CATALYSTS ...... 236

10.6.1 HYDROGEN SELECTIVITY ...... 238

10.6.2 LIGHT OLEFIN SELECTIVITY ...... 240

10.6.3 ISOMERISATION ...... 243

10.6.4 LIGHT CATALYTIC NAPHTHA, LIGHT CYCLE OIL AND COKE SELECTIVITY ...... 245

10.7 SUMMARY ...... 249

11 CONCLUSIONS ...... 251

REFERENCES ...... 254

xi LIST OF TABLES

Table 1: Major products and yields from Caltex's fluid catalytic cracking unit ...... 12

Table 2: Main types of cracking catalyst deactivation ...... 50

Table 3: Specification and application of gases used ...... 71

Table 4: Chemicals used and their application ...... 73

Table 5: Chemical and physical properties of fresh and equilibrium catalysts used ...... 81

Table 6: Standard reaction characteristics of the fresh and equilibrium catalysts ...... 82

Table 7: Crude oil types and mass percentage present in feedstock as sampled ...... 86

Table 8: Temperature program used for the calcining of ammonia pre- treated equilibrium catalyst...... 88

Table 9: Temperature program used for the calcining of nickel and vanadium pre-treated catalysts ...... 89

Table 10: Types, methods and quantities of catalyst pre-treatment applied ...... 91

Table 11: PIONA analysis of light catalytic naphtha (LCN) ...... 105

Table 12: Physical and chemical properties of the industrial and forecasted feedstock ...... 107

Table 13: Carbon on catalyst as a function of hydrocarbon pre- treatment type ...... 111

Table 14: Effect of hydrocarbon pre-treatment on hydrogen and dry gas selectivity when cracking reduced crude ...... 114

xii Table 15: Product selectivity comparison between untreated e-cat and reduced crude pre-treated catalyst when cracking reduced crude ...... 118

Table 16: Paraffin to olefin ratio of light hydrocarbon products as a function of hydrocarbon pre-treatment ...... 119

Table 17: Reaction product selectivity of light catalytic naphtha cracking before and after heat pre-treatment of the catalyst...... 130

Table 18: Reaction product selectivities of decane and hexadecane cracking before and after heat pre-treatment of catalyst...... 141

Table 19: Product selectivity of 2,2,4 tri-methyl cracking over untreated equilibrium and heat pre-treated catalyst ...... 148

Table 20: Temperature programmed desorption results of untreated equilibrium and ammonia pre-treated catalysts averaged over 5 runs ...... 159

Table 21: Effect of ammonia pre-treatment on product selectivities when cracking reduced crude ...... 168

Table 22: Comparison of selectivity shifts for hydrocarbon and ammonia pre-treated catalysts when cracking reduced crude ...... 170

Table 23: Shifts in branched to linear paraffin ratios as a function of conversion on ammonia pre-treated catalysts when cracking squalane ...... 189

Table 24: lso-butane and iso-butene yields for the cracking of squalane on untreated equilibrium and ammonia pre-treated catalysts ..... 197

Table 25: Shifts in C2:C4 and C3:C4 ratios as a function of ammonia pre-treatment when cracking squalane ...... 198

Table 26: Surface area measurements for treated and un-treated catalysts ...... 213

xiii Table 27: Product selectivities for squalane cracking over sand and nickel doped sand ...... 227

Table 28: Product selectivities for squalane cracking over a single and two stage sand / catalyst system ...... 229

Table 29: Paraffin to olefin ratio for light hydrocarbons on a single and two stage sand / catalyst system in the cracking of squalane .... 231

xiv LIST OF FIGURES

Figure 1: Caltex Lytton refinery flow diagram ...... 7

Figure 2: Line diagram of Caltex Lytton's UOP side by side fluid catalytic cracking unit ...... 1o

Figure 3: Photograph of microreactor apparatus used ...... 75

Figure 4: Line diagram of microreactor set up used ...... 76

Figure 5: Photograph of reactor tube showing feed inlet, preheat section and catalyst...... 78

Figure 6: Photograph of Caltex Lytton's fluid catalytic cracking unit. 102

Figure 7: Cracking of reduced crude on equilibrium catalyst and the effect of hydrocarbon pre-treatment ...... 109

Figure 8: Coke selectivity of reduced crude cracking as a function of hydrocarbon pre-treatment type ...... 112

Figure 9: Effect of hydrocarbon pre-treatment on light cycle oil selectivity when cracking reduced crude ...... 116

Figure 10: Activity of equilibrium catalyst as a function of heat pre- treatment when cracking reduced crude at 500 °C ...... 126

Figure 11: Temperature programmed desorption plot of untreated equilibrium catalyst...... 156

Figure 12: Temperature programmed desorption plot of ammonia pre- treated equilibrium catalyst...... 157

Figure 13: Activity loss as a function of ammonia pre-treatment level when cracking reduced crude ...... 161

Figure 14: Reaction product distillation from reduced crude cracking on untreated equilibrium and ammonia pre-treated catalyst...... 166

Figure 15: Propylene plus butylenes selectivity as a function of ammonia pre-treatment when cracking reduced crude ...... 172

xv Figure 16: Conversion of 2,2,4 tri-methyl pentane when cracked over heat pre-treated and ammonia pre-treated cracking catalyst compared with untreated equilibrium catalyst...... 17 4

Figure 17: Structure of squalane (branched C30 compound) ...... 178

Figure 18: Activity loss when cracking squalane on heat, hydrocarbon and ammonia pre-treated catalysts compared to untreated equilibrium catalyst ...... 180

Figure 19: Total product selectivity of squalane cracking over untreated equilibrium and ammonia pre-treated catalyst at 39 mass% conversion ...... 183

Figure 20: Total product selectivity of squalane cracking over untreated equilibrium and ammonia pre-treated catalyst at 65 mass% conversion ...... 185

Figure 21: Light paraffin selectivity of squalane cracking over untreated equilibrium and ammonia pre-treated catalyst at 65 mass% conversion ...... 187

Figure 22: Branched to straight paraffin ratio at 65 mass% conversion, as a function of ammonia pre-treatment in the cracking of squalane ...... 188

Figure 23: Total olefin selectivity of squalane cracking at 65 mass% conversion as a function of ammonia pre-treatment...... 192

Figure 24: Branched to straight olefin ratio at 65 mass% conversion, as a function of ammonia pre-treatment in the cracking of squalane194

Figure 25: Scanning electron micrograph of equilibrium catalyst at a magnification of 100 ...... 204

Figure 26: Pictures of untreated equilibrium catalyst particles showing breakdown of the catalyst structure ...... 205

xvi Figure 27: Pictures of untreated equilibrium and 20,000 ppm nickel impregnated catalysts at 1OOO times magnification ...... 206

Figure 28: Pictures of an untreated equilibrium and a 5,000 ppm vanadium pre-treated catalyst particle at 20,000 times magnification ...... 207

Figure 29: Nickel concentration on untreated equilibrium catalyst and 20,000 ppm nickel pre-treated catalyst...... 208

Figure 30: Concentration of vanadium on the untreated equilibrium and 5,000 ppm vanadium pre-treated catalyst ...... 209

Figure 31: Aluminium concentration for the untreated equilibrium catalyst and 20,000 ppm nickel pre-treated catalyst...... 210

Figure 32: XRD analysis of fresh, e-cat and 20,000ppm Ni impregnated catalysts ...... 212

Figure 33: Conversion of squalane over nickel pre-treated catalysts215

Figure 34: Hydrogen selectivity of nickel pre-treated catalysts in the cracking of squalane ...... 217

Figure 35: Methane selectivity of squalane cracking as a function of nickel pre-treatment ...... 218

Figure 36: Hydrogen to methane ratio of squalane cracking as a function of nickel pre-treatment ...... 219

Figure 37: 3-methyl-1-butene selectivity of squalane cracking over nickel pre-treated catalysts ...... 221

Figure 38: Coke yield from squalane cracking on nickel pre-treated catalysts ...... 222

Figure 39: Light cycle oil yield of squalane cracking as a function of nickel pre-treatment...... 224

Figure 40: Conversion of squalane over vanadium pre-treated equilibrium catalyst ...... 233

xvii Figure 41: Methane selectivity of vanadium pre-treated equilibrium catalysts when cracking squalane ...... 235

Figure 42: Conversion of squalane over pre-coked equilibrium catalysts ...... 237

Figure 43: Hydrogen selectivity of squalane cracking over pre-coked catalysts ...... 239

Figure 44: Light olefin selectivity of squalane cracking over pre-coked catalysts ...... 240

Figure 45: lso-butene selectivity of squalane cracking over pre-coked catalysts ...... 241

Figure 46: Trans-2-butene selectivity of squalane cracking over pre- coked catalysts ...... 242

Figure 47: Branched to linear paraffin ratio for C4 and CS as a function of catalyst pre-coking in the cracking of squalane ...... 243

Figure 48: Total branched to linear olefin ratio of squalane cracking as a function of catalyst pre-coking ...... 244

Figure 49: Light catalytic naphtha selectivity of squalane cracking as a function of pre-coking of equilibrium catalysts ...... 245

Figure 50: Coke selectivity of squalane cracking as a function of pre- coking of equilibrium catalysts ...... 247

Figure 51: Light cycle oil selectivity of squalane cracking as a function of pre-coking of equilibrium catalysts ...... 248

xviii 1 INTRODUCTION

The focus on the optimisation of the petroleum refinery has magnified considerably as the demand for transport fuels increases and environmental restrictions tighten. One of the major processes used in petroleum refining is fluid catalytic cracking. In some Australian refineries, such as Caltex's Lytton refinery, the fluid catalytic cracker processes almost half of the total crude intake into the refinery. As fluid catalytic cracking is the largest chemical process in a refinery, even slight changes in catalyst activity and product selectivity can result in large swings in profitability.

Several variables have a large impact on the profitability of a fluid catalytic cracking unit. Among these are product demand, hardware constraints, crude type, and catalyst type. Typically, product demand is dependent on market fluctuations and, as such, is not directly controllable. Hardware modifications usually require large amounts of capital expenditure and are usually high risk because of the difficulty in modelling their impact. Thus they are difficult to justify. Crude type is also somewhat indirectly controlled, refineries often requiring certain types of crude to be processed, depending on hardware constraints. However, within the band of processable crude types, the levels of heavy metals contamination and the propensity to form coke can affect their purchasing prospects. The catalyst is the heart of the fluid catalytic cracker and is most easily affected by crude types, and changes in reaction conditions.

The deactivation of a fluid catalytic cracking catalyst can change the activity and selectivity of the catalyst, thus resulting in a less than optimal product slate. As a result, it has become increasingly important to understand the impact that feed contaminants such as metals and

1 the propensity to form coke from the feed have on the catalyst and ultimately on the product selectivity. From this, the catalyst can be reformulated to provide better performance.

Unfortunately, the amount of research conducted relating fluid catalytic cracking to commercial operations is extremely limited relative to the amount of information available on the cracking of pure compounds on laboratory catalysts.

This thesis is a study of product selectivity changes that occur when a commercially obtained fluid catalytic cracking catalyst becomes deactivated as a result of coke, nickel and vanadium deposition. It will seek to draw conclusions on the effect that these contaminants have on the ability of the catalyst to crack certain feedstocks. The study will also explore the possibility of improving the refinery product selectivity through in-situ catalyst modifications.

Benefits arising from such a study include a deeper understanding of the deactivation of fluid catalytic cracking catalysts in industry, recommendations for improved catalyst reformulations, better crude purchases, and ultimately improved product selectivity.

2 2 LITERATURE REVIEW

This review is a study of the literature available on the process of catalytic cracking. The emphasis is placed on the deactivation of fluid catalytic cracking catalysts through coking and metals deactivation.

2.1 CATALYST OVERVIEW

A catalyst, as defined by Ostwald [1] (circa 1900), is "a substance that alters the velocity of a chemical reaction without appearing in the end products". According to the Macquarie Dictionary a catalyst is a substance that causes, or increases the rate of a chemical reaction, which is not itself changed permanently by the reaction [2]. In other words, a catalyst is a substance that changes the kinetics but not the thermodynamics of a chemical reaction. Changes in reaction rate produced by a catalyst are usually positive, corresponding to acceleration. By increasing the velocity of a desired reaction relative to unwanted reactions, the formation of a desired product can be maximised compared with undesirable by-products.

Catalysts can contain a very wide range of chemicals represented by both elements and compounds, especially halides, oxides, and sulfides. The earliest applications of catalysts were in the production of wine, vinegar, and soap. The first inorganic applications were introduced in the middle of the eighteenth century, and were rapidly expanded. Catalytic production of organic chemicals followed. Finally, the production of liquid hydrocarbon fuels used in energy generation rapidly grew into the largest, by far, of all industrial catalytic applications.

3 Nowadays, catalysts are used in most processes leading to the production of industrial chemicals, fuels, pharmaceuticals, and to the suppression and destruction of environmental pollutants. About 90% of industrial chemical processes are catalytic, corresponding to 20% of all commercial products manufactured. The value of the heterogeneous catalyst market is estimated at US$5 billion per year and is expected to grow to US$6.5 billion by the end of 2001. The world-wide value of fuels and chemicals produced by catalytic reactions is about US$2.4 trillion per year, more than the gross national product of many nations. The cost of catalysts is normally a relatively small part of processing costs and is estimated at 0.1 % of the value of fuels produced and about 0.22% of chemicals produced. In liquid fuels production, fluid catalytic cracking (FCC) is the dominant user of catalysts with about 40% of the total [3].

Hydrocarbon catalysis was first introduced by Bergius in 1913 for liquefaction of coal by hydrogenation. Subsequently Fischer and Tropsch synthesised hydrocarbon fuels from carbon monoxide and hydrogen. Catalytic cracking of crude petroleum, the largest catalytic process, was introduced industrially by Houdry in 1936. At about this time catalytic polymerisation of olefins to liquid polymers also began, followed shortly by production of C8 hydrocarbons by alkylation. Catalytic reforming (the conversion of paraffins and cycloparaffins (naphthenes) to aromatic compounds) was used commercially in the United States and Germany during World War II. Catalytic reforming has developed into the second largest industrial catalytic process. Hydrodesulfurisation was developed primarily to remove objectionable sulfur compounds and, as a result of increasing environmental concern, has recently increased in importance [3].

4 Apart from distillation, processes in a petroleum refinery are mainly catalytic, and include fluid catalytic cracking, reforming, desulfurisation, isomerisation, polymerisation, and alkylation.

2.2 REFINING AND CATALYSIS

In petroleum refining, both physical and chemical means are used to convert crude oils into fuels. Crude oils are characterised by their high carbon and hydrogen contents, even though heteroatoms such as sulfur, oxygen, and nitrogen are present. In most crude oils, traces of metals such as nickel, vanadium, iron and even arsenic are found. Crude oils consist of thousands of individual compounds: paraffins, naphthenes, and aromatics with molecular weights ranging from 16, for methane, up to several thousands for asphaltenes.

In a typical refinery, crude oil is distilled into four major fractions in a crude unit or multiple crude units. The fractions are known as wide range straight-run naphtha, kerosene Uet fuel), straight-run diesel (SRO), and reduced crude. These four fractions need to be further processed so that the final products will meet specifications. Refiners must also satisfy the relative market demands for the principal fuels (gasoline, kerosene, diesel fuel and fuel oil). For example the market demand for gasoline in Australia is about 40% of the total crude intake, and since the straight run naphtha fraction comprises only 15-25% of crude oil, the remaining must be made up from other processes such as fluid catalytic cracking and catalytic reforming. Further, the chemical components of the natural product consist primarily of relatively low octane number alkanes and naphthenes. The naphtha fraction is not of sufficient octane quality to power modern vehicles and it needs to be structurally modified using catalysts. Typically refiners have to cope with

5 10-20 specifications, such as octane quality and flash point, for each product.

From the distillation unit, the wide-range naphtha fraction is split into two streams, light straight-run naphtha and heavy straight-run naphtha. The light straight-run naphtha proceeds to an isomerisation unit, where the hydrocarbons are structurally reorganised so that the octane is increased, using a platinum based catalyst. The heavy-straight run naphtha is processed in a reformer, which increases the octane rating, by around 45 numbers, by dehydrogenating and isomerising the hydrocarbons using a supported platinum, rhenium and tin catalyst. Products from both these units are used as gasoline blendstocks.

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ru.d C The kerosene and straight run diesel (SRO) fractions do not require significant molecular rearrangement by catalytic processes. The kerosene is drawn off as jet fuel and the SRO is used as an automotive diesel fuel (AOF) blendstock.

Since the market demands for gasoline and diesel fuel are so high, methods to convert higher molecular weight (C20+) hydrocarbons to gasoline and light cycle oil (a diesel component) are attractive. The primary unit for accomplishing such a task is a fluid catalytic cracking unit (FCCU), which reduces the average molecular weight of the reduced crude by breaking the carbon-carbon bonds using a zeolite catalyst.

The first application of acid catalysis was realised in petroleum processing by Eugene Houdry with his introduction of catalytic cracking. Houdry discovered that the conversion of hydrocarbon molecules to lower molecular weight fractions was catalysed by a variety of acid treated natural clays. Further, and more importantly, Houdry found that the coke by-product of this catalytic cracking could be burned off, resulting in restored cracking catalyst activity [4]. A fixed-bed commercial process utilising the principle of catalytic cracking followed by regeneration was used from 1936 through 1941.

In 1942, the Standard Oil Company of New Jersey commissioned the first fluid-bed catalytic cracking (FCC) unit. Since then the fluid catalytic cracking process has grown to dominate the catalytic cracking field [5].

Today the fluid catalytic cracking process is widely used with over 350 units consuming approximately 1,400 tons per day of cracking catalyst [6]. Fluid catalytic cracking has gained such a pre-eminent position probably because the unit hardware is relatively inexpensive to construct and operate (compared to a non-atmospheric unit) and is

8 surprisingly flexible. Fluid catalytic cracking has the capability to meet ever-changing product, environmental, and operational demands quickly and profitably. However, the chemical mechanisms of catalytic cracking are extremely complex. On-going research is still required to generate subtle but very profitable improvements in the yields of the most desired products.

Porous microspheroidal catalysts with an average diameter of approximately 60 µm are necessary for fluidisation and for effective contact between catalyst and oil. The FCC process involves fluidising the catalyst, which is intimately mixed in a stream of vaporised hydrocarbon feedstock and steam. Contact between the active sites of the fluidised catalyst and the hydrocarbon at temperatures of 500°C - 530°C leads to desirable and controllable molecular weight reduction ("cracking") as well as to the formation of coke on the catalyst surface. The catalyst is continuously transported from the reactor to the stripper to the regenerator and back to the reactor. Adsorbed but largely unreacted hydrocarbons are removed ("stripped") from the catalyst surface in the stripper while the regenerator functions to continually restore most of the catalyst activity by combusting the coke formed on the catalyst surface during the cracking reactions. Figure 2 shows the basics of a UOP side by side FCCU, similar to that used by the Caltex Lytton Refinery.

9 Figure 2: Line diagram of Caltex Lytton's UOP side by side fluid catalytic cracking unit

Reactor effluent to r;:::::======-__.• main fractionator

Stripping steam

Atomising steam "'-.....

Secondary feed _) injection point --+-

Air Air blower

Feed pre heater

Acceleration steam

10 Some recent developments in fluid catalytic cracking include hardware changes such as efficient feed / catalyst mixing devices, short contact time cracking, elimination of post-riser thermal cracking, and resid cracking hardware. [5). Process optimisation and advanced control are at the forefront of today's applied computer technology. Catalytic cracker emissions have been cut by orders of magnitude through hardware, catalyst, and operational changes [5). ·

Fluid catalytic cracking uses 500,000 tons per year of catalyst worldwide [6], mainly as small spheres (40-100 µm) containing Y zeolite, often exchanged with rare earths and dispersed in a silica alumina matrix. While still relying on the unique Y zeolite commercialised 25 years ago, FCC catalyst systems have become more sophisticated, blending numerous functional components, custom tailored for each refiner's needs.

Catalytic cracking units are quite large, some processing over 16,000 tons per day of crude (>6,000,000 tons per year) [6). A multitude of catalysts exist for the process. Catalyst changes can and are made often, and the feedstocks to the unit are varied in quality depending on the most favourable economics of crude oil and refined products prices.

The major products of a fluid catalytic cracking unit are dry gas (H2, C1, C2s), propane and propylene (C3s), butanes and butylenes (C4s), light catalytic naphtha (LCN), light cycle oil (LCO), catalytic cracker bottoms and coke. Table 1 lists the major products and typical yields from Caltex's fluid catalytic cracking unit.

11 Table 1: Major products and yields from Caltex's fluid catalytic cracking unit

Product Yield (mass%) H2 0.2 C1 0.7 C2s 1.3 C3 paraffin 1.1 C3 olefin 3.9 C4 paraffins 3.5 C4 olefins 5.1 Light catalytic naphtha 49.2 Light cycle oil 25.2 Cracker bottoms 4.8 Coke 4.9 Total 100.0

Although the dry gas is not a saleable product (and thus is undesirable in large quantities), it can be used as fuel for the refinery. The C3s are blended into LPG or are sometimes sent to a polymerisation unit while the C4s are blended into autogas or sent to an alkylation unit. The products of both the polymerisation and alkylation units are used as gasoline blendstocks.

The light cycle oil must be hydrotreated before being blended with straight run diesel to make automotive diesel fuel (ADF). Hydrotreating the light cycle oil serves to remove sulphur and nitrogen compounds and hydrogenates the olefins, thereby increasing storage stability. The cetane number of the light cycle oil is also increased, favourable for

12 diesel engine performance. The catalytic cracker bottoms are clarified to remove catalyst fines (thus becoming known as clarified oil (CLO)), which is then used as fuel oil for marine diesel engines and various other uses.

Over the past decade in Australia, there has been a decline in the demand for fuel oil and a stagnation in the demand for gasoline, while the demand for liquefied petroleum gas (LPG) and diesel fuel has increased. Since a catalytic cracker is most suited to producing gasoline, most refineries have needed to hardware modify their catalytic crackers to produce more diesel or have built a catalytic hydrocracking unit. Reformulations of the catalytic cracking catalyst, by careful research, is another means by which the yield of LPG and diesel may be increased.

The four major manufacturers of fluid catalytic cracking catalysts are Grace Davison, Engelhard Corporation, Akzo Chemicals Inc., and CCIC (Catalysts and Chemicals Industries Company Ltd.). They produce at a total rate of approximately 1400 tons per day to supply an increasing number of FCC units, which, in turn, are processing an ever-increasing variety of feedstocks. New FCC unit designs and construction, particularly in the Pacific Rim area, testify to continuing growth in FCC [7]. Similarly, new FCC catalyst compositions are continually evolving to meet these new process and feedstock needs [8].

From the standpoint of the fluid catalytic cracking catalysts, three primary catalyst properties or components control the vast majority of the hydrocarbon conversions which occur [9]:

1. The primary zeolite (usually type Y, occasionally ZSM-5) unit cell size (its Si/Al ratio), 2. The presence or absence of a catalytically active matrix,

13 3. The presence or absence of a small pore, shape selective molecular sieve.

Compared to the other major catalytic hydrocarbon processes used in petroleum refining, catalytic cracking is unique in at least three aspects:

1. The process is operated so that the catalyst is continuously regenerated. 2. A wide variety of feedstocks can be effectively converted. 3. Hundreds of catalyst types designed specifically to enhance yields and quality of fuels for the varying needs of refiners are available.

Difficulties such as reactor limitations, poor product selectivity and catalyst deactivation are encountered in FCC when trying to convert feeds with low hydrogen to carbon (H/C) ratio, high Conradson carbon (residual carbon in oil), and high metals (Ni, V and Fe).

Overcoming these difficulties will be met only through substantial contributions from academic and applied research. Such successes will continue to enhance the economics and competitiveness of petroleum refiners.

2.2.1 FUEL QUALITY

The successful conversion of reduced crude involves either adding hydrogen, removing carbon or a combination of both, each route having its own drawbacks in that adding hydrogen is an expensive process and carbon removal processes produce a by-product that has a very low or even negative commercial margin.

14 As a result, feedstock hydrogen to carbon (H/C) ratio is extremely important in influencing yield structures from the FCCU. The effect of feedstock hydrogen content on conversion level, gasoline yield and coke yield over the range 11 .8 to 13.3 wt% hydrogen can be quite dramatic [1 O]. Increases in hydrogen content lead to less deleterious coke formation while enhancing conversion levels and gasoline yields.

Although globally gasoline is the most important transportation fuel, in Australia and Europe there is a strong demand for diesel fuel. The progressive increase in diesel oil consumption is influenced largely by progress in diesel engine technology, which has been considerable during the last decade, and the relative importance of trucks compared to passenger cars. In addition there has been an increase in the demand for LPG as more cars are converted to use the higher economy fuel. These increases are offset by a marked reduction in the demand for fuel oil and a slight decrease in gasoline consumption primarily due to a large improvement in the efficiency of the SI (spark ignition) engine.

Gasoline quality is mainly characterised by the research octane number (RON) and the motor octane number (MON). The higher the octane rating of a fuel, the higher the potential energy stored in the fuel. RON is obtained through the presence of aromatics and ethers, whereas MON is built up with isoparaffins and ethers and is reduced by the presence of alkenes. Although a high octane component, the production of benzene in gasoline is undesirable as it is considered a carcinogen. Sulphur compounds are also undesirable as they lead to SOx emissions and they inhibit the activity of catalysts for treating exhaust gases. For the manufacture of gasoline, low molecular weight hydrogen deficient molecules or highly branched molecules are considered best.

Diesel oil is primarily characterised by the cetane number, which is linked to the presence of alkanes. A higher cetane number is

15 considered better in terms of its usefulness as diesel fuel. The cetane number is decreased mainly by the presence of bi- and poly-nuclear aromatics. Once again sulphur compounds are undesirable. For diesel products, contrary to gasoline products, hydrogen rich molecules are the best.

Thus hydrotreated straight-run diesel oils have good cetane numbers, whereas light cycle oil (LCO) from the fluid catalytic cracker often does not meet the required sulfur content, cetane number and the low aromatics content required. Indeed, the light cycle oil streams mainly contain condensed aromatic rings, which have poor cetane numbers. Hence carbon removal processes such as FCC require the addition of large amounts of hydrogen in downstream processes to saturate the alkenes and/or the aromatics that are generated and to desulfurise and denitrogenate the products. These hydrogen addition processes, such as hydrotreating, require hydrogen as a feedstock, which is usually obtained from the reformer unit.

2.3 THE CHEMISTRY OF CATALYTIC CRACKING

Catalytic cracking is a generic term that describes a process involving many complex reactions. Although some thermal cracking reactions occur, most cracking is catalytic. This section provides a discussion of the chemistry of cracking (both thermal and catalytic).

The mechanisms of catalytic cracking over solid acid catalysts continue to generate a vigorous discussion even after 80 years of extensive research in this field [11-13]. Three main questions remain unresolved:

• The nature of the active sites on the catalyst surface

16 • The nature of transient reaction intermediates formed in the reactions between the sites and substrates (linear and branched alkanes) • The chemical mechanism of the cracking reaction, the C-C bond fission

As there are many hypotheses presented on the chemistry of catalytic cracking, some being contradictory in nature, only the major theories that are fairly widely regarded as being correct will be presented in this section.

2.3.1 THERMAL CRACKING

Cracking of a hydrocarbon involves the breaking of a carbon to carbon bond. Before the advent of catalytic cracking catalysts, thermal cracking was the primary process used by petroleum refiners to upgrade low value feedstocks. Even today, refiners employ thermal processes such as "delayed coking" and "visbreaking" for cracking of residual hydrocarbons.

Thermal cracking is a function of feedstock composition, pressure, temperature, and time. The reaction occurs when hydrocarbons in the absence of a catalyst are exposed to temperatures above 425°C. Commercial processes typically use temperatures below 650°C. The initial step in the chemistry of thermal cracking is the formation of free radicals [14] that are formed by the splitting of a C-C bond or a C-H bond of the feedstock hydrocarbon.

17 2.3.1.1 Thermal Cracking Reactions

The energy required to break the carbon-hydrogen bond is greater than that for breaking the carbon-carbon bonds, so the severing of the carbon-carbon bonds occurs more readily. The strength of the carbon hydrogen bond for a hydrogen attached to a primary, secondary or tertiary carbon atom decreases in the following order:

primary > secondary > tertiary

On the other hand, the energy required to break primary, secondary and tertiary carbon-carbon bonding varies only slightly, so there is little discrimination as to which bond will be broken [15]. However, a double carbon-carbon bond and a single bond adjacent to a double bond is more stable than one farther away from the double bond. For example, in 1-butene (shown below), the carbon-carbon bond f3 to the double bond (second bond away) will break more easily than the bond a to the double bond [16].

a f3 C=C-C-C

The reaction usually involves free radicals, which are extremely reactive. Free radicals can undergo alpha scission, beta scission, hydride transfer and/or polymerisation [14]. Beta-scission is cracking that occurs two bonds away from the unpaired electron. Beta-scission produces an olefin and a primary free radical. [17].

A free radical can also react with a hydrocarbon molecule by abstracting a hydrogen atom to produce a free radical of that hydrocarbon and a molecule that relates to the original free radical.

18 Free radicals do not undergo isomerisation by migration of an alkyl group. Consequently, no increase in branching of hydrocarbons in thermally cracked products occurs beyond that of the original branching in the feedstock itself [18].

Thermal cracking causes a reduction in molecular size of the original reactants, leading to the formation of a final product that is rich in C1 and C2 as well as a fair amount of alpha olefins. One of the drawbacks of thermal cracking is that a high percentage of the olefins that are formed during intermediate reactions polymerise and condense directly to coke. The variety of reactions possible during thermal cracking produces a wide array of products. A large increase in selectivity is achieved by catalytic cracking leading to a vastly different product spectrum to that of thermal cracking.

2.3.2 CATAL YT/C CRACKING

Before discussing the mechanisms of catalytic cracking, it is appropriate to provide a brief historical review of fluid catalytic cracking (FCC) catalyst and process development.

2.3.2.1 FCC Catalyst Development

The first commercial fluidised cracking catalyst was an acid-treated natural clay. Later, synthetic silica-alumina materials containing 10 to 15 percent alumina replaced natural clay catalysts. Synthetic silica-alumina catalysts were more stable and yielded superior products. In the mid- 1950s, alumina-silica catalysts containing 25 percent alumina came into use because of their higher stability. These synthetic catalysts were amorphous, meaning that their structure consisted of a random array of silica and alumina tetrahedrally connected. Some minor improvements

19 in yields and selectivity were achieved by switching to catalysts such as magnesia-silica and alumina-zirconia-silica. [14].

The use of X and Y zeolites in FCC catalysts in the early 1960s was a major advance in catalytic cracking. The addition of these zeolites substantially increased catalyst activity and selectivity. (Zeolites are about 1OOO times more active than amorphous silica-alumina catalysts. The higher activity comes mainly from greater strength and regular organisation of the active sites in the zeolites. ). The product distribution from a zeolite-containing catalyst is different from the product distribution produced by an amorphous silica-alumina catalyst.

Zeolites are crystalline alumina-silicates having a regular pore structure. Their basic building blocks are silica and alumina tetrahedra. Each tetrahedron consists of silicon or aluminium atoms at the centre of the tetrahedron with oxygen atoms at the corners. Because silicon and aluminium are in a +4 and +3 oxidation state respectively, there is a net charge of -1 which must be balanced by a cation to maintain electrical neutrality [14].

The activity and selectivity of zeolites significantly depends on the type of cations that occupy the zeolite structure. FCC zeolites are synthesised in an alkaline environment such as sodium hydroxide. The soda Y zeolites have little or no hydrothermal stability. However, these alkali cations can easily be exchanged and the acidity of the zeolite's active site is enhanced upon ion exchanging the sodium with cations such as hydrogen or rare earth ions. The most widely used rare earth compounds are lanthanum (La3+) and cerium (Ce3+).

20 2.3.2.2 Active Sites

Active sites usually occupy only a small fraction of the total surface of a cracking catalyst and are acidic in nature. These acid sites are both Bronsted and Lewis type. A catalyst can have any combination of strong or weak Bronsted sites and strong or weak Lewis sites. The acidic properties depend on several parameters, including method of preparation, dehydration temperature, silica-to alumina ratio, and the ratio of Bronsted to Lewis acid sites.

Bronsted's definition of an acid is a substance capable of donating a proton. A Bronsted acid is a traditional hydrogen donor acid, such as hydrochloric acid and sulfuric acid. Bronsted sites are strong protic acidic species which either protonate alkane molecules to nonclassical carbonium ions CnH+ 2n+3 with pentacoordinated carbon atoms [19-21] or directly protolyse C-C bonds in alkanes and produce smaller alkanes and carbenium ions [21-27]. Bronsted acids are usually regarded as the active species in olefin cracking [28-29]; even relatively weak Bronsted acids can protonate olefin molecules to form carbenium ions [30].

Lewis' definition of an acid is a substance that accepts a pair of electrons. Lewis acids may not have hydrogen in their structure but they are still acids. The classical example is aluminium chloride. Aluminium chloride in water will react with hydroxyl groups, causing a drop in solution pH. Being very strong aprotic acidic species with vacant orbitals, Lewis sites are capable of removing H- from alkane molecules and converting them into carbenium ions CnH+ 2n+1 [11-13].

21 In the case of isoalkanes, both Lewis and Bronsted centres are believed primarily to attack tertiary C-H bonds in isoalkanes; the structures of the carbocations of branched alkanes being [11-13]:

R-CH2-C+(R")-CH2-R' carbenium ion R-CH2-CH2 +(R")-CH2-R' carbonium ion

2.3.2.3 Mechanism of Catalytic Cracking Reactions

The main aim of a catalytic cracking reaction is to reduce the molecular weight of a large alkane I isoalkane to form smaller alkanes and olefins. Under usual cracking conditions (temperatures over 400°C) in the presence of a cracking catalyst, any large alkane or olefin produces a large number of various light products. The complexity of the product mixtures greatly hinders elucidation of the cracking mechanism. A number of theories have been proposed to explain the formation of small alkanes and olefins.

Initiation in the Cracking of Paraffins

A lengthy ongoing argument has persisted as to how the cracking of paraffins is initiated.

There is general agreement that, for catalytic cracking to be initiated, a carbocation must be formed [18]. Carbocation is a generic term for a positively-charged carbon ion in a hydrocarbon. Carbocations can further be subdivided into carbenium and carbonium ions.

22 Carbenium and carbonium ions may be represented in the following manner [18):

Carbenium Ion: CR3 + Carbonium Ion: C~H+

"R" represents either an alkyl group or a hydrogen atom. The carbenium ion is tricoordinated and the carbonium ion is pentacoordinated.

What does not seem to be resolved is whether catalytic cracking is initiated by the formation of carbenium ions or carbonium ions and whether Bronsted or Lewis sites are responsible for the initiation. Several hypotheses for the initiation step in catalytic cracking of paraffins have been proposed [31-34).

Some propose that the initiation of paraffin cracking can occur at either Bronsted or Lewis acid centres on the catalyst surface [35). Although the catalyst has a fixed ratio of Bronsted to Lewis acid sites at ambient temperature, the catalyst may be pre-treated at a high temperature just prior the reaction to increase the proportion of Lewis sites on the catalyst surface [36). Presumably this is due to the loss of a water molecule from two Bronsted sites, resulting in the formation of one Lewis site [37). This hypothesis has been utilised by some researchers to propose that cracking is initiated via the Lewis site mechanism, in which a carbenium ion is formed by the abstraction of a hydride ion from a saturated hydrocarbon by a strong Lewis acid site: a tricoordinated aluminium species [38-40). However, other researchers propose that there is clear evidence for the participation of Bronsted sites in cracking reactions [41-42]. They found that by subjecting a catalyst to pre-treatment temperatures above 500°C, the cracking activity was lost for linear paraffin cracking. Thus they concluded that

23 Lewis sites do not contribute to cracking selectivity by processes independent of the presence of Bronsted sites.

A carbenium ion may be readily formed on Bronsted sites from an olefin by the addition of a proton to the double bond or, more rarely, via the abstraction of a hydride ion from a paraffin by a strong Bronsted proton. This latter process requires the formation of hydrogen as an initial product [17, 43]. For many years analysis for hydrogen was difficult, and as a consequence, often neglected. It is therefore not surprising that several cracking mechanism hypotheses postulate that the initial carbenium ions are formed only by the protonation of olefins generated either by thermal cracking or present in the feed as an impurity [44].

Abbot and Wojciechowski [35], theorise that the selectivity and kinetics of catalytic cracking of a range of paraffin feedstocks can be explained by assuming that cracking occurs only at Bronsted sites. They propose that the cracking reaction of a paraffin is initiated through the formation of a pentacoordinated carbonium ion. Recently, it has been demonstrated that such an ion can be formed on the alkane itself by protonation, if a sufficiently strong Bronsted proton is available [13].

Generally, the reactivity for protolytic cleavage of paraffins found in heterogeneous catalytic cracking is in the order [45]:

Despite a lack of unanimity, it is now widely accepted that the initiating event in the cracking of paraffins is the protolysis of a feed molecule by a strong Bronsted acid on the catalyst surface [46]:

CnH2(n+1 l + H+s- ~ CnH+ 2(n+1 )+1 s· ~ CiH2i+2 +Cn-iH+2(n-i)+1 s· - where "S" represents the catalyst surface

24 The intermediate step in the above process is the formation of a carbonium ion, which then permits a paraffin and a carbenium ion to form by the dissociation of this transitional species. A carbenium ion was earlier thought to desorb quickly as an olefin, perhaps after skeletal rearrangement, but without significant interaction with the feed. Such a two-step cracking mechanism of necessity results in a one-to-one ratio of paraffins to olefins in the products. As data accumulated, showing that in many cases there is an excess of paraffins over olefins [47-49], the idea of hydrogen transfer began to be discussed.

