Catalytic Hydroconversion of Diphenylmethane with Unsupported Mos2

Catalytic Hydroconversion of Diphenylmethane with Unsupported MoS2

Ross S. Kukard

B.Sc. Eng. Chem. (hons), The University of Cape Town, 2006 M.Sc. Eng. Chem., The University of Cape Town, 2009

A THESIS SUBMITTED IN PARTIAL FULFILLMENT OFTHEREQUIREMENTSFORTHEDEGREEOF

Doctor of Philosophy

THE FACULTY OF GRADUATE AND POSTDOCTORAL STUDIES (Chemical and Biological Engineering)

The University of British Columbia (Vancouver)

September 2014

The mechanism by which hydroconversion catalysts promote residue conversion and coke suppression is unclear. Several theories are proposed in the literature but these have all been opposed, usually due to their lack of controlled mechanistic studies. A promising catalyst for residue hydroconversion is unsupported MoS2. This catalyst is effective but expensive and deactivates during the reaction. Model compound studies were needed to elucidate the mechanism of MoS2 catalysis in hydroconversion reactions, how this relates to residue hydroconversion and hence propose deactivation mechanisms and regeneration methodologies. Model compound screening in a commercially available stirred slurry-phase batch reactor identiﬁed diphenylmethane (DPM) as a suitable model reagent. Ex- periments were conducted at industrially applicable conditions of 445◦C, 13.8 MPa

H2 and catalyst loadings of 0 - 1800 ppm Mo (introduced as Mo octoate which formed the MoS2 active phase in-situ). Slow heat-up rates and wall catalysis, however, made this reactor unsuitable for detailed mechanistic studies. A novel mixed slurry-phase micro-reactor system was designed using externally applied vortex mixing and removable glass-inserts to allow for greater analytical resolution and determination of the thermocatalytic mechanism. Deactivated MoS2 catalysts, as coke-catalyst agglomerates recovered from residue hydroconversion studies [1], were evaluated using the DPM testing methodology and a deactivation mechanism proposed.

It was determined that the unsupported MoS2 crystallites hydrogenate the DPM feed to cyclohexylmethylbenzene (CHMB) which undergoes thermolysis to short chain hydrocarbon radicals. These short chain radicals stabilise, by radical addition or radical disproportionation, other radicals in the system by a chain stabili-

ii sation reaction, itself promoted by catalytic hydrogenation (for instance of oleﬁns formed during disproportionation). Deactivation of unsupported MoS2 in residue hydroconversion was proposed to be due to the formation of an unreactive, porous carbonaceous structure upon which the otherwise unaltered catalyst particles become supported. The pores physically exclude larger species, such as asphaltenes, from reaching the active sites. Inter-recycle solvent extraction to remove coke precursors was proposed to inhibit deactivation in residue hydroconversion whilst mechanical and chemical size reduction were suggested for breaking the porous structure and re-exposing the

MoS2 crystallites.

iii Preface

All of the work presented henceforth was conducted in the Department of Chemi- cal and Biological Engineering at the University of British Columbia, Point Grey campus. I, Ross S. Kukard, was the lead investigator of this work, responsible for all major areas of concept formation, micro-reactor design, construction and commissioning and all experimentation, data collection and analysis as well as preparation of this thesis. Kevin J. Smith was the supervisor of this research, involved throughout the project in concept formation and thesis edits. Hooman Rezaei commis- sioned the stirred batch reactor described in Section B.2.1 which I used to conduct the experiments and collect the data in Section 4.1. He also conducted the residue hydroconversion experiments to generate deactivated catalyst samples and comparison data (Section 4.2.3). The mechanical workshop in the Department of Chemi- cal and Biological Engineering, led by Doug Yuen, constructed the micro-reactor enclosures described in Section B.3.1 using designs that I prepared. I presented data from Section 4.1 at the 62nd Canadian Chemical Engineer- ing Conference (Vancouver, BC, Canada), 2012 and data from Section 4.2 at the 23rd Canadian Symposium on Catalysis (Edmonton, AB, Canada), 2014. Kevin J. Smith was involved in the preparation of these presentations.

iv Table of Contents

Abstract...... ii

Preface ...... iv

TableofContents ...... v

ListofTables...... xi

List of Figures ...... xiii

Nomenclature ...... xx

Acknowledgements ...... xxix

Dedication ...... xxx

1 Introduction ...... 1

2 LiteratureReview ...... 5 2.1 Current Hydroconversion Technology ...... 5 2.2 Residue Processing Technologies ...... 6 2.2.1 Carbon Rejection ...... 7 2.2.2 Hydroconversion ...... 8 2.2.3 Slurry-Phase Catalytic Hydroconversion ...... 10 2.3 Catalyst Testing ...... 12 2.3.1 Heavy Oil and Residue Oil Studies ...... 12

v 2.3.2 Model Compound Studies ...... 13 2.4 Micro-Reactors for Catalyst Testing ...... 16 2.4.1 Advantages and Disadvantages of Micro-Reactors . . . . 16 2.4.2 Micro-Reactors in Hydroconversion Studies ...... 17 2.5 Catalyst Activity and Deactivation ...... 18 2.5.1 Catalyst Selection ...... 19 2.5.2 Molybdenum Disulphide ...... 20 2.5.3 Processes of Deactivation ...... 26 2.5.4 Catalyst Regeneration Methodologies ...... 28 2.6 Summary of Findings from the Literature ...... 29

3 Experimental...... 30 3.1 Experimental Objectives and Programme ...... 30 3.1.1 Interpretation of Questions to Objectives ...... 31 3.1.2 Experimental Programme ...... 35 3.2 Experimental Apparatus and Supplies ...... 39 3.2.1 Reaction and Analytical Supplies ...... 39 3.2.2 Reactors and Conditions ...... 40 3.3 Analytical Equipment and Data Analysis ...... 53 3.3.1 Gas Product Analysis ...... 53 3.3.2 Liquid Product Analysis ...... 54 3.3.3 Solid Product Analysis ...... 54

4 ExperimentalResults ...... 55 4.1 Stirred Batch Reactor ...... 55 4.1.1 Model Compound Screening ...... 55 4.1.2 Benzene, Toluene and Decalin Blanks ...... 59 4.1.3 Diphenylmethane Studies ...... 60 4.2 BatchMicro-reactor...... 83 4.2.1 Inclined Stainless Steel Micro-Reactor ...... 83 4.2.2 Vertical Stainless Steel Micro-Reactor ...... 88 4.2.3 Glass Insert Micro-Reactor ...... 101

vi 5 DiscussionofExperimentalResults ...... 140 5.1 Model Compound Evaluation ...... 140 5.1.1 Model Compound Screening ...... 140 5.1.2 Diphenylmethane Studies ...... 142 5.1.3 Summary of Model Compound Evaluation ...... 155 5.2 Novel Reactor System Design and Testing ...... 157 5.2.1 Inclined Stainless Steel Micro-Reactor ...... 158 5.2.2 Vertical Stainless Steel Micro-Reactor ...... 162 5.2.3 Unmixed Glass Insert Micro-Reactor ...... 169 5.2.4 Mixed Glass Insert Micro-Reactor ...... 171 5.2.5 Summary of Micro-Reactor System Design and Testing . 183 5.3 Catalyst Study and Deactivation Investigation ...... 185

5.3.1 Active MoS2 ...... 185

5.3.2 Deactivated Coke-MoS2 Agglomerate ...... 187

5.3.3 Mechanism of MoS2 Deactivation in Residue Hydrocon- version ...... 192

5.3.4 Summary of MoS2 Activity and Deactivation ...... 195

6 Conclusions...... 197

References ...... 201

Appendices ...... 218

A Catalyst Deactivation and Regeneration ...... 219 A.1 Processes of Catalyst Deactivation ...... 219 A.1.1 Fouling ...... 219 A.1.2 Poisoning...... 222 A.1.3 Others...... 227 A.2 Catalyst Regeneration Processes ...... 230

B Experimental Apparatus and Procedures ...... 233 B.1 Detailed Experimental Programme ...... 233 B.2 Batch Reactor Speciﬁcations and Operation ...... 234

vii B.2.1 Description and Speciﬁcations ...... 241 B.2.2 Operating Procedure ...... 244 B.2.3 Safety Considerations ...... 255 B.3 Micro-Reactor Design, Development and Operation ...... 257 B.3.1 Design and Development ...... 257 B.3.2 Operating Procedure ...... 276 B.3.3 Safety Considerations ...... 287

C Analytical Apparatus, Procedures and Data Analysis ...... 289 C.1 GasProductAnalysis ...... 289 C.1.1 Analytical Equipment and Procedures ...... 289 C.1.2 Calibration, Data Acquisition, Analysis and Interpretation 292 C.2 Liquid Product Analysis ...... 296 C.2.1 Analytical Equipment and Procedures ...... 296 C.2.2 Data Acquisition, Analysis and Interpretation ...... 301 C.2.3 Calibration, Analysis and Experimental Uncertainty . . . 307 C.3 Solid Product Analysis ...... 318 C.3.1 Analytical Equipment and Procedures ...... 318 C.3.2 Calibration, Data Acquisition, Analysis and Interpretation 320

D CalibrationData...... 333 D.1 GasChromatograph...... 333 D.1.1 HP5980A Calibration Results ...... 333 D.1.2 Shimadzu GC-14B Calibration Results ...... 333 D.2 GCMS-QP2010 Gas Chromatograph - Mass Spectrometer Liquid Calibration Results ...... 335 D.2.1 Benzene...... 335 D.2.2 Toluene...... 337 D.2.3 Diphenylmethane ...... 339 D.2.4 Diphenylethane ...... 341 D.2.5 Diphenylpropane ...... 343

viii E DetailedExperimentalResults ...... 345 E.1 Stirred Batch Reactor ...... 345 E.1.1 Model Compound Screening ...... 345 E.1.2 Diluted Diphenylmethane ...... 346 E.1.3 Undiluted Diphenylmethane ...... 347 E.2 Stainless Steel Batch Micro-reactors ...... 356 E.2.1 Inclined Stainless Steel Micro-Reactor ...... 356 E.2.2 Vertical Stainless Steel Micro-Reactor ...... 356 E.3 Glass Insert Batch Micro-reactor ...... 356 E.3.1 Comparison with Stainless Steel Micro-Reactor ...... 356 E.3.2 Visual Mixing Studies ...... 356 E.3.3 Comparison of Liquid Loading Volumes ...... 356 E.3.4 Thermocouple Wall Activity ...... 356 E.3.5 EffectofMixingSpeed...... 375 E.3.6 Optimum Mixing Speed Evaluation ...... 375 E.3.7 Spent Residue Hydroconversion Catalyst Evaluation . . . 376

F DataProcessingandAnalysis...... 385 F.1 Data Acquisition and Analysis ...... 385 F.2 Gas and Liquid Mass Balances ...... 397 F.2.1 LiquidProduct ...... 397 F.2.2 GasProduct...... 398 F.3 Kinetic Analyses ...... 402 F.3.1 FirstOrder ...... 402 F.3.2 SecondOrder...... 403 F.3.3 TheArrheniusLaw...... 404 F.4 Thermodynamic Simulations ...... 405 F.5 Phase Density and Composition Simulations ...... 407 F.6 Physical Property Simulations ...... 408 F.7 Hydrogen Solubility and Diffusivity ...... 408 F.7.1 Hydrogen Solubility Simulations ...... 408 F.7.2 Hydrogen Dissolution Rate ...... 426 F.7.3 Hydrogen Diffusion ...... 427

ix F.8 Area:Volume Ratios ...... 431 F.8.1 Gas-Liquid Interfacial Area ...... 435 F.9 Coke Solubility ...... 436

x List of Tables

Table 2.1 Comparison of physical properties of various crude oils and residua...... 6 Table 2.2 Comparison of ﬂuidised bed and slurry-phase hydroconversion reactorperformance...... 11

Table 3.1 Summary of experimental programme...... 36 Table 3.2 Stirred batch reactor operating conditions ...... 43 Table 3.3 Micro-reactor operating conditions ...... 49