Catalytic Cracking Reactions

Once formed in the initial step, the organic intermediates of catalytic cracking - carbonium and carbenium ions - can undergo a number of different reactions. The nature and strength of the catalyst acid sites are critical in determining the stability of the ions, their residence time on the surface, and the type of reactions that the ion will undergo. Three dominant reactions that occur in the catalytic cracking process are:

• Cracking of a carbon-carbon bond • lsomerisation • Hydrogen transfer

25 Carbon-Carbon Bond Cracking Reactions

Branched Paraffin Cracking

The literature reports that the cracking of branched paraffin carbenium ions occurs via monomolecular ~ C-C bond scission [11-13]. There is a preference for beta-scission because the energy required to break this bond is lower than that needed to break the adjacent C-C bonds, as mentioned above in the thermal cracking section. The initial products of beta-scission are an olefin and a new carbenium ion [50]. The newly formed carbenium ion will then continue as a series of chain reactions. Small adsorbed carbenium ions such as a four-carbon or five-carbon species can then react with another paraffin feed molecule. The positive charge is transferred, leading to a product paraffin molecule and a new carbenium ion, which, in turn, can undergo cracking. Cracking does not eliminate the positive charge; it stays until two ions combine. The smaller ions are more stable and will not crack. They have a longer life and finally transfer their charge to a big molecule. However, the extent of this propagation sequence is not known. It is also possible that desorption of the residual ion as an olefin occurs after each cracking event, so that the original Bronsted site is reconstituted without any chain of reactions involving residual carbenium ions.

Reaction equation (1) shows the cracking of a 2,4-dimethylpentane carbenium ion:

(1)

In this case, the reaction produces an olefin (isobutene) and a small secondary carbenium ion which eventually abstracts H- either from a conjugated base on the catalyst surface or from another alkane

26 molecule and is converted into an alkane (propane). The p-scission reaction has been regarded as the principal feature of iso-paraffin and olefin cracking reactions for nearly 40 years [51]. Because p-scission is monomolecular, cracking is endothermic. Consequently, the cracking rate is favoured by high temperatures and is not equilibrium limited.

Long-chain hydrocarbons are more reactive than short-chain hydrocarbons; therefore, the rate of the cracking reactions decreases with decreasing chain length to the point that it is not possible to form stable carbenium ions.

Paraffin Cracking

The carbenium-ion mechanism for branched paraffin cracking has difficulties in explaining the cracking of linear and mono-branched alkanes. For example, the p-C-C bond scission of a mono-branched carbenium ion as shown in equation (2) (compared to reaction (1 )) produces a primary carbenium ion, an extremely endothermic reaction [52):

To avoid this difficulty, an alternative cracking route for such carbenium ions was proposed [53), which includes a charge shift and the formation of a secondary carbenium ion prior to the C-C bond scission, illustrated in equation (3):

(CH3)2C+ -CH2-CHrCH3 ~ (CH3)2CH-CH2-C+H-CH3 ~ (CH3)2C+H + CH2=CH-CH3 (3)

27 Although the formation of a small secondary carbenium ion in reaction (3) is less endothermic than that of the primary carbenium ion in reaction (2), the equilibrium concentration of the secondary ion via reaction (3) is many orders of magnitude lower than that of the tertiary ion. This makes the cracking route in reaction (3) extremely slow [52), and much slower than the direct cracking of linear paraffins via protolysis. However, the reaction of linear paraffins is inhibited by the competitive adsorption of product olefins, which also enhances the chain mechanism route.

In contrast to linear ions, the carbenium ion formed by hydride ion abstraction from a branched feed molecule undergoes beta-scission more rapidly than does the cleavage of the branched feed molecule through protolysis.

In catalytic cracking, it is apparent from the evidence cited that the feed molecule can be converted by either a monomolecular initiation reaction, (presumably involving a pristine Bronsted acid site) or a bimolecular reaction involving a carbenium ion adsorbed on a Bronsted base. As soon as carbenium ions are formed on the surface via the protolytic cracking responsible for initiation, bimolecular reactions begin to take part in a competition for the conversion of reactants. Thus, instead of the positively charged proton of a pristine Bronsted site acting as the agent for reaction, a larger positive moiety - the carbenium ion - now plays this role.

The simple mechanism which proposes that monomolecular protolysis yields a paraffin which is released to the gas phase and a carbenium ion which is more or less promptly desorbed as an olefin, appears to be inadequate. This inadequacy stems from the fact that the paraffin to olefin ratio of the products of alkane cracking is greater than unity.

28 Instead a bimolecular process of disproportionation, can explain the results from alkane cracking.

It can be seen that monomolecular initiation is not responsible for any process that could be described as hydrogen transfer. On the other hand, all bimolecular processes involve the transfer of hydrogen between feed molecules and carbenium ions. However, they do not all involve the simple transfer of hydride ions. The term "hydrogen transfer" may be misleading; "disproportionation" is a more general name for the processes involved in most "hydrogen transfer" reactions [54].

There exists the possibility of a number of conversion reactions involving the disproportionation of a feed molecule with the resident carbenium ion. These disproportionation reactions can occur between the insertion of the proton from a Bronsted acid site into a paraffin molecule during initiation by protolysis and the recovery of a proton from a carbenium ion in a termination reaction involving the desorption of an olefin.

This disproportionation would account for the fact that cracking produces an excess of paraffins over olefins [55-57]. Monomolecular protolytic cracking produces olefins and paraffins in a ratio of 1: 1. On the other hand, bimolecular disproportionation, involving the transfer of hydride ions (or carbon containing moeties with a negative charge), between a carbenium ion and a feed molecule produces only paraffin molecules. No matter how many times it is repeated on the given site, no olefins are produced at all. Thus a chain mechanism is described, involving monomolecular initiation, disproportionation of ions with the feed, and the desorption of ions to form olefins. There is a limit to how low the paraffin:olefin ratio can be. The number of p-cracking events that an average carbenium ion in the reaction can undergo before it becomes too small to crack determines this value.

29 The disproportionation reactions between surface carbenium ions and feed molecules leads to fresh carbenium ions, most of which are of the same configuration as those resulting from initiation reactions. However, some ions may be new, and their disproportionation reactions must also be included in the mechanism. All the ions present will either desorb as olefins or enter into succeeding disproportionation reactions, obeying the same rules as those that apply to ions derived from initiation reactions. This proceeds until a carbenium ion manages to desorb before it combines with a feed molecule from the gas phase. If desorption takes place, an olefin is released to the gas phase, a chain is terminated, a pristine Bronstad site is reconstituted, and the rejuvenated site becomes available to carry out a new initiation reaction and start a new chain of events. Studies of solvated ions show that, in general, smaller ions are favoured over larger ions. [58] The tendency of small ions to produce more olefins, than the larger ions, need not indicate that more such small ions are formed; only that they desorb rather than disproportionate.

Using a range of hydrocarbons, it was shown that, as the number of carbon atoms in the hydrocarbon molecule increased, the activation energy for cracking decreased [59]. Hence the rate of cracking of normal paraffins increases as the chain length increases to a maximum

(at n-C16) and then declines. Olefins crack more readily than paraffins since olefins are more easily converted to carbocations. lsoparaffins and naphthenes also crack faster than n-paraffins, which in turn crack faster than aromatics. Aromatic rings do not crack easily. Fluid cracking catalysts can be easily deactivated by polynuclear aromatics that block the acid sites of the catalyst [50].

Cumming and Wojciechowski [54] propose that the chain initiation (by protolysis), propagation (by disproportionation), transfer (by the ~-

30 cracking of an ion), and termination (by the desorption of an olefin) constitutes a general scheme for the chain of interdependent processes which takes place in the catalytic cracking of all hydrocarbon molecules.

/somerisation Reactions

In catalytic cracking, carbocations can easily rearrange to form tertiary ions, a relatively low energy reaction when compared with cracking reactions. Tertiary ions are more stable than secondary and primary ions. These ions can then crack to produce branched molecules.

Some of the advantages of isomerisation are as follows:

• Higher octane quality for petrol blendstocks • Higher-value chemical and oxygenate feedstocks • Lower cloud point for diesel fuel blendstock

Hydrogen transfer reactions

Little information is reported in the literature regarding the mechanistic details of hydrogen transfer. Both free-radical and ionic processes have been postulated [60].

Hydrogen transfer, or more correctly hydride transfer, is a bimolecular reaction in which one reactant is an olefin. An example of hydrogen transfer is the reaction of two olefins. Both olefins would have to be adsorbed on the active sites and the sites would have to be close together for these reactions to take place. In the reaction one of the

31 olefins becomes a paraffin and the other becomes a cyclo-olefin, so hydrogen is moved from one to another. The cyclo-olefin can then undergo further hydrogen transfer with another olefin to yield a cyclo di olefin. The cyclo di-olefin, being relatively unstable, will quickly rearrange to form an aromatic, which is extremely stable. Therefore, the hydrogen transfer of olefins can result in the formation of paraffins and aromatics. Naphthenic compounds are also hydrogen donors and can react with olefins to produce paraffins and aromatics (61 ).

A rare-earth-exchanged zeolite indirectly increases hydrogen transfer reactions. In simple terms, the rare earth forms bridges between two to three acid sites in the catalyst framework. In doing so, the rare earth basically protects those acid sites from being ejected from the framework. Results confirm the incidence of site density, suggesting that paired sites are needed for hydrogen transfer to take place. Because hydrogen transfer is promoted from adjacent acid sites, bridging these sites with rare earth oxides promotes hydrogen transfer reactions (14).

When hydrogen transfer reactions occur, the product spectrum shows a shift towards a lower concentration of olefins. Since olefins are the precursors for secondary cracking reactions, it can be deduced that hydrogen transfer reactions indirectly reduce "overcracking" of the gasoline. A catalyst with more hydrogen transfer characteristics will cause the net heat of reaction to be less endothermic.

Some of the drawbacks of hydrogen transfer reactions are:

• Lower gasoline octane quality • Lower concentration of light olefins in the LPG

32 • Higher concentration of aromatics in the gasoline and light cycle oil • Lower concentration of olefins in the front end of gasoline

There is general agreement about the main factors governing hydrogen transfer in cracking reactions. Though expressed through different approaches (such as the measurement of diminished olefin adsorption [62], the comparison between higher isomerisation or cracking to hydrogen transfer selectivity rates, or decreased hydrogen transfer selectivities [63]) various researchers have focused on site density as the controlling factor. It has been stressed for Y zeolites that the higher the dealumination, the lower the site density and the more intense the site isolation, all causing bimolecular hydrogen transfer reactions to decline [64].

An alternative view was given by Lombardo et al. [65]. By testing for acid sites in zeolites with ammonia poisoning and neopentane conversion, they postulated that hydrogen transfer reaction feasibility would be influenced by the time the reacting species remains on the active site. Increases in both acid site density and strength would favour secondary reactions such as hydrogen transfer. According to Peters et al. [66], their proposition would imply a negative relationship between hydrogen transfer and acid strength. The higher the acid strength, the higher the turnover numbers and the less the hydrogen transfer. Accordingly - and since acid strength increases with site isolation, a consequence of increasing dealumination [63] - it was suggested that both site density and time on site influence hydrogen transfer rates [66].

Cracking reactions of saturated hydrocarbons have been observed to be accompanied by formation of molecular hydrogen on solid acid catalysts at elevated temperatures [35]. For reactions of paraffins and cycloparaffins on solid acids, including zeolites, three distinct sources of

33 hydrogen have been identified. Two can be attributed to processes involving free radical thermal cracking [67], and reactions leading to coke and aromatics [68]. The third process involves interaction of the hydrocarbon with Bronsted sites on the catalyst surface and appears to proceed through protonation of the reactant molecule [35].

Typically in catalytic cracking, there is an excess of paraffins over olefins which has been noted by several authors [55-57]. An explanation for the excess is the formation of aromatics and coke which, being dehydrogenation reactions, make hydrogen available to saturate olefin molecules [24]. This is true, but the process is a minor contributor of products, involves multiple steps, and is not the direct hydrogenation of olefins than is sometimes envisaged. Most importantly, very little hydrogen is available for transfer from this source.

Cumming and Wojciechowski propose that hydrogen redistribution is thought to be the correct explanation for the puzzling excess of paraffins over olefins [54]. They show that the excess of paraffins over olefins is due not to "hydrogen transfer" but to a process which involves chains of consecutive reactions involving the carbenium ions left on the catalyst surface by the initiating protolysis reaction.

In the light of this proposed chain mechanism, increased hydrogen transfer is indicative of more disproportionation activity. This, in turn, is due to either more carbenium ions or more reactive carbenium ions on the surface - phenomena that may be responsible for coke and catalyst decay via ion-ion disproportionations.

As hydrogen transfer is a bimolecular reaction, any change in the concentration of adsorbed carbenium ions will affect this reaction. As a result, acid strength will also influence hydrogen transfer reactions. For

34 instance, a strongly acidic catalyst has been found to encourage hydrogen transfer reactions and coke formation [69].

The redistribution of hydrogen in the initial stages of catalytic cracking is mainly due to chain-propagating reactions between the feed and carbenium ions. At higher conversions, product olefin molecules can react with carbenium ions, contributing to more hydrogen shuffling and paraffin formation.

It can be seen that there is one sure indicator of "true" hydrogen transfer reactions, as opposed to those which are the result of the saturation of a product olefin by hydrogen available from the formation of coke. Since the mechanism outlined above requires that the number of moles of paraffins in the product must be equal to the number of moles of feed converted, we expect to observe a molar selectivity of 1 for the total of the product paraffins in the initial products of reaction. This has been found to be true under normal cracking conditions for all the C6 isomers. In cases where the total initial paraffin selectivity is greater than 1 , the excess paraffins must come from the saturation of product olefins by the dehydrogenation of coke.

2.3.2.4 Coke Formation and Catalyst Deactivation

Cracking, isomerisation, and hydrogen transfer reactions account for the majority of reactions occurring in catalytic cracking. There are other associated reactions that play an important role in unit operation. Two prominent reactions are dehydrogenation and coking. Under ideal catalytic cracking conditions, i.e., "clean" feedstock and catalyst with no contaminant metals, the yield of hydrogen is quite low. Dehydrogenation reactions, and the resulting formation of molecular

35 hydrogen, are noted to proceed significantly if the catalyst is contaminated with metals such as nickel and vanadium.

Catalytic cracking of hydrocarbon molecules yields a carbonaceous residue called coke. The chemistry of coke formation is very complex and not well understood. Similar to hydrogen transfer reactions, the formation of catalytic coke is a "bimolecular" reaction product and proceeds via carbenium ions or free radicals. In theory, coke yield should increase as hydrogen transfer rate is increased. Multi-ring aromatics and unsaturates are the principal coke-forming compounds [70]. Unsaturates such as olefins, diolefins, and multi-ring polycyclic olefins are very reactive and can quickly undergo dehydrogenation and polymerisation to form coke.

The decay of catalyst activity is due to the deactivation of individual sites by inactive and undesorbable carbenium ions. These are probably due to a preceding formation of unsaturated carbenium ions that can bridge adjacent sites on the catalyst. These ion bridges have been shown to reduce two pristine sites to zero activity in the cracking of small "model" compounds and result in second-order decay kinetics [54]. The more complex ions present in commercial feedstock cracking can deactivate sites one at a time by the elimination of a sufficient fraction of a complex carbenium ion, originating from a complex feed molecule, to produce an inactive species residing on one site. This mode of decay can dominate in the cracking of complex feeds, resulting in first-order, or exponential, decay.

36 2.3.2.5 Summary of Catalytic Cracking Reactions

In light of the above literature, the following conclusions can be made about catalytic cracking [54):

A. Strong Bronsted sites are associated with: 1 . High rates of protolysis 2. Weak conjugate Bronsted bases; and therefore with: a. Long surface residence times for ions b. High hydrogen transfer c. Rapid coke formation d. Rapid catalyst decay by second-order processes

B. Weak Bronsted sites are associated with: 1 . Low rates of protolysis 2. Strong Bronsted bases; and therefore with: a. Low hydrogen transfer b. Low paraffin:olefin ratios in the products c. Low coke make d. Slow catalyst decay by second-order processes

C. High site density is associated with: 1. Closely spaced sites 2. Many sites per unit of catalyst; and therefore with: a. High initial reaction rates b. High rates of decay by second-order processes

D. The presence of steric constraints is associated with: 1. Low rates of diffusion into the catalyst 2. Steric restrictions on the transition states in the pores; and therefore with: a. Low rates of bimolecular interactions in larger species

37 b. Low hydrogen transfer c. High orders of catalyst decay due to pore plugging

2.4 CRACKING CATALYST STRUCTURE AND PROPERTIES

Today's fluid catalytic cracking (FCC) catalysts are complex mixtures of functional components customised to meet a particular set of economic objectives and hardware constraints. Although complex mixtures, all cracking catalysts have at least one thing in common - a solid surface which contains acid sites of sufficient strength to cause the formation of a carbocation when a feed hydrocarbon molecule interacts with the surface. FCC catalyst research and development is directed towards producing catalyst formulations which improve the catalyst activity as well as its selectivity during cracking, i.e. its ability to form gasoline, light cycle oil and lighter products with minimum amounts of coke and slurry formation.

Historically, the first generation of catalysts was obtained by using natural materials, which were sometimes treated (acid-washed clays) and modified to obtain particles with the proper shape and size. For years, cracking catalysts were produced and used in the form of pellets and beads, but Exxon elected to utilise a powdered catalyst for their fluid catalytic cracking unit (FCCU). Collaborative efforts with the Massachusetts Institute of Technology (MIT) in the field of fluidisation ultimately led [71] to the development of micro-spheroidal catalysts in 1948. FCC catalysts progressed from these low alumina spray-dried clays to high alumina synthetic silica-alumina catalysts in the mid- 1950s.

Perhaps the biggest milestone in FCC catalyst technology occurred in the early 1960s when researchers from Mobil pioneered the introduction

38 of a modified zeolite X for petroleum cracking [72, 73]. The superior hydrothermal stability of these zeolite based catalysts increased the relative cracking activity (compared to the amorphous alumina catalysts) by several orders of magnitude [72]. In addition to the significant improvement in gasoline yield, major changes in the gasoline composition were observed, since the increased hydrogen transfer ability of zeolite-based catalysts (compared to amorphous alumina) also gave rise to a more stable gasoline. Typically, the sequence of reactions required to form a more stable gasoline involved naphthenes plus olefins to give aromatics plus paraffins.

The primary active ingredient of modern FCC catalysts is usually some form of Y zeolite (REY or USY), often modified by a series of chemical and physical processes to optimise a particular cracking reaction. The zeolite is usually embedded in an inorganic oxide matrix. The zeolite provides most of the cracking activity while the matrix possesses physical as well as some catalytic functions [50]. Within the terminology presently used in the FCC catalyst world, all the non-zeolitic components are grouped under the often confusing heading "matrix". This stems from an earlier time when a catalyst's activity was derived solely from the zeolite component. Modern FCC catalysts, however, utilise several multifunctional "matrix" materials to enhance critical elements such as bottoms upgrading and metals resistance. Examples of typical FCC catalyst components are shown below.

Zeolite: Y, USY, ZSM-5 Filler: Clay Additive: Alumina, Silica / Alumina Binder: Silica Sol, Alumina Sol, Alumina Gel, Silica/ Alumina Gel, Modified Clay

39 The role of the matrix (filler, binder and additive) is to bind the zeolite crystals together in a microspherical catalyst particle hard enough to withstand intraparticle and reactor wall collisions in a commercial catalytic cracking unit [74]. In addition, the matrix has a large number of mesopores (>20 A) which provides a medium for the diffusion of feedstock molecules and cracked products. It also provides a means for heat transfer during reaction and regeneration, thereby protecting the zeolite structure from structural damage by increasing its hydrothermal stability [7 4-76].

The resulting catalyst is a complex mixture or "assembly" of active and non-active materials, which together form the physical FCC particle with its architectural form and strength, and it's distinct catalytic properties.

2.4. 1 ZEOLITES

All modern cracking catalysts contain a zeolite. Zeolites are microporous crystalline solids with cages of molecular dimension that selectively accept or reject certain reactants or products, on the basis of their shape or the shape of a transition state between reactants and products. The principle of shape selectivity is considered to be an important principle in the control of selectivity in heterogeneous catalysis. The introduction of shape selectivity was pioneered by a research team at Mobil led by P. B. Weisz, that developed many of the principles of catalyst accessibility.

Zeolites, or crystalline aluminosilicates are silicates with an open framework, regular structure, and apertures of molecular dimensions based on a tetrahedron of four oxygen ions surrounding a smaller silicon or aluminium ion, thus making the internal surface available for adsorption and catalysis [77]. The tetrahedra are arranged to form the

40 porous framework of the zeolite crystal structure, whose pore diameters are determined by the number of oxygen ions arranged in regular rings in the crystal structure. Zeolite Y has channels about 7 .5 A in diameter connecting cavities about 13 A in diameter in a three-dimensional network which permits diffusion of hydrocarbon molecules into the interior of the crystal [78]. This may inhibit the reactions of larger molecules. Thus most gas oil cracking has been shown to occur on or near the outer surface of the Y zeolite crystals [79]. ZSM-5 is another type of zeolite with a straight channel and an interconnecting sinusoidal channel, which gives rise to more shape selective reactions [78], in particular the formation of gasoline and light olefins at the expense of hydrocarbons boiling in the cycle oil range.

Zeolites exhibit a rather narrow distribution of acid-site strengths and their catalytic activity is strongly influenced by the nature of the cation occupying the exchange sites on the structural framework [80]. For the Y zeolites employed in FCC catalysts, the acid sites necessary for carbocation formation are introduced via the exchange of rare earth or ammonium ions for the sodium ions incorporated during zeolite synthesis. The complex series of reactions, which occur when a hydrocarbon molecule encounters a zeolite, are dependent upon steric factors and on the strength of the acid sites present. For example, modifications to acid site strength and density, by changing the silica to alumina ratio or the distribution of framework atoms, lead to changes in catalyst activity, stability and selectivity [81, 82].

41 2.4.2 MATRIX

Matrix components may also provide some catalytic properties, and the proper choice can provide favourable pore structure, bottoms upgrading, coke selectivity, metals resistance and passivation.

Matrix cracking activity is especially required in cracking applications where distillation residue is mixed in with the feed. The large higher boiling point hydrocarbons in the residue (b.p. >540 °C) cannot physically enter the internal structure of the zeolite, thus an active matrix with large mesopores is required to pre-crack these molecules into smaller fragments before they enter the zeolite, where they are further cracked into more valuable products.

With the increasing demand for transportation fuels and the declining demand for heavy fuel oil, most of today's FCC catalysts are designed to include matrix with enhanced bottoms upgrading cracking activity [83]. The active "matrix" systems are often synthetic multi-functional alumina and/or silica-alumina systems with pores in the mesopore (30 to 500 A) range or macropore (500+ A) range that may have been chemically or physically modified to enhance activity. They may also serve as building blocks for pore architecture design and often have the capability to activate or deactivate metals depending on the custom design of the metal support interaction.

Since pre-cracking of large hydrocarbons (i.e. residue) is expected to be catalysed by the surface acid sites of the matrix, matrix activity is sometimes erroneously viewed synonymously with matrix surface area. More correctly, the activity of the matrix should be related to the site density. [71 ]

42 The component parts of the matrix may be categorised as follows:

• A catalytically active portion containing larger pores than the primary zeolite active component.

• A binder which also may be catalytically active but functions primarily to bind the components of the catalyst into an attrition resistant -60-micrometer microsphere; and

• A kaolinite filler which functions to density the microsphere and introduces a large-pore macro-system which is believed to facilitate access of large asphaltene composites (2.5 nm and larger) to active portions of the matrix or active sites associated with the exterior structure of the zeolite component.

2.4.3 CRACKING CATALYST PROPERTIES

2.4.3.1 Activity and Selectivity

Activity refers to the overall conversion of the feedstock or the amount of reactant in contact with the catalyst under a given set of conditions converted to all products. Selectivity means the efficiency in catalysing a desired transformation. It is usually expressed by a factor representing the amount of desired product formed, divided by the amount of reactants converted. Selectivity often decreases as the percentage of converted charge increases leading to unwanted products. Thus in industrial practice, a compromise is frequently made between selectivity and conversion. Another measure of selectivity is in product quality, e.g., the octane number of gasoline [4].

43 The pioneering work at Mobil clearly showed the dramatic improvement in both activity and selectivity of zeolites over amorphous catalysts, even though they both contain the same types of acid sites [72, 84, 85]. Acid catalyst activity is a function not only of the total number of acid sites but also their strength and accessibility. Product selectivity will be determined more by the distribution of acid site strengths, in other words the range of acid site strengths available, and their proximity to one another. Strong acid sites have been found to be necessary for catalytic cracking reactions [60, 86]. Zeolites can therefore be thought of as solid acids with the active sites being the acid sites.

The activity and selectivity of a catalyst are also affected by its pore structure. In a very slow reaction, molecules can diffuse through a pore system to the centre of a catalyst particle before they react, and the entire internal surface area will be used. In a fast reaction, diffusion is slower than the conversion and only the outer surface area and pore mouths of the catalyst will be used.

2.4.3.2 Stability

Stability refers to the ability to retain initial activity and selectivity over the catalysts' lifetime. Activity loss can be reversible as in coking or may be permanent as a result of overheating (frequently in the presence of steam) which causes sintering with loss of surface area, or by poisons that deactivate active centers. Loss of selectivity is usually caused by chemical transformation in the catalyst, frequently by addition of poisons, which cause unwanted reactions.

44 2.4.3.3 Accessibility

Catalyst accessibility is also very important. The open structure of crystalline zeolite Y includes three-dimensional channels or pores. These pores allow easy access to the acid sites for the majority of feedstock molecules and ready removal of cracked products, thereby minimising undesirable secondary cracking products, which tend to be light gases outside of the gasoline range. Zeolites also have high surface areas as a result of their open framework structures. Consequently it is possible to have many accessible acid sites.

The size and geometry of the zeolitic pores are at the molecular level and can therefore determine the reactants that have access to the many acid sites distributed throughout the internal surface. The nature of the products is similarly impacted. This process is known as shape selective cracking and the reactant molecules have been subject to molecular sieving.

One of the advantages of the pores in the zeolite molecule is their open, porous and uniform structure which allows easy accessibility to the active sites. However, FCC catalysts consist of zeolite crystals bound together with several matrix components into a microsphere, typically with 50-70 micrometers average particle size. The porosity of the matrix and hence the whole FCC particle is critical to ensure accessibility to the active catalytic sites.

Accessibility can be impacted by coking. Coking is the formation of carbonaceous by-products that become adsorbed on the catalyst surface and cannot be cracked into desirable products. Coking can lead to deactivation of the catalyst by several routes, one of which is the physical blocking of the pores.

45 2.4.3.4Acidity

Alumina shows no Bronsted acidity although, after calcination at relatively high temperature, it exhibits Lewis acidity. After the removal of a hydroxyl group, this leaves exposed an aluminium cation capable of accepting an electron pair. Silica shows neither Bronsted nor Lewis acidity, whereas silica-alumina shows both. The Bronsted sites are converted to Lewis sites on calcination and back to Bronsted sites upon the addition of water.

2.4.3.5 Framework Aluminium

It is now generally accepted that the activity and selectivity of Y zeolites are determined by an interplay of framework aluminium and non framework aluminium (NFA) species [87]. Other factors which influence Y zeolite behaviour include the silica to alumina ratio and the aluminium topography, i.e. the aluminium distribution over the zeolite crystal surface and pore morphology.

Since zeolite frameworks collapse in the presence of steam via rapid dealumination and FCC units use high temperature steam to regenerate the catalysts, stabilised frameworks (produced via controlled dealumination of the framework prior to the introduction into the FCC unit) are extremely important.

Treating zeolites such as Y with steam has been shown to remove framework aluminium (Alt) creating defect sites which are slowly filled by migrating silicon atoms [88-91]. The ultra-stable Y (USY) zeolite formed has much higher thermal and hydrothermal stability. It also contains nonframework alumina species, unless treated to minimise / remove them - typically with an acidic ion exchange process. The

46 nonframework alumina, is a complex mixture of species, some of which contain Lewis acid sites, others may enhance the Bronsted acidity. However it is still not clear whether it plays a significant role in FCC catalysis, other than blocking the zeolitic channels [92].

The number of Alt in zeolite Y is directly proportional to the number of Bronsted acid sites. Thus as the number of Alt decreases during dealumination it would seem reasonable to expect the acidity and catalytic activity to decrease. In fact the acidity and activity initially increase as aluminium atoms are removed from the framework [93]. A maximum activity for cumene alkylation has been observed at about 20- 30 Alt per unit cell (far lower than the 48-56 Alt per unit cell found in as synthesised Y zeolite) [80]. This corresponds to a range of 5.4-8.6 Si/Al framework atoms.

To explain these observations the concept of "next nearest neighbours" (NNN) has been introduced [94]. The idea is that the strongest acid sites are at those Alt atoms which have no nearest or next nearest Alt neighbour tetrahedral (T) atoms. In other words they have only Si T atoms surrounding them, excluding the framework oxygens. Lowenstein's rule already requires that the nearest neighbour atoms cannot be Alt, as no zeolite framework has been observed to have adjacent Alt T-atoms [95]. In the case of zeolite Y it has been calculated that the point at which all Alt are isolated, i.e. with no NNN Alt atoms, corresponds to a ratio of about 7 Si/Al framework atoms [96, 97]. Thus, as Alt are removed from the zeolite framework, the acidity should increase until all Alt are isolated with no NNN aluminium atoms. All sites should then have the same strength, but the overall acidity will decrease as more Alt are removed.

In practice things are not so simple. Some acid sites still appear more active than others. Poisoning experiments, with bases such as NaOH or

47 NH3, indicate that only a few acid sites have to be poisoned to markedly reduce the catalytic activity [65, 98). Furthermore, the displaced framework aluminium may play a role in the overall acidity and activity. As discussed earlier, there is considerable controversy as to whether the nonframework alumina species enhance the acidity, possibly by the generation of "superacid" Bronsted sites [99, 100), or poison the strongest Bronsted acid sites by cation interaction [86, 101), or lead to increased coke production due to their Lewis acidity [60), or only block the zeolitic channels, reducing the catalyst activity but having no effect on the product selectivity [92).

The value of the unit cell size can be used to calculate the number of alumina tetrahedra present by several methods and the known distribution of sites with n-NNN can be used to calculate ratios of strong acid sites to weak acid sites [9, 102). Calculations of this type show that gasoline-oriented catalysts, which equilibrate in use at a unit cell size of 24.50 ± 0.05, would favour hydrogen transfer reactions over cracking reactions. This combination of predominantly weak sites with relatively few strong sites gives low olefin yields and relatively high coke yields due to extensive hydrogen transfer. [3] However, gasoline yields and conversion level would be expected to be high due to the high number of relatively weak sites available.

Conversely, unit cell size (UCS) equilibration at low values (<24.35 A) gives a zeolite with relatively few but predominantly strong sites. Here the value of the cracking to hydrogen transfer ratio is high and the catalyst would give low coke yields with relatively less gasoline. However, more olefins would be observed in the gasoline and lighter

(C3 and C4 olefins) range due to the overall suppression of hydrogen transfer relative to cracking. Catalysts using zeolites which equilibrate in the 24.30 A ± 0.05 range are appropriate as gasoline octane

48 enhancement catalysts or as catalysts which enhance light olefin yields for alkylate or ether manufacture (reformulated gasoline).

Catalysts containing zeolites which equilibrate at intermediate U.C.S. levels (e.g., 24.40 ± 0.05 A) are useful in resid cracking applications where a blend of high activity (and stability) along with good coke selectivity is needed. In this instance, site strength and number is balanced so that the "difficult to crack" portions of the resid feed can be converted but with minimal quantities of catalytic coke formed. Since the resid feed itself contains high concentrations of coke-forming molecules (condensed ring aromatics, asphaltenes, Conradson carbon residue), this can be advantageous.

Thus, catalytic cracking activity is a function not only of the total number and type of acid sites, but also the density and distribution of acid site strengths. It is also dependent on the accessibility of the acid sites, since the reactant molecule has to be able to reach the required active site to react. To make matters more complicated the geometry around the active site has been proposed to play a significant role in the acid site strength [103]. The larger the T-0-T angle (T= tetrahedral atom) (i.e. Si-0-AI angle), the stronger the Bronsted acidity.

2.5 DEACTIVATION

The classic definition of a catalyst is a substance which alters the rate at which a chemical reaction occurs, but is itself unchanged at the end of the reaction. It is a practical reality, however, that catalysts deactivate over time. Catalyst life may be as short as a few seconds, as in fluid catalytic cracking (FCC), or as long as several years for ammonia synthesis but, inevitably, the catalyst will need regeneration or replacement [104].

49 As mentioned in Section 4, there are four desirable properties of a catalyst in fluid catalytic cracking: activity, selectivity, stability and accessibility. The FCC world would be a lot more transparent if these four indices would move in the same direction upon changing a single process variable e.g. temperature. However this is not the case and compromises have to be made all the time. The problem is compounded by the fact that the catalyst undergoes rapid deactivation, which affects all four indices in different ways [105].

Deactivation of a cracking catalyst will almost certainly reduce the activity. However deactivation of the catalyst may or may not affect the stability, accessibility or selectivity. Thus no general rule can be given for the trade-off between activity and selectivity. When a gain in selectivity is achieved by selective poisoning of the catalyst, it will almost certainly be paid for by a decrease in activity [105].

The three main types of catalyst deactivation are poisoning, fouling, or thermal degradation (sintering). Table 2 below summarises the deactivation types.

Table 2: Main types of cracking catalyst deactivation

Deactivation Reversible Irreversible Catalyst Aging Hydrothermal Catalyst Poisoning Coke, N, s, 0 Na, V, Ni, etc. (Polars) Catalyst Fouling Coke deposits Metal deposits

Deactivation can be reversible such as fouling with coke and polycyclic aromatics [106, 107], or poisoning by basic and polar molecules e.g.

50 nitrogen compounds which are readily adsorbed on to the catalyst acidic sites. This leads to an instantaneous but reversible deactivation [108, 109].

Deactivation can also be irreversible such as that caused by sintering [104]. Irreversible catalyst poisons (or deposits) can even influence the catalyst during the first passage through the reactor, but are not (easily) removed during the stripping and/or regeneration stages. Examples are the heavy metals in industrial feedstocks such as vanadium and nickel and other poisons such as alkali components, iron and copper [11 O, 111 ].

Deactivation due to physical causes, such as loss of surface area from sintering, abrasion, etc, are normally a minor problem. Chemical causes are more serious and can be considered as poisoning of the active acid sites.

The major impurities associated with industrial feedstocks are species containing basic nitrogen, basic metals such as sodium, and aromatic compounds of nickel, vanadium and iron. Sodium, as brine, is typically removed from the feed via a desalter prior to the FCC unit. Failure of this unit can have disastrous effects on FCC catalysts. The basic nitrogen content of the feed varies tremendously and can be reduced by pre-treatment in extreme cases, but lower levels can be handled by catalyst design. Nickel, vanadium and iron generally require special additives or advanced catalyst design [104].

Implicitly one would suggest that catalyst activity decreases linearly with the concentration of poison. This is far from the case. Because only a few active sites are involved in catalysis, it is not surprising that catalysts can be poisoned by only small amounts of strongly adsorbed molecules. Hence deactivation is often very rapid in the presence of

51 only small amounts of a catalyst poison. In a few cases, poisons can act favourably by removing least-active sites or sites responsible for loss of product by unwanted secondary reactions. In the latter case the rate of disappearance of reactant may not drop and the rate of appearance of product may even increase upon the introduction of the poison. Deactivation of FCC catalysts does not only yield a drop in activity, but usually also a change in selectivity.

2.5.1 THERMAL DEACTIVATION

Thermal deactivation is a general term referring to both thermal and hydrothermal forms of catalyst deactivation. The thermal deactivation of a catalyst is a permanent deactivation that occurs at very high temperatures. Activity loss from thermal effects is caused by the reorganisation of the catalyst structure at these high temperatures. Thermal deactivation effects are usually more severe than those caused by pure hydrothermal (steam) deactivation and start to occur at about one third to one half the melting point of the catalyst [18]. Hydrothermal deactivation, also a permanent loss of catalyst activity, is the more prevalent mechanism for activity loss during the regeneration cycle. Hydrothermal deactivation results from zeolite dealumination and subsequent crystallinity loss, as well as pore size shifts and surface area loss in the matrix [18].

2.5.2 POISONING

For zeolites, site poisoning and pore blockage are the most commonly observed deactivation mechanisms [112-119]. Site poisoning is due to irreversible adsorption of catalysts poisons on the active sites. For

52 active site poisoning, there are three limiting models, uniform site poisoning, selective site poisoning and pore mouth poisoning [120].

In the uniform site poisoning model, deactivation is assumed to occur uniformly throughout the catalyst particle and the poison deactivates all sites identically. As uniform poisoning is likely to occur when the diffusion rate of the poison in the zeolite is large compared to the rate constant of the poisoning reaction, it is rarely observed in catalytic cracking. [121 ]. The uniform poisoning model predicts that the activity would decrease linearly with loss in the number of acid sites.

Selective site-poisoning applies if the active sites are inhomogeneous, i.e., some sites are more active than others and the poison deactivates different sites differently. Super acidic sites, for example, have been proposed in zeolites with non-framework alumina [122-130]. Alternatively, the acid strength is thought to be determined by the silica to alumina ratio where the acid strength of isolated acid sites, i.e., those with no second nearest neighbour alumina, have higher acid strengths and catalytic activities than those non-isolated acid sites [131-133]. If these more active sites are selectively poisoned, then rapid deactivation occurs after only a small fraction of the total active sites are poisoned. Alternatively, selective coking might occur preferentially at one type of acid site, e.g., Bronsted or Lewis. In the selective poisoning model, the catalytic activity decreases more rapidly than the total number of active sites poisoned, resulting in a change in the overall acid strength or acid type distribution upon deactivation.

In pore mouth poisoning, all acid sites are assumed to be uniformly distributed but the rate of reaction is much more rapid than the rate of diffusion. Thus the area around the pore mouth quickly becomes deactivated by direct poisoning, further reducing diffusion into the pore. As with the selective site poisoning model, the pore mouth poisoning

53 model predicts that the activity will decrease more rapidly than loss in the number of active sites [134]. Although in FCC the rate of reaction is much more rapid than the rate of diffusion, cracking catalysts usually do not have a uniform distribution of active sites [122-130]. Thus it is unlikely that fluid catalytic cracking catalysts will undergo his type of poisoning.