Table 4.1 Conversion results for model compound screening ...... 56 Table 4.2 Major products for model compound hydroconversion . . . . . 58 Table 4.3 Gaseous products for diphenylmethane hydroconversion . . . . 59 Table 4.4 Benzene and toluene blank test conversions ...... 60 Table 4.5 Major products for benzene blank tests ...... 60 Table 4.6 Major products for toluene blank tests ...... 61 Table 4.7 Major products for decalin blank test ...... 61 Table 4.8 Major products for undiluted diphenylmethane hydroconversion 69 Table 4.9 Structures of hydroconversion side-products ...... 70 Table 4.10 Comparison of heating rate effect on diphenylmethane hydroconversion ...... 86 Table 4.11 Major products for diphenylmethane hydroconversion in inclined micro-reactor...... 92 Table 4.12 Coefﬁcients for diphenylmethane hydroconversion kinetic models in vertical micro-reactor ...... 94

xi Table 4.13 Major gaseous products for diphenylmethane hydroconversion in vertical micro-reactor ...... 96 Table 4.14 Coefﬁcients for diphenylmethane hydroconversion kinetic models in glass insert micro-reactor ...... 101 Table 4.15 Comparison of liquid volume effect on diphenylmethane hydroconversion ...... 111 Table 4.16 Scanning electron microscopy with energy dispersive X-ray quan- tiﬁcation of solids from diphenylmethane hydroconversion after 4h ...... 119 Table 4.17 Major constituents of isom./cond. lump for diphenylmethane hydroconversion in glass insert micro-reactor at 150 µL . . . . 134 Table 4.18 Major gaseous products for diphenylmethane hydroconversion in glass insert micro-reactor ...... 135 Table 4.19 Results from coke-catalyst agglomerate evaluation by diphenylmethane hydroconversion ...... 138 Table 4.20 Major gaseous products for coke-catalyst agglomerate evaluation by diphenylmethane hydroconversion ...... 139

Table 5.1 Physical property simulations for diphenylmethane hydroconversion at various conditions ...... 177 Table 5.2 Solubility of coke-catalyst agglomerates in diphenylmethane . 192

xii List of Figures

Figure 2.1 Thermal cracking of diphenylmethane ...... 16

Figure 2.2 Rendering of arbitrary 5-layer stack of MoS2 ...... 21 Figure 2.3 Thermal decomposition of diphenylmethane ...... 23 Figure 2.4 Thermocatalytic decomposition of diphenylmethane . . . . . 25 Figure 2.5 Surface reactions for hydrogen and diphenylmethane . . . . . 26 Figure 2.6 Summary of main thermocatalytic decomposition mechanisms of diphenylmethane ...... 27

Figure 3.1 Thermal cracking of DPM, DPE and DPP ...... 33 Figure 3.2 Decahydronaphthalene structure...... 33 Figure 3.3 Reactor heating profile comparison ...... 34 Figure 3.4 Process flow diagram of stirred batch reactor ...... 44 Figure 3.5 Laboratory implementation of stirred batch reactor ...... 45 Figure 3.6 Process flow diagram of micro-reactor ...... 50 Figure 3.7 Laboratory implementation of micro-reactor ...... 51

Figure 4.1 Conversion results for diphenylpropane hydroconversion . . . 57 Figure 4.2 Conversion results for diphenylmethane hydroconversion with sigmoidaltrends ...... 62 Figure 4.3 Benzene molar yield for diphenylmethane hydroconversion . . 64 Figure 4.4 Toluene molar yield for diphenylmethane hydroconversion . . 64 Figure 4.5 Benzene:toluene molar ratio for diphenylmethane hydroconversion ...... 65

xiii Figure 4.6 Conversion results for undiluted diphenylmethane hydroconversion with sigmoidal trends ...... 67 Figure 4.7 Conversion with reaction temperature results for undiluted diphenylmethane hydroconversion ...... 67 Figure 4.8 Logarithmic conversion with inverse reaction temperature for undiluted diphenylmethane hydroconversion ...... 68 Figure 4.9 Benzene molar yield for undiluted diphenylmethane hydroconversion ...... 72 Figure 4.10 Toluene molar yield for undiluted diphenylmethane hydroconversion ...... 72 Figure 4.11 Cyclohexylmethylbenzene molar yield for undiluted diphenylmethane hydroconversion ...... 73 Figure 4.12 Mass yield of other cracking products for undiluted diphenylmethane hydroconversion ...... 73 Figure 4.13 Mass yield of isomerisation and condensation products for undiluted diphenylmethane hydroconversion ...... 74 Figure 4.14 Benzene:toluene molar ratio for undiluted diphenylmethane hydroconversion...... 74 Figure 4.15 Benzene:toluene molar ratio against catalyst loading for undiluted diphenylmethane hydroconversion ...... 75 Figure 4.16 Pressure change for undiluted diphenylmethane hydroconversion 76 Figure 4.17 X-ray diffractogram for 600 ppm Mo in undiluted diphenylmethane hydroconversion ...... 78 Figure 4.18 X-ray diffractogram for 1800 ppm Mo in undiluted diphenylmethane hydroconversion ...... 79 Figure 4.19 Transmission electron microscopy image for 600 ppm Mo in undiluted diphenylmethane hydroconversion ...... 80 Figure 4.20 Transmission electron microscopy image for 1800 ppm Mo in undiluted diphenylmethane hydroconversion ...... 81

Figure 4.21 MoS2 sheet size distribution for 600 ppm Mo in undiluted diphenylmethane hydroconversion ...... 82

Figure 4.22 MoS2 stack height distribution for 600 ppm Mo in undiluted diphenylmethane hydroconversion ...... 82

xiv Figure 4.23 Conversion results to study wall activation during diphenylmethane hydroconversion ...... 84 Figure 4.24 Conversion results comparing wall activity during diphenylmethane hydroconversion ...... 85 Figure 4.25 Conversion results for diphenylmethane hydroconversion in inclined micro-reactor ...... 87 Figure 4.26 Benzene molar yield for diphenylmethane hydroconversion in inclined micro-reactor ...... 88 Figure 4.27 Toluene molar yield for diphenylmethane hydroconversion in inclined micro-reactor ...... 89 Figure 4.28 Cyclohexylmethylbenzene molar yield for diphenylmethane hydroconversion in inclined micro-reactor ...... 89 Figure 4.29 Mass yield of other cracking products for diphenylmethane hydroconversion in inclined micro-reactor ...... 90 Figure 4.30 Mass yield of isomerisation and condensation products for diphenylmethane hydroconversion in inclined micro-reactor . . 90 Figure 4.31 Benzene:toluene molar ratio for diphenylmethane hydroconversion in inclined micro-reactor ...... 91 Figure 4.32 Conversion results for diphenylmethane hydroconversion in vertical micro-reactor ...... 93 Figure 4.33 Benzene molar yield for diphenylmethane hydroconversion in vertical micro-reactor ...... 95 Figure 4.34 Toluene molar yield for diphenylmethane hydroconversion in vertical micro-reactor ...... 96 Figure 4.35 Cyclohexylmethylbenzene molar yield for diphenylmethane hydroconversion in vertical micro-reactor ...... 97 Figure 4.36 Fluorene molar yield for diphenylmethane hydroconversion in vertical micro-reactor ...... 97 Figure 4.37 Benzene:toluene molar ratio for diphenylmethane hydroconversion in vertical micro-reactor ...... 98 Figure 4.38 Gas chromatograms for diphenylmethane hydroconversion in vertical micro-reactor with 0 ppm Mo ...... 99