2.5.3 FOULING

Catalyst deactivation by fouling (commonly referred to as pore blockage) is also very prevalent in fluid catalytic cracking. The predominant reason for this is that zeolite pores are only slightly larger than the reactant molecules, and only a few atoms of carbon or other foulant may be required to effectively block pores. One type of pore blockage is pore mouth plugging. In this model, essentially all foulant is deposited near the pore mouths and very few of the active sites are actually covered by the foulant. The activity declines more rapidly than the number of deactivated sites and the diffusion rate decreases as the pores become increasingly blocked. This model does not assume or require direct poisoning of acid sites. The relationship between cracking activity and measured acidity depends on whether access of the acidity probe molecule is also restricted [134].

In addition to pore mouth plugging, pore blockage could also occur further inside the crystal, or the coke could block channels or intersections deeper inside the crystal. For a zeolite with a three dimensional channel network, like H-USY, the effect of a blocked pore would be to poison only those sites within the immediate supercage. As a result, the activity would be proportional to the number of remaining and accessible active sites [134].

54 Of the potential models for deactivation in catalytic cracking, the selective site poisoning and pore mouth plugging model are the most likely models accounting for a large decrease in activity with a small decrease in the total number of unpoisoned acid sites.

The selectivity of a catalyst is usually changed by the types of deactivation listed above. A selectivity change caused by a change in acid density, acid strength distribution, or conversion can be classified as an intrinsic selectivity change, while a selectivity change caused by a change in catalyst pore structure can be classified as a shape selectivity change [135].

Although there are some researchers in catalysis who prefer not to become involved in studying deactivation, the process is critical to overall performance. Due to the complexity of industrial feedstocks and the difficulty of executing experiments with tight material balances, fundamental understanding of catalytic cracking has lagged far behind that of other catalysed processes. To date, most of the published studies on catalytic cracking have been aimed at understanding the behaviour of pure light hydrocarbons over model catalysts. Data of an industrial nature is scarce.

2.5.4 COKING

Deactivation can be due to impurities in the feed or from the actual cracking reaction [104]. Associated with the latter is coke formation. In fact the primary cause of cracking catalyst deactivation is the accumulation of "coke" on the catalyst surface [136]. The term "coke" is a generic description, designed to describe a variety of carbonaceous deposits, _typically high molecular weight hydrocarbon residues [137]. The chemical nature of these carbonaceous deposits may vary widely,

55 from graphitic carbon to condensed polymers. Coking of the catalyst in catalytic cracking continues to plague both the researcher and the plant operator and as such coke formation on silica-alumina catalysts has become of particular interest [138].

The mechanism of coke formation is complex and only partially understood. Different reactants can form different types of coke on the surface of a given catalyst and, on different catalysts, the same reactants can give different cokes [104].

There are four types of coke associated with FCC catalytic cracking [104]: 1. Conradson coke which is inherent to the feed and typically less than 5% of the total coke. 2. Contaminant coke, which is produced from metal contaminants in the feed that are deposited on the catalyst. Nickel and vanadium are the major contaminant metals contributing to this type of coke formation. The overall contribution to the total coke is dependent on the feed used. 3. Catalyst-to-oil coke which results from poor stripping of useful products from the catalyst after the cracking reaction, and which then undergo further secondary cracking reactions. 4. Catalytic coke which is the coke that is formed as an inevitable part of the cracking process.

It should be stressed that coke is not "all evil". In fact in a fluid catalytic cracking unit, a certain amount of coke on the catalyst is necessary to maintain the heat balance between the exothermic regeneration reaction and the endothermic cracking reaction. The ideal amount of coke on catalyst will be determined by the particular design constraints of each FCC unit. Coke is also known to be a source of hydrogen for the saturation of olefins in the cracking reactions, leading to a more

56 stable gasoline. This has been demonstrated by monitoring the ratio of carbon to hydrogen in the coke, which increased as conversion increased, while product olefinicity decreased [139].

2.5.4.1 Coke Formation

The mechanism for coke formation is still unknown but high molecular weight compounds are known to be built up by a series of polymerisation-condensation reactions, especially from alkenes [140- 142] and from polyaromatics [142-147]. In the case of alkenes it is due to their rapid transformation through carbenium ion bimolecular reactions (oligomerisation, alkylation and hydrogen transfer) while for polyaromatics it is due to their slow diffusion in the pores owing to the strong adsorption of these basic molecules on the acid sites. Subsequent dehydrogenation/ordering reactions may lead to coke, whether or not the surface is catalytic [148]. Cyclisation reactions along with ~-elimination and hydride transfer can lead to polymers, polycyclics, and graphitic species. [104] Coke formation occurs slowly from alkanes, naphthenes, and monoaromatics, whose transformation into alkenes and polyaromatics is slow. The formation of these coke maker molecules is then the limiting step of coke formation [149].

The major parameters affecting the formation of coke are the temperature, acidity and pore structure of the catalyst as well as the type of hydrocarbon. The acidity and pore structure influence both the reactions involved in the formation of coke molecules and their retention. Acidic catalysts are known to favour coking [17]. Thus the stronger the acid sites the faster the reactions and the slower the diffusion of basic intermediates hence the faster the coke formation. The density of the acid sites also has an effect on coke formation.

57 Higher site density means more bimolecular reactions, which, in turn, promotes the formation of coke [150].

The size of the space available near the active acid sites has contrary effects on the rate of coke formation: the narrower the space the stronger the steric constraints in the formation of the bulky intermediates of bimolecular reactions (negative effect) but also the greater the concentration effect (marked positive effect on bimolecular reactions). However, at high reaction temperature, the main effect of the pore structure concerns the retention of coke molecules in the pores. The greater the difference between the sizes of cavities (or channel intersections) and of their pore apertures, the easier the steric blockage of coke molecules and the faster the coking [149].

Different hydrocarbons will cause more or less coke formation. Menon [151] concluded that hydrocarbons with molecular structures similar to those of cyclopentadiene, indene, phenanthrene, indan and fluorene are notorious coke-makers. Cooper and Trimm [152] studied the tendency to coke formation for different hydrocarbons on a platinum alumina (Pt / Al 20 3) catalyst and found the following relation: methylcyclopentane > 3-methylpentane - n-hexane > benzene > cyclohexane. Hydrocarbons that have a tendency to form a lot of coke are called coke precursors.

2.5.4.2 Coke Deactivation

During alkane cracking, H-USY zeolite is deactivated by the deposition of carbonaceous deposits (coke). Coke formation on zeolites has been the subject of recent reviews [113, 114] and much research has focused on the mechanisms by which coke causes deactivation. Most postulated deactivation mechanisms fall into two categories, coke

58 deposition directly on active sites, or coke deposited elsewhere but blocking access to active sites [112-113). These two broad categories can be viewed as site poisoning models and pore blockage models, respectively. Site poisoning [112, 115), pore blockage [116) or a combination of the two [117-119, 153] have been suggested as the cause of loss of zeolite activity in several reactions as discussed above in Section 2.5.2 and 2.5.3.

Experimentally, the initial coke deactivation of H-USY is rapid and then decreases more slowly with increasing time-on-stream. Generally, the loss in catalytic activity is greater than the loss in the number of acid sites. Correlating the activity with the amount of coke deposited on the catalyst shows the activity to decease rapidly with the first deposition of coke and then more slowly as the content of coke increases. It is not known whether this initial, rapid deactivation is due to inhomogeneity of the active sites, e.g., sites of higher activity are poisoned first, or is due to other effects such as pore blocking, a decrease in diffusion rate, etc. Various deactivation models can be distinguished, however, if the pore size distribution, molecular diffusion rate, acid site concentration and acid strength distribution are compared for a coke-free and a deactivated catalyst [134].

Coke is a typical example of a reversible catalyst poison. The deactivation influence of coke depends very much on the nature of the coke, its structure and morphology and the exact location of its deposition on the catalyst surface [151, 154, 155]. Coke formation follows the adsorption of coke precursors on the catalyst surface. The adsorption depends on the strength of the interaction and the volatility of the species [110, 111 ].

59 Site Poisoning

Coke in cracking catalysts results in greater catalyst deactivation than can be accounted for by the loss in the total number of acid sites. Cracking of hydrocarbons and deposition of coke occur on active sites near the external crystal surfaces, while active sites in the inner portion of the zeolite particles do not initially participate in the reaction and therefore are not affected by coke. As the coke shell builds up and the active sites near the surface are poisoned, the reaction zone proceeds toward the interior of the zeolite crystal. The rapid loss in catalyst activity is due in part to the longer diffusional path in the coked catalyst [134]. Eventually the coke can block the catalyst pores.

In catalytic cracking the rate of reaction is generally rapid compared to the rate of diffusion and the cracking reactions occur near the external surface of the zeolite crystal with acid sites at the particle interior contributing little to the observed conversion. Coke poisons the active sites near the zeolite surface, thus, increases the diffusion path of the reactant and results in a rapid loss in activity with the loss of a small number of acid sites [134].

When the active sites are covered, the activity can fall fast even with relatively moderate coking. It is important that the coke accumulation occurs slowly so that the pore mouth is not blocked, cutting off the rest of the pore. When coke is formed on a catalyst, the degree of deactivation can vary considerably for different reactions. The coke may block some sites on the catalyst more than others [156] or alter the diffusion rates [156], and cause changes in selectivity.

With site coverage, one coke molecule poisons one active site. However one can associate to this mode of deactivation the inhibition of activity due to a competition for adsorption on the acid sites between

60 reactant and coke molecules. The activity decrease is lower than with site coverage.

A study by Hopkins et al. [134] observed that, during the cracking of hexane on H-USY, small amounts of coke formed produced a large decrease in activity; however, few acid sites were poisoned. Compared to a fresh USY zeolite, the coke deactivated H-USY did not show significant changes in the acid strength distribution, number of acid sites, fraction of Bronsted and Lewis acid sites, pore size distribution, or hexane diffusivity that could account for the large decrease in the cracking activity. Therefore, taken as a whole, their data did not support the selective poisoning deactivation model where highly active sites are preferentially poisoned [134]. They proposed that the rapid loss in activity resulting from a small decrease in the number of acid sites is consistent with a site poisoning deactivation model for a diffusion-limited reaction, i.e., the rate of hexane reaction is greater than the rate of diffusion. For a diffusion limited reaction, the initial hexane cracking and coke deposition occurs primarily at active sites near the external surface of the crystal with little reaction at the particle interior. As deactivation proceeds, the thickness of the coked region grows and hexane diffuses further into the zeolite particle before undergoing a catalytic reaction [134].

Pore Blocking

Pore blockage has generally a more pronounced deactivating effect than site coverage [155, 157, 158]. Indeed one coke molecule can block the access to more than one active site. The higher the active site density the more pronounced this deactivating effect of pore blockage [157-159]. To this mode of deactivation one can associate limitation by coke molecules of the access of the reactants to the pores.

61 Deactivation under diffusional limitation by pore blocking as a result of coking is commonly observed especially when there is a loss of activity much greater than would be expected from the amount of foulant deposited [160, 161]. Since micro porosity is inherent to cracking catalysts, the deposition of foulant at the entrance to or along the length of the pore system results in loss of surface area in excess of that covered by the foulant [162].

Further, if the desired reaction is faster than coke formation and mass transfer is also fast, then coke will be deposited relatively uniformly throughout the pore system. However, if the coke-forming reactions are fast, as is the case with catalytic cracking, coke formation will be concentrated at the entrance to the pore [104].

Pore blockage can result in bigger differences in selectivity of various catalysts, because of changes in pore architecture [109, 163]. The catalysts, which are relatively less accessible for large hydrocarbons, will be more sensitive to pore mouth plugging and blocking [157].

Selectivity Changes

As a result of coke formation in cracking catalysts, the selectivity of the catalyst will change. This is a result of intrinsic selectivity and shape selectivity changes caused by coke formation [135].

Forrisier and Bernard [161] showed that, at constant conversion, the formation of coke decreases the gasoline and coke selectivity while it increases the gas selectivity.

62 Distillate mode FCC units normally operate at low severity for maximisation of the middle-distillate (cycle oil) yield. These units are designed with partial combustion regenerators and operate at lower temperature. Consequently, the coke on the spent catalyst does not burn completely in the regenerator.

The effect of CRC (carbon on regenerated catalyst) on FCC performance has been studied by many authors. In general, it was found that, for zeolite catalysts, there is significant activity and selectivity loss when compared to an amorphous catalyst [164]. Ritter [165] studied the residual coke on a zeolite catalyst and concluded that the active sites of the zeolite were quickly attacked by coke precursors resulting in reduced conversion and gasoline yield and increases in gas and coke yield. He also observed that an increase in carbon on regenerated catalyst produced gasoline with a higher octane number. Venuto and Habib [164] also reported a similar observation. Ritter and Creighton [166] and Ritter [165] hypothesised that the yield of light cycle oil (LCO, boiling range 216-370 °C) could be increased by increasing CRC level, but no quantitative effect was shown. Mandal [167] showed that there was an optimum coke on regenerated catalyst level of around 0.2 - 0.3 wt % for which the distillate yield was at a maximum. Controlling the CRC level on catalyst for maximising middle-distillate yield opens up another direction for improving the plant performance.

The pre-coking of H-USY zeolite has been shown to reduce the number of acid sites, but the decrease was substantially less than the decrease in cracking activity. For example, with 0.55 wt.% carbon, the relative cracking activity was 52% (i.e. a decrease of 48%), while the relative number of acid sites was 80% (a decrease of only 20%). At the highest coke level (4.4% carbon), the relative n-hexane cracking activity was only 6%, while the relative number of acid sites was measured to be 84%. It was reported that relative number of acid sites remains nearly

63 constant at approximately 80% for all pre-coke levels [134]. This may be explained by theorising that the probe molecule used in the method of measurement was small enough so as to not be affected by pore restricting coke, and thus counted acid sites that reactant molecules did not have access to.

2.5.5 METAL CONTAMINANTS

In recent years the addition of distillation residue to catalytic cracking feedstocks has increased the level of contaminate metals on cracking catalysts. The metals of primary concern are nickel and vanadium. At near atmospheric conditions, these metals, when deposited on the catalyst, catalyse dehydrogenation reactions during the cracking process to increase gas and coke yields at the expense of gasoline, cycle oil and other valuable products. The effects of these metals have been recognised for some time [168]. Numerous strategies to deal with the deleterious effects have been developed. These include the addition of nickel passivation agents, hydrotreating to remove metals from the residue FCC feed, and development of metals-resistant catalysts. The extent to which these various strategies have been successful varies with the individual refiner and the operating philosophy and goals.

Vanadium and nickel are both extremely deleterious and are deposited on the cracking catalyst as their host molecule is converted to lighter products and coke. Nickel forms a metal/metal oxide site on the catalyst surface, which promotes dehydrogenation reactions, causing light gas

(H 2) and coke formation [169]. Vanadium has an additional poisoning function in that it forms, in the presence of process steam, a mobile species of vanadium (+5) oxide, which severely deactivates the Y zeolite structure by forming a eutectic with it [170]. While catalyst formulation changes of the primary catalyst can improve nickel and

64 vanadium tolerance and the use of additives can passivate nickel and vanadium, these compounds can still be deleterious when present on the catalyst in excess of 3000 ppm. The metals are normally present in the feedstocks used in Australian refineries at levels of up to 7 or 8 ppm, but will quickly accumulate on the catalyst.

The deactivation of cracking catalysts by feed metals is caused by: pore structure modifications, the activity that these metals show toward secondary cracking reactions, and the irreversible destruction that metals like vanadium (and copper) can cause on the catalyst cracking component (the zeolite).

However, major economic incentives exist for the processing of distillation residue containing high levels of nickel and vanadium in catalytic crackers. The quantity of residue, which can be fed to current units is limited by the cracking catalyst's ability rapidly to deactivate the deposited metals and thus prevent excessive coke and gas formation. It has been found that nickel can occupy several types of sites on the cracking catalyst surface. Some nickel sites have been found to be more active for coke and gas production than others. Vanadium was found to interact with nickel in a manner, which actually inhibits the deactivation behaviour of nickel.

FCC catalyst deactivation by vanadium and nickel can occur via various mechanisms and rates [108, 109, 163, 171-174]. The formation of metal silicates and/or aluminates have been proposed [173, 175-177], which seem to form more easily by reduction and oxidation cycles [178].

65 Basically nickel and vanadium can influence the catalyst via four main reactions [179]:

Mechanism Potency (1) Destruction or neutralisation of catalyst active sites(V > Ni) (2) Dehydrogenation reactions leading to coke and gas formation (Ni > V) (3) Oxidation promotion, leading to a higher C02/CO ratio in the regenerator [180] (Ni > V) (4) Pore mouth blockage (Ni> V)

Nickel is not really an active site poison. Instead it acts as a dehydrogenation catalyst, especially if widely dispersed throughout the catalyst, and generates undesirable coke and hydrogen [104].

Studies show that vanadium causes the destruction of the zeolite in the FCC catalysts by a mechanism of acid attack or solid-solid transformation, as well as additional dealumination of the zeolite framework in the presence of steam and at high temperature. While these effects result in the reduction in crystallinity (zeolite Y content), specific area and unit cell size of the Y zeolite, the reduction in activity was the most pronounced effect as the amount of vanadium in the catalysts was increased.

The average metals levels on equilibrium fluidised cracking catalysts (FCC) have been steadily increasing over recent years, and the physicochemical properties of metal-resistant FCC remains an important topic to refiners. It is known that nickel (like iron) does not quench cracking activity, via catalyst deterioration, even when present at levels up to 3 wt.%. In contrast, above 0.5 wt.% vanadium loading, catalyst properties deteriorate rapidly and, at 1 wt.% vanadium, a near

66 total loss of surface area and (useful) cracking activity is expected [173].

Dealumination is more severe in HY-type materials. The zeolite vanadium resistance seems to decrease when rare earth ions are present and to increase with increasing extra-framework alumina (generated during steam aging). At the high temperatures used for FCC regeneration, oxycations of vanadium (V02+ or V02+) could attack the Al-0-Si bond in H-Y and cause lattice collapse. The amorphous product recrystallises then into mullite and the excess silica forms tridymite, while excess vanadium forms a supported V20 5 like phase. In the presence of steam, hydroxy vanadyl ions and other protonated species resulting from the interaction of V compounds (such as V20 5) with steam, could accelerate alumina removal from the faujasite structure, thus facilitating mullite formation [173].

In rare earth stabilised Y zeolites (such as REHY and CREY crystals), it is believed that Ce4+ ions, present as an oxycerium complex, undergo a redox reaction with oxyvanadyl cations (V02+), and form a stable orthovanadate. Removal of other charge-compensating cations (such as Na+ ions) in the form of vanadates further destabilises the crystal lattice, thus promoting zeolite destruction. Vanadate formation, together with framework dealumination resulting from the acidity generated by the hydrolysis of vanadium compounds during steaming, is believed responsible for the greater ease with which vanadium impurities can destabilise rare earth containing faujasite crystals [173].

When contaminated with vanadium, part of the catalyst surface appears to lose some of its roughness during steaming at 760 °C as viewed by a scanning electron microscope (SEM). The smaller pores (slits) appear blocked, and stacking of what is believed to be V20s layers has also been observed. Pore blockage and crystallinity losses are probably the

67 two main causes of the drastic reduction in surface area and cracking activity suffered by FCC in the presence of vanadium [173).

Khouw et al. [181] report that catalysts contaminated to high vanadium levels are still capable of converting light feeds, but not heavier feeds.

Development of more metals-tolerant cracking catalysts would be aided by further research and subsequent gaining of knowledge as to the effect nickel and vanadium has on cracking reactions especially with respect to industrial catalysts and feedstocks.

68 3 PROJECT OBJECTIVES

The main objective of this thesis is to investigate the effect various types of deactivation and catalyst pre-treatment methods have on the cracking of different feedstocks on an industrial fluid catalytic cracking catalyst with the aim of increasing the efficiency of the fluid catalytic cracking unit.

The primary focus of the project is on maintaining an industrial perspective and as such, the thesis will not delve into deep discussions on individual cracking reactions. Rather it will look to establish theories on the pre-treatment and deactivation of fluid catalytic cracking catalysts and what effects this has on the activity, product selectivity and reaction pathways of the catalyst.

There are two major streams of investigation in this project. The first is the pre-treatment of the catalyst with a light hydrocarbon to enhance product selectivity and the second is the study of common forms of deactivation affecting catalytic cracking processes worldwide.

The study will use an industrially equilibrated fluid catalytic cracking catalyst obtained from the Caltex Lytton Refinery. Amongst other feedstocks, an industrial feedstock will also be used and is sourced from the Caltex Lytton Refinery.

In the first section of the study, a light hydrocarbon will be used to pre treat the catalyst with the expectation that it will produce more light cycle oil, which in the current economic climate is a higher valued product than gasoline. The pre-treatment is expected to selectively deactivate the strongest of the active sites, thereby producing a "milder" catalyst with a more uniform distribution of active sites for the cracking

69 of the main feedstock. The investigation will then seek to explain the results.

The second section of the study will look at deactivating the catalyst through nickel, vanadium and coke impregnation. These contaminants are known to be quite detrimental to the catalyst and a greater understanding of their impact on the catalytic cracking process will allow for a more profitable and efficient operation.

70 4 EXPERIMENTAL

4.1 MATERIALS

4.1.1 GASES

The gases used in this project were supplied by BOC Australia. The gases were dried and purified by flowing through in-line gas purifiers packed with drierite and molecular sieve SA obtained from Alltech Australia.

Table 3 lists the gases used in this project, their purity and application.

Table 3: Specification and application of gases used

Gas Specification Application Helium High Purity Carrier gas for HACH Carle (>99.99%) Series 400 AGC gas chromatograph Carrier gas for Hewlett Packard (HP) 5890 gas chromatograph Carrier gas for Perkin Elmer (PE) Autosystem gas chromatograph Carrier gas for temperature programmed desorption (TPD)

Hydrogen High Purity Flame ionisation detector (FID) (>99.99%) feed gas for HP 5890 gas chromatograph

71 Flame ionisation detector (FID) feed gas for PE Autosystem gas chromatograph Air Inst. Grade Flame ionisation detector (FID) (>99.9%) feed gas for HP 5890 gas chromatograph Flame ionisation detector (FID) feed gas for PE Autosystem gas chromatograph Nitrogen High Purity Inert gas for blanketing catalyst (>99.99%) in reactor and for catalyst stripping. Carrier gas for HACH Carle Series 400 AGC gas chromatograph

Carbon Food Grade Sub ambient cooling media for Dioxide (>99%) HP 5890 gas chromatograph Sub ambient cooling media for PE Autosystem gas chromatograph Beta mixture 2.09 mol% H2 in air Standard calibrating mixture for HACH Carle Series 400 AGC gas chromatograph Ammonia High Purity Adsorbent probe molecule for (>99.99%) temperature programmed desorption (TPD)

72 4.1.2 CHEMICALS

Table 4 gives a list of chemicals used in this project and their applications.

Table 4: Chemicals used and their application

Chemical Application Toluene Solvent for catalyst metals impregnation. Hydrocarbon for pre-coking of catalyst. Nickel naphthenate Dopant for metals impregnation of cracking catalyst Vanadium naphthenate Dopant for metals impregnation of cracking catalyst Ammonia solution (28 mass%) Base for catalyst alkalising Light catalytic naphtha (LCN) Refinery hydrocarbon stream for the in-situ pre-coking of cracking catalyst

2,2,4 tri-methyl pentane (iso- Microreactor catalyst testing octane) feedstock n-Hexadecane (nC16) Microreactor catalyst testing feedstock

Squalane (branched C30 Microreactor catalyst testing hydrocarbon) feedstock Atmospheric distillation residue Microreactor catalyst testing (industrial feedstock) feedstock SORT (simulated distillation Standard for the calibration of retention time) standard. Mixture Hewlett Packard and Perkin of n-paraffins (C5-C40) Elmer gas chromatographs

73 4.2 ACTIVITY AND SELECTIVITY TESTING

4.2.1 REACTOR AND FURNACE SYSTEM

The principal tool used for these studies was the Caltex Lytton cracking microreactor, which is routinely used to help Caltex select the most profitable fluid catalytic cracking (FCC) catalyst for refinery service. These in-house studies have previously given Caltex a competitive advantage in Australia's fluid catalytic cracking field.

The catalysts and feeds were reacted in a modified MAT (microactivity test) fixed bed reactor obtained from the Grace Catalyst Corporation [182, 183]. A picture of the system is given in Figure 3 and a line diagram of the system is shown in Figure 4.

74 Figure 3: Photograph of microreactor apparatus used

I .

75 Figure 3: Photograph of microreactor apparatus used

75 The furnace consists of three zones each individually controlled by Shinko FCS temperature controllers. The temperature setpoints of the controllers, used for the experimentation, from top to bottom were 538°C, 510 °C, and 503°C giving a catalyst bed temperature of 500°C ±1 °C. The catalyst bed temperature was measured using a calibrated Fluke 52 K/J thermometer.

The glass reactors used were obtained from DC Scientific Glass in the United States. They are made from borosilicate glass and have an inner diameter of 1 .5 cm. The reactor is packed with glass wool, catalyst, and glass beads before being inserted into the furnace.

The method of packing the reactor is as follows. Firstly 0.1 g of glass wool is inserted into the bottom of the reactor. This is to prevent the catalyst from flowing out through the bottom of the reactor. Then a measured amount of catalyst ranging from 0.1 - 5g is accurately weighed out using a Mettler AE 200 analytical balance and poured in a free flowing manner into the reactor. Another wad of glass wool (0.2g) is placed in the reactor on top of the catalyst to separate the catalyst from the glass beads. Finally 1Og of 1 mm glass beads are added to the reactor. Their purpose is to preheat and evenly distribute the feedstock as it is injected prior to contact with the catalyst. A packed reactor is illustrated in Figure 5.

77 Figure 5: Photograph of reactor tube showing feed inlet, preheat section and catalyst.

78 The feed tube, preheater tube and thermocouple, collectively referred to as a "dead man", are inserted into the reactor with an "o" ring spacer. The dead man is clamped to the reactor so as to form an airtight seal.

Once the reactor is placed in the furnace, a 2 ml syringe, containing an accurately weighed amount of feed, is attached via a three-way valve to the feed line of the dead man. A nitrogen purge line is also attached to the three-way valve. Nitrogen is allowed to flow freely through the reactor while it equilibrates to the reaction temperature.

The rate of nitrogen flow is controlled at 50 ml/min using a Brooks 0152i flow meter and controller.

The syringe housing is attached to a syringe driver capable of different rates of delivery. For the purposes of this thesis, the rate of feed delivery was kept constant at 1.5 ml/min. To achieve different levels of feed conversion, the catalyst to oil ratio was changed by altering the amount of catalyst or by changing the feed injection time.

4.2.2 PRODUCT COLLECTION SYSTEM

The products of the cracking reaction (gas, liquid, and solid) must all be collected for accurate mass balancing, yield analysis and product determination.

A liquid product collection flask was attached via an "o" ring and clip to the bottom of the reactor. The flask was placed in a cold bath of saturated sodium chloride solution, which had been chilled to a temperature of -16 °C. The sodium chloride was required to prevent the freezing of the water. The outlet of the liquid product collection flask was connected to a cotton wool trap, which served to trap any escaping

79 liquid vapour. The cotton wool trap was joined to a gas burette via a 70 cm glass transfer line.

The gas burette was also filled with a saturated sodium chloride solution. In this instance the sodium chloride was required to minimise the amount of hydrocarbons in the product mix lost via dissolution in the water. A glass bulb attached to the burette, with a flexible piece of hosing, served to equilibrate collected gases to atmospheric pressure.

4.2.3 CATALYST

The catalyst used in this project was obtained from the Caltex Refinery at Lytton in Queensland, Australia. The sample was taken from the FCC unit as equilibrium catalyst (e-cat) on the ih of November 1997. The equilibrium catalyst is predominantly Access 908 ABP from Akzo Nobel. In the unit this catalyst would have experienced cracking temperatures of between 500 °C and 515 °C and regeneration temperatures of around 730 °C in the presence of air and steam.

The catalyst was calcined in a Carbolite Furnace (CSF 1100) for 20 hours at 650 °C to drive off water and remaining coke on catalyst. The catalyst was then cooled and stored in a desiccator. This calcining step was performed prior to additional pre-treatment of the catalyst.

Table 5 and Table 6 shows the basic chemical and physical properties and the standard reaction characteristics of Access 908 ABP as sampled from the fresh catalyst hopper and the regenerator (e-cat) on the ?1h of November 1997.

80 Table 5: Chemical and physical properties of fresh and equilibrium catalysts used

Catalyst grade Fresh Catalyst Equilibrium Catalyst Catalyst type Akzo Access - Akzo Access - 908ABP 908ABP Chemical Analyses

A'203 44.3 wt% 45.6 wt% Rare Earth 2.39 wt% 2.22 wt% Loss on Ignition * 12.5 wt% Na 0.21 wt% 0.44 wt% Sb - 1653 ppm Ni - 5236 ppm V - 792 ppm Cu - 87 ppm Fe - 1.01 wt% Carbon - 0.39 wt% Physical Analyses Particle Size Distribution 0-19 µm 1 wt% Owt%

19-38 µm 10 wt% 2wt%

38-75 µm 36wt% 40wt%

75-106 µm 26wt% 30wt% 106 µm + 27wt% 28wt% Average Particle Size 76 µm 82 µm

Surface Area 270 mL/g ** 103 mL/g Meso Surface Area Not measured 57 mL/g

81 Apparent Bulk Density 0.67 glee** 0.73 glee Pore Volume 0.40 cclg ** 0.39 cclg Micro Pore Volume Not measured 0.021 cclg

Notes:* LOI is the weight% of volatile material in the catalyst lost after heating to 1OOO cc ** Sample calcined for 1 hour at 566 cc

Table 6: Standard reaction characteristics of the fresh and equilibrium catalysts.

Catalyst grade Fresh Catalyst Equilibrium Catalyst Catalyst type Akzo Access - Akzo Access - 908ABP 908ABP MAT conversion * 85wt% 65wt% Hydrogen Factor 2.9 Carbon Factor 1.8

Notes: * Micro Activity Test: Catalyst pre-treatment: 3 hours @ 566 cc and 2 hours at 788 cc, 100% steam, 1 atmosphere

Explanations for the above chemical, physical and catalytic properties appear below.

The total surface area of a catalyst is made up of the matrix and zeolite (crystalline silica alumina) surface areas. For a given catalyst, surface area is related to activity, and a decrease in surface area will inevitably result in a decrease in activity [184]. Higher matrix surface areas combined with active alumina results in increased conversion of the

82 heavy distillate residue. The comparison of the fresh catalyst and equilibrium catalyst surface area shows a differential of 167 m2/g with a corresponding loss in MAT activity of 20. Thus it can be seen that for an accurate representation of a commercial operation, the equilibrium catalyst must be used over the fresh catalyst. This has often not been the case in the past with many researchers opting for pure laboratory prepared catalysts.

The mesa surface area is the surface area of pores > 20 A (zeolite pores) and < 100 A (macropores) in diameter. Mesopore formation occurs in the zeolite crystal during the dealumination step and is important in improving zeolite accessibility. The pore size of the zeolite is 7-8 A and its surface area may be estimated by total surface area minus - mesa surface area. The surface area is measured using micro porisimetry instruments utilising various suitably sized probe molecules.

Fluid catalytic cracking catalyst is typically >95% silica alumina. Rare earth is exchanged into the zeolite to improve the ability of the catalyst to withstand hydrothermal deactivation. The rare earth on the Akzo Access catalyst is predominantly lanthanum (-69 mass% of total rare earth used) with the remainder being cerium and neodymium. As rare earths also enhance hydrogen transfer reactions, they alter the cracking selectivity of the catalyst. Hence a catalyst low in rare earth will produce more LPG (liquefied petroleum gas) of a less saturated nature than a catalyst high in rare earth. The selectivity to gasoline (petrol of boiling point range 18 °c - 218 °C) will be lower, however the octane number (RON and MON) will be higher. Low rare earth catalysts are higher in surface area than high rare earth catalysts [184].

Sodium is measured because of its detrimental effect on cracking catalysts. It is an alkali metal, which neutralises acid cracking sites. Sodium will also cause the zeolite to lose its crystal structure, which

83 leads to increased deactivation. The effect of sodium, vanadium and high temperature can lead to very rapid zeolite destruction. Potassium and calcium have the same effect on acidity but must be present in much higher levels [184]. Typically, in a refinery, the crude oil is passed through a desalter prior to entering the catalytic cracker. This serves to remove a large majority of the sodium and sodium levels on the equilibrium catalyst are usually not a problem.

Antimony is added in liquid form to the reduced crude prior to reaction. It is a nickel passivator and as such, it serves to reduce the dehydrogenating effects of nickel on the catalyst. The passivator forms nickel antimonites by reaction with nickel in the reactor and nickel oxide in the regenerator [184].

Nickel is a contaminant in the feedstock that deposits on the catalyst and which catalyses dehydrogenation reactions [184]. Levels as high as 13,000 ppm have been observed in some South-East Asian refineries. Nickel at such levels needs to be heavily passivated with antimony.

Vanadium is also a feed contaminant, which deposits on the catalyst. It deactivates the zeolite by forming eutectics with it, a phenomena much increased by sodium and temperature. It also has a dehydrogenation activity 20-25% that of nickel [184].

The Micro Activity Test (MAT) is a catalyst activity test using standard conditions and standard feedstock. The test employs a technique very similar to that used for the catalyst / feed reactions in this thesis. In particular, the standard feedstock is injected over the catalyst and the products are analysed for conversion. The micro activity test serves to separate catalyst effects from feed and process changes, thereby providing a reasonable comparison between different catalysts. The MAT conversion for each catalyst is defined as 100 minus the weight

84 percent of products boiling above 221 °C, which can also be written as 100 minus the fraction of products boiling in the clarified oil and light cycle oil ranges. The kinetic conversion is defined as MAT conversion / (100 - MAT conversion) and is also used as a measure of the catalysts activity. The hydrogen factor and the coke factor indicate the propensity of the catalyst to form hydrogen and coke respectively. The hydrogen selectivity is represented by the gas factor, which is defined as the hydrogen to methane weight percent yields ratio. The coke selectivity is represented by the coke factor, which is calculated by dividing the weight percent coke yield by the kinetic conversion [185).

Apparent bulk density is a standardised measurement of catalyst density. It is determined by pouring the sample into a funnel and letting it free fall into a graduated cylinder. The bulk density is then calculated from the volume and weight of the free fallen sample. Changes in the apparent bulk density of an equilibrium catalyst can be used to detect operational problems in an FCCU; for example an increase in apparent bulk density for a given equilibrium catalyst coupled with a decrease in surface area is indicative of thermal deactivation.

Pore volume is the void volume of the catalyst microsphere and is measured using probe molecules and porisimetry instruments.

4.2.4 INDUSTRIAL FEEDSTOCK

In addition to the pure hydrocarbon feedstocks used in this thesis such as 2,2,4 tri-methyl pentane, n-hexadecane and squalane, an industrial feedstock was also used to closely represent commercial operation. The feedstock was carefully selected by looking at the five year feed forecast for Caltex's Lytton refinery and matching, as close as possible,

85 the average of each critical property with a feedstock that was to be processed by the refinery.

The feedstock selected was processed by the refinery on the 2ih of September, 1996. A sample of this feedstock was obtained. It had an initial boiling point of 173 °C and a final boiling point of 750 °c.

The breakdown of the crude types present in the feedstock mix are listed in Table 7.

Table 7: Crude oil types and mass percentage present in feedstock as sampled

Crude Oil Type Percentage in feed (mass%) Ardjuna 46.8 Bach Ho 0.5 Belida 0.2 Clarified Slurry Oil 0.1 Cossack 12.8 Jackson 2.5 Kerapu 3.5 Kutubu 11.9 Northwest Shelf Condensate 3.7 Skua 18.0

86 4.2.5 PRE-TREATMENT OF CATALYSTS

In the experimentation of this thesis, both the feedstock and the catalyst were changed I modified in order to evaluate the system under investigation. The catalyst was modified in many ways including in-situ coking, in-situ and ex-situ ammonia pre-treatment, in-situ heat pre treatment, metals pre-treatment, ex-situ pre-coking, and metals impregnation on a sand I catalyst split system.

4.2.5.1 In-Situ Coking

As the name implies, in-situ methods of pre-treatment occur whilst the catalyst is in the reactor. The equilibrium catalyst was loaded as per the previously described method and a preliminary injection of feed is made over the catalyst at 500 °C. The feedstock used for this preliminary injection was varied between light catalytic naphtha and reduced crude to investigate whether different preliminary feedstocks made a difference to the catalyst's selectivity.

4.2.5.2 In-Situ Ammonia Pre-Treatment

Again for this method of pre-treatment the equilibrium catalyst was pre loaded into the reactor. A concentrated ammonia solution (25%) was injected over the catalyst in varying quantities from 0.2 ml to 2.5 ml. This was performed as a preliminary investigation into the effects of an alkaline solution on the acidity of the catalyst and hence the product selectivity. Once it was determined that the in-situ ammonia pre treatment affected the selectivity of the catalyst, the ex-situ method of ammonia pre-treatment was used as a more reliable and repeatable method.

87 4.2.5.3 Ex-Situ Ammonia Pre-Treatment

Ex-situ ammonia pre-treating of the catalyst was also performed using a 25 mass% solution of ammonia (sp. gr. 0.91) from BDH Chemicals.

The equilibrium catalyst was soaked in the ammonia solution prior to being placed in an 80 °C oven for 16 hours. The catalyst was then removed and placed in a Carbolite muffle furnace for calcining using the program in Table 8.

Table 8: Temperature program used for the calcining of ammonia pre treated equilibrium catalyst.

Function Temperature (°C) Time (hours) Hold 250 °C 1.5 hours Ramp and Hold 650 °C 5.5 hours

The catalyst was then stored in a desiccator and used for all experiments requiring the ammonia pre-treated catalyst.