xv Figure 4.39 Gas chromatograms for diphenylmethane hydroconversion in vertical micro-reactor with 1800 ppm Mo ...... 100 Figure 4.40 Comparison of conversion results for diphenylmethane hydroconversion in stainless steel and glass insert micro-reactors . . 102 Figure 4.41 Benzene molar yield for diphenylmethane hydroconversion in glass insert micro-reactor ...... 103 Figure 4.42 Toluene molar yield for diphenylmethane hydroconversion in glass insert micro-reactor ...... 104 Figure 4.43 Cyclohexylmethylbenzene molar yield for diphenylmethane hydroconversion in glass insert micro-reactor ...... 104 Figure 4.44 Mass yield of isomerisation and condensation products for diphenylmethane hydroconversion in glass insert micro-reactor . 105 Figure 4.45 Benzene:toluene molar ratio for diphenylmethane hydroconversion in glass insert micro-reactor ...... 105 Figure 4.46 Mixing of 400 µL diphenylmethane hydroconversion reaction product in glass mock-up ...... 106 Figure 4.47 Effect of mixing speed on “vortex” height in glass mock-up for diphenylmethane hydroconversion reaction product ...... 107 Figure 4.48 Mixing of 150 µL diphenylmethane hydroconversion reaction product in glass mock-up ...... 108 Figure 4.49 Mixing of 400 µL diphenylmethane hydroconversion reaction product in glass mock-up with thermocouple ...... 109 Figure 4.50 Mixing of 150 µL diphenylmethane hydroconversion reaction product in glass mock-up with thermocouple ...... 110 Figure 4.51 Conversion results to study thermocouple wall activation during diphenylmethane hydroconversion ...... 112 Figure 4.52 Benzene and toluene molar yields for thermocouple wall activation during diphenylmethane hydroconversion ...... 113 Figure 4.53 Benzene:toluene molar ratio for thermocouple wall activation during diphenylmethane hydroconversion ...... 113 Figure 4.54 Cyclohexylmethylbenzene molar yield for thermocouple wall activation during diphenylmethane hydroconversion ...... 114

xvi Figure 4.55 Mass yield of isomerisation and condensation products for thermocouple wall activation during diphenylmethane hydroconversion ...... 114 Figure 4.56 Inﬂuence of mixing speed on conversion results for diphenylmethane hydroconversion in glass insert micro-reactors . . . . 116 Figure 4.57 Benzene molar yield for diphenylmethane hydroconversion in glass insert micro-reactor ...... 116 Figure 4.58 Toluene molar yield for diphenylmethane hydroconversion in glass insert micro-reactor ...... 117 Figure 4.59 Benzene:toluene molar ratio for diphenylmethane hydroconversion in glass insert micro-reactor ...... 117 Figure 4.60 Cyclohexylmethylbenzene molar yield for diphenylmethane hydroconversion in glass insert micro-reactor ...... 118 Figure 4.61 Transmission electron microscopy of solids from diphenylmethane hydroconversion at 0 RPM ...... 121 Figure 4.62 Transmission electron microscopy of solids from diphenylmethane hydroconversion at 2000 RPM ...... 122 Figure 4.63 Scanning electron microscopy with energy dispersive X-ray image of solids from diphenylmethane hydroconversion after 4 h123 Figure 4.64 Field emission scanning electron microscopy images of solids from diphenylmethane hydroconversion at 0 RPM ...... 124 Figure 4.65 Field emission scanning electron microscopy images of solids from diphenylmethane hydroconversion at 2000 RPM . . . . 125 Figure 4.66 Field emission scanning electron microscopy images of solids from diphenylmethane hydroconversion at 2250 RPM with usual heat-upmixing ...... 126 Figure 4.67 Field emission scanning electron microscopy images of solids from diphenylmethane hydroconversion at 2250 RPM without heat-upmixing ...... 127 Figure 4.68 Angled ﬁeld emission scanning electron microscopy images of solids from diphenylmethane hydroconversion at 2250 RPM . 128 Figure 4.69 Conversion results for diphenylmethane hydroconversion in glass insert micro-reactor at 2000 RPM ...... 129

xvii Figure 4.70 Benzene molar yield for diphenylmethane hydroconversion in glass insert micro-reactor at 150 µL...... 130 Figure 4.71 Toluene molar yield for diphenylmethane hydroconversion in glass insert micro-reactor at 150 µL...... 131 Figure 4.72 Benzene:toluene molar ratio for diphenylmethane hydroconversion in glass insert micro-reactor at 150 µL ...... 131 Figure 4.73 Cyclohexylmethylbenzene molar yield for diphenylmethane hydroconversion in glass insert micro-reactor at 150 µL . . . . . 132 Figure 4.74 Mass yield of isomerisation and condensation products for diphenylmethane hydroconversion in glass insert micro-reactor at 150 µL...... 133

Figure 5.1 Thermocatalytic decomposition products of benzene and toluene144 Figure 5.2 Thermodynamic simulations of benzene and toluene decomposition to methane ...... 144 Figure 5.3 Thermodynamic simulations of phenyl and benzyl radical decomposition...... 147 Figure 5.4 Phase simulations of reaction mixture ...... 153 Figure 5.5 Formation of ﬂuorene and hexahydroﬂuorene from diphenylmethane...... 154 Figure 5.6 Proposed thermocatalytic decomposition mechanism of diphenylmethane in stirred batch reactor ...... 156 Figure 5.7 Comparison of thermolysis mechanisms of cyclohexylmethylbenzene...... 160 Figure 5.8 Proposed thermocatalytic decomposition mechanism of diphenylmethane in inclined micro-reactor ...... 163

Figure 5.9 Pressure drop to study H2 dissolution in diphenylmethane . . 165

Figure 5.10 Modeled concentration proﬁles to study H2 diffusion through diphenylmethane ...... 166 Figure 5.11 Comparison of H* abstraction mechanisms from diphenylmethane ...... 167 Figure 5.12 Comparison of phenyl radical stabilisation mechanisms . . . . 169

xviii Figure 5.13 MoS2 crystallite formation and movement in glass insert micro- reactor...... 182 Figure 5.14 Proposed thermocatalytic decomposition mechanism of diphenylmethane...... 188 Figure 5.15 Asphaltene model compound bibenzyl-cholestane ...... 190 Figure 5.16 Proposed mechanism for unsupported catalyst deactivation in residue hydroconversion ...... 195

xix Nomenclature

Roman Symbols

∆ T Gr Gibbs free energy of reaction at temperature, T, in kJ/mol ∆ T Hr Enthalpy of reaction at temperature, T, in Ha, kcal/mol or kJ/mol

A : V Ratio of reactor and internals surface area to the volume of liquid reaction mixture, cm−1 aext Linear ﬁt parameters for response factor calibrations, dimensionless aint Linear ﬁt parameters for internal standards, dimensionless

Agc Peak area from gas chromatography analysis, counts.time b Peak broadening factor for the Scherrer equation using the FWHM method, radians bint Linear ﬁt parameters for internal standards, dimensionless

BT Benzene:toluene molar ratio, dimensionless

C Mass concentration, wt% c Mass composition, fractional

C′ Molar concentration, mol%

Convx.y Conversion factor from x to y units, (kcal/mol)/Ha or kJ/kcal

xx D Mass diffusivity in calculation of the Sherwood number (see Section 5.2.4), m2/s

DH Vessel diameter in calculation of the Reynolds number (see Section 5.2.4), m

Ex.Total Total energy of species x from Accelrys Materials Studio simulations, Ha

Ea Activation energy, J/mol

F Spherical packing factor, dimensionless h Height, m

Hx Enthalpy of species x, Ha, kcal/mol or kJ/mol

Hcx Enthalpy correction to total energy for species x from Accelrys Mate- rials Studio simulations, Ha

ID Inner diameter

IQ Interquartile range

K Scherrer constant, 0.76 used for MoS2, dimensionless k Reaction rate coefﬁcient, h−1

′ −1 −1 k Reaction rate coefficient modified to account for catalyst loading, h .wt%cat k′′ Reaction rate coefficient for a constant time experiment, h−1

KSh Mass transfer coefﬁcient in calculation of the Sherwood number (see Section 5.2.4), m/s

L Length, m

LSh Characteristic length in calculation of the Sherwood number (see Section 5.2.4), m m Mass, g

xxi Mr Molar mass, g/mol n Number of moles, mol

OD Outer diameter

P Pressure, Pa p Crystallite size from the Scherrer equation, nm

Q Molar ratio, dimensionless

Q′ Volumetric ratio, dimensionless

Qx Q1 and Q3 being the ﬁrst and third quartiles respectively

R Ideal gas constant, 8.3144621 J/mol.K

′ r Reaction rate on a mass basis, mol/h.wt%cat

Re Reynolds number (see Section 5.2.4), dimensionless

Rgc Response factor for gas chromatography analysis, counts s Standard deviation, dimensionless

Sc Schmidt number (see Section 5.2.4), dimensionless

Sh Sherwood number (see Section 5.2.4), dimensionless

SS SSresid is the sum of squares of the residuals and SStotal is the total sum of squares (the difference between the experimental and the average)

T Temperature, K t Time, s

V Volume, m3 v Mean ﬂuid velocity in calculation of the Reynolds number (see Section 5.2.4), m/s

xxii 3 VBP.i Molar volume of species i at its normal boiling point, m /mol

X Mass conversion, fractional

X′ Mass conversion, wt%

Y Molar yield, molproduct formed/molreagent consumed

′ Y Mass yield, gproduct formed/greagent consumed

Greek Symbols

β Integral breadth for use in the Scherrer equation (see Section C.3.2), radians

η Wilke-Chang parameter (see Section F.7)

λ Wavelength in X-ray diffraction analysis, nm

µ Dynamic viscosity, kg/m.s

ν Kinematic viscosity, deﬁned as µ/ρ, m2/s

φ Mass concentration ratio, dimensionless

φ ′ GCMS area ratio, dimensionless

φ ′′ Volumetric ratio, dimensionless

φ ∗ Solvent association parameter for Wilke-Chang correlation (see Section F.7)

ρ Mass density, kg/m3

θ Bragg angle of basal peak for use in the Scherrer equation (see Section C.3.2), radians

Subscripts

.in At the beginning

.out At the end

xxiii added Mixed into solution when diluting samples for analysis atm Atmosphere

Bal Balance

Bead Relating to the glass beads used as spacers in the glass insert micro- reactor

Benz Benzene cal From or relating to calibration, e.g. φcal is the calibrated φ value cat Catalyst

Dec Decalin dil Desired dilution for loading

DPE Diphenylethane

DPM Diphenylmethane est Estimated (used for estimating DPM conversion for sample dilution) ext Not using and internal standard, only a response factor

ﬂow Flow rate

G.L Gas to liquid (ratio)

GLTotal Total gas and liquid

GTotal Total gas

H2 Hydrogen i One in a series of samples insert Relating to the glass insert int Internal standard

xxiv LTotal Total liquid

Meth Methane

Mo Molebdynum model Model compound mol.CS2.Mo Moles of CS2 with respect to Mo

O2 Oxygen ppm Desired loading pre Precursor concentration

Prod Total product

Prop Propane pur Purity of stock species resid Residue

RTotal Total reactor rxn Reaction samp From sample shell Relating to the stainless steel shell of the glass insert micro-reactor

Targ Target species (the species of interest when performing calculations)

Tol Toluene

Deﬁnitions

Agilent Agilent Technologies Inc.

Brooks Brooks Instrument

xxv Coke Solid deposits, both carbonaceous and metallic (metal crystallites for example) in nature formed during residue hydroconversion reactions which are, by deﬁnition insoluble in toluene at 20◦C and 1 atm

Conversion The percentage of model compound consumed during the reaction with respect to that initially loaded into the reaction (see Section C.2.3)

Inner fence The inner interquartile boundaries, deﬁned as Q1 − 1.5 × IQ < x < Q3 + 1.5 × IQ, used for outlier identiﬁcation (see Section C.2.3)

OMEGA OMEGA Engineering Inc.

Outer fence The outer interquartile boundaries, deﬁned as Q1−3×IQ < x < Q3+ 3 × IQ, used for outlier identiﬁcation (see Section C.2.3)

Parr Parr Instrument Company

Residue Speciﬁcally vacuum residue, fraction of oil-derived hydrocarbons boiling above 525◦C

Shimadzu Shimadzu Scientiﬁc Instruments Inc., a division of Shimadzu Corpo- ration

Trace Species in the GCMS chromatogram representing less than 0.25 area% on a reagent- and diluent-free basis

Yield The ratio, molar-based or mass-based, of a particular product or group of products with respect to the model compound consumed during the reaction (see Section C.2.3)

Acronyms

A:V Ratio of reactor and internals surface area to the volume of liquid reaction mixture

API American Petroleum Institute gravity is a comparison of the density of 141.5 an oil with respect to water, being deﬁned as: ◦API = − SGat15.6◦C 131.5

xxvi B:T Benzene:Toluene molar ratio

BBCh Bibenzyl-cholestane. An asphaltene model compound proposed by Alshareef et al. [2] and discussed in Section 5.3.2.