4.2.5.4 In-Situ Heat Pre-Treatment

Heat pre-treatment of the catalyst was performed in-situ. The equilibrium catalyst was placed in the reactor and with nitrogen flowing at 50 ml/min, the furnace temperature was increased from 500 °C to 600 °C, held briefly, then returned to 500 °C for reaction at this temperature. The purpose of this method of pre-treatment was an effort to try and shift the Bronsted to Lewis site ratio. As mentioned previously in the literature search some researchers report that increasing the

88 temperature of a cracking catalyst under nitrogen serves to increase the Lewis to Bronsted ratio of acid sites.

4.2.5.5 Metals Impregnation

The metals pre-treatment of catalysts was carried out using the Modified Mitchell method [265].

The catalyst was pre-treated with nickel and vanadium metals. For each metal pre-treatment, the equilibrium catalyst was soaked in a solution of the appropriate metal naphthenate. The nickel naphthenate solution was 6.35 mass% nickel in toluene. The vanadium naphthenate solution was 0.872 mass% vanadium in toluene. Enough additional toluene was added to just form a slurry. The catalyst was then placed in an 80 °C oven for 16 hours to drive off the excess toluene. The catalyst was removed from the oven and placed in a Carbolite muffle furnace. A program was initiated as shown in Table 9.

Table 9: Temperature program used for the calcining of nickel and vanadium pre-treated catalysts.

Function Temperature (°C) Time (hours) Hold 150 °C 0.5 hours Ramp and Hold 250 °C 3.5 hours Ramp and Hold 650 °C 5 hours

The catalysts were then used in experiments requiring nickel or vanadium impregnated catalysts.

89 4.2.5.6 Metals Impregnation on a Sand/Catalyst Split System

For this system the feedstock was reacted over metals impregnated sand in a single or two stage set-up.

Since the experiments utilising this system were designed to be directly compared to the metals impregnated equilibrium catalyst, the metals impregnation of the sand was conducted in exactly the same fashion as described above.

In the single stage set-up, metals impregnated sand was placed in the reactor and feedstock reacted over it. The purpose of these experiments was to ascertain what effects the actual metals had on the feedstock independent of any cracking reactions. The control experiment used was the reaction of feedstock over untreated sand. This gave an indication of extent of thermal cracking on the starting feedstock. For the single stage sand reactions, the catalyst to oil ratio was calculated to be the ratio of sand to feedstock.

In the two stage set-up, metals impregnated sand was placed in the reactor and above it, untreated equilibrium catalyst was also measured into the reactor. The purpose of this system was to evaluate whether the metals affected the selectivity in the same way when they are separate to the catalyst surface as compared to the metal impregnated equilibrium catalyst.

4.2.5. 7 Ex-Situ Pre-Coking

The ex-situ pre-coking of the equilibrium catalyst was performed by Akzo Nobel using a short contact time reactor and toluene as a feedstock for rapid, uniform coking. Since the in-situ method of pre-

90 coking is not very repeatable, the ex-situ method was used to evenly distribute coke on a good volume of catalyst. The time that the catalyst remained in the short contact time reactor, referred to as time on stream was varied to achieve different amounts of coking. Three levels of coke on catalyst were achieved, 1.4 mass%, 2.5 mass% and 3.4 mass% as measured by a Leco carbon analyser.

4.2.5.8 Summary of Catalyst Pre-Treatment Methods.

The types, methods and levels of catalyst pre-treatment are summarised in Table 10. Their analysis is discussed in the results and discussion section.

Table 10: Types, methods and quantities of catalyst pre-treatment applied

Type Method Amount Coke ln-situ pre-coking with LCN As defined by catalyst in fixed bed at 500 °C to oil ratio Coke ln-situ pre-coking with LCN As defined by catalyst in fixed bed at 600 °C to oil ratio Coke ln-situ pre-coking with As defined by catalyst reduced crude in fixed bed to oil ratio at 500 °C Ammonia ln-situ pre-treatment using 0.2 ml 25% ammonia solution Ammonia ln-situ pre-treatment using 2.5 ml 25% ammonia solution

Ammonia Ex-situ pre-treatment using 5% NH 3 in solution saturation by ammonia

91 solution

Ammonia Ex-situ pre-treatment using 25% NH 3 in solution saturation by ammonia solution

Heat In Situ heat pre-treatment 600 °C under N2 Nickel Modified Mitchell method 5.0 mass% (5,000 ppm) Nickel Modified Mitchell method 10.0 mass% (10,000 ppm) Nickel Modified Mitchell method 20.0 mass% (20,000 ppm) Vanadium Modified Mitchell method 2.5 mass% (2,500 ppm) Vanadium Modified Mitchell method 5.0 mass% (5,000 ppm) Coke Pre-coke in short contact 1.4 mass% time reactor Coke Pre-coke in short contact 2.5 mass% time reactor Coke Pre-coke in short contact 3.4 mass% time reactor Nickel Modified Mitchell method 10 mass% (10,000 impregnated ppm) sand

As mentioned previously several feedstocks were used in the experimentation. These included 2,2,4 tri-methyl pentane, n-decane, n hexadecane, squalane (branched C30 compound), light catalytic naphtha and the refinery industrial feedstock (reduced crude).

92 The experiments were designed to simulate as closely as possible the real industrial process. As a result if a pre-treated catalyst requires more catalyst (increased residence time) to achieve the same conversion, then this will also occur in the industrial situation. Therefore the results obtained are reflective of the "real" process.

4.2.6 PRODUCT ANALYSIS

The feeds were reacted over each catalyst type at varying catalyst to oil ratios in order to generate product curves and selectivity information and similar conversions. After each reaction the products were collected (gas, liquid and solid) and were analysed using various gas chromatographs and a Leco carbon analyser for the coke determination. The methods of analysis are outlined below.

4.2.6.1 Gas

The fraction of product collected in the gas burette was analysed using a Hach Carle gas chromatograph and a Perkin Elmer Autosystem gas chromatograph.

The Hach Carle Series 400 AGC gas chromatograph uses a series of columns, molecular sieves and valving to separate the hydrocarbon gases. Specifically, it uses in combination, a 7 ft column packed with molecular sieve SA, 6 ft Haysep Q, 3½ ft CW15 400 Porasil C, 2 ft 1500 Carbowax B152EE CW1540 on Chrompaw, and a PTA 3.5.

The chromatograph has the benefit of being able to analyse for hydrogen content in the gas by employing a hydrogen transfer tube. The quantity of nitrogen and oxygen can also be determined using this

93 chromatograph. The determination of oxygen in the gas sample gives an indication of whether the system is leaking or not. The set of results for a particular reaction are discarded if the oxygen reading is too high, since it implies a possible loss of products.

The Carle was controlled by desktop computer utilising Perkin Elmer Nelson's Turbochrom Navigator Ver. 4.1 software via a Perkin Elmer 900 series interface.

A standard of 2.09 mol% hydrogen in air was used to calibrate the Carle gas chromatograph. This also serves to provide a means for mass balancing the gas fraction of the cracked products.

The Perkin Elmer Autosystem employed a fused silica column of type BP1. The length of the column is 50m with i.d. of 0.32 mm and 1 um film thickness. The detector used is a flame ionisation detector (FID), which allows the detection of hydrocarbons as small as methane. The carrier gas pressure was set at 65 kPa.

The PE Autosystem was also controlled by the Turbochrom Navigator software via a Perkin Elmer Link 600 series interface.

4.2.6.2 Liquid

The liquid portion of the products as trapped in the cold trap were analysed using a Hewlett Packard 5890 gas chromatograph and the Perkin Elmer Autosystem.

The Hewlett Packard 5890 employs a stainless steel column of 6 feet in length and inner diameter of 0.085 in. The support is a chromsorb K/HP of mesh range 80/100 with liquid phase of OV-101 present at a level of

94 10%. The HP 5890 uses a flame ionisation detector for detection of hydrocarbons. A flame ionisation detector response factor of 1.0 was used to calculate weight percent selectivities for all hydrocarbon species.

The temperature program used was as follows: Initial temp O °C hold for 1.5 min, increase at a rate of 15 °C per min to 360 °C and hold for 1O min. The injector temperature and detector temperature were both set at 370 °C. The helium carrier gas was set at a rate of 35 ml/min and confirmed using a calibrated flow meter.

Results from the HP 5890 were processed using a Hewlett Packard 3393A Integrator. The calculation for boiling ranges was performed within the integrator using an ASTM method for simulated distillation.

The Perkin Elmer used the same column as specified previously. The injector and detector were both set at 340 °C. The carrier gas pressure was set at 85 kPa. The temperature program was as follows: Initial temp -12 °C, hold for 3 min, increase at 2 °C per min to 80 °C then increase at 1O °C per min to 31 O °C and hold for 30 min. The total run time for the program was 102 mins.

4.2.6.3 Solid

The coke on the catalyst and glass beads was quantified using a Leco carbon analyser. This step was necessary to ensure a complete mass balance of the reaction.

In the study, conversion is defined as: all cracked products divided by (all cracked products plus unconverted feed), unless otherwise noted.

95 4.2. 7 CATALYST CHARACTER/SA TION

4.2.7.1 Temperature Programmed Desorption

Temperature programmed desorption measurements were made, using a rig set up at the University of New South Wales, to determine the acidity of the catalyst. These were done by mixing the catalyst with alpha alumina in a 1 :9 ratio and placing the catalyst in a quartz reactor. The catalyst was then saturated with a measured amount of ammonia. It is assumed that the amount of ammonia used for the saturation is proportional to the acidity of the catalyst. The catalyst was then heated to 800 °C on a programmed temperature ramp. The desorption of ammonia from the catalyst was measured using a thermal conductivity detector. All ammonia remaining on the catalyst at 800 °C was assumed to be adsorbed only on the most acidic sites.

IR spectroscopic measurements were made on some of the treated catalysts however analysis was extremely difficult, especially when dealing with the ammonia pre-treated catalysts. This often led to poor or inconclusive results. Therefore the results from these experiments were not used in the thesis discussion.

4.2.7.2 SEM / EDAX Machine and Method

To investigate the adequacy of the metals impregnation, analyses such as scanning electron microscopy (SEM) and elemental analysis by SEM EDAX were performed.

The sample preparation for analysis by SEM involved attaching some catalyst particles to 10mm circular aluminium "stubs" using double sided carbon tape.

96 Specimens that are inherently conductive do not need to be coated. However cracking catalyst samples are non-conducting and will usually charge up as electrons from the beam accumulate on the surface. Charged areas will deflect electrons emerging from the surface and will distort the image, making it unsightly or even unacceptably distorted. All insulating specimens are best prepared by coating the specimen with a conducting material.

Gold was chosen to coat the catalyst samples. This is because gold is a good conductor, produces a strong secondary electron signal (has a good stopping power for electrons), is non-crystalline, and is unreactive. A 20 nm layer of gold was applied to the catalyst surface by sputtering from a target in a cold discharge using a Polaron sputter coater.

After coating, the samples were placed, individually, in a Cambridge S- 360 scanning electron microscope and viewed at different magnifications, usually at 20 kV.

The sample preparation for elemental analysis was a little more involved and required the setting of the catalyst in a resin base. A sample of each catalyst was placed in an empty plastic cylindrical container with a removable bottom. A mounting epoxy resin for embedding the catalyst was then added to the cylinder. Once the epoxy had hardened, immobilising the catalyst on the bottom surface, the sample was removed from the cylinder and the surface of interest was ground down on wet and dry sand paper to liberate the catalyst from the epoxy. The process was conducted in different steps, with different grades of paper and monitoring the surface with an optical microscope. Finally, with the help of a rotating microcloth, the sample surface was polished with a 1O µm diamond paste and then final polished with a

97 0.25 µm diamond paste. Using this technique, the catalyst particles were sectioned, revealing the inner surface, which was then analysed using elemental analysis.

The samples were analysed in a Cameca SX50 microprobe, which uses wavelength dispersive and energy dispersive X-ray spectrometers to analyse the X-ray spectra emitted by electron-irradiated specimens.

4.2.8 STATISTICAL SIGNIFICANCE OF RESULTS

Due to the nature of the equipment used to perform the cracking reactions in this thesis, it was necessary to take great care in the experimental phase of the work. Each point that was used as a representative result in figures, tables, and discussion was obtained by an average of about ten experimental runs, which differed by no more than ± 2%. This was done by ensuring that each experimental run had a mass balance of 100% ± 2%. If this was not the case, the result was not reported nor was it used in the averaging process.

98 5 HYDROCARBON ON CATALYST PRE-TREATMENT

5.1 INTRODUCTION

Many researchers that study catalytic cracking use catalysts that are of a pure zeolitic nature. This is done in order to simplify the product spectrum. Zeolites are synthetic crystalline materials with an ordered tetrahedral structure of silica and alumina ions. Although zeolites are responsible for most of the cracking activity in a catalyst [184], thereby providing useful cracking information, their selectivities and changes in selectivity are not reflective of commercial operation.

This is because commercial catalysts, along with a zeolite, contain complex binders and a matrix. Binders are used in the fluid catalytic cracking catalyst to bind the matrix and zeolite components into a single homogeneous particle. The matrix forms the non-zeolite part of a cracking catalyst. It serves both physical and catalytic functions. Physical functions include providing support for the zeolite, particle integrity and attrition resistance, acting as a heat transfer medium and allowing the free flow of feed and products into and out of the microsphere. Catalytic functions include heavy oil upgrading and resistance to poisons. Both these materials together with proprietary formulations make commercial catalysts much more complex, with respect to reaction pathways, than the pure zeolites.

Despite this, very few researchers have incorporated the use of commercial catalysts in their studies and fewer still have examined equilibrium catalysts direct from a commercially operated unit [186]. This would indicate that an area of catalytic cracking, essential to the overall picture of the field, has not yet been investigated. Thus the

99 purpose of this thesis was partly to investigate catalytic cracking from the other end of the spectrum, not from the purist individual chemical reactions point of view but rather from the practical, commercial application of catalytic cracking. Thus the thesis will look at how catalysts can be affected in commercial operation and to study shifts in activity and selectivity to ascertain why they occur.

It was decided that, to model the process and reactions in industry, it was necessary to use a commercial catalyst that had been equilibrated in an industrial operation. The catalyst used in this study was obtained from Caltex's Lytton Refinery in Queensland and served to shed more light on the cracking of feedstocks over an industrial catalyst. The catalyst sampled was Akzo Nobel's Access - 908 ABP fluid catalytic cracking catalyst, which was still being used at the refinery, for the fluid catalytic cracking operation, at the time of writing.

Naturally the catalyst must be sampled from the fluid catalytic cracking unit after regeneration and before reaction, so that the catalyst does not become heavily coked. Typically this is done from the e-cat line. The catalyst taken from this section of the reactor is termed equilibrium catalyst. Equilibrium catalyst (e-cat) derives its term from the fact that the catalyst has become "equilibrated" with the system. This "equilibrated" catalyst may contain catalyst that has just been added to the system through to catalyst that may have been in the system for over 200 days. Equilibrium catalyst differs from fresh catalyst predominantly in metals levels, which accrue from the contaminant metals in the feedstock, and surface area reduction due to age. Both these properties affect the conversion and selectivity capabilities of the catalyst. The basic chemical and physical properties of Access 908 ABP as sampled from the fresh catalyst hopper and the regenerator (equilibrium catalyst) on the same day are displayed in the section on Experimental Technique.

100 In the commercial operation of a fluid catalytic cracking unit (FCCU), equilibrium catalyst enters the reactor from the regenerator via the e-cat line and slide valve. The equilibrium catalyst enters the bottom of the reactor, known as the riser, at around 700 °C. The catalyst is propelled up the vertical riser by a steam injection point at the bottom of the riser. About 2 m from the bottom of the riser, the catalyst is contacted with feed which is introduced to the riser via feed injectors at a preheat temperature of 230 °C. The cracking reaction takes place. At the top of the riser, the feed, catalyst and product mixture of around 500 °Center into a riser termination device (RTD). This device is used to separate the hydrocarbons from the catalyst. A picture of the Caltex Lytton Refinery's fluid catalytic cracking unit is shown in Figure 6 and a schematic of the unit is shown in Figure 2 (Literature Search).

101 Figure 6: Photograph of Caltex Lytton's fluid catalytic cracking unit.

102 5.2 PRE-INJECTION OF LIGHT CATALYTIC NAPHTHA

One of the aims of this study was to evaluate whether the addition of a light hydrocarbon stream to the bottom of the riser would increase the light cycle oil (LCO I diesel blendstock) yield which, in the current economic climate, is a higher valued product than light catalytic naphtha (gasoline).

Presently, all fluid catalytic cracking units in Australia inject only steam into the bottom of the riser to lift the catalyst. It was hypothesised that the addition of a light hydrocarbon fraction such as light catalytic naphtha could deposit a little coke on the catalyst thereby changing the selectivity of the catalyst. The expected mechanism was that the coke would be deposited on the most active cracking sites, thereby preventing the overcracking of the feedstock into lighter components and thus increasing the yield of light cycle oil (boiling range -220 °C - -370 °C) relative to light catalytic naphtha (C5 °C - -220 °C).

The expected coke formation from the cracking of light catalytic naphtha may be deposited on the catalyst in a number of ways. Both thermal and catalytic cracking will lead to the formation of coke on the catalyst.

The thermal cracking of hydrocarbons, also known as pyrolysis, produces non-selective coke. This leads to a uniform blanketing effect on the catalyst, and ultimately to a reduction in overall conversion. The rate of pyrolysis is greatly increased at temperatures above 500 °C. Certain materials will produce more coke than others. Multi ringed aromatics such as anthracene and asphaltenes will produce more coke than straight chain hydrocarbons such as decane.

103 If the coke was catalytic coke, meaning that it formed as a product of the reaction process, one would expect that it would cover the catalyst at the site of the reaction first (namely the acid sites) thereby deactivating the catalyst in a selective fashion. The preliminary study was intended to assess if pre-treating the FCC catalyst with a light hydrocarbon results in a change in conversion and if so, whether the deactivation by pre-coking was selective or non-selective.

The injection of naphtha was expected to produce a multiple-fold effect. Certainly the effect discussed above was expected, combined with the fact that the injection of light catalytic naphtha meant that a blendstream of relatively less value (than light cycle oil) was being consumed to hopefully form liquefied petroleum gas (LPG), another high value product.

The injection of light catalytic naphtha into the bottom of the riser would possibly simulate steam cracking of naphtha. Steam cracking is a non catalytic process whereby steam plus naphtha at temperatures between 700 °c and 900 °c will result in the production of light olefins. With the presence of the fluid catalytic cracking catalyst the cracking should be more severe and more selective.

5.2.1 LIGHT CATALYTIC NAPHTHA CHARACTERISATION

Light catalytic naphtha is predominantly paraffinic and olefinic with some degree of aromaticity. A PIONA analysis was conducted on the light catalytic naphtha using a gas chromatograph at Caltex's Kurnell refinery in Sydney. PIONA is an abbreviation for Paraffins, Isa-paraffins, Olefins, Naphthenes and Aromatics. The results appear in the table below.

104 Table 11: PIONA analysis of light catalytic naphtha (LCN)

Hydrocarbon Group Quantity in LCN (Mass%) Normal Paraffins 10.55 lso Paraffins 25.79 Naphthenes 7.85 Normal Olefins 13.37 lso Olefins 15.80 Cyclic Olefins 2.97 Total Olefins 32.14 Aromatics 14.65 Polynaphthenes 0.33 B.P. >200 °C 8.70 Total 100.00 C/H ratio 6.0260

5.2.2 INDUSTRIAL FEEDSTOCK CHARACTERISATION

Feed is also an important consideration when trying to model industrial processes. Although thousands of papers have been written describing the reaction of model compounds, there were very few papers found reporting on the cracking of an industrial feedstock on a commercial catalyst [187] and no papers found describing the cracking of an industrial feedstock over a pre-coked industrial catalyst. In order to most closely model the process, it was decided to use an industrial feedstock from the Caltex Lytton refinery. The feedstock chosen was representative of the refinery's five-year feedstock forecast in terms of distillation and critical cracking properties.

105 The critical properties used were specific gravity, Conradson carbon and UOP-K factor. Conradson carbon is the amount of residual carbon in an oil as measured after a destructive distillation. A high Conradson carbon level in an FCC feedstock produces high delta coke and increases the regenerator temperature. The UOP-K factor is a measure of the aromaticity of an FCC feed. Feeds with UOP-K factors above 12 are paraffinic and below 11 .5 are considered aromatic. The formula for UOP-K is the CABP divided by SG, where CABP refers to the cubic average boiling point. FCC feedstocks with a lower UOP-K value are harder to crack since they are more aromatic in nature.

Both the specific gravity and distillation properties are important parameters in the crackability of the feedstock. If the distillation end point of a feedstock is high and the specific gravity is low, then the feed is likely to be more paraffinic and therefore easier to crack. If the distillation end point is low and the density, relatively high, then the feed is most likely aromatic, which is harder to crack.

The physical properties of the sampled feedstock mix and the forecasted feedstock are displayed in Table 12.

106 Table 12: Physical and chemical properties of the industrial and forecasted feedstock

Physical property 5 yr forecast Feed 27/9/96 Specific gravity 0.9002 0.9024 Con carbon (wt%) 2.55 2.61 UOP-K 12.01 11.97 Hydrogen wt% 12.85 N/A Carbon wt% 86.2 N/A Nickel (ppm) 2.5 2.8 Vanadium (ppm) 0.4 0.6 Total nitrogen (ppm) 865 943 Basic nitrogen (ppm) 302 317 Sulfur (wt%) 0.13 0.15 Sodium ppm 1.2 0.9

Initial boiling point 0 c 172.8 167.3 5% point °C 257.2 254.1 50% point °C 405.0 407.9 95% point °C 627.8 632.9 Endpoint °C 750.0 755.6

The above table shows the similarity between the forecasted feedstock and the sampled feedstock, and thus provides an excellent base for representing commercial operation.

107 5.2.3 STANDARD AND PRELIMINARY REACTIONS

The first step in this study was to establish a set of results from a standard set of reactions of the industrial feedstock on the equilibrium catalyst.

There are many ways to compare results from catalytic cracking reactions. In industry the most commonly used representation is that of conversion as a function of catalyst to oil ratio. Since conversion is a measure of activity, one can easily measure the difference in activity of two catalysts if compared at the same catalyst to oil ratio. This is also known as comparing at constant catalyst to oil ratio. Since it is necessary to gain an adequate understanding of the catalyst and quite often necessary to determine what catalyst to oil ratio will lead to the same conversion between catalysts, a series of reactions must be performed at varying catalyst to oil ratios [188].

Once an adequate set of reactions has been performed, it will allow the comparison of the product selectivities of the different catalysts at constant conversion. This is important because ultimately it will determine which catalysts form the most of a single product at a given conversion.

The second step was to determine whether pre-treating the catalyst with a light hydrocarbon stream does in fact affect the activity and/or selectivity of the catalyst, thereby affecting the product yields. The initial experiments involved a preliminary reaction of light catalytic naphtha (LCN) over the equilibrium catalyst at 500 °C prior to reacting industrial feedstock over the now pre-treated catalyst.

108 Figure 7 shows the relationship between conversion and catalyst-to-oil ratio (defined as mass/mass) for the cracking of reduced crude at 500

°ᆰC and the effect that pre-treating the catalyst has on the activity.

Figure 7: Cracking of reduced crude on equilibrium catalyst and the effect of hydrocarbon pre-treatment

-0~ (/) (/) 75 ro E -C 0 ·en.... 65 Q) > C X Untreated equilibrium catalyst , 0 ~-L1 () + LCN pre-treated at 500 C 55 ~---· A LCN pre-treated at 600 C O Reduced crude pre-treated A. 45 0 5 10 15 20 25 Catalyst to oil ratio

Conversion in the above graph refers to 100 - LCO - CLO. That is 100 mass% minus the fraction boiling in the light cycle oil range (220-372

°ᆰC) in mass% minus the fraction of the product boiling in the clarified oil range (372+ °C) in mass%.

As expected the catalyst that was pre-treated with the light catalytic naphtha at 500 °c showed a reduction in conversion when compared with the untreated equilibrium catalyst. At a catalyst to oil ratio of about 7, the conversion of the industrial feedstock on the untreated catalyst

109 was -74 mass% while on the pre-treated catalyst the conversion was reduced to -68 mass%.

This is probably due to coke forming on the catalyst as a result of the pre-treatment with light catalytic naphtha and hence reducing the activity of the catalyst. Although this graph shows that the activity of the catalyst has been affected (indicating some deposition of coke on the catalyst), it does not convey any information about possible changes in selectivity.

Other experiments were performed including pre-treating the catalyst with light catalytic naphtha at 600 °C and also pre-treating the catalyst with reduced crude. It was expected that both of these pre-treatments would deposit more coke on the catalyst. Pre-treating at higher temperatures with light catalytic naphtha increases the severity of both thermal and catalytic cracking, thereby forming more coke on the catalyst. In addition pre-treating with reduced crude, which contains larger branched chain paraffins (which are more reactive than shorter paraffins) was also expected to produce more coke on the catalyst.

The results from these experiment sets were plotted on the same graph. The graph shows that, as expected, the conversion of the industrial feedstock, reduced crude, is reduced further by pre-treating the catalyst with light catalytic naphtha at 600 °C and also by pre treating the catalyst with reduced crude. Both these methods of pre treatment resulted in a greater loss of conversion when compared to light catalytic naphtha pre-treatment at 500 °C. As discussed above, this was expected, as both methods would produce greater amounts of coke on the catalyst surface. This was confirmed by measuring the coke on catalyst after each method of pre-treatment (prior to subsequent reaction) as shown in Table 13.

110 Table 13: Carbon on catalyst as a function of hydrocarbon pre treatment type

Method of pre-treatment Carbon on pre-treated catalyst LCN at 500°C 0.3 mass% LCN at 600 °C 0.7 mass% Reduced crude at 500 °c 0.9 mass%

5.2.3.1 Coke Selectivity on Hydrocarbon Pre-Treated Catalyst

An interesting product yield is total coke. It could be hypothesised that, because the pre-treated catalysts are less active, the coke on catalyst after reaction would be less, since cracking would be less severe - unless there are some other reactions leading to coke formation that are enhanced as a result of the pre-coking.

Figure 8 shows that in fact the coke on catalyst as a result of cracking increases for the pre-treated catalysts.

111 Figure 8: Coke selectivity of reduced crude cracking as a function of hydrocarbon pre-treatment type

11 • 0 10

-~0 Cl) 9 Cl) ro E 8 -Q) ~ 0 7 () Q 6 X Untreated equilibrium catalyst • LCN pre-treated at 500 C 5 ---·---·---·- A LCN pre-treated at 600 C Lo Reduced crude pre-treated ------T-----~- 4 50 55 60 65 70 75 80 Conversion (mass%)

It is interesting to note that although the reduced crude pre-treating only deposits slightly more coke on the catalyst, it promotes coke formation much more so than the light catalytic naphtha pre-treatment. For example at a conversion of about 62 mass%, the reduced crude pre treated catalyst produces almost one mass% extra coke when compared with the light catalytic naphtha pre-treated catalyst at 500 °C. This indicates that either the extra small amount of deposited coke promotes increased coke selectivity or that the coke from the reduced crude pre-treatment is different in its composition, how it deposits on the catalyst or a combination of the above.

112 The hypothesis that the increased coke selectivity is due to the reduced crude coke being different in composition or depositing in a different way is reinforced when viewing the total coke yield for the light catalytic naphtha pre-treated catalysts. That is, at the same conversion for all tested conversions, the light catalytic naphtha pre-treated catalyst forms a similar amount of total coke regardless whether the pre-treatment is carried out at 500 °C or 600 °c. This shows that, although the difference in coke levels on the two light catalytic naphtha pre-treated catalysts was measured to be 0.4 mass% (double that of the difference between the reduced crude pre-treated catalyst and the light catalytic naphtha pre-treated at 600 °C catalyst) there is no appreciable difference in the total coke selectivity.

Thus it can be concluded that the hydrocarbon type used for the pre coking does affect the coke selectivity of the pre-treated catalyst.

Coke is considered to be a hydrogen deficient residue. To that end, it would make sense that there must be more hydrogen in the remaining products, either as molecular hydrogen or as hydrogen rich compounds. Molecular hydrogen can be directly measured, and the paraffin to olefin ratio of similar carbon number groups gives a good indication of an increase or decrease in hydrogen saturated compounds.

5.2.3.2 Hydrogen and Dry Gas Selectivity on Hydrocarbon Pre Treated Catalysts

Hydrogen may be produced by the protolysis of the tertiary C-H bond [35, 56, 189]. By tabulating the hydrogen yield for the equilibrium (untreated) catalyst and the reduced crude pre-treated catalyst at similar conversions, it can be seen that the hydrogen yield is greater for

113 the pre-treated catalyst. This information is shown in Table 14 at two different conversions (-62 mass% and -65 mass%).

Table 14: Effect of hydrocarbon pre-treatment on hydrogen and dry gas selectivity when cracking reduced crude

Conversion Hydrogen Dry gas (mass%) (mass%) (mass%) Untreated e-cat 65.6 0.40 1.71 Pre-treated with 65.1 0.44 1.88 reduced crude Untreated e-cat 62.0 0.35 1.55 Pre-treated with 61.9 0.39 1.63 reduced crude

Dry gas refers to the total yield of hydrogen, methane, ethane and ethene. The increase of these yields also correlates well with the increase in the coke selectivity of the hydrocarbon pre-treated catalysts.

On close examination of the yields, it can be noted that the increase in the yield of hydrogen is about 10 mass% over that of the untreated equilibrium catalyst. This does not account for the -35 mass% increase in coke yield of the reduced crude pre-treated catalyst over the untreated equilibrium catalyst. This would indicate that not all the excess hydrogen ends up as molecular hydrogen, but some would also be present in the hydrocarbon products. This can be observed by comparing paraffin to olefin ratios of the products. This is discussed later in this chapter.

114 The above findings already show that not only is the activity of the catalyst reduced when pre-treated with a hydrocarbon stream, but that the selectivity is also affected. This is evidence that the coke deposited by the pre-treatment reaction does not just uniformly cover the catalyst, but rather selectively blocks certain acid sites. Although the selectivity change of hydrogen and coke yield is important, what is of most concern is the selectivity of usable products, namely light cycle oil, light catalytic naphtha and light gases. As discussed previously, the Australian market is under pressure to produce more and more light cycle oil ( diesel fuel constituent) at the expense of light catalytic naphtha. Let us examine what the hydrocarbon pre-treatment has done to the light cycle oil selectivity.

5.2.3.3 Light Cycle Oil Selectivity on Hydrocarbon Pre-Treated Catalysts

It was hypothesised that the hydrocarbon pre-treatment would either uniformly deactivate the acid sites, which has been found not to occur, or more probably would selectively deactivate the stronger sites, thereby reducing the severity of the cracking and reduce the overcracking of light cycle oil into lighter components. While we do see a reduction in the catalyst activity we do not see an increase in the light cycle oil yield. Instead there is an unexpected finding. By comparing at constant conversion, (which gives an indication of change in selectivity independent of activity changes) it can be seen that the yield of light cycle oil (LCO) is reduced for the pre-treated catalysts when compared with the untreated equilibrium catalyst. Figure 9 shows the reduction in light cycle oil, at constant conversion, when the catalyst has been pre treated with a hydrocarbon.

115 Figure 9: Effect of hydrocarbon pre-treatment on light cycle oil selectivity when cracking reduced crude

35 -r------x Untreated equilibrium catalyst • LCN pre-treated at 500 C 30 ~~--~~---"',_---~---j A LCN pre-treated at 600 C o Reduced crude pre-treated

I -~ OU) 25 CJ) ro E -0 20 () ...J

10 +------,.------__. 45 55 65 75 85 Conversion (mass%)

Another interesting observation that can be made from Figure 9 is that the light cycle oil yield is reduced by -2.5 mass% at conversions between 60 and 80 mass%, despite the method of pre-treatment. This is interesting because it indicates that there are only a limited amount of sites affecting the light cycle oil selectivity that are being affected by the pre-coking. Once these sites have been coked, any additional coking does not result in increased selectivity shifts. As these results show that the same shifts in selectivity are observed independent of the method of coking, it can be inferred that the type and method of coking does not bear significant impact on the cracking mechanism that is being influenced.

116 The above finding seems to be in contradiction to the previously discussed coke selectivity, where it was found that the type of coking does bear significant impact on the coke formation mechanism. It can only be concluded that these mechanisms are independent of each other.

A possible explanation for the reduction in light cycle oil yield is that the hydrocarbon pre-treatment of the equilibrium catalyst has caused the deactivation of sites responsible for the formation of light cycle oil. This could be due to a multitude of reasons including a weaker sites hypothesis which says that the pre-treated catalyst is not able to crack off large light cycle oil molecules and, instead, cracks off smaller hydrocarbons boiling in the LPG and light catalytic naphtha range. Alternatively, the formation of stronger overall acidity hypothesis suggests that the pre-treated catalyst actually increases the over cracking to light catalytic naphtha and lighter hydrocarbons. The actual mechanism cannot be determined from the above information and more analysis and additional experimentation is required to clarify the situation.

An alternative hypothesis is that the sites that are deactivated from the pre-treatment are indeed sites that promote the polymerisation of light hydrocarbons to form light cycle oil. This hypothesis is addressed later in the thesis when discussing the cracking of feed hydrocarbons lighter than light cycle oil over pre-treated catalysts and analysing the product selectivity to determine whether more or less material is formed that has a boiling point higher that the starting feedstock.

Naturally the final explanation for the decreased selectivity of light cycle oil on the pre-treated catalyst may be a combination of the above two mechanisms, or may be due to an as yet unpredicted mechanism.

117 As mentioned above, one possible explanation is that perhaps the same amount of light cycle oil is being formed initially on the pre-treated catalyst as compared with the untreated catalyst, but that this fraction is being further cracked (or overcracked) to the gasoline fraction and/or lighter fractions. Indeed it can be observed that, for these sets of experiments, accompanying the decrease in light cycle oil selectivity was an increase in hydrogen and dry gas yields (as discussed above) as well as a slight increase in light olefins. The iso-butane yield decreased while the yield of C3 and C4 normal paraffin yields remained unchanged. This can be seen in Table 15 below (all yields are in mass%).

Table 15: Product selectivity comparison between untreated e-cat and reduced crude pre-treated catalyst when cracking reduced crude

Reduced Reduced Catalyst Untreated Untreated crude pre- crude pre- type e-cat e-cat treated treated Conversion 62.00 61.94 65.61 65.05 (mass%) H2 0.35 0.39 0.40 0.44 Dry gas 1.53 1.63 1.71 1.88 C3 0.72 0.70 0.76 0.77 C3= 3.24 3.48 3.53 3.78 nC4 0.43 0.40 0.44 0.42 iC4 2.36 1.76 2.50 2.12 C4= 4.11 4.32 4.08 4.36 LCN 43.16 42.36 45.33 43.18 LCO 28.81 27.58 26.54 25.20 CLO 9.19 10.47 7.85 9.75 Coke 6.43 8.07 7.26 8.64

118 5.2.3.4 Paraffin to Olefin Ratio Shifts

This leads into the analysis of paraffin to olefin ratio of the products. It can be quite difficult to measure the paraffin to olefin ratio of larger hydrocarbons. This is due to the vast range of isomers formed and the necessity to identify each of these so that they can be classed into a certain carbon number group. However the paraffin to olefin ratio of smaller hydrocarbons can provide some useful information. Not only can they reveal where excess hydrogen may end up, but they can also indicate changes in the extent of hydrogen transfer reactions.

The paraffin to olefin ratio of light hydrocarbons is observed to decrease on the pre-treated catalyst, when compared to that of the untreated equilibrium catalyst, as shown in Table 16 (ratios are displayed as mass% I mass%).

Table 16: Paraffin to olefin ratio of light hydrocarbon products as a function of hydrocarbon pre-treatment

Reduced Reduced Untreated Untreated Catalyst crude pre- crude pre- equilibrium equilibrium treated treated Conversion 62.00 61.94 65.61 65.05 (mass%) C3s 0.22 0.20 0.22 0.20 C4s 0.68 0.50 0.72 0.58

The decrease in paraffin to olefin ratio of light hydrocarbons on the pre treated catalysts indicates that these light hydrocarbons are on average

119 more hydrogen deficient than the light hydrocarbon products of the equilibrium catalyst. This is contrary to what was expected when considering the increase in coke selectivity discussed above. However from this, assuming a hydrogen balance across the system, it can be deduced that the excess hydrogen must end up in the heavier hydrocarbons.

As mentioned above, changes in the paraffin to olefin ratio also indicates a change in the extent of hydrogen transfer reactions. A decrease in the paraffin to olefin ratio indicates a decrease in the hydrogen transfer reactions. As the thesis progresses, the hydrogen transfer mechanism is increasingly discussed as it appears to be a critical factor in selectivity changes.

There are several different mechanisms by which the coke may be acting on the catalyst in the hydrocarbon pre-treatment of the catalysts. One possible mechanism is that the coking of the catalyst may be affecting the pores in the crystalline structure such that the reactant or intermediate molecules are approaching the spatial dimensions of the pore, thereby inducing shape selective catalysis [190]. It is noted in various reports that shape selectivity affects the extent of hydrogen transfer reactions. Haag et al. [191] report that this is due to the difficulty of accommodating the required bimolecular transition state, and not because of any large differences in the diffusivities of the molecular species involved.

Corma et al. [192], Abbot and Wojciechowski [193], and others [20, 191, 194, 195] report that hydrogen transfer and bimolecular processes in general have been shown to be suppressed in the small pores of H ZSM-5 in comparison to H-Y and La-Y cracking catalysts. This may reinforce the argument that the hydrocarbon pre-treatment is leading to

120 the build-up of coke around the pore mouth such that the pore dimensions become reduced.

Indeed the results from the hydrocarbon pre-treating of the catalyst indicate that the selectivity is similar to that of H-ZSM-5 in that on H ZSM-5 zeolites (with smaller pores) the cracking is mainly due to monomolecular reactions, whereas hydride transfer requires a bimolecular transition state and would be sterically hindered. Thus on Y zeolite, or on the equilibrium catalyst, with larger pores, the bimolecular mechanism is favoured, leading to higher paraffin to olefin ratios for the light hydrocarbons.