BBP 4-Benzylbiphenyl, see Table 4.9

BPR Back-Pressure Regulator

CHMB Cyclohexylmethylbenzene, see Table 4.9

DNM Di(1-naphthyl)methane

DPE Diphenylethane

DPM Diphenylmethane

DPP Diphenylpropane

EPB 1-Ethyl-2-(1-phenylethyl)-benzene, see Table 4.9

ETB 1,1’,1”-(1-Ethanyl-2-ylidene)tris-benzene, see Table 4.9

FID Flame Ionisation Detector

FWHM Full width at half maximum height. A technique used to determine the peak broadening factor for use in the Scherrer equation (see Section C.3.2)

GC Gas Chromatograph

GCMS Gas Chromatography - Mass Spectroscopy

HexF 1,2,3,4,4a,9a-Hexahydrofluorene, see Table 4.9

HP Hewlett Packard

IUPAC International Union of Pure and Applied Chemists

LEL Lower Explosive Limit

LHHW Langmuir-Hinshelwood-Hougen-Watson kinetics

xxvii MBP One of several methyl-substituted biphenyl species as shown in Table 4.9

MFC Mass Flow Controller

MFM Mass Flow Meter

MS Mass Spectroscopy

MTP 2-Methyl-1,1,1-triphenyl-propane, see Table 4.9

STP Standard Temperature and Pressure, deﬁned as 273.15 K 100 kPa

TC Thermocouple

TCD Thermal Conductivity Detector

XRD X-ray Diffraction

xxviii Acknowledgements

The author would like to gratefully acknowledge the efforts and support of all of the individuals and organisations involved in the development of this research endeavour, speciﬁcally:

Prof. Kevin J. Smith for the opportunity to conduct this research, support in determining the direction and focus for this endeavour, and guidance in the development of this report.

Prof. Naoko Ellis, Prof. Marek Pawlik and Prof. Paul Watkinson for their role as supervisory committee and the invaluable information, assistance and direction they have provided.

Hooman Rezaei for assistance operating the stirred batch reactor, assistance with analyses and preparation of spent coke-catalyst samples.

Sharhzad Jooya Ardakani for assistance with experimental analyses.

Victoria Whiffen, Shahin Goodarznia, Farnaz Sotoodeh and Mina Alyani for general assistance in the lab.

Richard Ryoo, Doug Yuen and all of the other “behind-the-scenes” administra- tive staff without whose tireless efforts research could not be conducted.

The author would also like to thank UOP LLC, Alberta Energy Research Insti- tute (AERI) and National Science and Engineering Research Council (NSERC) of Canada for ﬁnancial support for this project.

xxix Dedication

This thesis is dedicated to all those who helped me reach this point in my life: family, friends and colleagues. In particular those who weren’t even around to see it begin ... I’m looking at you, Charlotte!

xxx Chapter 1

Introduction

As the supply of lighter, easier to process crude oil declines, oil refineries are shifting to heavier, more contaminated crude oils. These heavier oils pose numerous challenges to the refineries. Foremost are the elevated yields of atmospheric and vacuum distillation residua obtained from their processing. This residue is valuable as it may be converted to liquid fuel products. Unfortunately, heavier oils contain higher levels of contaminants such as S, N and metals which are concentrated in the residue. These contaminants complicate residue processing and pose an environmental hazard if not recovered and disposed of appropriately. This research aims to contribute toward an effective residue hydroconversion technology by studying the activity and deactivation of an unsupported molybdenum sulphide catalyst in a slurry-phase hydroconversion reactor. Slurry-phase residue hydroconversion reactors are favoured for their high levels of residue conversion and high yields of valuable liquid products. Such systems also exhibit reduced “coke” yields (solid carbonaceous and metallic deposits) and increased contaminant removal compared to alternatives such as fixed-bed reactors. These latter two points result in less environmentally damaging solid waste being generated. MoS2-based catalysts offer improved residue hydroconversion performance over, for example, FeS-based alternatives. Unfortunately, such MoS2 catalysts are also more expensive with Mo metal being several hundred times more expensive than Fe. To be economically feasible, deactivated MoS2 catalysts must be regenerated and recycled. A thorough understanding of the mechanism by which

1 such a catalyst functions and the process by which it deactivates is needed so that, in future, an efﬁcient regeneration methodology can be developed.

Background Oil refineries have two main options for the processing of residue oil feedstocks: carbon rejection and hydrogen addition. Hydrogen addition, performed in the presence of a suitable catalyst, is generally considered to be more efficient, affording improved conversion, yield of the desired high value liquid products and contaminant removal. Hydrogen addition may be performed in a fixed bed, fluidised bed or slurry-phase reactor. Slurry-phase reactors, whilst not as widely implemented in industrial hydroconversion applications as the other two systems, have been shown to offer various benefits over competing technologies. Many of these slurry operations use Fe-based catalysts on a “once-through” basis due to the low cost and ease of supply of such material. Research into alternative unsupported metal catalysts has led to the development of better performing Mo-based options. Such catalysts, however, demand prolonged operation through recycle and reuse to be economically feasible, a process that requires a thorough understanding of the deactivation of the catalyst together with an efficient regeneration methodology. Unfortunately, whilst the deactivation of supported metal catalysts (such as those utilised in fixed and fluidised bed hydroconversion applications) has been extensively studied and effective regeneration techniques developed, there remains relatively limited research regarding the deactivation of the unsupported metal catalysts of interest for slurry-phase hydroconversion. It is believed that the deactivation of unsupported Mo-based catalysts is the result of some interaction between the active MoS2 phase and coke formed during reaction. Unfortunately, neither the nature of this interaction nor the mechanism by which the active catalyst functions, is well understood with research on the topic being limited to conceptual models. Understanding these processes is key to regeneration for, if the deactivation could somehow be inhibited, reversed or otherwise “reset” to regain the functionality of the active catalyst between recycles, the catalyst may be used repeatedly with a reduced degradation in performance.

2 Hypotheses The hypotheses of this research project may be stated as follows:

1. A mechanistic understanding of hydroconversion affected by active and de-

activated unsupported MoS2 catalysts will allow for the deduction of the process of deactivation of such catalysts in residue hydroprocessing reactions.

2. The deactivation is due to the morphology of the coke, which precipitates during the reaction and agglomerates with the catalyst particles, changing with continued recycling of the coke-catalyst solid material recovered after the reaction. The hardening of the coke results in the formation of a solid, unreactive “support” for the catalyst and subsequent physical deactivation.

Key Questions These hypotheses lead to four major questions to be addressed in this investigation as presented below, together with minor questions where necessary to clarify the speciﬁc concepts of interest.

1. Can model compound experiments provide sufﬁcient information to under-

stand the hydrocarbon reaction mechanism over the MoS2 catalyst?

(a) What model compounds suitably represent heavy oil residua feedstocks? (b) Are these model compounds simple enough to provide the analytical resolution for mechanistic studies?

2. Does the reaction environment (condition response time and stability, reactor material, mixing regime, etc.) affect catalyst performance?

(a) What effect does each factor have and to what degree? (b) Can such effects be overcome through creative reactor engineering?

3. What are the hydroconversion reaction mechanisms associated with fresh

and deactivated MoS2 catalysts?

3 (a) Do the hydroconversion mechanisms, proposed and opposed in the literature, accurately represent the reaction as determined through model compound studies? (b) How does the process of deactivation affect the catalytic mechanism? (c) Can the mechanism of catalyst deactivation be determined and prevention or regeneration methodologies proposed?

Scope of Investigation The primary aims of this study were to:

• Develop a model compound testing methodology to accurately and efﬁ- ciently test unsupported catalyst performance.

• Determine the mechanism whereby MoS2 affects a catalytic hydroconversion reaction under appropriate conditions.

• Compare the mechanism of active and deactivated MoS2 catalysts and hence elucidate the mechanism of deactivation.

• Propose feasible preventative or regenerative techniques.

4 Chapter 2

Literature Review

To understand the advancements in knowledge which this investigation fulﬁlls, it is necessary to examine the technology and knowledge currently available in the associated ﬁelds, the research surrounding such technologies and those aspects which are perceived to require further study. To this end, the current “state of the art” is presented as per published literature.

2.1 Current Hydroconversion Technology The oil processing industry has an inevitably ﬁnite feedstock. This limitation has manifested in a decline in the quality of the crude oils being processed. Higher value feedstocks such as lighter “sweet” crude oils are steadily being replaced by heavier oils [3–7] or even alternative types of feeds altogether (such as bitumen derived from oilsands) [5]. Lighter feedstocks are desirable for they contain higher concentrations of species in the gasoline and distillate fuels range and reduced contaminant levels, factors which make them easier to process and generally result in higher yields of more valuable products and reduced waste. Heavier feedstocks, characterised by higher average boiling points, are often referred to as “sour” as they contain high concentrations of S, N and O together with metals such as V and Ni [3]. Not only does the processing of these heavier feeds usually result in lower yields of the desired lighter products, but the higher concentrations of the various contaminants makes this processing more difﬁcult and hence more expensive.

5 The use of these less-than-ideal feedstocks have their preliminary influences in the atmospheric and vacuum distillation systems, usually the first processing steps applied to a feedstock during oil refining for, with their higher average boiling points, such feedstocks result in far higher yields of distillation bottoms or residue as shown in Table 2.1. Not only are these residua more difficult to process than lighter streams due to their inherent physical properties (such as their vastly higher densities [presented in Table 2.1 as ◦API, an inverse of density] and their tendency to precipitate solids for instance) but the contaminants are often associated with the heavier species in these bottoms streams and are hence concentrated in the residue oil (as may be seen from the sulphur and metals contents in Table 2.1). It is thus essential from both an economic perspective, due to the higher proportion of the product such residua represent when processing heavier feedstocks, and from an environmental perspective, due to the increased amounts of contaminant materials present both in these streams and in the overall process, that technologies for the efficient refining of such residua be developed and implemented. In the context of this study, the focus is not on the processing of a heavy oil feedstock but rather the troublesome vacuum residue resulting from its distillation or from the use of bituminous oilsands feedstocks. Table 2.1: Comparison of physical properties of various crude oils and residua.

Property Oils Residua Light crude Cold Lake Athabasca Cold Lake Athabasca API gravity (◦API) 38 10 9 2.1 2.1 Viscosity (m2/s at 40◦C) 5 5000 7000 - - Sulphur (wt%) 0.5 4.4 4.9 6.15 6.18 Metals1 (ppm) 22 220 280 470 490 Vacuum residue2 (vol%) 38 10 9 100 3 100 3 Adapted from Gray [3]. 1 - Parts per million by mass. 2 - Liquid volume percent boiling above 525◦C. 3 - By deﬁnition, these residua are vacuum residue.

2.2 Residue Processing Technologies Simply put, it is desired to convert such vacuum residue or bitumen feeds into more valuable lighter products whilst simultaneously removing the contaminant species [3, 7], a procedure complicated by the physical properties of the feed and the ten-

6 dency of some of their constituent compounds (namely asphaltenes) to precipitate as difﬁcult to handle sediment or solid deposits [3] (referred to as “coke”). The properties that make these streams difﬁcult to process in terms of their physical properties arise from elevated carbon:hydrogen (C:H) ratios which in turn are due to the high concentrations of various aromatic and other unsaturated species. Two primary routes exist for the conversion of such feeds, both serving to reduce the C:H ratio, hence resulting in a decline in the viscosity, boiling point and solid formation tendencies of the feed. These routes involve either reducing the amount of carbon or increasing the hydrogen, termed “carbon rejection” and “hydroconversion” respectively [3, 4, 8, 9]. Both processes are carried out at elevated temperatures which results in the thermal (radical) cracking of larger hydrocarbon molecules in the heavy feed, further reducing the viscosity and boiling points, but potentially forming hard carbonaceous deposits (as part of the coke) due to radical condensation reactions [2, 9–16].