The results of Wielers et al. [196] show that, with decreasing pore dimensions, cracking proceeds less and less by the bimolecular route. At the same time, there is a decrease in the paraffin to olefin ratio. They reason that propagation reactions, which produce the excess paraffins and hence the high paraffin to olefin ratio, are slowed down by the shape selective effects of smaller pores. Wielers et al. [196] also reported that, as the conversion level increases, bimolecular reactions were reported to become more important and the "cracking mechanism ratio" decreased. Bimolecular reactions in general become more prominent at higher conversions. The presence of olefins, which form carbenium ions on Bronsted sites, increases the probability of disproportionation reactions and reduces the concentration of pristine sites available for protolysis. The observed changes in activity and selectivity with conversion are then to be expected.

5.3 SUMMARY

From the preceding results and discussion it can be concluded that the hydrocarbon pre-treating of the catalyst caused a reduction in catalyst

121 activity when compared with the untreated equilibrium catalyst. In addition a shift in the product selectivity was also observed, such as the increase in hydrogen and coke formation and, more importantly, the unexpected decrease in light cycle oil formation. These effects are attributed to the formation of coke on the catalyst as a result of the hydrocarbon pre-treatment. As the amount of coke on the catalyst was increased, the effects became more pronounced.

Two hypotheses to explain how the coke interacts with the catalyst have been formed. Firstly, that the reduction in the catalyst activity from the hydrocarbon pre-treating of the equilibrium catalyst may be a result of a decrease in overall site density, which would lead to a reduction in hydrogen transfer reactions, thereby shifting the selectivity. Alternatively, or in addition, the shift in selectivity could be a result of pore restriction or pore blocking by coke and/or the active sites are being covered with coke. These hypotheses indicate that only a small amount of coke would be necessary to prevent a large number of active sites from contacting the feed, resulting in a large decrease in activity for a small amount of coke, as was observed from the results.

Three hypotheses were developed to explain the decrease in light cycle oil selectivity as a result of the hydrocarbon pre-treatment of the catalyst:

1 . Deactivation of sites responsible for the formation of light cycle oil from reduced crude 2. Formation of stronger sites that promote the overcracking of light cycle oil into light catalytic naphtha and lighter components 3. Deactivation of sites that promote polymerisation of light hydrocarbons to form light cycle oil

122 Further experimentation would be necessary to unravel which hypotheses are valid and which are not. The first port of call was to deactivate the catalyst without physical modification.

There are a few different ways of pre-treating the catalyst so as to affect the acidity without depositing carbonaceous material on it. One of these ways involves pre-treating the catalyst with heat and under nitrogen. As discussed in the literature review, the heat pre-treatment of the cracking catalyst serves to shift the Bronsted to Lewis site ratio. Another method involves pre-treating the catalyst with an alkali to neutralise the acid sites. Both these methods do not affect the spatial dimensions of the pore and thus will be addressed in the following chapters in an attempt to clarify the hydrocarbon pre-treatment observations.

123 6 HEAT PRE-TREATMENT OF CRACKING CATALYST

6.1 INTRODUCTION

Two major hypotheses arose in the last chapter to explain the activity and selectivity change when pre-treating the catalyst with a hydrocarbon. One hypothesis deals with the coke affecting the spatial dimensions of the pore, while the other shows that the site density may be reduced, thereby affecting the activity and the selectivity. In light of the above, it was decided that a method of affecting the sites without changing the spatial dimensions of the pore needed to be evaluated. Modifying the ratio of Bronsted to Lewis sites has been found to have an impact on the product selectivity. The literature [197] reports that as the temperature of a cracking catalyst is increased, a molecule of water can be released from two Bronsted sites, thereby forming a Lewis site. ------·------This was determined to be a valuable method of modifying the catalyst active sites without altering the pore size. This method has not been used before to evaluate the cracking of industrial feedstocks over industrial catalysts, nor has it been used to evaluate whether hydrocarbon pre-treatment of a cracking catalyst affects its shape selectivity.

Several experiments involved this method of pre-treatment. For the purposes of this thesis this method of catalyst pre-treatment will be termed heat pre-treatment and was accomplished by increasing the temperature of the catalyst to 600 °C and equilibrating with water under nitrogen, before reacting the feedstock over the catalyst at 500 °C as was done with the hydrocarbon pre-treatment experiments. Similarly in the standard reaction the catalyst was equilibrated with water at 500 °C, before reacting the feedstock over the catalyst at 500 °C. In this way the

124 ratio of Bronsted to Lewis sites is related to the equilibration temperatures of 600 °C and 500 °C. It was envisaged that these experiments would serve to clarify whether the shifts that were seen with the hydrocarbon pre-treatment can still be observed independent of pore fouling effects. This study also attempts to clarify whether Lewis sites participate in the cracking reaction, which has been a subject of much contention. The relative importance of Bronsted and Lewis acid sites on solid acid catalysts during cracking processes has been debated for many years [198-201] and the present study provides evidence to help resolve this issue for reaction of paraffins on zeolites.

6.2 ACTIVITY LOSS

Upon heat pre-treatment of the catalyst, reduced crude (cracker feed) was reacted over the catalyst at 500 °C. The initial results showed similar shifts to the coking pre-treatment. The conversion of the reduced crude was less over the heat pre-treated catalyst when compared with the untreated equilibrium catalyst at the same catalyst to oil ratio. At constant conversion, the light cycle oil yield was reduced while the hydrogen and light catalytic naphtha yields increased as was found with the hydrocarbon pre-treatment experiments. Figure 10 shows the reduction in conversion as a result of the heat pre-treatment.

125 Figure 10: Activity of equilibrium catalyst as a function of heat pre treatment when cracking reduced crude at 500 °c

90 -,------~------

~ 80 0 (/) (/) ro E -c: 70 -~0 Q) ~ X Untreated equilibrium catalyst i ()0 60 -+----+---1------j x • !•Catalyst equilibrated with water Lat 600 C

50 -----~---....,..-----.----.----""' 0 5 10 15 20 25 Catalyst to oil ratio

At a catalyst to oil ratio of five, the heat pre-treated catalyst shows a reduction in conversion of about 4 mass%. This compares with the light catalytic naphtha pre-treatment at 500 °C which shows a reduction in conversion of 8 mass% at the same catalyst to oil ratio.

Previous studies show that for the reaction of C6 paraffins on H-Y zeolite, initiation of cracking is followed by an acceleration in rate as product olefins are adsorbed at Bronsted sites [35, 202]. Thus it might be expected in the current experiments, where there is a shift towards higher Lewis to Bronsted site ratios after the heat pre-treatment, that this kinetic phenomenon would be reflected in concurrent changes in product selectivity. In particular, an increase in the proportion of the total reaction proceeding via interaction with Lewis sites (i.e. adsorbed

126 carbenium ions) should be seen if it is assumed that cracking occurs through different pathways on the two types of active site. Alternatively it can be viewed that there is a decrease in the proportion of the total reaction proceeding via interaction with Bronsted sites.

If a correlation is made with the previous experiments involving the hydrocarbon pre-treatment of the equilibrium catalyst, it can be hypothesised that, since the selectivity shifts are directionally the same, the proportion of Bronsted sites is being reduced relative to the amount of Lewis sites. In comparing with the literature, an interesting observation can be made. Abbot [203] reports that initiation of cracking occurs at Bronsted sites on H-Y and amorphous silica alumina. Further, no direct evidence was found for participation of Lewis acid sites on the catalyst framework itself. However, following initiation, reactions on these catalysts are accelerated through a chain process occurring at Lewis sites generated by adsorption of product olefins at Bronsted sites. The resulting change in the dominant cracking mechanisms were found by Abbot [203] to be reflected in the product selectivity by a decrease in formation of molecular hydrogen as conversion increases. This is not seen in the hydrocarbon pre-treatment experiments, or in the heat pre treatment experiments. Instead, an increase in the formation of molecular hydrogen is seen for the cracking of industrial feedstock on hydrocarbon and heat pre-treated catalysts. It can be concluded that this increase in molecular hydrogen came from the decrease in paraffins relative to olefins for light hydrocarbons or from the increase in coke selectivity.

Other literature reports that the Lewis acid sites could withdraw electron density from a neighbouring framework lattice oxygen of protonic sites, thereby enhancing the strength of the Bronsted acid sites [92, 122]. The formation of these enhanced Bronsted sites has been used to explain observed increases in cracking activity [92, 204, 205]. Further, the

127 strong Lewis acid sites of extra framework alumina (AIEF) promote hydride abstraction and help initiate the cracking reaction [202]. There is conflicting data regarding this latter proposal, since extraction of AIEF from H-USY was found by Carma et al. [100] to lower the gas oil cracking activity, while on the other hand, the same extraction was found by Bamwenda et al. [206] to increase the cracking activity of 2- methylpentane.

For the experiments that were conducted in this study with respect to hydrocarbon pre-treatment and heat pre-treatment, there were no increases observed in cracking activity due to "enhanced" Bronsted site strength. However the results do concur with those found by Carma et al. [100] with regard to a decrease in the catalyst activity.

Two questions that need to be answered are:

1. Is there polymerisation occurring, other than the formation of coke and 2. What effect (if any) do olefins have on the activity and selectivity.

The issue of polymerisation became evident in the previous chapter dealing with the hydrocarbon pre-treatment of the catalyst. It was difficult to determine whether the reduction in light cycle oil was a matter of the light cycle oil over-cracking into lighter compounds (hence the increase in light hydrocarbons and light catalytic naphtha), or due to active sites, that promote the light cycle oil formation via polymerisation, being deactivated. It is not known what effect if any the presence of olefins has on this particular system. As mentioned earlier in the chapter, Abbot and Wojciechowski [35, 202] showed that for the reaction of C6 paraffins on H-Y zeolite, initiation of cracking is followed by an acceleration in rate as product olefins are adsorbed at Bronsted

128 sites. It would be interesting to see if the presence of olefins in the feedstock exaggerated the selectivity shifts previously observed.

6.3 CRACKING OF LIGHT CATALYTIC NAPHTHA

One feedstock of an industrial nature that contains olefins and is lighter than light cycle oil is light catalytic naphtha, as used in the hydrocarbon pre-treatment experiments. This made it an excellent feedstock choice for answering the two questions. Light catalytic naphtha is a product of the fluid catalytic cracking process and as a result contains a significant proportion of olefins (32.14 mass% as measured by the PIONA gas chromatography method). Further it is the first product cut lighter than the light cycle oil cut. It was proposed that the light catalytic naphtha would be cracked over the equilibrium catalyst and also cracked over the heat pre-treated equilibrium catalyst as per the previous experiments.

Hopefully the results of the cracking of light catalytic naphtha would provide a two-fold benefit. Firstly, since it contains olefins, it would give a measure of whether olefin presence in the feed exaggerated the shifts observed. Secondly it would indicate whether polymerisation, other than coke formation, occurs during the reaction.

Indeed in the cracking of light catalytic naphtha, for the same catalyst to-oil ratio, there was a reduction in conversion of about 6 mass%. Conversion to products in this case is defined as 100 mass% minus the mass percentage of light catalytic naphtha remaining in the product. These results correlate well with the previous experiments where a marked decline was also reported when the catalyst was pre-treated at 600 °C. Once again it is believed that the effect has been attributed to the loss of water from the catalyst, accompanied by the conversion of

129 Bronsted sites into Lewis acid sites. This is used as evidence to support the idea that Bronsted rather than Lewis acid sites are dominant in initiating the cracking process.

Although several runs were conducted, two in particular were extracted because of their similarity in conversion. The comparison at similar conversions of all products allows the quick identification and quantification of selectivity shifts. The two runs chosen are shown in Table 17. Once again the conversion was determined as 100 minus the light catalytic naphtha (LCN) yield (mass%). The yields (mass%) of individual products were then each normalised by calculating total yield of products (excluding LCN) as 100 mass%.

Table 17: Reaction product selectivity of light catalytic naphtha cracking before and after heat pre-treatment of the catalyst.

Untreated catalyst Heat pre-treated catalyst Conversion (mass%) 49.9 49.9 H2 0.75 2.10 Dry Gas 4.12 15.08 C3 2.42 2.36 C3= 11.69 12.94 NC4 6.92 6.59 iC4 10.38 7.71 C4= 9.81 13.79 LCN 50.09 50.06 LCO 12.43 7.27 CLO 1.72 0.98 Coke 39.76 31.21

130 6.3.1 HYDROGEN AND DRY GAS SELECTIVITY

Starting with the hydrogen yield it can be seen that there is an incredibly large increase in the selectivity of hydrogen on the heat pre-treated catalyst when cracking light catalytic naphtha. The increase in hydrogen selectivity is 180%, compared with the 10% increase that was seen when cracking reduced crude over a catalyst that had been hydrocarbon pre-treated with reduced crude at 500 °c. Dehydrogenation is the undesirable extraction of hydrogen from cracked products. It is known to be promoted by nickel, and to a lesser extent, other metal contaminants on the catalyst. It seems that the cracking of a partially olefinic compound on a heat pre-treated catalyst has greatly increased the dehydrogenation reactions. Thus it can be concluded that the increase of Lewis sites relative to Bronsted sites catalyses dehydrogenation reactions, which are easily promoted by the presence of olefins.

There is also an unusually large selectivity shift towards more dry gas (comprising methane, ethane and ethene) in the order of 260%. It is believed that this is synonymous with decreased hydrogen transfer reactions. A decrease in the hydrogen transfer reactions would increase the amount of olefins in the product spectrum and also increase the amount of dry gas. It would seem that the increase in Lewis sites relative to Bronsted sites has led to the decrease in hydrogen transfer reactions. As discussed in the literature review, hydrogen transfer reactions are heavily dependent on the proximity of the active sites. The more distant the active sites are from each other, the less hydrogen transfer will occur. As the heat pre-treatment of the catalysts converts two Bronsted sites into one Lewis site, the distance between active sites is increased, and therefore a decrease in hydrogen transfer reactions results.

131 6.3.2 OLEFIN SELECTIVITY

Once again it can be seen that the olefins have increased relative to the paraffins leading to a decrease in the paraffin to olefin ratio. The olefins show an increase in yield while the normal paraffins show no change, or a slight decrease while the iso-paraffin shows a marked decrease. This can also be explained in terms of hydrogen transfer reactions, whereby there is a decrease in these reactions leading to a greater proportion of olefins relative to paraffins. The detailed aspects of hydrogen transfer are addressed later in the thesis.

Confirming the above observation is the increased selectivity of C4 olefins on the heat pre-treated catalyst, which seems to form at the expense of the formation of C4 paraffins, in particular iso butane. This is explained using the hydrogen transfer theory as discussed previously.

From the above results, it can be seen that the presence of olefins in the feedstock greatly exaggerates the shifts seen from the heat pre treatment of catalysts when compared with cracking a feedstock that has negligible paraffins in it. This satisfies the first question asked.

The second question asks two things. Firstly are there polymerisation reactions occurring in the cracking of hydrocarbons over an industrial catalyst and secondly if polymerisation is occurring, does the increase of Lewis sites relative to Bronsted sites assist or hinder this reaction.

The results show that most definitely there is polymerisation (other than coke formation) occurring as indicated by the fact that there are significant amounts of light cycle oil and clarified oil forming as a result of the cracking reaction, whether it be on the untreated equilibrium catalyst or the heat pre-treated catalyst. So it can be predicted for the

132 earlier experiments involving the cracking of reduced crude that light cycle oil is being formed as a result of the cracking of the reduced crude, as well as a result of the lighter hydrocarbons polymerising.

Further it can be seen that upon heat pre-treatment of the catalyst, there is a reduction in the formation of light cycle oil and clarified oil via polymerisation reactions. There is a reduction of 42% and 43% for the formation of light cycle oil and clarified oil respectively. This compares with the 4.5% reduction in light cycle oil yield for the reduced crude pre treated catalyst. This provides further evidence that the presence of olefins in the feedstock exaggerates those effects seen when there is a negligible amount of olefins in the feed. The above findings also indicate that Bronsted sites are responsible for polymerisation reactions. The natural progression of polymerisation reactions appears to follow the sequence: light catalytic naphtha to light cycle oil to clarified oil to coke.

There are literature results that suggest an active role of Lewis acid sites of extraframework aluminium species in the deactivation process. In most of these studies, the conclusions are based on differences in the coke material or differences in the cracked gas phase products between zeolite samples containing different amounts of Lewis acid sites. Abbot and Guerzoni [207] compared H-mordenite samples that contained different ratios of Bronsted to Lewis acid sites prepared by pre-treating the zeolite to different temperatures. They found that the alkane/alkene ratio in the products of n-octane cracking was above unity and higher for the sample with a higher fraction of Lewis acid sites. Since a larger alkane/alkene ratio in the product stream implies the formation of more unsaturated adsorbed species, their work suggests that Lewis acid sites accelerate the chain process of cracked products on the catalyst that leads to coke formation. Similarly, they found a correlation of Lewis acidity with a decline in hydrogen production, which they interpreted as evidence that Lewis acid sites

133 accelerate the chain process of cracked products on the surface that leads to coke formation [203].

The above has not been found to be the case in the experiments that have been performed so far in this thesis. The results gained so far have shown that quite the opposite is true, that the alkane to alkene ratio decreases for the catalyst that has a higher fraction of Lewis acid sites. Indeed for C3 on the untreated catalyst (Table 17), the paraffin to olefin ratio is 0.207, and drops to 0.182 for the heat pre-treated catalyst. For C4s the difference becomes more pronounced with 1.764 decreasing to 1.037. For strictly theoretical cracking of a paraffin, one should observe one paraffin and one olefin forming. Thus the paraffin to olefin ratio of the products is unity initially. However the matter becomes substantially more complex as the feed complexity increases (especially with the introduction of olefins), the time on stream increases (which induces cracking of the initially formed products), and the catalyst complexity increases (which can promote or hinder many reactions including hydrogen transfer reactions).

In the above paraffin to olefin ratio comparisons, it can be seen that for C3s the ratio is below unity, yet for C4s the ratio is above unity. This is an interesting observation since it shows that for the C3s on the untreated catalyst, hydrogen transfer is not a prominent reaction, yet for the C4s, hydrogen transfer and isomerisation (which is an easily catalysed reaction) is evident. Looking at the heat pre-treated catalyst, the C4s show a greater reduction in hydrogen transfer reactions has occurred along with a large reduction in isomerisation reactions. It is likely that the reason the C3s show an increase in dehydrogenation reactions is that they are coke precursors already on their way to forming coke via polymerisation.

134 In other studies, the presence of Lewis acid sites has also been shown to affect the coke content of a deactivated zeolite. Moljord et al. [208), reported that a Y zeolite with a large amount of extra framework alumina (AIEF) contained more coke after deactivation than one with much less AIEF, and increased the rate of deactivation [205). The results from the initial experiments in this thesis also show the same to be true, that more Lewis sites increases the coke selectivity. However, the same observations were not found for the cracking of a feedstock containing olefins.

Flego et al. [209) reported that spectroscopic investigation of the coke formed on LaH-Y-FAU catalyst showed that, on a catalyst sample that had been activated to a higher temperature (and thus contained more Lewis acid sites), the coke involved a higher degree of unsaturation. This is also hypothesised to be true for the current experiments, which show a large increase in hydrogen yield, probably from the dehydrogenation of the light hydrocarbons (precursors for coke formation) and from the coke itself.

Thus it would be expected that heat pre-treatment would increase the yield of coke. However the results of the light catalytic naphtha cracking on heat pre-treated catalysts shows the opposite, in that the yield of coke is decreased.

Kuehne et al. [21 OJ, investigated the cracking activities and the acidic properties of H-USY and (H-NH4 )-USY while cracking 2-methylpentane. These two samples differed in that the strong Lewis acid sites of (H NH4 )-USY were poisoned by ammonia. The results showed that the activities of the two samples were very similar, that the cracking rate constant of the (H-NH4 )-USY sample is identical to that of the H-USY sample, and that the deactivation rate constants and the product distributions were very close for the two samples. In other words, the

135 strong Lewis acid sites were found to have very little effect on the cracking activity or the deactivation characteristics of the H-USY zeolite. Thus it was concluded that that the Lewis sites had little to do with the initiation of catalytic cracking.

This thesis has also hypothesised that Lewis sites have little to do with the initiation or propagation of catalytic cracking. The decrease in activity of the heat pre-treated catalyst (catalyst with more Lewis sites) that has been seen in the experiments performed (but were not seen in Kuehne's results) can be explained by the fact that, in this system, the quantity of Bronsted sites have decreased when they are converted to Lewis sites as a result of the heat pre-treatment. In Kuehne's system there was no decrease in the amount of Bronsted sites, and hence no decrease in activity.

Kuehne et al. [21 OJ further hypothesised that their data shows that strong acid sites (either Bronsted or Lewis acid sites) are not important in the cracking reaction. This does not seem to be true for the system examined here. In this thesis it is hypothesised that strong Bronsted sites are important in the cracking reaction. This is further evaluated later in the thesis particularly in the section dealing with the ammonia pre-treating of the catalyst.

It is possible that the different conclusions from these investigations are due to the fact that the measurements did not make direct comparison among zeolite samples that were otherwise identical but differed only in the presence of Lewis acid sites. In some cases, different types of zeolites were used [211) whereas in others, the zeolites differed in the extent of leaching [205, 206, 208) or concentration of Bronsted acid sites [207, 209). In these investigations, other factors (such as pore size or pore blockage by AIEF species) may contribute to the differences in

136 deactivation among samples, in addition to the presence of Lewis acidity.

6.4SUMMARY

In summary, it is seen that heat pre-treatment reduced the activity of the catalyst while increasing the yield of hydrogen and dry gas and of butenes, largely at the cost of loss of yield of isobutene, light cycle oil and coke. Further the paraffin to olefin ratio was found to drop upon heat pre-treatment of the catalyst. These observations are largely explained by a decrease in hydrogen transfer reactions and a decrease in polymerisation reactions due to the decrease in Bronsted sites relative to Lewis sites. In particular the conversion of two Bronsted sites into one Lewis site as a result of the heat pre-treatment decreases the site density, which is hypothesised to promote hydrogen transfer reactions. Further, it is concluded that Lewis sites do not play an appreciable part in the catalytic cracking mechanism, whether in initiation or propagation.

It was also found that the shifts are more exaggerated using light catalytic naphtha as feed, especially the hydrogen and dry gas yields, when compared with the experiments involving a negligible amount of olefins in the feedstock. This may be attributable directly to the presence of olefins, which enhance the catalytic cracking process.

The results from this study show that trends in the product spectrum were the same with the reduced crude feed over the pre-coked catalyst as with the heat pre-treated catalyst. One exception is the yield of coke, which increased for the hydrocarbon pre-treated samples and the heat pre-treated samples, but decreased for the heat pre-treated catalyst when reacting light catalytic naphtha over the catalyst. This is believed

137 to be due to the decreased site density on the heat pre-treated catalyst. The decreased site density will reduce the amount of polymerisation reactions that are required to form the coke. It is believed that this was not observed with the reduced crude feedstock since this feedstock contained no olefins, and in addition had much larger molecules in particular large paraffins, which are known to coke easily.

The results show that the light cycle oil and clarified oil yields are both reduced for the heat pre-treated catalyst. This indicates that there has been a reduction in polymerisation reactions due to the decrease of Bronsted sites. The decrease in Bronsted sites means that there is a reduction in the site density. Site density is critical for hydrogen transfer reactions and polymerisation reactions. A decrease in the site density means that these reactions are more unlikely to occur and products arising from these reactions are less likely to be formed.

The above findings can be correlated with the observations made from the cracking of reduced crude over hydrocarbon or heat pre-treated catalyst, where it was observed that light cycle oil selectivity was also reduced. The above conclusion can be used to confirm one of the initial hypotheses, that the observed reduction in light cycle oil selectivity is partially due to less light hydrocarbons being polymerised back to light cycle oil and heavier molecules. In addition it is possible that the light cycle oil is being overcracked into lighter fractions and that in the first instance, less light cycle oil is being formed from the reduced crude.

Since both the reduced crude and the light catalytic naphtha feedstocks are very complex in nature and therefore produce an incredibly complex product spectrum, it was decided that it would be beneficial to examine some model pure compounds, while still using the industrially sourced catalyst. This would have the advantage of determining if the selectivity shifts seen with the reduced crude and light catalytic naphtha are also

138 evident with a pure compound or are just the result of some contamination in the industrial feedstocks. Industrial feedstocks are contaminated with heavy metals (such as nickel and vanadium) and can affect the product selectivity during cracking by poisoning or deactivating the catalyst. The effect of both these heavy metals are examined later in the thesis along with the effects of coking. An added benefit of using pure model compounds is that they would serve to provide direct correlation with the literature as most previous studies have focused on pure compounds.

139 7 CRACKING OF PURE COMPOUNDS

7.1 INTRODUCTION

In the previous chapter it was determined that it would be beneficial to examine the cracking of some pure model compounds. Several different pure hydrocarbons have been tested. The first studied was the cracking of n-decane and n-hexadecane, as these compounds have been extensively studied in the literature.

7.2 CRACKING OF DECANE AND HEXADECANE

N-decane and the n-hexadecane were individually cracked over the industrial equilibrium catalyst as per the method used in the previous chapter. The feeds were also cracked over heat pre-treated equilibrium catalyst.

As expected, the cracking of n-decane and n-hexadecane over the heat pre-treated catalyst exhibited a decrease in conversion when compared with the conversion over the untreated equilibrium catalyst. This is due to the reduction in Bronsted sites relative to Lewis sites, as Bronsted sites are hypothesised to be the primary cracking sites.

Other previous investigations also have concluded that there is a close correlation between cracking activity for n-octane on H-mordenite at 400 °C and the number of Bronsted sites on the catalyst [212] by examining the influence of ion exchange on activity. Similar correlations between cracking activity for linear alkanes and Bronsted sites have been reported for other catalysts [19,213,214].

140 As in the previous chapter, the most effective way quantitatively to display the large product spectrum is in a table rather than graphical format. For each feedstock, two runs with similar conversions were selected and the results are presented in Table 18. For the most effective comparison with the previous work, the conversion in the cracking of n-decane was calculated as 100 minus the proportion of the products boiling in the light catalytic naphtha range. For the cracking of n-hexadecane, the conversion was calculated as 100 minus the proportion of the products boiling in the light cycle oil range minus the proportion boiling in the light catalytic naphtha range. The two different definitions of conversion in the above two cases is due to the fact that the two compounds have different molecular sizes, with the n-decane and the n-hexadecane boiling in the light catalytic naphtha and light cycle oil ranges respectively. The boiling ranges (light catalytic naphtha, light cycle oil and clarified oil), although not commonly used in the study of pure hydrocarbon cracking have been used in this thesis to allow comparison with the cracking of industrial feedstocks.

Table 18: Reaction product selectivities of decane and hexadecane cracking before and after heat pre-treatment of catalyst.

Decane Hexadecane Yield Untreated Heat Untreated Heat (mass%) pre-treated pre-treated Conversion 15.8 15.8 29.3 29.3 H2 0.016 0.019 0.188 0.245 Dry Gas 1.91 1.64 2.28 2.18 C3 6.25 6.38 5.41 5.26 C3= 12.64 13.52 22.58 22.99 NC4 6.57 6.90 7.11 6.77 IC4 11.70 11.53 18.97 17.48

141 C4= 14.70 16.64 32.07 32.23 LCN LCO 8.21 5.19 CLO 0 0 2.83 2.08 Coke 38.00 38.20 8.56 10.77

7.2.1 SHIFTS IN PRODUCT SELECTIVITY

Initial observation of the results shows similar selectivity shifts to those seen earlier with the hydrocarbon and heat pre-treatment. As anticipated the shifts were not as pronounced as those from the light catalytic naphtha cracking due to the decreased reactivity of straight paraffins, enforcing the hypothesis that the presence of olefins enhances the shifts exhibited between the heat pre-treated and the untreated catalysts.

Both the decane and hexadecane exhibit a reduction in light cycle oil and clarified oil formation respectively on the heat pre-treated catalyst. As discussed earlier this is due to a reduction in polymerisation reactions due to the decrease in acid site density as some of the Bronsted sites are converted to Lewis sites in the heat pre-treatment process.

Interestingly, the coke selectivity increases on the heat pre-treated catalysts when cracking decane and more so when cracking hexadecane. This is opposite to those observations made during the cracking of light catalytic naphtha on the heat pre-treated catalyst, but the same as the observation made in the reduced crude cracking experiments. It has been determined that this is due to the fact that no olefins are present in these feedstocks. It is hypothesised that the coke

142 formed from the cracking of the n-paraffins on the heat pre-treated catalysts is less dehydrogenated than the coke formed in the cracking of light catalytic naphtha on the same catalyst.

This assumption can be made by calculating the increase in molecular hydrogen yield relative to coke selectivity shifts. In the cracking of hexadecane over the untreated and heat pre-treated catalyst, there is an increase in hydrogen yield of 30 mass%, with a corresponding increase in coke yield of 26 mass%. In contrast, in the cracking of light catalytic naphtha over the untreated and heat pre-treated catalyst, there is an increase in hydrogen selectivity of 180 mass% and a decrease in coke selectivity of 21.5 mass%. Thus, assuming that the extra hydrogen comes from the formation of dehydrogenated species (in particular coke) it can be hypothesised that the coke formed in the cracking of the light catalytic naphtha is much more dehydrogenated than that formed during the cracking of hexadecane.

In comparing the selectivity shifts of the light olefins, it can be seen that the shifts exhibited in the decane and hexadecane cracking, although quantitatively much less, are directionally the same as those observed in the cracking of light catalytic naphtha on the heat pre-treated catalyst. The same can be said of the isobutane yields.

In the previous chapter the paraffin to olefin ratio was discussed. Abbot and Wojciechowski [35] made mention of the paraffin to olefin ratio in their studies on the cracking. They noted that the paraffin to olefin ratio was close to unity for the light hydrocarbons in the cracking of n heptane. If the paraffin to olefin ratios of the light hydrocarbons are compared for the cracking of decane, hexadecane and light catalytic naphtha, an interesting trend can be observed. For the cracking of these feedstocks on the untreated equilibrium catalyst, the C3 paraffin to olefin ratios are 0.49 (decane), 0.24 (hexadecane), and 0.21 (LCN).

143 A slight, however statistically significant, decrease in these ratios is seen on the cracking of the feedstocks on the heat pre-treated catalyst, 0.47 (decane), 0.23 (hexadecane), and 0.18 (LCN). From this it can be determined that the extent of hydrogen transfer reactions decreases as chain length increases and as number of Bronsted sites decrease.

The same decrease in paraffin to olefin ratio as the Bronsted sites are reduced can be made for the C4 products. The paraffin to olefin ratio for the cracking of n-decane before and after the heat pre-treatment is 1.24 and 1.11 respectively. As expected, these ratios decrease for the cracking of the n-hexadecane yielding 0.81 and 0.75 for the paraffin to olefin ratio before and after heat pre-treatment. This enforces the hypothesis that hydrogen transfer reactions decrease as Bronsted sites are converted to Lewis sites (reduction in acid site density) as a result of the heat pre-treatment.

In comparison with the literature, the effects of the heat pre-treatment on the cracking of hexadecane were very similar in response to that found by Abbot and Guerzoni in their experiments with n-octane [207].

Abbot and Guerzoni [207] report on the catalytic reactions of n-octane on H-mordenite at 400 °C. Evidence from changes in activity and selectivity with catalyst pre-treatment temperature show that the dominant cracking processes leading to formation of acyclic alkanes and alkenes take place on Bronsted sites. Formation of aromatics and coke, associated with hydrogen transfer and cyclisation processes, are however, enhanced by the presence of Lewis sites on the catalyst surface. This finding of increased coke formation by Abbot and Guerzoni [207] is in accordance with the findings presented above, however it is not believed to be due to the reason given by Abbot and Guerzoni as discussed above [207].

144 As with the experiments with n-decane and n-hexadecane, Abbot and Guerzoni [207] were not able to correlate ratios of branched to linear alkane cracking fragments with the ratio of Lewis to Bronsted sites present when cracking n-octane over H-mordenite. However, they found, as this thesis does, that ratios of alkanes and alkenes can be correlated with changes in the ratio of Bronsted to Lewis sites present. However, this correlation cannot be further extended to give a ratio of cracking processes via protolytic and ~-scission mechanisms.

Brait et al. [215] reported that dehydrogenation can be catalysed by Lewis acid sites and rapidly deactivates with time on stream. The above results show an increase in hydrogen yield indicating, in accordance with Brait et al., an increase in dehydrogenation reactions.

Brait et al. [216] also reported that the product distribution of n hexadecane cracking could be fully explained by reaction pathways identical to those of n-hexane conversion (i.e. protolytic cracking, dehydrogenation and hydrogen transfer). The detailed discussion of reaction pathways is beyond the scope of this thesis. However, given that the yields of n-C3 and n-C4 differ in the direction of the shifts when cracking decane and hexadecane over untreated and heat pre-treated equilibrium catalysts, it would seem unlikely that the reaction pathways of a smaller hydrocarbon can explain the product distribution of a larger hydrocarbon.

Brait et al. [215] reported that octane numbers were proportional to iso/n-paraffin ratio and the paraffin/olefin ratio of n-hexadecane cracking. No investigation was carried out of the octane numbers of the product. However in measuring the paraffin to olefin ratios and the iso to n-paraffin ratios of the decane, hexadecane and light catalytic naphtha cracking, it can be seen that, directionally, they move in the same direction, thus certainly not contradicting the findings of Brait et al.

145 [215]. Brait et al. [215] also found that the rate of conversion to isoalkanes, which they believe is characteristic of hydride transfer, is directly correlated to the amount of coke formed. Although the results of the light catalytic naphtha cracking indicate that the above may be correct (with the yield of iC4 decreasing by 26 mass% and the yield of coke decreasing by a comparable 22 mass% over the heat pre-treated catalyst), the results of the decane and hexadecane tell a different story. With these two feedstocks, the results show an inverse relationship between iso-alkane selectivity and coke yield. Thus it would indicate that the correlation between isoalkane and coke formation is entirely feedstock and/or catalyst dependent. In the cracking of decane and hexadecane the inverse relationship is possibly because of the change in the catalyst acidity. This alters the extent of hydride transfer reactions and thus results in the inverse relationship.

As demonstrated by Chen et al. [86], the ratio of the Bronsted to Lewis acid sites is much smaller for the USY than for the Y zeolite and decreases rapidly as a function of the severity of the steam treatment. Consequently fewer paraffins and more olefins are formed as primary products over the USY than over the Y zeolite due to the higher Lewis acidity of the former zeolite. This is direct confirmation of the observations made in this thesis with regards to the cracking of hydrocarbons over a heat pre-treated catalyst.

In summary, the results indicate that, directionally, most of the shifts are the same as those discussed previously. It is seen that there was a reduction in polymerisation reactions when cracking the n-decane and the n-hexadecane on the heat pre-treated catalyst as was seen in the cracking of reduced crude and light catalytic naphtha on hydrocarbon and heat pre-treated catalysts. This was due to the reduction in site density as a result of the heat pre-treatment converting Bronsted sites to Lewis sites. In addition, the paraffin to olefin ratio was also seen to

146 decrease, as in the reduced crude and light catalytic naphtha cracking, on the heat pre-treated catalyst. This indicated a reduction in hydrogen transfer reactions, also a result of the decrease in acid site density. Probably the most notable difference is that the shifts exhibited by these normal paraffins are quantitatively much less than those observed with the light catalytic naphtha. This enforces the hypothesis that the presence of olefins enhances the shifts exhibited between the pre treated and untreated catalysts.

7.3 TRI-METHYL PENTANE CRACKING

The quantity of reaction products from the cracking of n-decane and n hexadecane was found to still be quite complex in nature. It was decided that the cracking of a smaller hydrocarbon would partially alleviate this problem, since the amount of possible reaction products is largely reduced, and therefore would provide better information on selectivity shifts. The hydrocarbon chosen was 2,2,4 tri-methyl pentane. This molecule is also highly branched thereby simulating the branched components in industrial feedstocks (reduced crude).

The tri-methyl pentane was cracked over the equilibrium catalyst and the heat pre-treated catalyst as per the previous experiments. At the same catalyst to oil ratio, there was a large reduction in the conversion of tri-methyl pentane of about 8 mass%. The increase in hydrogen yield was also quite pronounced, as in the light catalytic naphtha experiments. Interestingly, the cracking of tri-methyl pentane on a heat pre-treated catalyst, when compared to the cracking on the untreated catalyst, showed a reduction in coke yield. This was directionally the same as that noted for the cracking of light catalytic naphtha. Directionally all other shifts were the same as those observed with the light catalytic naphtha cracking but, despite the large reduction in

147 conversion, were quite small in magnitude, much like the normal paraffin results.

Table 19: Product selectivity of 2,2,4 tri-methyl cracking over untreated equilibrium and heat pre-treated catalyst

Untreated equilibrium Heat pre-treated Yield (mass%) catalyst catalyst Conversion 21.89 20.93 H2 0.014 0.040 Dry gas 0.48 0.53 C3 0.44 0.48 C3= 3.65 4.32 nC4 0.54 0.53 iC4 39.90 38.72 C4= 33.42 36.85 Coke 21.54 13.56

Abbot and Wojciechowski [24) studied the cracking of 2-methylbutane and 2,2,4 tri-methyl pentane in steady state flow experiments over an H-Y catalyst. Molecular hydrogen was formed as an initial product from the former, but not the latter, nor did it appear in the cracking of n hexane or of cyclopentane. This shows a contradiction to the results in this thesis, where molecular hydrogen is certainly observed to have formed as a product from the cracking of 2,2,4 tri-methyl pentane as well as other paraffins. Furthermore it has been observed that the yield of molecular hydrogen varies depending on the pre-treatment method. One possible reason for the discrepancy is that the studies by Abbot and Wojciechowski [24) were conducted at low times on stream (TOS).

148 This was done in an effort to establish the primary products of catalytic cracking. The experiments conducted in this thesis are performed at relatively higher times on stream in order to simulate the industrial process more realistically.

Lombardo et al. [65] have found molecular hydrogen to be a primary product with isobutane as have other researchers such as van den Berg et al. [217] and Pine et al. [64). However Lombardo et al. [22, 65, 218) were not able to observe the formation of molecular hydrogen as a primary product in the cracking of neopentane. The above researchers each concluded that molecular hydrogen is formed as a primary product only when the structure of the feed molecule contains a tertiary C-H bond (except when it is highly shielded by multiple branching). They also report that the same chemistry occurs on relatively weak acids such as silica alumina, albeit at much higher temperatures and that the tendency to liberate molecular hydrogen increases with temperature in the primary reaction steps. Thus it would seem that in the cracking of 2,2,4 tri-methyl pentane, hydrogen is not formed as a primary product due to the high shielding of the tertiary C-H bond. This confirms the results of Abbot and Wojciechowski [24] while also providing an explanation of the results in this thesis.

The experiments performed in this thesis with iC8, nC10 and nC16 all show yields for molecular hydrogen. However it is hypothesised that this may be due to the secondary cracking of primary reaction products containing tertiary C-H bonds rather than the formation of molecular hydrogen as a primary cracking product for these hydrocarbons.

In the cracking of 2,2,4 tri-methyl pentane, in this thesis, comparable amounts of isobutane and isobutene were found to be the principal hydrocarbon products. This is to be predicted for protonolysis of the central C-C bond to yield a paraffin and the most stable carbenium ion

149 in the primary reaction step. As mentioned above, researchers Lombardo et al. [65), van den Berg et al. [217), and Pine et al. [64) have reported that molecular hydrogen is formed as a product from the cracking of iso-butane. Thus it can be seen that molecular hydrogen can be formed as a secondary product from the cracking of 2,2,4 tri methyl pentane.

Since an increase in hydrogen is observed for the cracking of 2,2,4 tri methyl pentane on the heat pre-treated catalyst, it can be concluded that the conversion of Bronsted sites to Lewis sites promotes the formation of primary products that are molecular hydrogen precursors and/or promotes the secondary reactions that form molecular hydrogen from the precursors. Lombardo et al. [65] believe that a primary protonolysis must always occur in the cracking of pure paraffins. This is the initiation step required to maintain the carbenium ion chain reactions involving hydrogen transfer, even when the latter become dominant [219]. On the heat pre-treated catalyst, a decrease in hydrogen transfer reactions was observed. Thus it is hypothesised that there is a relative excess of primary protonolysis reactions occurring on the heat pre treated catalysts, lending itself as precursors to increased formation of molecular hydrogen.

It was also observed in the cracking of 2,2,4 tri-methyl pentane, that there is an absence of volatile aromatic species in the products. This matches literature reports [24, 65).

150 7.4SUMMARY

In summary, the heat pre-treatment of the industrial equilibrium catalyst led to a reduction in activity of the catalyst, decreasing the conversion of the feedstocks: reduced crude, light catalytic naphtha, n-decane, n hexadecane and 2,2,4 tri-methyl pentane. The deactivation of the catalyst was not uniform (which would be represented by an equal yield reduction of all products) but rather selective. Between feedstocks, almost all shifts in selectivity were directionally similar, differing only quantitatively.

Further, the selectivity shifts observed by changing the Bronsted to Lewis site ratio through heat pre-treatment were almost identical directionally to those observed from pre-treating the catalyst with a hydrocarbon. This reinforces the hypothesis that:

1. Bronsted sites are the major contributing sites for catalytic cracking, 2. Hydrocarbon pre-treatment (through coking) deactivates the major contributing sites for catalytic cracking, and 3. Lewis sites, even if coked by hydrocarbon pre-treatment, have little effect on the product yield.

As an extension to the above conclusions, it can be hypothesised that the pre-treatment of the catalyst with a hydrocarbon deactivates certain sites selectively, rather than by uniform coke deposition or pore plugging. The heat pre-treatment experiments certainly appear to confirm this.

151 Other conclusions made from the heat pre-treatment experiments are that:

1. The presence of olefins in the feedstock exaggerates the shifts observed (this was seen in the cracking of light catalytic naphtha) due to the increased reactivity of olefins, 2. The heat pre-treatment of the catalyst reduces the active site density thereby reducing hydrogen transfer and polymerisation reactions (this was seen in the reduced paraffin to olefin ratios and reduced formation of heavy hydrocarbons), and 3. The mechanism for the formation of coke is extremely complex, therefore not allowing for the easy determination of shifts in coke selectivity.

Since the heat pre-treatment and hydrocarbon pre-treatment caused a selective deactivation of the industrial equilibrium catalyst through reduction of active sites, it was expected that the poisoning of the catalyst with an alkali would be expected to have the same effect. It was decided to initiate a study into the alkali pre-treatment of the industrial catalyst using ammonia as the alkali.

152 8 AMMONIA PRE-TREATMENT OF CATALYST

8.1 INTRODUCTION

In the previous chapters, two forms of catalyst pre-treatment were used, hydrocarbon and heat. Both these methods had advantages and disadvantages. The advantage of the hydrocarbon pre-treatment was that it closely represented the commercial operation, or potential commercial direction of a refinery, in particular Caltex's Lytton refinery. The disadvantages of this particular method of deactivation was that it was conducted in-situ on each sample run and therefore was quite difficult to replicate. Further, it was a complex system in terms of reaction pathways, and therefore very difficult to analyse. For example, it was not determinable whether the coking from the hydrocarbon pre treatment imparted a chemical or physical effect or both to the cracking catalyst. Therefore, it was decided to deactivate the catalyst through heat pre-treatment. The advantage of this method was that it imparted negligible physical effects to the catalyst, thereby clarifying the chemical / physical problem that was observed with the hydrocarbon pre treatment. However the disadvantage of this system was that it also was conducted in-situ and therefore difficult to replicate. Furthermore, it was extremely time consuming as the catalyst needed to be carefully cooled in the reactor / furnace so as to not rehydrate the catalyst and prevent physical damage to the catalyst by introducing a large temperature differential.

It was envisaged that the advantages of the ammonia pre-treatment was that it could would introduce negligible physical effects to the catalyst, and could be performed ex-situ, thereby allowing a greater quantity of catalyst to be pre-treated. Further ammonia has been used

153 in numerous studies [65, 220] to selectively neutralise the acid sites without fouling or blocking the catalyst pores, thus providing a good background with which to compare results.

8.2 PRELIMINARY EXPERIMENTS

The aim of the ammonia pre-treatment studies was first to check if the ammonia pre-treatment would lead directionally to the same shifts as those seen with the hydrocarbon and heat pre-treatment and then to proceed with further ammonia pre-treatment experiments to help explain the hypotheses generated in the previous chapters.

Initially the ammonia was injected in-situ over the untreated equilibrium catalyst at a temperature of 500 °C under a blanket of nitrogen. The concentration of the solution used was 28 mass% ammonia and the initial injection quantities were 0.2 ml and 2.5 ml. Once it was determined that a measurable effect from the ammonia pre-treatment could be observed, it was decided to move to an ex-situ method of pre treatment, thereby improving repeatability and efficiency.

The ex-situ method of ammonia pre-treatment involved soaking the industrial equilibrium catalyst in the ammonia solution (28 mass%) and allowing it to sit overnight. The excess ammonia was then evaporated in an 80 °C oven for 16 hours. The catalyst was then placed in a Carbolite muffle furnace and was heated to 250 °C for 1.5 hours and then calcined at 650 °C for 5.5 hours. It was hypothesised that this would only allow the most strongly bonded ammonia to remain adsorbed on the catalyst. Following this and prior to use, the catalyst samples (treated and untreated) were allowed to equilibrate with room conditions.

154 When ammonia (NH 3) reacts with a Bronsted site, a stable NH/ ion is produced together with its conjugate SiAIO- ion. Sodium (Na+) on the catalyst accomplishes the same result. It is hypothesised that this will reduce the steady state concentration (lifetime) of metastable carbenium ions and the overall activity of the catalyst and will induce non-linear poisoning effects of the catalyst.

Lombardo et al. [65] report that at 500 °C (under an inert atmosphere) ammonia is retained on about 4% of total number of lattice alumina. They also report that the poisoning of only -10% of sites with ammonia will eliminate all activity in a cracking catalyst. This would indicate that either these 10% of sites (possibly the strongest in acidity) are critical to the initiation and/or propagation of sites or that sites other than Bronsted sites are required to initiate the chains. They also report that poisoning sites is 10 times more effective than removing sites by dealumination. Presumably this is due to the selective poisoning of the strongest sites first, whereas site removal by dealumination is non selective.

In order to establish what percentage of sites have been poisoned after ammonia pre-treatment and calcining, a temperature programmed desorption analysis was conducted on the pre-treated catalyst and compared with the untreated equilibrium catalyst.

8.3 TEMPERATURE PROGRAMMED DESORPTION ANALYSIS

In the temperature programmed desorption analysis, the catalyst sample is mixed with aAl20 3 in a 1 :9 ratio. This mixture is then placed in a glass reactor similar to the one used in the microactivity test and then positioned in a cylindrical furnace. Ammonia gas is then introduced into the system at room temperature at a flowrate of 50 ml/min.

155 Breakthrough of the ammonia gas can be observed using a thermal conductivity detector (TCD). The gas is turned off and the furnace temperature is then ramped to 800 °C at a rate of 30 °C/min. Peaks in the TCD signal (recorded in mV) indicate the desorption of the ammonia. From the intersection with the temperature plot, the temperature of the desorption can be calculated.

Figure 11 and 12 show the temperature programmed desorption plots of the untreated and ammonia pre-treated catalysts respectively.

Figure 11: Temperature programmed desorption plot of untreated equilibrium catalyst.

-TCD Signal/ mV 800 1.2 --Temperature / °C 700 1.0 600 0.8 > e 500 ~ ..._ 0.6 3 'O "tu 400 ~ &> 0.4 ~ ci5 C: 300 cil 0 ..._ 0.2 0 ~ 200 (j 0.0 100 -0.2 0 0 1000 2000 3000 4000 Time/ s

156 Figure 12: Temperature programmed desorption plot of ammonia pre treated equilibrium catalyst

-TCD Signal/ mV 800 1.2 --Temperature/ °C

1.0

600 0.8 > e ~ 0.6 3 --;; "'O - 400 (1) 0.4 -~rn i 0 et 0.2 0 ~ 200 -n 0.0

-0.2 ----.---,--..----,---.---.----,.----,.....----1 0 0 1000 2000 3000 4000 Time Is

The major desorption peaks appear to be near to 100 °C and 720 °C. There also appears to be a shoulder at around 300 °C. With some of the samples, the ammonia was adsorbed and desorbed over the bed several times. Upon each extra adsorption, the adsorption capacity of the catalyst decreased. This indicates that at each adsorption / desorption cycle, some ammonia was being retained on the catalyst.

Desorption temperatures are an indication of acid site strength. Interestingly, the ammonia pre-treated catalyst shows higher desorption temperatures than the untreated catalyst. In the above two plots, the desorption temperatures of the untreated catalyst are 90 °C, 281 °C, and 700 °C and the desorption temperatures of the ammonia pre treated catalyst are 95 °C, 340 °C and 724 °C. This seems a little unusual since it would be expected that the ammonia would be retained on the untreated catalyst longer, since it would have stronger sites remaining. However, it is hypothesised that although this may be the

157 case the gaseous ammonia on the ammonia pre-treated catalyst bonds stronger to the already existing NH4 + ions.

The areas under the desorption curve are indicative of the total amount of ammonia desorbed in the analysis. It can be observed that the area under the desorption curve is larger for the ammonia pre-treated catalyst. This indicates that more ammonia is desorbed presumably because some ammonia from the pre-treatment step is also desorbing. One noticeable difference in the two plots is the desorption peak near 700 °C. For the ammonia pre-treated catalyst, the peak at 724 °C is larger than for the untreated catalyst peak at 700 °C. This shows that:

1. Some of the ammonia adsorbed from the pre-treatment step is desorbing, and 2. This desorption is predominantly occurring from the stronger acid sites. Thus this confirms that the remaining ammonia from the pre-treatment stage has been adsorbed at the strongest acid sites.

At least five temperature programmed desorption runs were conducted for each catalyst type. Table 20 shows the average of the results observed from the temperature programmed desorption measurements.

158 Table 20: Temperature programmed desorption results of untreated equilibrium and ammonia pre-treated catalysts averaged over 5 runs.

Untreated Ammonia pre-treated

Adsorption of 2.02 X 10-4 1.81x10-4 ammonia (mol/g)

Desorption of 2.08 X 10-4 3.46 X 10-4 ammonia (mol/g) 1st Desorption peak 92 °C 106 °C 2nd Desorption peak 281 °C 313 °C 3rd Desorption peak 695 °C 750 °C

The additional ammonia desorbed from the untreated catalyst compared to that adsorbed is within the error of the experiment. This difference is possibly due to non-ammonia molecules triggering the detector, however largely the difference is negligible.

As mentioned above, the extra ammonia desorbed from the pre-treated catalyst is presumably ammonia from the pre-treatment process. As is to be expected, less gaseous ammonia could be adsorbed on the ammonia-pre-treated catalyst when compared with the untreated catalyst. From the difference in these two quantities, it can be estimated what proportion of sites contained adsorbed ammonia from the pre treatment process. Assuming that the gaseous ammonia and the ammonia solution are adsorbed as a monomolecular layer on the acid sites, it is estimated that approximately 10% of the acid sites have been covered from the ammonia pre-treatment process.

159 Further analysis of the results shows that there is more ammonia desorbed from the ammonia pre-treated catalyst than adsorbed on the untreated catalyst. This indicates that the saturation of the catalyst with ammonia solution is more effective than gaseous dosing and suggests that the range of acid sites onto which ammonia can adsorb from solution is not identical to those available via gas-phase adsorption. As a result, it can be predicted that the ammonia from the pre-treatment is adsorbed on a higher proportion of sites than the 10% initially calculated.

8.4 REDUCED CRUDE CRACKING

The first set of experiments with the ammonia pre-treated catalyst involved the cracking of reduced crude (as had been used for the experiments in previous chapters). The reduced crude was cracked over three different ammonia pre-treated catalysts, two in situ involving 0.2 ml and 2.5 ml ammonia (28%) solution and one ex situ involving soaking in ammonia solution overnight as described above.

As expected the pre-treatment of the catalyst with ammonia resulted in a reduction in feedstock conversion. Figure 13 shows this reduction in conversion from the three different ammonia pre-treatment methods.

160 Figure 13: Activity loss as a function of ammonia pre-treatment level when cracking reduced crude

-0~ ~ 70 ro E -C: 0 60 ·1n.... a> > C: 50 0 0 I >

It can be seen that with the in situ method of ammonia pre-treatment, the addition of more ammonia led to a greater reduction in conversion. With the addition of 0.2 ml ammonia in situ a decrease in conversion of nearly 3 mass% occurs for the catalyst to oil ratios between 3 and 9. The addition of 2.5 ml ammonia in situ results in a conversion loss of about 5-6 mass% also occurring across the catalyst to oil range of 3 and 9. The relationship between quantity of ammonia and conversion loss does not seem to be linear. This indicates that for this method of pre-treatment, the catalyst quickly reaches a saturation point for ammonia and that the sites critical for the cracking process are the first to be affected.

161 For the ammonia pre-treatment by soaking and calcination, a much larger decrease in conversion resulted. At a catalyst to oil ratio of 4, the conversion loss was almost 25 mass% and, at a catalyst to oil ratio of 9, the loss was about 16 mass%. The fact that the ex-situ method of ammonia pre-treatment showed a much greater reduction in conversion is interesting since the catalyst was calcined at 650 °C for 5.5 hours in air at atmospheric conditions. According to Lombardo et al. [65] this would be expected to result in near complete desorption, with ammonia molecules associated with Bronsted sites of weaker acidity being preferentially removed. They found that treatment with 0 2 at 500 °C completely restored the activity of an ammonia pre-treated cracking catalyst. This certainly was not observed with the experiments performed in this thesis using the ammonia / calcined catalyst.

Hildebrandt and Skala [221] found that temperatures of up to 600 °C are required to remove ammonia from the strongest sites of a (NH4) USY catalyst. They also found that it was these strong sites that correlated with activity for the cracking of light Pennsylvania gas oil.

From the findings of Lombardo et al. [65] and Hildebrandt and Skala [221] it can be assumed that any ammonia remaining on the equilibrium catalyst after calcining at 600 °C for 5.5 hours must be adsorbed on only the strongest acid sites. Since the ammonia deactivates the strongest acid sites so selectively and efficiently, it is envisaged that this will greatly exaggerate any effect seen from the hydrocarbon or heat pre-treatment experiments.

Lombardo et al. [65] report that, for the cracking of neopentane on H mordenite and H-ZSM-5 zeolites (silica to alumina ratio of 7 to 8), the ammonia retained on the catalyst after 500 °C treatment was only about 4% of the total number of lattice alumina, but that this was still enough to reduce the conversion to about one third of its unpoisoned value.

162 They further hypothesise that the lethal dose corresponds to 1O ± 3% of the total number of the possible acid sites (the number of aluminium atoms/g). In the case of H-Y catalyst, they report that amounts of ammonia corresponding to only 3% of the maximum number of po~sible sites was sufficient to eliminate activity.

Although the present experiments certainly show that only small residual amounts of ammonia on an industrial USY catalyst can affect the conversion or reduced crude, it does not seem to have as severe an effect as that reported by Lombardo et al. [65). Instead it can be calculated that poisoning greater than 10% of the sites (as estimated from the temperature programmed desorption experiments) only reduces the conversion to about two thirds of its unpoisoned value. It can certainly be seen that the poisoning of 10% of sites does not eliminate the activity of the catalyst, let alone the poisoning of only 3% of the sites as was seen with Lombardo et al.'s H-Y catalyst [65].

However in accordance with the results of Lombardo et al. [65] a linear relationship between number of poisoned sites and conversion loss was not seen. It is hypothesised that this may be attributed to the concentration of strong Bronsted sites able to crack small paraffin molecules being only a small fraction of the total on the catalyst and hence only small amounts of ammonia are required to affect the conversion.

Gorte and co-workers [222) have studied the interaction of alcohols and amines with the acid sites of several zeolites. They have reported that one molecule of base per framework alumina was strongly chemisorbed, independent of zeolite structure or aluminium concentration. However, the present results show that only small amounts of ammonia are required substantially to reduce the conversion, in other words indicating that a majority of the sites are too

163 weak to affect carbenium ion catalysis, and too weak to adsorb and retain ammonia molecules. Thus it is believed that the present system does not concur with the findings of Gorte et al. [222].

Since the poisoning of relatively few sites on the catalyst leads to such a marked reduction in activity, it may be hypothesised that sites other than Bronsted sites may be responsible for initiating the cracking reaction. As discussed extensively in the literature review section, opinion differs widely as to the nature of the primary process. Brenner and Emmett [189] postulated intrinsic dehydrogenation sites, while McVicker et al. [223] suggested that strong electron deficient sites are present which effect a radical decomposition of the substrate, a view accepted by Bizreh and Gates [224]. Marczewski [225] and Zholobenko et al. [226], on the other hand, hold Lewis acid sites responsible. This last opinion is not in accord with the results observed with the catalyst heat pre-treatment experiments. However, these ideas do have one thing in common, viz., that sites other than Bronsted acid may be required to initiate the chains. If these were small in number and chemisorb ammonia very strongly, any of these concepts would afford an explanation for the results observed.

Lombardo et al. [65], found that water, although less basic than ammonia, was nearly as effective in reducing the activity. They also report that the water pre-treatment of the catalysts is qualitatively similar to ammonia pre-treatment in selectivity changes. Brait et al. [215] also report that the addition of water decreased the rates of all reaction pathways in the cracking of n-hexane due to competitive adsorption. Although not part of the initial experimental design of this thesis, experiments with water pre-treatment were conducted to ascertain the above findings. In the experiments, water was injected over the catalyst in-situ prior to passing reduced crude feedstock over the catalyst.

164 Varying amounts of water was used (between 1 ml and 5 ml) however there was a negligible loss of activity or change in selectivity.

It is believed that the water pre-treated catalysts do not exhibit the same loss of conversion or shifts in selectivity simply because the water is not chemisorbed as strongly to the acid sites as with the ammonia pre treatment. Thus it cannot block reactions that proceed from these sites.

The possibility that physical effects of the ammonia pre-treatment playing a role in the reactions should not be overlooked. The cracking catalyst used is essentially a molecular sieve whose pores are of molecular dimensions. Could the adsorption of ammonia, by physical means, have a critical effect on the reaction rates? Lombardo et al. [65] report that this is unlikely for the following reasons. They found that the quantity of ammonia required to almost eliminate the cracking catalyst's activity did not differ widely for the several different catalyst structures tested. In fact, they found that slightly more ammonia was required for a mordenite, which has a unidimensional pore system than for the three dimensional Y-zeolites where pore blocking is more difficult. Similarly, the results for H-ZSM-5, where molecular traffic control may be expected, were comparable in terms of ammonia required to eliminate activity. Diffusional effects should become more pronounced as the reaction rate is increased, yet the mordenites were many times more active than H-Y which had similar activity as H-ZSM-5. Thus it is assumed that physical effects of ammonia pre-treatment are negligible, certainly when compared with the possible physical effects of coking due to hydrocarbon pre-treatment.

165 8.4.1 REACTION PRODUCT DISTILLATION

Another way of viewing the overall shift in the conversion of a cracking catalyst is to plot the cumulative mass % of products as a function of their boiling point. Although this method of interpretation is conceptually slightly more difficult, it does allow the quick observation of change in conversion. Figure 14 shows the cumulative mass % of products as a function of boiling point for the cracking of reduced crude over the untreated catalyst and over the ammonia pre-treated catalyst.

Figure 14: Reaction product distillation from reduced crude cracking on untreated equilibrium and ammonia pre-treated catalyst.

600

500

(.) -0 400 ...... -C: ·5 CL 300 Ol C: ·5 200 co

100 1~x-U-nt-re-ated~~ilib-rium-catalyst-~~-- I ~-~~i::~- 1

I • Amronia pre-treated by soaking e~ 0 0 20 40 60 80 100 Cumulative mass%

It can be observed that for the ammonia pre-treated catalyst, there is a larger proportion of heavy (high boiling point) material remaining than in the untreated catalyst, indicating an overall decrease in conversion. Although difficult to observe, certain shifts in selectivity can be noticed.

166 For example at -80 °C the ammonia pre-treated catalyst shows a horizontal section on the plot indicating that a lot of material boiling at 80 °C is formed. For the untreated catalyst the slope of the plot at this point is much greater indicating that comparatively less 80 °C material is formed. In the range 140 °C to 240 °C, the ammonia pre-treated catalyst shows a greater slope than the untreated catalyst indicating that less material is formed in this range on the ammonia pre-treated catalyst. This range corresponds to the boiling range for light cycle oil, confirming the results seen earlier with the hydrocarbon and heat pre treatment.

8.4.2 PRODUCT SELECTIVITY SHIFTS OF EX-SITU AMMONIA PRE TREATED CATALYST

To further investigate whether ammonia pre-treatment of the catalyst affects the catalyst in a similar way to the hydrocarbon and heat pre treatment methods, an analysis of product selectivity is required. It was decided that the best sets of results to compare with the untreated equilibrium catalyst were those from the ammonia pre-treated catalyst that had been soaked in ammonia ex-situ, rather than the catalysts that had been pre-treated with ammonia in-situ. This was because the ammonia soaked catalyst had shown the greatest reduction in conversion with the cracking of reduced crude. Thus it was expected to produce the most exaggerated shifts, thereby clearly indicating a change in cracking product selectivity.

Reduced crude was cracked over the untreated and ammonia pre treated catalysts at 500 °Cat varying catalyst to oil ratios.

The most effective way to display the results is in tabular form rather than graphical due to the large amount of products. However, all

167 selectivity shifts observed in the ammonia pre-treatment experiments are directionally consistent for each product throughout the tested catalyst to oil range. Table 21 shows the dramatic shifts in selectivity from the catalyst pre-treated by soaking in ammonia at constant conversion. The comparison is made with the untreated equilibrium catalyst at about 69 mass% conversion of reduced crude. Conversion is defined as 100 minus the fraction of product boiling in the clarified oil (CLO) range. All yields have been normalised for conversion and are in mass%.

Table 21: Effect of ammonia pre-treatment on product selectivities when cracking reduced crude

Ammonia Yield (mass%) Untreated pre-treated H2 0.347 0.571 Dry Gas 1.61 2.29 C3 0.76 0.93 C3= 3.42 5.99 nC4 0.44 0.46 iC4 2.40 2.59 C4= 4.49 7.89 LCN 45.77 40.04 LCO 34.67 26.36 Coke 6.42 13.45

The results reveal that directionally the shifts were very similar to those observed by the cracking of reduced crude over the hydrocarbon pre treated catalysts. The most notable exception is the decrease in light catalytic naphtha yield on the ammonia pre-treated catalyst compared

168 with the increase seen with the hydrocarbon pre-treatment. This is theorised to be a result of a large decrease in site density. It is hypothesised that the site density on the ammonia pre-treated catalyst has become so low (relative to the hydrocarbon pre-treated catalyst) that:

1. Polymerisation reactions have decreased resulting in a large decrease in light catalytic naphtha yield and light cycle oil yield as a result of light hydrocarbon polymerisation, and 2. Large feedstock molecules are not being cracked to form light catalytic naphtha, as they had been in the hydrocarbon pre treatment experiments, but rather even smaller fragments (light hydrocarbons) are being cracked off due to an overall decrease in site strength and site density.

The increased yield of coke is a direct result of an increase in dehydrogenation reactions, thus forming more coke and more hydrogen deficient species such as the light olefins.

The other notable difference in the ammonia pre-treated selectivity shifts, when compared to the hydrocarbon pre-treated shifts is their magnitude. Table 22 shows the difference in magnitude of shift for a few of the yields.

169 Table 22: Comparison of selectivity shifts for hydrocarbon and ammonia pre-treated catalysts when cracking reduced crude

Hydrocarbon pre- Ammonia pre-treated treated Conversion 8% decrease 22% decrease Hydrogen 10% increase 65% increase Dry Gas 10% increase 42% increase LCO 12% decrease 24 % decrease Coke 28% increase 110% increase

If a comparison is made between the product yield shift differences, it can be seen that the difference in pre-treatment methods does not equally affect all reaction pathways. For example, on the ammonia pre treated catalyst, a decrease of 24% in light cycle oil yield is observed compared with a decrease of only 12% on the hydrocarbon pre-treated catalyst. This represents a two fold change in the reaction pathway taken to form light cycle oil. However in the comparison of coke yield an almost four fold change can be observed with the ammonia pre-treated catalyst showing an increase of 110% compared to a 28% increase on the hydrocarbon pre-treated catalyst.

Thus it can be concluded that while the ammonia pre-treatment does increase the severity of the deactivation and does exaggerate the shifts observed with the hydrocarbon pre-treatment, it does not affect all reaction pathways equally. In the above table it can be seen that polymerisation reactions (responsible for light cycle oil formation) are not affected as much as dehydrogenation reactions (responsible for dry gas and coke formation).

170 Lombardo et al. [65] found that the ammonia pre-treatment of a zeolite repressed the formation of polymeric surface residues. It is assumed that the term polymeric surface residues refers to coke. This is not observed in the experiments conducted with the ammonia pre-treated equilibrium catalysts in this study. However, it may be said that the major contributing reaction pathway for the formation of coke in these experiments is via dehydrogenation rather than polymerisation. This is not to say that no polymerisation occurs, but rather that there is a more prominent reaction pathway. Thus it can be concluded that in the current experiments, there may in fact be a reduction in the formation of polymeric surface residues, if it is assumed that the term refers to coke formed via a polymerisation process.

As mentioned previously the analysis of light olefins in the product yield is important, since shifts in this product indicate a change in hydrogen transfer reactions, an important part of the cracking process. In addition light olefins such as propylene and butylenes are important from a commercial standpoint in that they are required for production of liquefied petroleum gas or autogas as well as being important feedstocks for polymerisation and alkylation units.

Figure 15 shows the large increase in light olefins (C3 + C4 olefins) observed on the ammonia pre-treated equilibrium catalyst when compared to the untreated equilibrium catalyst while cracking reduced crude.

171 Figure 15: Propylene plus butylenes selectivity as a function of ammonia pre-treatment when cracking reduced crude

-0~ en en 12 ro E -II 10 "d" () + II 8 --l - M I ()

6 x Untreated equilibrium catalyst • Ammonia pre-treated by soaking ex-situ ---T------~ 4 55 60 65 70 75 80 Conversion (mass%)

It may be hypothesised that the increase in olefinic compounds relative to paraffinic compounds is due to a decrease in secondary reactions. This would occur via the following chain of reactions. The Bronsted sites attack the carbon-carbon bonds via Bscission cracking. Carbenium ions are formed which may vary in their stability, and consequently their lifetime on the catalyst surface. This feature is postulated to be a determining factor in controlling the extent of secondary reactions. When the ion lifetime is sufficiently long, substantial concentrations of compounds can be present simultaneously in the catalyst bed, which can oligomerise leading to secondary reactions, the formation of paraffinic products. The increase of olefinic products indicates that oligomerisation may be hindered, meaning that there is a reduction in secondary reactions. However, if this were solely the case, the yield of

172 paraffins should decrease, but from Table 22, although they are also seen to increase, the increase is not as substantial as for olefins. Thus it is believed that the olefins do not originate solely from this process, but rather from a process involving the cracking of small fragments from large feedstock molecules, which then undergo severe dehydrogenation to produce light olefins. This theory explains the slight increase in light paraffins and also the increases in dry gas and coke.

As mentioned above, the most notable difference in the shifts observed, when compared with the hydrocarbon pre-treatment experiments, was that for light catalytic naphtha selectivity (Table 22). Instead of increasing as was to be expected as a result of the hydrocarbon pre treatment experiments, the yield of light catalytic naphtha decreased but, in comparison with the decrease in light cycle oil, the decrease was relatively less. From this outcome and keeping in mind that the light hydrocarbons increased overall, it can be hypothesised that the severe pre-treatment of the catalyst with ammonia soaking has resulted in a catalyst capable only of cracking off the small side chains of the branched components in the feedstock, rather than large light catalytic naphtha compounds or the large light cycle oil compounds. This shows that the increased production of light gases and light catalytic naphtha as seen for the hydrocarbon pre-treatment experiments is not due to overcracking of the products, as was first hypothesised, but rather that the molecules that are the most easy to crack off first do just that.

The smaller shifts observed with the straight chain hydrocarbon conversion over the heat pre-treated catalysts are in agreement with this.

173 8.4.3 PURE HYDROCARBON CRACKING ON AMMONIA PRE TREATED CATALYSTS

As the products of reduced crude are quite difficult to analyse and since they contain metals and basic nitrogen (which affect conversion and selectivity) it was once again appropriate to direct the focus on the cracking of pure, more simplistic compounds. As had been tested previously, the cracking of 2,2,4 tri-methyl pentane and hexadecane was performed again, but this time on the ammonia pre-treated catalyst.

Figure 16 shows the marked decrease in conversion when cracking the tri-methyl pentane over an ammonia pre-treated catalyst.

Figure 16: Conversion of 2, 2, 4 tri-methyl pentane when cracked over heat pre-treated and ammonia pre-treated cracking catalyst compared with untreated equilibrium catalyst.

50 __ __JI__ X Untreated equilibrium catalyst + Heat pre-treated ' 40 ------j A Ammonia pre-treated by soaking ex-situ I -;:R. I -0 (/) (/) I ro E 30 -C: .Q (/) '- Q) 20 > C: 0 (.) I I 10 - X I I • IA fa A.. I 0 ~ I 0 5 10 15 20 25 30 35 40 Catalyst to oil ratio

174 As can be seen in the above graph, despite increasing the catalyst to oil ratio to over 35, it was very hard to crack the 2,2,4 tri-methyl pentane to a conversion of greater than about 6 mass% on the ammonia pre treated catalyst. This is due to the difficulty in initiating the cracking of low carbon chain length molecules such as tri-methyl pentane when compared with the relative ease of initiating cracking with the much larger molecules found in reduced crude. As a result it was not possible to derive information about changes in yield selectivities, as there were no similar conversions for comparison with the cracking on the untreated catalyst. Suffice to say that the ammonia pre-treated catalyst produced a very large decrease in conversion when compared to the heat pre-treated and untreated equilibrium catalysts.

As mentioned, n-hexadecane was also cracked over the ammonia pre treated catalyst, but again some of the yields were so small that any observation of shifts in selectivity were prone to large error.

8.5 SUMMARY

The pre-treatment of the catalyst with liquid ammonia by soaking resulted in a severe decrease in conversion. It was confirmed that there is not a linear relationship between proportion of sites poisoned and loss in conversion. This is hypothesised to be a result of strong Bronsted sites being only a small fraction of the total number of acid sites and being deactivated by the adsorption of ammonia. Once again a shift in selectivity was observed with increases in hydrogen, light olefins and coke with decreases in light catalytic naphtha and light cycle oil. The decrease in light cycle oil was determined to be a result of a decrease in polymerisation reactions (thus less light hydrocarbons polymerising to form light cycle oil) and that the large feed molecules are not being cracked to form light cycle oil, rather forming smaller

175 hydrocarbons. Both of these effects are due to a decrease in site strength and site density with the strength of the catalyst so reduced that only low energy reactions are catalysed. Thus the ammonia pre treated catalyst is capable of only cracking off the small side chains of the branched components of the feedstock, and as such the hypothesis that the reduced light cycle oil yield was a result of overcracking can be ruled out.

Although the ammonia pre-treatment was found to affect the selectivity of the catalyst, it was determined that it does not affect all reaction pathways equally. For example, it was found that polymerisation reactions were not as affected as hydrogen transfer reactions, which rely on high site density.

A new hydrocarbon needed to be found for further experimentation with the ammonia pre-treated catalyst. It should have characteristics similar to that of the average of the compounds in reduced crude, readily show shifts in selectivity and be pure so as not to affect the catalyst through contamination or poisoning, except by its own coking as a result of catalytic cracking. Squalane was selected for testing as discussed in the next section.

176 9 CRACKING OF SQUALANE

9.1 INTRODUCTION

Considerable work has gone into the study of small-paraffin molecules in catalytic cracking. Recent studies show that the cracking of paraffins have focused on molecules with seven or less carbon atoms [33, 78, 227]. By studying a small hydrocarbon molecule, some of the complexities inherent in the catalytic cracking process can be avoided, and results are generally easier to analyse than those from higher molecular weight reactants, mainly because of the limited number of possible reactions (and thus products) in such cases. However the higher molecular weight reactants are the major components of industrial gas-oil feedstocks and therefore small hydrocarbon test compounds do not form a good representation of industrial operation.

Furthermore, several researchers studying the reactions of pure hydrocarbons have concentrated on the behaviour of linear paraffins under cracking conditions. Studies on branched isomers, an important component of industrial feedstocks, have been reported to a lesser extent [20, 24, 35, 228-230]. Molecular dimensions of the feed paraffin are also important in the consideration of shape selectivity [228, 229], while intrinsic reactivity can be significantly influenced by the presence of a hydrogen atom bonded to a tertiary carbon atom [35, 230].

The above examples show how important feedstocks are to the cracking reaction and how the overall product selectivity can be affected by the feedstock. Further it is noted that the cracking behaviour of small molecules is not typical of larger molecules or industrial feedstocks.

177 Ignoring this fact may have allowed some damaging misconceptions to arise.

At the other end of the scale it can be quickly seen that the cracking of industrial feedstocks is complex and contamination such as metals and basic nitrogen in the feedstock can affect the catalyst through poisoning. Therefore an alternative feedstock needs to be found, one that is representative of industrial feedstocks yet without contaminants.

To find a compound that is representative of an industrial feedstock is quite tricky. Reduced crude consists of a large proportion of branched paraffins. In addition cyclo paraffins and aromatics are present. Since cyclo paraffins and aromatics are much harder to crack, and therefore do not play a large part in cracking reactions, it was decided to concentrate on a compound that was quite large and branched in nature. This will ensure that cracking occurs quite readily and that any shifts in selectivity will therefore be quite pronounced. The hydrocarbon that suited this description was squalane.

Squalane is a branched hydrocarbon of 30 carbons in chain length, thus satisfying the requirement to be similar to the average of the industrial feedstock used so far in this thesis. Furthermore it is available as a pure compound, thus satisfying the requirement to be as void as possible of contamination.

The structure of squalane appears in Figure 17 below.

Figure 17: Structure of squalane (branched C30 compound)

178 The formula weight of squalane is 422.83. Its boiling point is 176 °C at 0.05 mm of Hg and its melting point is -38 °C.

In reviewing the literature, there appears to have been no previous work done that involves squalane as a test molecule for catalytic cracking on pure zeolites, let alone on pre-treated commercial catalysts. Thus the work done on the cracking of squalane over an industrial catalyst (both untreated and pre-treated) appears to be the first of its kind.

9.2 SUITABILITY OF SQUALANE AS A REPRESENTATIVE TEST MOLECULE

It was important to repeat the full range of experiments conducted with the other previously used hydrocarbons to investigate whether the same observations could be made of the squalane and that it was in fact a satisfactory substitute for the industrial feedstock.

Thus, the squalane was cracked over the hydrocarbon, heat and ammonia pre-treated catalysts. As expected, there was a reduction in conversion for all the methods of pre-treatment as shown in Figure 18.

179 Figure 18: Activity loss when cracking squalane on heat, hydrocarbon and ammonia pre-treated catalysts compared to untreated equilibrium catalyst

-~0 (/') (/') ctl 60 ._E C 0 ·u5 i... Q) 40 > C 0 o~_l_. __ ~ (.) i X Untreated equilibrium catalyst 20 -to---a....,....-;-r------1 ! • Heat pre-treated llit,. Reduced crude pre-treated 0------'""'OAmmon,ia pre-treated by ~aking ex-situ1 0 2 4 6 8 10 Catalyst to oil ratio

The ammonia pre-treated catalyst showed the greatest reduction in conversion of over 70 mass% at a catalyst to oil ratio of 2.5. The heat pre-treated and hydrocarbon (reduced crude) pre-treated catalysts showed a decrease in conversion of -13 mass% and -28 mass% respectively at the same catalyst to oil ratio of 2.5. The reason that the ammonia pre-treated catalyst showed the largest decrease in conversion is due to the ammonia blocking the largest percentage of strong acidic sites selectively through adsorption.

180 9.2.1 CRACKING OF SQUALANE ON AN AMMONIA I HEAT PRE TREATED CATALYST

Another variation of the squalane / catalyst system was tested whereby ammonia pre-treated catalyst was placed in the reactor and then heat pre-treated in situ (as per the procedure for all the heat pre-treated experiments) i.e. that the ammonia pre-treated catalyst was heated in situ to 600 °C under a nitrogen blanket and then equilibrated at 500 °C before cracking. This essentially represented a catalyst pre-treated with ammonia plus heat pre-treatment.

What was expected to occur from this was a further reduction in conversion over just the ammonia pre-treated catalyst. However this did not occur. Although the conversion reduction was substantial (when compared to the untreated equilibrium catalyst), a negligible further decrease in conversion was observed over the ammonia pre-treated catalyst. This indicates that the sites required for initiation and propagation of the cracking reaction were already deactivated by the ammonia pre-treated catalyst, and further shifting of Bronsted to Lewis sites via the heat pre-treatment did not serve to enhance this deactivation. In other words, the ammonia pre-treatment was an extension of the heat pre-treatment alone. The heat pre-treatment did not add any further measureable deactivation to the catalyst that was not already accomplished with the ammonia pre-treatment. This is significant because it enforces the hypothesis that the primary cracking site is the Bronsted site and that the ammonia pre-treatment adequately deactivates this site.

Taking the hypothesis that Bronsted sites of zeolites cleave C-H and C C bonds of paraffins as they do in superacid systems [19, 24, 231, 232, 233], it can be suggested [125, 234] that the formation of such supersites requires the interaction of aluminium ions expelled from the

181 lattice with Bronsted hydroxyl groups. Ammonia might either poison such sites selectively, or interfere with the Lewis-Bronsted interaction by reacting with the Bronsted / Lewis centre.

Since the squalane had responded as expected on the variously pre treated catalyst, there was no reason not to proceed with further analysis of the results. As has previously been done, a breakdown of the total product selectivity is required to determine shifts within yields, thereby allowing a greater insight into the reaction changes occurring with the catalyst. It was decided that the results from the squalane cracking over the ammonia pre-treated catalyst would provide the most information as selectivity shifts were envisaged to be the largest of any of the pre-treatment methods.

9.2.2 PRODUCT SELECTIVITY SHIFTS

It was decided to compare results of the squalane on ammonia pre treated catalyst with that on the untreated equilibrium catalyst at 39 mass% conversion. From Figure 18 this can be seen as a good overlap of conversions, where a result for each catalyst had been obtained.

A comprehensive method of analysing overall shift in selectivity as a function of carbon number is to plot the total yield (mass%) of each carbon number. This plot can be seen in Figure 19 for the cracking of squalane on the ammonia pre-treated and untreated equilibrium catalyst.

182 Figure 19: Total product selectivity of squalane cracking over untreated equilibrium and ammonia pre-treated catalyst at 39 mass% conversion

~~---~----~ ----*"- Untreated equilibrium catalyst 10 ~------+-¼------, -+-Amronia pre-treated by soaking ex-situ :

8 - --0~ Cl) Cl) ro E 6 -"'O Q) 5= 4 -

0 0 5 10 15 20 25 Carbon number

Figure 19 shows that poisoning of the most active sites with ammonia results in a higher propensity towards formation of smaller cracking fragments in the range C3 to C5 at the expense of larger fragments at C7 and above. This is in direct accordance with those results observed for the cracking of reduced crude over ammonia pre-treated catalysts. In those experiments an increase in light hydrocarbons below C5 was observed over the ammonia pre-treated catalysts. Further that the yields of light catalytic naphtha (boiling range starting at C5) and light cycle oil decreased over the ammonia pre-treated catalyst.

The above result is a significant observation, as it would be expected that, after poisoning with ammonia, the remaining Bronsted sites would be generally of lower acidity compared to those present on the untreated equilibrium catalyst. Previously Carma et al. [232], suggested

183 that, for the cracking of n-dodecane on H-Y, amorphous silica-alumina, mordenite and ZSM-5 (different catalysts to the one used in this study) a greater abundance of smaller cracking fragments was associated with a greater abundance of stronger acid sites on a particular catalyst. This seems to contradict the results seen with the earlier systems and also with the cracking of squalane on the ammonia pre-treated catalyst. This is discussed in greater detail later in the chapter.

The above graph shows that at C6 the yields are similar, but from C7 onwards the yield is quite diminished for the pre-treated catalyst, hence resulting in lower yields of light catalytic naphtha and light cycle oil.

The comparison of the selectivities as a function of carbon number changes when the conversion is increased from 39 mass% to 65 mass% as shown in Figure 20.

184 Figure 20: Total product selectivity of squalane cracking over untreated equilibrium and ammonia pre-treated catalyst at 65 mass% conversion

35 ------

"C a, 15 ~ 10

0 l---.Z:....------i--=~;;;;a;...... ~~.... ----! 0 5 10 15 20 Carbon Number

To achieve this increased conversion the catalyst to oil ratio needs to be higher. Thus it would be expected that any shifts observed at the higher conversion would be more pronounced.

Here it can be seen that the light gases on the ammonia pre-treated catalyst have increased even more relative to the untreated catalyst. At 65 mass% conversion there is a relative increase of 64% in C4s, whereas at 39 mass% conversion there is only a 38% relative increase over the untreated catalyst.

This noticeable increase in total hydrocarbon yield for C3, C4 and C5 species may, in part, be compared with the findings of Abbot and

185 Wojciechowski [35) in their experiments with mixtures of H-Y and ZSM- 5 zeolite over H-Y zeolite individually. They report that a shift in selectivity towards lower carbon number for cracking over the smaller pored H-ZSM-5 may be attributed to the low site density of this zeolite [35), an attribute that is hypothesised in this thesis to be a critical factor in catalytic cracking selectivity. lnfrared investigations by Hedge et al. [235] have shown that the Bronsted acid sites of H-ZSM-5, while less numerous because of its lower alumina content, are stronger than those of H-Y faujasite. Similarly, it was found that H-P zeolite has a smaller number but stronger Bronsted acid sites than H-Y zeolite [236). Bonetto et al. [237] have studied the performance of mixtures of USY and p zeolite in the cracking of gas oil at low conversions. They also found that ~ zeolite shifts the product distribution toward shorter hydrocarbons and produces more olefins than USY.

By separating hydrocarbon product distributions into paraffins and olefins, branched and unbranched species, the origin of changes in the product distribution may be traced.

Thus, a more detailed investigation of the

186 Figure 21: Light paraffin selectivity of squalane cracking over untreated equilibrium and ammonia pre-treated catalyst at 65 mass% conversion

-. 5 0~ (/) (/) 4 - co E -"O 3 Q) ~ 2

1 +------+--1------l1 -*-Untreated equilibrium catalyst -+-Ammonia pre-treated by soaking ex-situ 0 0 1 2 3 4 5 6 7 8 Carbon number

An increase of 0. 7 mass% is seen for C3 paraffins and an increase of 1.5 mass% is seen for C4 paraffins. In general an increase in the amount of normal paraffins may indicate an increase in hydrogen transfer reactions. However it is important to correlate this with changes in olefins selectivity. For there to be an increase in hydrogen transfer reactions, the yield of normal paraffins must increase relative to any changes in the selectivity of olefins. This is studied in detail later in the chapter.

187 9.2.3 SHIFTS IN EXTENT OF ISOMER/SAT/ON REACTIONS

Due to a reduction in branched paraffin yield at C4 and C5 over the ammonia pre-treated catalyst, the branched to straight paraffin ratio decreases over the pre-treated catalyst for light hydrocarbons, when compared to the untreated equilibrium catalyst as seen in Figure 22.

Figure 22: Branched to straight paraffin ratio at 65 mass% conversion, as a function of ammonia pre-treatment in the cracking of squalane

25 ------....------

g 20 ~ I,,_ ro a> 15 C .8 as 10 - .c (.) C: ~ co 5 --~~------1 ~Unt~eated equilibrium catalyst i I ~ Ammonia pre-treated by soaking ex-si~J 0------""' 3 4 5 6 Carbon Number

A decrease in branched to normal paraffin ratio of about 6 numbers can be observed for C4s and C5s when comparing the cracking of squalane on the ammonia pre-treated catalyst with the untreated equilibrium catalyst.

This ratio gives an indication of the degree of isomerisation. A reduction in this ratio indicates a decrease in isomerisation. Zhao et al. [25] and

188 Daage and Fajula [47] have shown that isomerisation is a process of lower activation energy. Since isomerisation is a relatively low energy reaction compared to cracking, the above results indicate that the ammonia pre-treatment is so severe that it has caused a slight decrease even in low energy reactions.

In the present case, the decrease in isomerisation was not as large for conversion at 39 mass% compared with that at 65 mass% as seen in Table 23. This is due to the fact that higher conversions (therefore higher catalyst to oil ratios) exaggerate the shifts observed.

Table 23: Shifts in branched to linear paraffin ratios as a function of conversion on ammonia pre-treated catalysts when cracking squalane

Untreated Ammonia Conversion Carbon Decrease equilibrium pre-treated (mass%) number in B:L (%) catalyst catalyst 39 mass% 4 9.24 6.39 31% conversion 5 16.23 12.74 22% 65 mass% 4 11.87 6.67 44% conversion 5 19.23 13.93 28%

Once again a correlation of the ammonia pre-treated catalyst can be made with ZSM-5. Rajagopalan and Young [238] and Biswas and Maxwell [239] noted that reduced branching occurred in the presence of the pentasil additive. Other literature reports that the ratio of branched to linear C4 and C5 paraffins is greater than unity for Y and USY faujasites, but is shifted to lower values for ZSM-5 and p zeolite. Similar behaviour was observed by Abbot [202] and Zainuddin et al. [240] over

189 ZSM-5 and H-Y faujasite. They theorise that there is less branching over ZSM-5 and B-zeolite because of their smaller pore sizes. It was suggested by Abbot and Wojciechowski [35] that branched paraffins can be generated as primary products via the rearrangements of the carbonium transition state ions before their cracking and that, due to their larger pores, the faujasites favour to a larger extent the formation of branched intermediates. Abbot and Guerzoni [241] noted that the ratio of branched to linear paraffins cannot be related to the Lewis to Bronsted ratio of the zeolite but rather is a result of the pore zeolite structure.

The results of this thesis would indicate that although the ammonia pre treated catalyst appears to approach the reaction pathways of ZSM-5, this is not for the same reasons as those outlined above (namely smaller pore sizes). Ammonia pre-treatment is believed to not affect the pore dimensions of the catalyst. However, the adsorption of ammonia on the active sites results in a decrease in site density. Thus it is believed that the decrease in isomerisation seen is a direct result of a decrease in site density, not pore size or structure. The higher the site density, the more likely isomerisation reactions are likely to occur due to the proximity of the sites. However, because isomerisation reactions are a relatively low energy reaction it is expected that these reactions will not be as hindered as those that are higher in energy such as cracking reactions or hydrogen transfer reactions. It has already been seen that there is a great decrease in cracking reactions through the large decrease in conversion. Is there also a large decrease in hydrogen transfer reactions?

190 9.2.4 SHIFTS IN EXTENT OF HYDROGEN TRANSFER REACTIONS

Hydrogen transfer plays an important role in the gas oil cracking process, due to its link with product quality and catalyst performance in FCC. Increases in hydrogen transfer reactions will increase the yield of gasoline, but it will decrease the octane value of the gasoline. Hydrogen transfer reduces the amount of olefins in the product through bimolecular hydrogen transfer, whereby reactive olefins and naphthenes are converted to more stable paraffins and aromatics [242]. Further hydrogen transfer from aromatics, coupled with condensation and polymerisation can lead to the formation of coke. It is assumed that these bimolecular hydrogen transfer reactions occur between molecules co-adsorbed at adjacent acid sites, located at aluminium atoms bonded to the same silicon atom. These aluminium atoms are called next nearest neighbours (NNN). Suarez et al. [243] concluded that hydrogen transfer selectivity correlated well with the number of paired aluminium sites (framework aluminium atoms with one NNN, (next nearest aluminium neighbour)). In Y zeolites the number of NNN can vary from zero to three.

It is envisaged that as the amount of adsorbed ammonia increases, the number of NNN approaches zero. At a certain point, the sites will become sufficiently remote from each other that they become essentially identical. As hydrogen transfer reactions depend on close proximity of acid sites, it is expected that there will be a large reduction in hydrogen transfer reactions as a result of ammonia pre-treatment.

Many different indices can be used to estimate hydrogen transfer properties in gas oil cracking on a given catalyst or series of catalysts, for example, the iso-butane yield [17], the olefin or coke selectivity [44], the ratio of butenes to butanes [244], the ratio of paraffins to olefins in different hydrocarbon fractions [237], or the yield of olefins in a given

191 hydrocarbon range (e.g. gasoline range [245] or C4-C7 range [246]). Some of these methods will be used below to determine the change in hydrogen transfer reactions occurring on the catalysts.

As mentioned above, an expected large reduction in hydrogen transfer reactions as a result of ammonia pre-treatment can be seen by an increase in the yield of olefins.

Indeed the largest shift in selectivity was that displayed by the light olefins and in particular C4 olefins, which increased by 11 mass% over the ammonia pre-treated catalyst when compared to the untreated equilibrium catalyst. Figure 23 shows the dramatic relative increase in olefins.

Figure 23: Total olefin selectivity of squalane cracking at 65 mass% conversion as a function of ammonia pre-treatment

25 ------...... ------.----.

-0~ Cl) Cl) 15 - ctl E -Cl) ;:;::::C: 10 Cl) 0

5 ~ Untreated equilibrium catalyst -+-Anrronia pre-treated by soaking ex-situ 0 ~-----"""'------....,- -::::;:.:;;;;-_-_-_~..,--_------~-~~__. 0 1 2 3 4 5 6 7 Carbon Number

192 The rapid "center" cracking of the reaction intermediates from squalane cracking generates C3 and C4 olefins and carbenium ions. The saturation of these olefins over the Bronsted sites via hydrogen transfer reactions as well as the desorption of carbenium ions from the Lewis sites will generate C3 or C4 paraffins [24 7].

However, the large increase in light olefin yields combined with the relatively small increase in light paraffin yields shows that the reaction responsible for the saturation of the olefins has been hindered. Since hydrogen transfer reactions are responsible for this process, and since hydrogen transfer reactions require pairs of neighbouring Bronsted sites, it can be concluded that the ammonia is effectively being adsorbed at the Bronsted sites, thereby reducing Bronsted site density. This theory is confirmed by the works of several researchers [201, 243, 248, 249] who report that hydrogen transfer reactions require pairs of neighbouring Bronsted sites and are responsible for the saturation of olefins generated during the primary cracking process.

It is hypothesised that the suppression of bimolecular hydrogen transfer reactions (as a result of a decrease in site density) occurs because the time that the hydrocarbon remains on the catalyst surface is reduced. Thus the hydrocarbon desorbs as an olefin before it has time to undergo a hydrogen transfer reaction to give an increase in olefins relative to paraffins. Similar results have been seen in the literature for

catalysts with high silica to alumina ratios, such as ZSM-5, ~ zeolite or catalysts that have been dealuminated. For example, for the individual Y zeolite, the paraffin to olefin ratio has been found to be close to 1, suggesting a high degree of hydrogen transfer, whilst on H-ZSM-5 it is reported to be approximately 0.5. The strong tendency for conversion of olefins to paraffins via hydrogen transfer on the faujasite is reduced by the presence of H-ZSM-5 [65, 248].

193 Guerzoni and Abbot [250] also found that the addition of H-ZSM-5 to H y during the cracking of n-hexadecane enhances the formation of olefins in the range C3 - CS, with concurrent suppression of hydrogen transfer processes. Further, that ratios of branched to linear isomers are decreased for paraffins by addition of the pentasil, while the reverse is observed for olefinic products. This has also been found to be the case on the ammonia pre-treated catalyst when cracking squalane and is shown for C4 and CS in Figure 24.

Figure 24: Branched to straight olefin ratio at 65 mass% conversion, as a function of ammonia pre-treatment in the cracking of squalane

2.5 0 :;::::; ro L.. 2 C ..:: Q) I 0 I L.. 1.5 4k ro >E-- Q) -C 0 1 "C- Q) .r:. (.) Untreated equilibrium catalyst C 0.5 """*"" ro L.. a::i I -+-Ammonia pre-treated by soaking ex-situ 0 3 4 5 6 Carbon Number

However, there is a report in the literature that is in direct contradiction with the results observed. Lombardo et al. [65] found that, in their experiments with ammonia pre-treated catalysts, the paraffin to olefin ratios increased with increasing conversion. They conclude that the

194 olefins do not desorb readily from the H-M and H-Y zeolites studied, whereas paraffins do. Instead they report that the olefins formed oligomerise on the catalyst surface and the oligomeric ions produced rearrange, disproportionate, crack (releasing paraffins), and undergo hydrogen transfer producing surface residue (coke). They seem to contradict their own findings by postulating that hydrogen transfer reaction feasibility is dependent on the time the reacting species remains on the active site, on high acid site density and on high acid site strength [65]. These hypotheses are in accordance with the findings in this thesis, which state that acid site density decreases with ammonia pre-treatment and thus hydrogen transfer reactions also decrease,

Other researchers [63, 97] concur, reporting that acid strength increases with site isolation, (usually a consequence of increasing dealumination) and that both site density and time on site influence hydrogen transfer rates [65].

There is a possibility that the increase in olefins over paraffins is a result of a relative increase in Lewis sites, and therefore should be addressed. The reaction pathway would occur via carbenium ions [196] and/or via nonclassical carbonium ions [251] generating an olefin and a lighter carbenium ion. As secondary reactions, the catalyst may stimulate oligomerisation of olefins to aromatics. Thus the higher Lewis acidity would lead to:

1. A larger proportion of olefins as primary products, and 2. The preservation of the primary olefins because of the lower rates of hydrogen transfer due to a decrease in site density.

Although an increase in these reactions is observed with the heat pre treatment experiments over the untreated equilibrium catalyst, a pronounced increase in these reactions with the heat pre-treatment

195 experiments over the ammonia pre-treatment experiments is not observed. Thus the excess production of olefins does not seem to occur via Lewis site cracking since the heat pre-treated experiments would be expected to show the largest increases because the amount of Lewis sites are increased as Bronsted sites are converted to Lewis sites. Thus it may be concluded that this is not the primary mechanism by which the excess in olefins occur, but rather a cause of a decrease in site density as discussed above.

The above conclusion is confirmed in part by the works of Pine et al. [64] who report that, in their experiments with Y-zeolite, the higher the dealumination, the lower the site density and the more intense the site isolation, the greater was the decline in bimolecular hydrogen transfer reactions.

Another method of observing changes in hydrogen transfer reactions is by calculating the iso-butane to iso-butene ratio. Parra et al. [252], using 13C labeled iso-butene reacting with cyclohexane, were able to confirm that iso-butane is directly formed from iso-butene by hydrogen transfer. Thus if the iso-butane to iso-butene ratio decreases, this represents a decrease in hydrogen transfer reactions. The iso-butane and iso-butene yields for the cracking of squalane on untreated equilibrium and ammonia pre-treated catalysts is shown in Table 24.

196 Table 24: /so-butane and iso-butene yields for the cracking of squalane on untreated equilibrium and ammonia pre-treated catalysts

Untreated equilibrium Ammonia pre-treated Yield (mass%) catalyst catalyst Iso-butane 4.63 5.60 Iso-butene 6.59 13.15 iC4:iC4= 0.70 0.43

Although a small increase in iso-butane is seen, the large increase in olefins indicates that the hydrogen transfer reactions are hindered. This is confirmed by viewing the iso-butane to iso-butene ratio, which is observed to decrease from 0.70 to 0.43 over the ammonia pre-treated catalyst compared to the untreated equilibrium catalyst when cracking squalane.

Ratios of C2:C4 and C3:C4 can also be used to indicate extent of hydrogen transfer. High ratios are taken to be indicative of cracking via pentacoordinated carbonium ions, while low values are taken to suggest that cracking is predominantly through the "classical" reaction mechanism, involving ~-scission and bimolecular hydrogen transfer [54]. Table 25 shows the shifts exhibited in the C2:C4 and C3:C4 ratios when cracking squalane over untreated and ammonia pre-treated catalyst.

197 Table 25: Shifts in C2:C4 and C3:C4 ratios as a function of ammonia pre-treatment when cracking squalane

Conversion Catalyst C2 to C3 ratio C3 to C4 ratio Untreated 0.043 0.357 equilibrium 39 mass% Ammonia 0.050 0.455 pre-treated Percentage shift 16% increase 22% increase Untreated 0.018 0.326 equilibrium 65 mass% Ammonia 0.022 0.506 pre-treated Percentage shift 22% increase 18% increase

Since the C2:C3 and C3:C4 ratio increases by approximately 20% for the cracking of squalane over ammonia pre-treated catalyst compared to untreated equilibrium catalyst at two different conversions, it can be concluded that there are less reactions proceeding via the "classical" reaction mechanism, and hence a reduction in bimolecular hydrogen transfer.

Observations from this study also show an increase in olefins accompanied by a smaller increase in hydrogen. These results from squalane cracking on ammonia pre-treated catalysts show that hydrogen transfer was markedly decreased when compared to the untreated equilibrium catalyst, as were isomerisation reactions - but only to a lesser extent. It was also shown that the ammonia pre-treated catalyst exhibited qualities very similar to those reported of ZSM-5 by

198 several researchers [237, 243, 253]. They note that for zeolites with comparable silica to alumina ratio, but with different structures, the rate of hydrogen transfer and isomerisation increased in the order H-ZSM-5, H-beta, RE-USY. However these researchers explained their observations as a consequence of decreasing transition state shape selectivity effects. The experiments with the squalane over ammonia pre-treated catalysts indicates that the changes in reaction rates do not occur by this effect, but rather occurs because of a decrease in site density and a decrease in the time the initial reaction product remains adsorbed on the catalyst surface. Because the site strength and density is reduced, the reaction product cannot remain adsorbed, and therefore cannot undergo hydrogen transfer or oligomerisation.

Because paraffins and olefins are produced from initiation / desorption, initiation / ~ scission, and hydride ion transfer I ~ scission cycles, the paraffin to olefin ratio is a function of the rates of these cycles. Since rates of hydride ion transfer reactions increase when conversion increases, increasing conversion favours hydride ion transfer I ~ scission cycles. These cycles produce olefins and paraffins at a 1:1 ratio, which is the same ratio produced by initiation / desorption cycles.

In contrast, the initiation / ~ scission cycles produce two olefins per iso paraffin molecule consumed. Assuming that significant hydrogen transfer does not occur to form coke or aromatics, the paraffin to olefin ratio for species with three or more carbon atoms should always be less than 1, in agreement with the conclusions reached by others [56].

The rates of oligomerisation reactions depend on the olefin concentration in the gas stream and the surface coverages by carbenium ions. Therefore, the rates of these processes also increase with conversion, but at a slower pace than the hydrogen transfer reactions.

199 9.3 SUMMARY

In this section, squalane was used as a test molecule for the investigation of cracking catalyst deactivation. Although it appears that this is the first use of squalane in cracking catalyst studies, it has proven to be an excellent molecule for the analysis of cracking reactions and is representative of industrial fee9stocks. As expected there was a reduction in the conversion of squalane to cracked products as a result of catalyst pre-treatment. In addition there was a change in selectivity with a higher propensity towards formation of smaller cracking fragments at the expense of hydrocarbons C7 and larger.

It was found that the pre-treatment of the catalyst with ammonia was so severe that it caused a slight decrease even in low energy reactions such as isomerisation. Although the pre-treated catalyst appeared to mimic observations seen for ZSM-5, it was concluded that it was not the result of the same reason (namely smaller pore sizes). Instead it was determined to be a result of a decrease in site density, also causing a dramatic decrease in hydrogen transfer reactions.

200 10 METALS AND COKE DEACTIVATION OF CATALYST

10.1 INTRODUCTION

The nature of crude oils available to Australian refineries is changing as native crudes run out. The trend is towards heavier, more aromatic feedstocks, which can be expected to be harder to crack and to produce more coke. In addition they contain a higher degree of metals contaminants such as nickel and vanadium. Although generalities have been made as to the effect of these changes on processing and products, there have been little or no systematic studies found relating feedstock impurities to product selectivities and the catalyst characteristics. This prompted investigation to consider the effects of masking the sites on a cracking catalyst through deposition of nickel, vanadium, and coke, all known to be present in cracker feed and to be deposited on the catalyst. It was envisaged that the theories generated thus far in the thesis would be used to unravel the observations seen with the metals and coke deactivation of industrial equilibrium catalyst.

The performance of FCC catalysts for processing heavy resid feedstocks is often controlled by the tolerance of the catalyst to metal contaminants. These contaminants (mostly nickel and vanadium) are deposited on the catalyst when metal containing porphyrin complexes are present in the FCC feed.

The first step of this section of the study was to deposit nickel and vanadium on the catalyst such that the effects would be representative of commercial operation. In commercial operation, catalysts go through a natural aging process, in that some of the catalyst particles may have been in the unit for several months, while other catalyst particles may

201 have been only in the unit for a day. As far as metals existing on the catalyst surface goes, some particles may have more or less metals contamination than other particles as a result of their time in the unit and what feedstocks they have seen.

One method of metals deposition that seems to come quite close to commercial operation is cyclic deactivation. As the name implies, it involves the cracking of a metals impregnated feedstock over the catalyst, then the regeneration of the catalyst and so on in a cyclic fashion. Thus the catalyst becomes aged in a similar way to the commercial operation. This method of metals deactivation is quite expensive and involves a fairly large quantity of catalyst and feedstock. It seems to be a most useful technique when trying to evaluate catalysts for commercial operation.

For the purposes of this thesis, the aim was to introduce metals on an already aged catalyst, simulating the introduction of a heavily contaminated feedstock to the operation. The technique used to deposit nickel and vanadium on the equilibrium catalyst was the modified Mitchell incipient wetness technique as described in the experimental section. In short this technique is great for simulating a large influx of contaminated feedstock on the catalyst and has been used quite extensively in the cracking catalyst world. It involves the soaking of the catalyst in the appropriate metal naphthenate solution prior to calcination in a muffle furnace.

The literature reports that there are no significant differences in metal distribution across the catalyst particles resulting from different deactivation procedures such as incipient wetness and cyclic deactivation techniques when compared to industrial catalysts [254]. Although the incipient wetness technique does result in the metals being slightly more active for dehydrogenation, all yield shifts are

202 directionally the same when compared to the cyclic deactivation technique and industrial experience [254]. Since this study is more interested in the qualitative shifts due to metals on the catalyst, this incipient wetness method of metals deposition was satisfactory.

10.2 CATALYST CHARACTERISATION BY SEM / XRD

To investigate the adequacy of the metals impregnation, analyses such as scanning electron microscopy (SEM) and elemental analysis by SEM EDAX were performed. In addition a measure of the catalyst's crystallinity was obtained by performing XRD analysis.

10.2.1 SCANNING ELECTRON MICROGRAPHS

The pictures in this section show scanning electron micrographs of the equilibrium catalyst and the metals impregnated catalysts.

203 Figure 25: Scanning electron micrograph of equilibrium catalyst at a magnification of 100

From the above picture it can be seen that the catalyst particles are roughly spherical in shape spanning on average -50 µm in diameter. It can be noticed that a few of the particles are not spherical, appearing to be fused together, or even broken. Further, the particles appear to have "prickles" or small nodules on them. It has been reported by Lappas et al. [255] that, in contrast, fresh catalyst particles are almost perfectly spherical and do not contain these nodules, being smooth in appearance.

These nodules are expected to form on the catalyst over time in the fluid catalytic cracking unit and are caused by metals depositing on the catalyst surface.

204 Figure 26: Pictures of untreated equilibrium catalyst particles showing breakdown of the catalyst structure

In the above two pictures, two equilibrium catalyst particles are shown at magnifications of 1500 and 603 times respectively. From these two pictures it can be seen that the structure of the equilibrium catalyst particles has been modified due to the severe conditions in the fluid catalytic cracking unit. As a result the surface area and thus acid sites of the particle are changed. Hence the product selectivity of these particles will not be the same as those particles that have not been affected. This illustrates the importance of carrying out research on industrial equilibrium catalysts and not just pure laboratory prepared zeolites.

205 Figure 27: Pictures of untreated equilibrium and 20, OOO ppm nickel impregnated catalysts at 1OOO times magnification

Figure 27 above shows an untreated equilibrium catalyst particle on the left at 1OOO times magnification and a 20 ,000 ppm nickel impregnated catalyst particle on the right at the same magnification. On the untreated equilibrium catalyst particle the nodules are clearly observed. Although a little difficult to see, the nickel impregnated particle shows an increase in these nodules, slightly finer in size but certainly increased in quantity. This confirms that there is some noticeable physical difference between the untreated and nickel pre-treated catalysts.

206 Figure 28: Pictures of an untreated equilibrium and a 5, OOO ppm vanadium pre-treated catalyst particle at 20, OOO times magnification

The left picture above shows an untreated equilibrium catalyst particle at 20,000 times magnification. The "hills" on the particle surface create the nodules observed in Figure 28 above. The picture on the right shows a catalyst particle impregnated with 5,000 ppm vanadium also at 20,000 times magnification. It can be clearly seen that there are extra nodules on the already undulating surface of the catalyst. The formation of these extra nodules stems from the impregnation of the catalyst with the vanadium, which has formed a eutectic with the catalyst, at high temperature, and become bonded to the surface.

The above pictures show that there is a visibly observable difference between the untreated equilibrium catalyst particles and the metals impregnated particles. However it is a little hard to determine from these pictures whether the metals sit on the outside of the catalyst or actually penetrate the interior of the catalyst and to what concentration. Therefore it was important to analyse the catalyst particles using an elemental technique.

207 10.2.2 ELEMENTAL ANALYSIS

Elemental analysis was conducted on the untreated equilibrium catalyst and the metals impregnated catalysts in order to determine the level of metals penetration and to visibly compare the difference in metals concentration.

The first two pictures show a comparison of nickel concentration on the untreated equilibrium catalyst and the 20,000 ppm nickel pre-treated catalyst.

Figure 29: Nickel concentration on untreated equilibrium catalyst and 20, OOO ppm nickel pre-treated catalyst

The picture on the left shows the concentration of nickel on the untreated equilibrium catalyst and as expected, is quite low. The concentration of nickel on the equilibrium catalyst was 5236 ppm as mentioned in the experimental section. The picture on the right shows the concentration of the nickel on the 20,000 ppm nickel pre-treated catalyst. This picture shows quite a high density of nickel particles as

208 would be expected. Due to the method of sample preparation (as described in the experimental section), some of the catalyst particles have actually been sectioned and therefore a view of their cross-section is possible. This was not the case for the scanning electron micrographs. Thus, it can be observed that on some of the particles a ring of nickel appears around the catalyst particle. This indicates a tendency towards higher concentrations of metal close to the external surface.

Figure 30: Concentration of vanadium on the untreated equilibrium and 5, OOO ppm vanadium pre-treated catalyst

This set of pictures shows the concentration of vanadium on the untreated equilibrium catalyst and the 5,000 ppm vanadium pre-treated catalyst. The concentration of vanadium on the untreated equilibrium catalyst is 792 ppm as written in the experimental section . The vanadium pre-treated catalyst analysis certainly shows an increase in vanadium concentration, however the tendency for metal concentration near the catalyst surface is not evident as it was for the nickel

209 impregnated catalyst. The vanadium pre-treatment appears to have impregnated the catalyst particle more uniformly than the nickel pre treatment. This may also be due to the reduced concentration of vanadium (5,000 ppm) compared to nickel (20,000 ppm).

Kesavan and Law [256] report that it is generally accepted that vanadium is deposited on specific sites in the interior of the catalyst. This appears to be verified by the above pictures.

The final set of pictures for the elemental analysis shows the concentration of aluminium for the untreated equilibrium catalyst and the 20,000 ppm nickel pre-treated catalyst.

Figure 31 : Aluminium concentration for the untreated equilibrium catalyst and 20, OOO ppm nickel pre-treated catalyst

These pictures show the apparent decrease in aluminium concentration over the 20,000 ppm nickel pre-treated catalyst when compared to the untreated equilibrium catalyst. On the nickel pre-treated catalyst, the

210 aluminium concentration is patchy indicating a decrease, which can be linked to the reduction in catalytic activity that was observed for the nickel impregnated catalysts.

The above pictures show that the concentration of metals is certainly increased when compared with the untreated equilibrium catalyst. In addition, as a result of metals pre-treatment, the concentration of aluminium is decreased, corresponding to a decrease in catalytic activity.

10.2.3 XRD ANALYSIS

XRD analysis was performed on the catalyst samples in order to ascertain if there were any changes in crystallinity as a result of the pre treatment of the catalyst samples. Changes in the crystallinity of a catalyst sample has the potential to alter the product selectivity of a catalyst. As a result, it was important to establish whether the changes observed using the pre-treated catalysts were the result of changes in active sites or were affected by changes in the physical properties of the catalyst.

Three catalyst samples were analysed by X-ray powder diffraction analysis by monochromatized CuKa radiation using a Philips X'Pert system. The samples analysed were the fresh catalyst sample, the equilibrium catalyst sample and the 20,000ppm Ni pre-treated equilibrium catalyst sample. The XRD spectra obtained appear in Figure 32 below.

211 Figure 32: XRD analysis of fresh, e-cat and 20, OOOppm Ni impregnated catalysts.

I 19 ~ 18 - E 17 - ~ 16 I 15 T y 14 - 13 - C O 12 U 11 N T s

10.00 20.00 30.00 40.00 50.00 60.00 2-Theta Angle (deg)

The data showed that the fresh catalyst was predominantly a well crystallised zeolite of the faujasite type (Na2Al2Si4012.8H20). The equilibrium catalyst and nickel pre-treated samples also contained peaks corresponding to zeolite but these were much lower and some were even missing indicating both a less well developed crystallinity of the zeolite and a much lower zeolite content in the samples. In addition both of these catalysts showed broad bands of X-ray amorphous material.

Thus the results indicate that both e-cat samples have both undergone some change to the original phase composition, probably as a result of the temperatures and conditions experienced in the fluid catalytic cracker. However, the results also show that the difference in crystallinty between the e-cat catalyst sample and the 20,000 ppm Ni pre-treated sample is negligible when compared with the fresh catalyst sample. This is as expected since the e-cat and the Ni pre-treated

212 catalysts all undergo the same experimental process and any difference experienced by the two samples is insignificant when compared with the average age of the catalyst particles from the fluid catalytic cracker. The "equilibration" of the catalyst in the fluid catalytic cracker is where the majority of crystallinity changes would take place.

10.2.4 SURFACE AREA MEASUREMENTS

The surface area of the fresh, equilibrium and metals impregnated catalysts was also measured in order to narrow down the possibilities of other factors affecting the shifts in selectivity. The results from the surface area experiments is tabulated in Table 26.

Table 26: Surface area measurements for treated and un-treated catalysts

Catalyst Surface Area (m2/g) Error (m 2/g) Fresh 252.3 5.4 Equilibrium 98.0 1.6 20,000 ppm Ni treated 94.1 1.6 5,000 ppm V treated 95.1 1.6

The results show that the equilibrium catalyst sample and both metals impregnated catalysts possess a large difference in surface area from the fresh catalyst as expected. As mentioned in the x-ray diffraction results, this is a result of the temperatures and conditions experienced in the fluid catalytic cracker.

213 The results also show that the difference in surface area between the e cat catalyst sample and the 20,000 ppm Ni and 5,000 ppm V pre treated samples is negligible. Again, this is expected since the e-cat and the Ni / V pre-treated catalysts all undergo the same experimental process and any difference experienced by the two samples is insignificant when compared with the average age of the catalyst particles from the fluid catalytic cracker.

As a result of the XRD measurements and the surface area measurements, it can be concluded that the e-cat and pre-treated samples used in the experimentation are almost identical and that changes in product selectivity are unlikely to be a result of changes to the physical properties.

10.3 NICKEL PRE-TREATED CATALYSTS

One of the major metal contaminants in industrial feedstocks is nickel. High levels of nickel in the feedstock can poison the cracking catalyst leading to changes in activity and product selectivity.

In order to systematically examine the effects nickel has on an industrial cracking catalyst, nickel naphthenate was deposited on the equilibrium catalyst at levels of 5,000 ppm, 10,000 ppm and 20,000 ppm total nickel. The equilibrium catalyst already contained 5,236 ppm Ni on it as a result of time spent in Caltex's fluid catalytic cracking unit.

Reactions of squalane were then carried out over the nickel pre-treated catalysts and compared with the untreated equilibrium catalysts.

In Figure 33 it can be seen that the addition of nickel to the equilibrium catalyst reduces the conversion for all catalyst to oil ratios tested.

214 Figure 33: Conversion of squalane over nickel pre-treated catalysts

100 -..------,------,

-0~ Cl) Cl) ctl E 60 -C: 0 ·u5 I I '- Q) 40 > C: j __ L~ 0 I x Untreated equilibrium catalyst 1 () • 5,000 ppm nickel pre-treated 20 h. 10,000 ppm nickel pre-treated o 20,000 ppm nickel pre-treated 0+------,.-----1---I 0.0 0.5 1.0 1.5 2.0 2.5 3.0 3.5 Catalyst to oil ratio

It can be observed that, beginning at a catalyst to oil ratio of about 2.0, only a slight increase in conversion can be observed for increasing catalyst to oil ratios. This was also observed to be the case with the ammonia pre-treatment experiments. It indicates that, at a certain level of catalyst to oil ratio, any additional catalyst does not assist in the cracking of the feedstock, which can be concluded to be the result of the poisoning of sites and the resulting decrease in site density. This is not to say that there is no cracking within the already cracked products, just no further conversion of feedstock into products.

This observation has also been made in practice at the Caltex Lytton refinery. At times when the process has reached high metals levels on the catalyst, the increasing of the catalyst to oil ratio does not appear to yield any appreciable increase in conversion and often the process

215 needs increased additions of fresh catalyst to the system to help the unit "recover".

10.3.1 HYDROGEN AND METHANE SELECTIVITY

The largest relative change in yields was observed for hydrogen selectivity as shown in Figure 34. It is interesting to note that if Figure 33 and Figure 34 are correlated, then as the catalyst to oil ratio increases past about 2.0, the extra catalyst seems to predominantly contribute to increases in hydrogen as seen by the sharp increase as conversion increases. Presumably this is due to the catalysis of dehydrogenation reactions. Thus it can be seen that, although no further feedstock is converted into products above a catalyst to oil ratio of about 2.0, there is still reactions proceeding, predominantly for the formation of hydrogen.

216 Figure 34: Hydrogen selectivity of nickel pre-treated catalysts in the cracking of squalane

0.35 ------

X Untreated equilibrium catalyst I 0

0.30 • 5,000 ppm nickel pre-treated f-1--~------ii A 10,000 ppm nickel pre-treated -?ft. 0.25 O 20,000 ppm nickel pre-treated 1-~------1 Cl) Cl) ro E 0.20 -C: Q) eo, 0.15 "'O f 0.10 X 0.05

0 20 40 60 80 100 120 Conversion (mass%)

This sharp rise can also be seen for methane, although the yield of methane is very similar between the differing concentrations of nickel pre-treated catalysts. This indicates that there is a certain concentration of pre-treatment that yields the same amount of methane. It can be seen that the methane yield increases for increasing catalyst to oil ratios, so it seems unlikely that the cause is due to the total amount of nickel present in the system, but rather that it is the concentration on the catalyst surface that catalyses the formation of methane. The selectivity of methane for the cracking of squalane over the nickel pre treated catalysts can be seen in Figure 35.

217 Figure 35: Methane selectivity of squalane cracking as a function of nickel pre-treatment

0.20 ------

0.16 • -;:f?_ -0 Cl) Cl) ro 0.12 E X -Q) C: ro 0.08 .c I -Q) ~~ ___L_ ~ XUntreated equilibrium catalyst l ~~----, +5,000 ppm nickel pre-treated ~

A. 10,000 ppm nickel pre-treated I' O 20,000 ppm nickel pre-treated 0.00 ...,______....,. ___ .,.

0 20 40 60 80 100 120 Conversion (mass%)

Since the yields of the nickel pre-treated catalysts are bunched together and clearly on a separate trend compared to the equilibrium catalyst (which also contains industrially deposited nickel), it would indicate that the nickel impregnation as a result of the modified Mitchell method does indeed show a difference in how it makes the catalyst behave when compared to the equilibrium catalyst with respect to dehydrogenation reactions. This is in accordance with the findings of Bendikssen et al. [254] who report that the modified Mitchell method does lend itself towards higher activity for dehydrogenation reactions, as mentioned above.

A more correct measure for the occurrence of dehydrogenation is the molar ratio of hydrogen to methane. This ratio is shown to significantly increase as a function of nickel contamination in Figure 36.

218 Figure 36: Hydrogen to methane ratio of squalane cracking as a function of nickel pre-treatment

4-----...,...---~------. X Untreated equilibrium catalyst 0 A 10,000 ppm nickel pre-treated :;:::::; O 20,000 ppm nickel pre-treated ~ 3 ~------11 Q) + 5,000 ppm nickel pre-treated C ro ..c_.. Q) E 2 .9 C Q) O> I e 1 "O i >, I _; I •I • • • X

0 20 40 60 80 100 Conversion

The above results confirm the hypothesis that the nickel pre-treated catalysts show increased dehydrogenation as the concentration of nickel on the catalyst is increased.

10.3.2 LIGHT OLEFIN SELECTIVITY

The yields for ethene and for propylene also show a sharp increase at higher conversions similar to that seen for hydrogen and methane reflecting the increasing amount of dehydrogenation reactions occurring as catalyst to oil ratios increase. An important question at this stage, that needs to be answered, is "how much of this effect is due to the blocking of active sites by the nickel, and how much is due to the

219 chemical effects of nickel?" This question can only be answered with further experimental work and is dealt with in the section on sand I catalyst split systems. The essence of this section is the cracking of squalane over a multistage catalyst system. The process involves the cracking of the squalane over equilibrium catalyst and then over nickel doped sand. This is intended to clarify what actual effects the nickel has on the process independent of the blocking or poisoning of acid sites.

For the light gases larger than 3 carbons in chain length, the increases seen above for the hydrogen, methane and propylene are not seen. Rather, the yield of olefins (both branched and straight chain) shows a decrease with increasing nickel concentration. Figure 37 depicts this decrease for 3 methyl butene.

220 Figure 37: 3-methyl-1-butene selectivity of squalane cracking over nickel pre-treated catalysts

0.4

x Untreated equilibrium catalyst 'cf!. f/) • 5,000 ppm nickel pre-treated f/) co 0.3 A 10,000 ppm nickel pre-treated E o 20,000 ppm nickel pre-treated -Q) C Q) ::, 0.2 .c- I T""" I >-..c Q) 0.1 -E ('I')

0.0 0 20 40 60 80 100 Conversion (mass%)

These results indicate that olefins with carbon numbers 4 or greater are remaining on the catalyst (as opposed to being desorbed) and are hypothesised to be the precursors for the manufacture of coke and hence decrease in yield selectivity.

This is confirmed by the findings of Groten et al. [257] who report that as the average surface species (carbenium ion) becomes more dehydrogenated, it also becomes less reactive both in bimolecular chain propagation (hydrogen transfer) with gas phase reactant molecules and in olefin formation by desorption.

Further, the above hypothesis is confirmed by observing the coke selectivity of the cracking of squalane on nickel pre-treated catalysts.

221 The coke selectivity was observed to increase for the increasing concentration of nickel. Figure 38 shows this.

Figure 38: Coke yield from squalane cracking on nickel pre-treated catalysts.

4------X Untreated equilibrium catalyst l

+ 5,000 ppm nickel pre-treated 1 3 I A 10,000 ppm nickel pre-treated I -'#. L__? 20,000 ppm nickel pre-treatedT ____ _I Cl) I~- ___ Cl) Cll E. 2 Q) ! ::it:. • 0 (.)

0 20 40 60 80 100 Conversion

This graph also depicts the sharp rise as seen for the hydrogen, methane and light gases, further confirming the hypothesis that olefins greater than 3 carbon numbers in length proceed to the formation of coke.

It can be concluded from the above findings that the variations in the surface contamination of nickel can be used to explain the hydrogen and coke selectivity differences in cracking catalyst samples. Further that because hydrogen and coke are both undesirable products of the cracking reaction (for a commercial operation), that the profitability of

222 the fluid catalytic cracking unit can be severely affected by the concentration of nickel in the feedstocks.

10.3.3 LIGHT CYCLE OIL SELECTIVITY

An interesting observation was made for the cracking of squalane over the nickel pre-treated equilibrium catalyst. At low conversions the yield of light cycle oil seems to increase as conversion increases. However, at high conversions, the yield of light cycle oil decreases as conversion increases. Thus the graph produced forms a horseshoe shape. This can be seen in Figure 39.

223 Figure 39: Light cycle oil yield of squalane cracking as a function of nickel pre-treatment.

13 ------..------. • 12

,-.. 11 cf:. ~ 10 ro ~ X -0 9 (.) _J 8 x Untreated equilibrium catalyst • 5,000 ppm nickel pre-treated 7 A 10,000 ppm nickel pre-treated o 20,000 ppm nickel pre-treated 6 -----...... -----,.------'! 0 20 40 60 80 100 Conversion

It is envisaged that the yield of light cycle oil increases as conversion increases until the point at which the light cycle oil starts to crack itself, forming smaller hydrocarbons and thereby decreasing in yield.

It has been decided to term this change in yield the overcracking point. For light cycle oil this occurs at around 60 mass% conversion. At this conversion, there exists a peak in light cycle oil yield. Beyond this point, the light cycle oil yield decreases indicating that the light cycle oil is being overcracked into lighter products.

In general the light cycle oil yield on the nickel pre-treated catalysts does show a lower selectivity than that of the untreated equilibrium catalyst. At a conversion of about 70 mass% a decrease in light cycle

224 oil yield of about 3 mass% can be seen over the catalyst pre-treated with 20,000 ppm nickel when compared to the untreated equilibrium catalyst. This is a significant finding as it would suggest that, in a commercial operation striving to make light cycle oil, profits could quickly be eroded if the nickel contamination on the catalyst is not kept to an operationally achievable minimum.

10.4 SAND/ EQUILIBRIUM CATALYST SPLIT SYSTEM TESTING

This section of testing involved cracking the squalane over a two stage catalyst system and was designed to answer the question "Are the selectivity shifts a result of the nickel on the catalyst blocking / poisoning the acid sites, or are they due to the chemical effects of the nickel independent of any interactions with the catalyst?"

In these experiments sand was doped with nickel using the modified Mitchell method. The doped sand was then placed in the glass reactor both with and without an amount of equilibrium catalyst placed on top. The squalane was then injected over the catalyst at 500 °c as per previous experiments.

This split catalyst design allows the evaluation of the cracking capacity of the actual metals, independent of any interaction with the equilibrium catalyst. It removes all effects associated with the poisoning or blockage of acid sites on the equilibrium catalyst so that the chemical effects of the nickel can be evaluated.

Comparing the cracking of squalane over just sand (at a sand to oil ratio of -4) with the cracking of squalane over equilibrium catalyst with 10,000 ppm nickel (at a catalyst to oil ratio of -2) shows, as expected, that the conversion is much reduced over the sand, with conversions of

225 3.88 and 61.84 mass% respectively. When squalane was cracked over nickel pre-treated sand ( containing -10,000 ppm nickel), the conversion slightly improved to 8.53 mass% again at a sand to catalyst ratio of -4, but was still significantly below that of the nickel impregnated equilibrium catalyst. This is no real surprise as the absence of the cracking catalyst is expected to produce a much decreased conversion since there are no active sites on the sand. The cracking over the sand essentially represents pure thermal cracking.

However it does serve a purpose to analyse the product selectivity of the sand against the nickel impregnated catalyst. In order to compare the yield selectivities of the above experiments, the yields from the squalane cracking over sand need to be normalised to the conversion of the squalane over equilibrium catalyst with 10,000 ppm nickel. The product selectivity of the above experiments is shown in Table 27 below. Only the yields from hydrogen to C4 compounds inclusive are shown since there was a negligible amount of higher molecular weight hydrocarbons formed from the cracking over sand. This very fact actually ties in quite well with the observations of cracking over heat and ammonia pre-treated catalysts in that, as the activity of those catalysts were reduced by pre-treatment, there was an observed shift towards lighter compounds with a noticeable decrease in high molecular weight hydrocarbons.

226 Table 27: Product selectivities for squalane cracking over sand and nickel doped sand

Catalyst type E-Cat w/ 10,000 Sand Sand + 10,000 ppm nickel ppm nickel Feed Squalane Squalane Squalane Cat/oil ratio 1.99 3.97 4.09 Normalised 61.84 61.03 61.41 conv. (mass%) H2 0.01 1.16 2.21 C1 0.11 4.43 1.12 C2 0.07 5.71 1.09 C2= 0.22 10.87 1.43 C3 0.27 15.70 3.44 C3= 3.81 3.73 0.82 iC4 3.14 0.40 0.17 iC4=/1C4 4.22 5.92 1.66 nC4 0.26 0.71 0.20 t2C4= 1.66 0.00 0.08 c2C4= 1.21 0.00 0.26 Coke 1.77 0.08 24.07

Since the conversion of the squalane over sand is so low, the normalisation of the yields can skew the results. However, in general observation, the cracking of squalane over the sand results in large quantities of dry gas, as expected from the thermal cracking. The presence of the nickel on the catalyst appears to reduce the thermal cracking load, with the exception of the hydrogen yield which increases. This is due to the dehydrogenating aspect of the nickel. The nickel also

227 results in a large yield of coke, presumably corresponding with the increase in hydrogen yield.

The above result of the nickel on sand cracking of squalane shows that it is not the nickel alone that causes the yields shown for the cracking of squalane on equilibrium catalyst impregnated with nickel. This is to be expected as there has been no catalytic cracking introduced on the nickel impregnated sand.

Since the conversion on the sand is so low, it also indicates that the nickel on the equilibrium catalyst should not appreciably participate in affecting the feedstock (squalane) whether the nickel is on equilibrium catalyst or on sand. In other words, while the feed remains uncracked the actual nickel component does not chemically modify the feedstock.

The above section only tells the first part of the story, the next section reveals the yields when squalane is cracked over the two stage system comprising equilibrium catalyst and nickel impregnated sand. Thus as the feedstock passes over the equilibrium catalyst and is cracked, the cracked products pass immediately over the 10,000 ppm nickel impregnated sand. These results are compared with the 10,000 ppm nickel impregnated equilibrium catalyst.

228 Table 28: Product selectivities for squalane cracking over a single and two stage sand I catalyst system

Catalyst type E-Cat w/ 10,000 ppm E-Cat + Sand w/ nickel 10,000 ppm nickel System Single stage system Two stage system Feed Squalane Squalane Cat/oil ratio 1.99 1.98 Normalised conv. 61.84 61.84 (mass%) H2 0.0129 0.0243 C1 0.1090 0.0953 C2 0.07 0.07 C2= 0.22 0.16 C3 0.27 0.07 C3= 3.81 3.67 iC4 3.14 3.29 iC4=/1C4 4.22 3.91 nC4 0.26 0.30 t2C4= 1.66 1.69 c2C4= 1.21 1.27 3m1C4= 0.16 0.08 2mC4 (iC5) 3.12 3.62 1C5= 0.34 0.38 2m1C4= 1.07 1.16 nC5 0.21 0.25 t2C5= 1.17 1.30 c2C5= 0.61 0.66 2m2C4= 2.32 2.54 C4m2C5= 0.00 0.16

229 2,3dimC4 0.23 0.29 t4m2C5= 0.09 0.11 2mC5 1.45 1.89 2m1C5= 0.21 0.26 3mC5 0.76 0.98 1C6= 0.28 0.54 c3C6= 0.26 0.00 t3C6= 0.28 0.31 2m2C5= 0.50 0.34 c3m2C5= 0.65 0.57 t2C6= 0.40 0.76 c2C6= 0.25 0.49 t3m2C5= 0.59 0.32 2,3dim2C4= 0.16 0.69 2mC6 1.64 1.91 3mC6 0.90 1.01 C7 olefins 3.46 2.44 2mC7 0.59 0.68 4mC7 0.66 0.73 3mC7 0.20 0.28 CB olefins 3.14 4.57 C9s 3.57 3.56 C10s 4.35 3.73 C11s 3.72 3.59 Coke 1.77 1.54

The above table shows that the yields of the two different systems are fairly similar. The two stage system showed a doubling of the hydrogen yield, yet this did not appear to be reflected in the coke yields as in the previous section. The results indicate that the two stage system mimics

230 the single stage system comprising the nickel impregnated equilibrium catalyst. This implies that the nickel has no physical or chemical effect on the equilibrium catalyst. In other words it would appear that the nickel on the equilibrium catalyst does not block or poison any acid sites, or that the sites that it does block or poison do not actively contribute to the cracking process. Instead it can be seen that the product selectivities stem from almost a purely chemical reaction between the cracked hydrocarbon products and the nickel to produce the results seen in both the single and two stage system.

To clarify this result, it is pertinent to look at other factors that may indicate a subtle shift in hydrogen transfer reactions or the like. As previously measured, the comparison of the paraffin to olefin ratio gives an indication of extent of hydrogen transfer reactions. The comparison of this ratio for C3 to C5 was conducted for both systems and the results are displayed in Table 29.

Table 29: Paraffin to olefin ratio for light hydrocarbons on a single and two stage sand I catalyst system in the cracking of squalane

Carbon number E-Cat E-Cat + Sand w/ 10,000 ppm nickel w/ 10,000 ppm nickel 3 0.07 0.02 4 0.48 0.52 5 0.59 0.63

The above table shows that there is negligible difference between the two systems indicating that there is no appreciable difference in extent of hydrogen transfer reactions.

231 Likewise the branched to linear paraffin ratio shows little difference with the ratio for CS paraffins at 14.81 and 14.32 for the single stage and two stage system respectively.

The results of the above analyses is conclusive in indicating that the effects of nickel contamination are purely a result of the nickel itself and not through any interaction it has with the equilibrium catalyst.

10.5 VANADIUM PRE-TREATED CATALYSTS

As mentioned previously, along with nickel contamination, vanadium is also renowned for being a severe catalyst poison [258]. Thus it was pertinent to investigate the effects of vanadium on the industrial catalyst. Two different concentrations 2,500 ppm and 5,000 ppm total vanadium were used for the vanadium pre-treatment of the equilibrium catalyst. The vanadium was deposited on the equilibrium catalyst using the modified Mitchell method described earlier. Squalane was used as the feedstock to measure the selectivity changes on the vanadium pre treated catalysts.

For the vanadium pre-treatment, most of the effects seen were similar to those seen with the nickel pre-treatment. For example a decrease in larger carbon chain length olefins and an increase in coke selectivity was observed with increasing vanadium concentration. Three of the more notable exceptions were the conversion and methane selectivities and the lack of response to varying concentrations of vanadium.

Firstly, observation of the conversion versus catalyst to oil ratio on the vanadium pre-treated catalysts reveals that, contrary to the observation made with the nickel pre-treatment, at low catalyst to oil ratio the

232 vanadium pre-treated catalysts actually appear to enhance the conversion, when compared with the untreated equilibrium catalyst. At a catalyst to oil ratio of about 0.7, the vanadium pre-treated catalysts show an increase in conversion of about 25 mass%. This can be seen in Figure 40 below.

Figure 40: Conversion of squalane over vanadium pre-treated equilibrium catalyst

_ 80 'cf2. Cl) Cl) E 60 -C: 0 ·u5 '-Q) 40 > C: 0 ()

0 +,------'! 0 2 3 Catalyst to oil ratio

This is an unusual observation and indicates that the vanadium itself is catalysing the conversion of squalane feedstock into smaller hydrocarbons. Further analysis of yields is required to see exactly what products are being made from this extra conversion. This will assist in determining what reactions are being catalysed. A possible explanation is that at low catalyst to oil ratios, a large proportion of the feedstock passes over already deactivated catalyst and that the deactivated catalyst, together with the vanadium promotes the quick conversion of

233 feedstock into hydrogen and coke (which are both registered as converted products), hence relatively higher conversion than over the untreated equilibrium catalyst.

Another point of note is the different shape of this curve relative to the curve of the untreated equilibrium catalyst which indicates once again, as with the nickel pre-treated catalysts, that the modified Mitchell method for metals impregnation does not do a very good job when it comes to modelling an industrial process.

It can also be observed that the concentration of vanadium deposited by the modified Mitchell method does not appreciably affect the amount of conversion for the levels of vanadium pre-treatment tested. This is clearly contrary to the observation made for the nickel pre-treatment. From this it can be concluded that the vanadium pre-treatment affects the cracking reactions differently to the nickel pre-treated catalysts.

The third notable difference from the nickel pre-treatment is the yield of methane. Figure 41 shows the yield of methane as a function of conversion. The characteristic sharp increase that was seen with the nickel contaminated catalyst can also be seen with the vanadium pre treatment. However what is different is that at lower conversions, the yield of methane on the vanadium pre-treated catalysts is actually much lower than that of the untreated equilibrium catalyst. At higher conversions, the yield of methane on the vanadium pre-treated catalysts approaches that of the equilibrium catalyst, finally exceeding the yield of the equilibrium catalyst for the highest conversions observed.

In comparing these results with those found for the nickel pre-treatment experiments, it can be concluded that vanadium also catalyses dehydrogenation reactions.

234 Figure 41: Methane selectivity of vanadium pre-treated equilibrium catalysts when cracking squalane

0.12 ------

I x Untreated equilibrium catalyst X 0.10 • 2,500 ppm vanadium pre-treated A 5,000 ppm vanadium pre-treated -0~ ~ 0.08 ro E '; 0.06 C: ro .c. a5 o.o4 ~ 0.02

0.00 +----...------0 20 40 60 80 100 Conversion (mass%)

The above results from the cracking of squalane on nickel and vanadium pre-treated catalysts lends itself to a few conclusions. Firstly that the process used to impregnate the equilibrium catalyst does not adequately simulate the industrial deposition of metals on the catalyst. This can be seen in most of the graphs where the trendlines of the artificially doped catalysts does not have the same shape as that of the equilibrium catalyst, which already contains levels of nickel and vanadium on it.

An alternative theory to this observation is that the extra concentration of metals on the equilibrium catalyst could be shifting the curve shape, as a result of overcoming a certain metals concentration, to dramatically change the ratio of sites deactivated by nickel and vanadium. This does seem to be a less likely explanation.

235 10.6 PRE-COKING OF CRACKING CATALYSTS

Catalyst deactivation has been the subject of intensive study over the past 50 years [120, 147, 259-263]. One particular aspect of deactivation, which is inherent in many catalytic processes and has been the subject of much debate is the deposition of coke occurring with the main reaction process.

It was decided to revisit the deactivation by pre-treatment with a hydrocarbon (pre-coking), as had been discussed earlier in the thesis, using a more systematic approach. Samples of equilibrium catalyst were sent to Akzo Nobel in the United States where coke was deposited on the catalyst, using a short contact time reactor, to three different carbon levels (1.4, 2.5 and 3.4 mass%). This style of reactor uses a larger amount of catalyst than the microreactor and therefore is useful for pre-coking a large amount of catalyst uniformly.

The pre-coked samples were then placed in the microreactor and squalane was cracked over them.

In viewing the conversion of squalane, the pre-coked catalysts show a marked deactivation of the catalyst even at what would be described to be fairly low coke levels.

The reduction in conversion can be seen in Figure 42.

236 Figure 42: Conversion of squa/ane over pre-coked equilibrium catalysts

100 ------

80 -0~ en en ro E 60 -C 0 ·u5 .... -- Q) 40 > I C 0 -- ~l~ x Untreated equilibrium catalyst] (_) I 20 -1.+-----=~------i • 1.4% coke pre-treated A 2.5% coke pre-treated o 3.4% coke pre-treated ! 0 ------! 0 2 4 6 8 10 12 Catalyst to oil ratio

It can be seen that for the pre-coked catalysts, as with the ammonia pre-treated catalysts, the amount of feed converted showed little response despite increasing the catalyst to oil ratios to almost 12. It is believed that this is a result of the coke blocking the most active sites on the catalyst and thereby significantly reducing the activity. It can be interpolated from the graph that as little as 3.4 mass% carbon on catalyst can reduce the activity of the catalyst by almost 80% at a catalyst to oil ratio of around 3 when compared with the untreated equilibrium catalyst.

This is confirmed in part by the findings of Figueiredo et al. [264] who studied the adsorption of propene and its conversion into coke on an FCC catalyst. They showed that the coke formation involved the strongest acid sites, capable of retaining 2,6 dimethylpyridine at 450 °C. The nature and the catalyst's reactivity were found to change with

237 deposition temperature. Hence the selectivity changes that they observed with the coked catalysts were concluded to be a result of coke being selectively deposited on the strongest sites first.

Naturally, a closer examination of the present results was required to see what selectivity changes if any were occurring on the pre-coked catalysts.

10. 6.1 HYDROGEN SELECTIVITY

In viewing the yield of hydrogen as a function of conversion, it can be seen that the pre-coked catalysts show a similar shaped hydrogen curve as that of the nickel pre-treated catalysts. However the curve is steeper and occurs at a lower conversion. Figure 43 shows this.

238 Figure 43: Hydrogen selectivity of squalane cracking over pre-coked catalysts

0.20 ------.------• X Untreated equilibrium catalyst I o 3.4% coke pre-treated 0.16 +----·------A 2.5% coke pre-treated -0~ • 1.4% coke pre-treated (J) (J) E 0.12 - A- - -C Q.) g> 0.08 .... X "C >, I 0.04

0.00 J.,--~--"""""!IIC====::::.__:_:...__J.. __ ___j ____J 0 20 40 60 80 100 120 Conversion (mass%)

If the hypothesis presented earlier for the cracking of squalane on ammonia pre-treated catalysts, regarding the need for paired alumina sites for hydrogen transfer, is accepted, then it is not surprising to observe hydrogen transfer sites to be more coke sensitive than isomerisation sites. As a result, when the catalyst becomes increasingly coked, the quantity of paired alumina sites significantly reduces lending itself to rapid increases in molecular hydrogen formation.

239 10.6.2 LIGHT OLEFIN SELECTIVITY

Unusually a reduction in total light olefins can also be seen as a function of pre-coking level. This is displayed in Figure 44.

Figure 44: Light olefin selectivity of squalane cracking over pre-coked catalysts

40 ------,-----~------~· I I x Untreated equilibrium catalyst !

-?f2. I • 1.4% coke pre-treated I ~ 30 ~ 11;. 2.5% coke pre-treated ro E [ o 3.4% coke pre-treated -co UI 20 N (..) -ff) C ~ 10 Q) 0

0 20 40 60 80 100 Conversion

This decrease in olefins occurs for both branched and straight chain olefins as demonstrated in Figures 44 and 45 below.

240 Figure 45: /so-butene selectivity of squa/ane cracking over pre-coked catalysts

1------6

-'?f2.. 5 Cl) Cl) CU E 4 - -(l) C 2 3 I :::J .c ,---c------1 __ _ I o 2 X Untreated equilibrium catalyst Cl) I 03.4% coke pre-treated '

1 +------~~~~------, A2.5% coke pre-treated + 1.4% coke pre-treated

0 20 40 60 80 100 Conversion (mass%)

241 Figure 46: Trans-2-butene selectivity of squalane cracking over pre coked catalysts

3.0

X ~ 2.5 ~ Cl) Cl) ro 2.0 E -Q) C 1.5 ...... Q) :::, ..c I N ------;;"------,7"--I_X_U_nt-re~ate-d-eq_u_ilibriu~m-c-at-aly-st-. I 1.0 Cl) C + 1.4% coke pre-treated ro ...... I... 0.5 A 2.5% coke pre-treated O 3.4% coke pre-treated ___J

0.0 0 20 40 60 80 100 120 Conversion (mass%)

It is believed that the reason for the decrease in olefins is due to the longer diffusion path as a result of pore plugging from the pre-coking. As a result the olefins do not readily desorb (unlike the observations made on the ammonia pre-treated catalyst), nor can they undergo hydrogen transfer reactions due to the increased site distance (also a result of the pre-coking), instead they remain on the catalyst and oligomerise to form coke. This theory forms the hypothesis that the pre coking of the catalyst affects the shape selectivity of the catalyst such that the new catalyst structure shows a significant effect on space demanding reactions, such as bimolecular reactions (hydrogen transfer).

242 10.6.3 ISOMER/SAT/ON

As observed with the earlier types of pre-treatment tested, the pre coking also showed a decrease in branched to linear paraffin ratio for light hydrocarbons. This confirmation can be seen in Figure 47.

Figure 47: Branched to linear paraffin ratio for C4 and CS as a function of catalyst pre-coking in the cracking of squalane

25 0 :.::; ro i... C: 20 iE ro roi... c. 15 roi... Q) C: -+-Untreated equilibrium catalyst - 10 0 """'*- 2.5% coke pre-treated "C- Q) .c. (.) 5 C: ro coi... 0 3 4 5 6 Carbon Number

Once again it is believed that this is due to a decrease in isomerisation reactions as discussed in the analysis of ammonia pre-treated catalysts. Although isomerisation is a low energy reaction, its progress is hindered by the pre-coked catalyst, primarily due to a reduction in site density and therefore a reduction in secondary reactions.

243 However, for the olefins, the opposite is true, where the branched to linear ratio is observed to increase as level of pre-coking increases. This can be seen in Figure 48.

Figure 48: Total branched to linear olefin ratio of squalane cracking as a function of catalyst pre-coking

2.5

u :;:::; ro .... 2.0 C ~ .;:: Q) ·~ 0 " .... 1.5 , ro , " Q) ~ -C ...... 0 1.0 "O Q) .r= x Untreated equilibrium catalyst (.) o 3.4% coke pre-treated C ....ro 0.5 A 2.5% coke pre-treated ~ CD + 1.4% coke pre-treated

0.0 0 20 40 60 80 100 Conversion

This observation ties in with the earlier hypotheses regarding increased olefin residence time on the pre-coked catalysts. As a result of the extra time that the olefins reside on the catalyst due to the longer diffusion path created by the pore plugging, and since hydrogen transfer reactions are more sensitive to coking than isomerisation reactions, then the amount of branched olefins increases relative to the linear olefins. In other words, the longer the olefins reside on the catalyst due to diffusional effects before desorbing, the more likely they are to isomerise than undergo hydrogen transfer, hence the increase in

244 branched to linear ratios for the pre-coked over untreated equilibrium catalysts.

10.6.4 LIGHT CATALYTIC NAPHTHA, LIGHT CYCLE OIL AND COKE SELECTIVITY

Figure 49 shows the yield of light catalytic naphtha as function of level of pre-coking.

Figure 49: Light catalytic naphtha selectivity of squalane cracking as a function of pre-coking of equilibrium catalysts

80 ------

60 ,-...

0~ Cl) Cl) CU E 40 -z (..) ...J ~ _ _l_ 20 X Untreated equilibrium catalyst' + 1.4% coke pre-treated A2.5% coke pre-treated O 3.4% coke pre-treated

0...------"!'------"""""'0 20 40 60 80 100 Conversion (mass%)

A decrease in light catalytic naphtha yield as the amount of pre-coking is increased is evident in the above graph. The results of this thesis indicate that it can be quite hard to predict the behaviour of the light catalytic naphtha yield with regard to catalyst pre-treatment type and

245 concentration. This is determined to be due to the fact that light catalytic naphtha is the middle component in the range of yields. It can be seen that, in most of the results, the amount of light hydrocarbons increases, yet the amount of light cycle oil decreases with increasing pre-treatment levels. Both these components straddle the light catalytic naphtha. Therefore, depending on the severity of the decrease in the light cycle oil compared with the increase in the light hydrocarbons, the yield of light cycle oil may increase or decrease with type of catalyst pre treatment.

However, it can be seen for this particular set of results with the pre coked catalysts, that the shift of LCN is directionally the same as that from the catalyst with ammonia pre-treating by soaking. This indicates that the pre-coking of the catalyst to these levels is very severe, more so than that of the initial hydrocarbon pre-treatment experiments (0.3, 0.7 and 0.9 mass% coke on catalyst) and heat pre-treatment experiments.

The results of the above pre-coking experiments indicates that poisoning and fouling must be carefully distinguished. Deactivation by poisoning refers to the adsorption of components onto active sites, so that they are poisoned and thereby inactive for the desired reaction(s). Implicit in this mechanism is that the small amount of adsorbed poison, whilst it may be sufficient to give a dramatic loss in reactivity, only accumulates onto the surface to form a very thin layer, most likely an incomplete monolayer, which then adds negligibly to the solid volume of the catalyst. In contrast, deactivation by fouling involves the deposition of significant amounts of foulant on the catalyst surface. These deposits can occupy a significant portion of the pore volume and thereby can begin to distort the pore structure. It is envisaged that the pre-coking of the catalyst poisons as well as fouls the catalyst surface. This

246 conclusion is made in comparing the results with those of the ammonia pre-treatment, which is envisaged just to poison the active sites.

Confirmation of the hypothesis that the olefins remain trapped on the pre-coked catalysts undergoing oligomerisation to coke would be seen by an increase in the coke selectivity as level of pre-coking treatment increases. This is indeed seen in Figure 50, which shows coke selectivity as a function of pre-coking concentration.

Figure 50: Coke selectivity of squalane cracking as a function of pre coking of equilibrium catalysts

X Untreated equilibrium catalyst

25 ~----+-=----' • 1.4% coke pre-treated A.2.5% coke pre-treated 20 O 3.4% coke pre-treated -~0 Cl) Cl) m E 15 -Q.) ~ 0 () 10

0 0 20 40 60 80 100 Conversion (mass%)

The shape of the plots representing coke selectivity indicate that coke formation occurs as a secondary reaction. This is in contrast to Figures 44 and 45 representing C4 olefin formation and also the formation of light cycle oil. The linearity of those graphs indicate that they are formed as a primary reaction process.

247 In examining the light cycle oil selectivity, the overcracking point can once again be seen as it was for the metals pre-treatment. However, due to the range of conversions examined, the upward slope of the overcracking point was not as evident as previously seen. This can be seen in Figure 51 below.

Figure 51: Light cycle oil selectivity of squalane cracking as a function of pre-coking of equilibrium catalysts

16 ------

-0~ en en ro E 8 -0 (.) ....J 4 x Untreated equilibrium catalyst + 1.4% coke pre-treated 1:i,. 2.5% coke pre-treated o 3.4% coke pre-treated 0 ...... ______;;;;;;=;;..-~=-;;;;-====~ 0 20 40 60 80 100 Conversion (mass%)

It can be assumed that at 0% conversion, the yield of light cycle oil will also be 0%. Thus it can be predicted that for the pre-coked catalysts, their curve shape will also be horseshoed. An interesting point is the position of the curves for the pre-coked catalysts. As the amount of pre coking increased, the curve was shifted towards lower conversion. This can be explained by the earlier hypothesis that, as the feedstock is cracked on the pre-coked catalysts, the products remain on the catalyst

248 surface for a longer residence time than on the equilibrium catalyst, thereby further cracking into smaller hydrocarbons. Hence as conversion increases, the yield of light cycle oil will ultimately decrease.

10.7 SUMMARY

A significant finding of this section was the inability of the modified Mitchell method for metals impregnation to adequately represent an industrial operation. Despite this some valuable findings were made. As expected the nickel and vanadium pre-treated catalysts showed an increase in hydrogen yields. This stems from the fact that both nickel and vanadium are dehydrogenation catalysts, nickel more so than vanadium.

Interestingly, contrary to the findings for the earlier methods of catalyst pre-treatment, a decrease in light olefins yield was seen for the metals impregnated catalysts. This was explained to be the result of the metals holding the light olefins, which quickly polymerised to form coke, hence the dramatic increase in coke selectivity.

The sand / equilibrium split system testing revealed a very interesting finding, that the nickel pre-treatment of the catalyst does not appear to affect the equilibrium catalyst either physically or chemically, with all selectivity shifts seen as a sole result of the interaction of cracked products with the nickel itself. Although an unexpected finding, it is possible that the result is not representative of industrial operation due to the method of nickel impregnation used (modified Mitchell).

The results of the pre-coking of the catalyst showed that the coke blocked the most active sites on the catalyst, thereby significantly reducing activity and altering selectivity. Again the hydrogen yield was

249 observed to increase, presumably the result of a decrease in hydrogen transfer reactions and the increased formation of coke. The light olefins were observed to decrease, believed to be the result of a longer diffusional path due to coke and hence the reluctant desorption of the olefins, resulting in an increased propensity to form coke. The branched olefins on the pre-coked catalyst were observed to increase, also due to the increased residence time and an increased propensity to isomerise than undergo hydrogen transfer.

250 11 CONCLUSIONS

This thesis reports on a thorough study into the deactivation of an industrial fluid catalytic cracking catalyst, using various deactivation methods and several different feedstocks with the aim of increasing the profitability of commercial operations within Australia.

One of the main aims of the study was to investigate the pre-treatment of the catalyst, by way of a light hydrocarbon, with the intention to produce more light cycle oil (a high valued product). The study revealed the opposite to be true with a decrease in light cycle oil selectivity.

With the ensuing investigation it became clear that this was a result of a decrease in catalyst acid site strength and site density. As a result the catalyst had an increased propensity to form lighter hydrocarbons by cracking small side chains from the large branched feedstock molecules.

All methods of catalyst pre-treatment, which involved hydrocarbon, heat, ammonia, nickel, vanadium and coke were shown to result in decreased conversion of feedstocks, increased formation of hydrogen and dry gas and a reduction in light cycle oil selectivity. In general most methods of pre-treatment also resulted in increased light olefin yields and increases in coke selectivity. The prediction of the shift in light catalytic naphtha was quite tricky and has been found to be dependent on severity of deactivation with the most severe deactivation showing a decrease in light catalytic naphtha.

Several reaction pathways were found to be affected by the pre treatment, but not all equally. Naturally the actual cracking reaction pathway was affected as seen by the decrease in conversion, but other

251 reaction pathways such as hydrogen transfer, isomerisation and polymerisation were also affected. Site density went a long way to answering why these reaction pathways were affected, with a decrease in site density resulting in decreases in bimolecular reactions (cracking and hydrogen transfer).

Site density was found to be crucial for hydrogen transfer and polymerisation reactions to occur. A decrease in the site density means that these reactions are more unlikely to occur and therefore products arising from these reactions are less likely to be formed, in particular paraffins and light cycle oil.

Throughout the investigation several interesting points were made:

1. Lewis sites do not appreciably participate in the initiation or propagation of catalytic cracking 2. The presence of olefins in the feedstock dramatically exaggerates any effects seen with straight or branched paraffins 3. Bronsted sites are the major contributing sites for catalytic cracking 4. Pre-treatment of the catalyst deactivates the major contributing sites for catalytic cracking and as a result the deactivation of the catalyst is not linear with only a small amount of pre-treatment required to significantly reduce the effectiveness of the catalyst

In the pre-treatment of the catalyst with metals, it was unexpectedly discovered that the selectivity shifts of nickel impregnated catalyst was not due to any physical or chemical effects that the metal had on the catalyst, but rather due to the interaction of cracked products with the nickel itself. The results of this unforeseen finding may be slightly tainted by the observation that the method used to impregnate the

252 metals on the catalyst (modified Mitchell) is not an accurate representation of an industrial operation.

In relating the findings of this study with commercial operation in Australia, it is strongly recommended that efforts be made to ensure the integrity of the fluid cracking catalyst during operation and not allow the contamination of the catalyst with coke and metals. This can be accomplished with a frequent monitoring program and proper crude selection. Further that the injection of a light hydrocarbon such as light catalytic naphtha into the bottom of the riser will result in decreased yields of light cycle oil, combined with increases in dry gas (including hydrogen), light olefins and coke.

In recommending future work, further investigation should be conducted with pure hydrocarbons on nickel and vanadium impregnated catalysts. The impregnation should be conducted using a cyclic deactivation technique, such that the doped catalysts will more closely represent industrially aged catalysts.

Further it would be most beneficial to study the effect that iron as a contaminant has on fluid catalytic cracking catalysts. This stems from the fairly destructive behaviour that has been observed to occur in the commercial operation of catalysts containing more than 1 mass% iron.

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