2.2.1 Carbon Rejection The carbon rejection or coking process, operated at elevated temperature and pressure (between approximately 450 to 565◦C and 1 to 20 bar, for processes such as visbreaking or fluid coking [8]) relies solely on thermally initiated radical reactions to both crack larger, higher boiling molecules into lighter species and to condense carbon-rich radical fragments into coke [3]. The removal of carbon as coke results in an overall reduction in the C:H ratio for the liquid species remaining, manifesting as a decline in the viscosity and average boiling point temperature [17, 18]. Unfortunately, the radical fragments which condense to form coke are often asphaltene molecules [2, 14, 19, 20]. These asphaltenes, defined as those species being insoluble in n-pentane or n-heptane [21], are known to contain disproportion- ately large amounts of the heteroatomic and metallic contaminant species [19–22], which are thus concentrated in the resulting coke. Whilst this is beneficial in that these species are being removed from the liquid product, it does however mean that the low value coke by-product, which may present up to 20 wt% of the final product [17], is heavily contaminated and thus represents a significant environmental hazard [23–25], with disposal costs thus associated with it. A greater financial

7 impact comes from the loss of valuable liquid products to such solid species, with coke formation thus representing a significant negative economic impact on the process. This mechanism of contaminant removal is also not the most efficient as it does not directly remove unwanted species. For a contaminant atom (such as a metal, sulphur or nitrogen) to be removed, it must be contained within a radical fragment which must then participate in a condensation reaction. Only this way are such contaminants captured in the solid coke which may be removed. Furthermore, as there is little control as to which of the fragments react (limited to temperature and residence time control), over-cracking to form lower value gaseous products is often a problem in the carbon rejection process [3, 8]. Despite these drawbacks, the low cost and simple operation of such processes means that they are profitable and hence common in commercial oil processing operations, manifesting as the visbreaking, delayed-coking, fluid-coking and flexi-coking systems [4, 17].

2.2.2 Hydroconversion Hydroconversion operating conditions vary greatly, with temperatures ranging from 370 to 450◦C and pressures from 7 to 27 bar [8, 17, 26], depending on the reactor type (ﬁxed bed, ﬂuidised bed or slurry-phase), catalyst type and feed. This process is often conducted in the presence of either a supported metal catalyst, such as

NiMo/Al2O3, or an unsupported metal catalyst, Fe or Mo for instance. The active metal phase is the metal sulphide, with the catalyst either being introduced as such a sulphide, or as a metal which is rapidly sulphided in-situ. Similarly to the carbon rejection process, cracking within a hydroconversion reactor occurs by radical reactions initiated by the elevated temperatures, with coke being formed by condensation reactions between radicals. The purpose of the catalyst in this system is to stabilise, “cap” or “quench” excess radicals, reducing both condensation reactions and over-cracking, resulting in lower yields of coke and gases. This is widely held to occur with the catalyst in some manner “activating” hydrogen dissolved in the residue oil to form free hydrogen radicals which then stabilise hydrocarbon radicals [3, 8, 9], inhibiting continued reaction and adding hydrogen to the molecules (resulting in an overall decrease in the C:H ratio). At present there are three main methodologies for the implementation of hy-

8 droconversion: fixed bed reactors, fluidised (or ebullated) bed reactors and slurry- phase reactors. For the processing of contaminated heavy feedstocks, fixed bed systems, despite being easy to implement and operate, are generally a poor choice as rapid coking and metals deposition at the entry to the catalyst bed results in a sharp increase in the pressure drop across the reactor, a decline in performance and frequent process shutdowns [4, 8, 18]. This may be avoided through the use of guard beds or guard reactors or by occasionally agitating the catalyst bed in some manner to break up deposits and agglomerates [4]. Furthermore, it is necessary to shut the reactor down in order to change the catalyst, although having multiple reactors running in parallel makes the impact of this action on overall plant operation less significant [4, 8]. Despite these factors, most commercial hydroconversion systems operate with fixed bed reactors, processing lighter, less contaminated feeds at milder conditions [18]. Fluidised bed reactors are a significant improvement on fixed bed systems specifically in terms of being able to remove and replace a portion of the catalyst charge as it becomes deactivated without interrupting the operation of the unit [4]. Such reactors, for instance the LC-Fining [3] and other fluidised-catalytic-cracking (or FCC) [4] reactors, are also able to handle far heavier and more contaminated feeds than fixed bed units as solids formed during the reaction, or even entrained in the feed itself, do not deposit on an immobile bed, increasing pressure drop and inhibiting performance as they would in a fixed-bed system [4]. Smaller catalyst particles may also be used, with the reduction in diffusion length increasing the overall reaction rate without adversely affecting the pressure drop [4]. Unfortunately, such catalysts must also be tailored to operation in a mechanically demanding environment, increasing their expense which, together with the additional complexity of operating a fluidised bed system, has limited the commercial applications of these reactors in residue hydroconversion processes [8, 18], with only LC-Fining and H-Oil being operated commercially [5]. Originally developed in Germany in the 1920’s, slurry-phase reactors have recently seen increased application for the hydroconversion of heavy feedstocks [8]. Combining the advantages of the fluidised bed systems (reduced pressure drop due to solid deposits and catalyst replacement without interruption) but allowing for smaller catalyst particles to be utilised (millimeter range of fluidised bed supported

9 catalysts down to nanometer sized unsupported catalyst particles), slurry-phase reactors are seen to offer high reaction rates and conversion together with reduced coke yields [26].

2.2.3 Slurry-Phase Catalytic Hydroconversion Numerous reviews have been published detailing the catalytic and operational factors of slurry-phase hydroconversion systems ([4, 5, 8, 26] for instance), and only a brief overview of those processes of commercial interest are provided here. The original implementation of slurry-phase hydroconversion technology makes use of low cost, single-use unsupported metal catalysts such as Fe, which may be introduced as iron sulphate (as per the CANMET process) [4, 8] or as an unrefined FeSx- containing mineral, usually pyrrhotite [5, 27], (as per the VEBA Combi Cracking (VCC) and HDH processes) [4, 5, 26]. These systems are operated at slightly more severe conditions than the conventional fixed and fluidised bed units, allowing for higher levels of conversion and contaminant removal to be achieved at the expense of elevated coke generation and increased catalyst deactivation [4, 8] with typical operating conditions and product specifications presented in Table 2.2. These negative effects are, however, not an issue as the unsupported Fe catalysts are disposed of after deactivation unlike the more expensive supported metal catalysts used in fixed and fluidised bed systems [5]. Of the many slurry-phase processes developed since the 1980’s, with the Micrometallic-coke (M-coke), ENI Slurry Technology

(EST), Super Oil Cracking (SOC), Intevep HDH and (HC)3 being added to those presented above [4, 8, 17, 18, 26, 28, 29], none have seen implementation beyond pilot plant scale [4, 17] due to lower proﬁtability as compared to thermal processes. VCC technology saw pilot scale operation of 4000 bbl/d, CANMET was run at some 5000 bbl/d and SOC was reported at 3500 bbl/d [4, 30]. Unfortu- nately, development of these technologies slowed by around 2000 with only ENI technology still being operated on pilot scale [30]. There are several reasons for this [17, 18, 31]: severe conditions, high hydrogen costs (due to hydrogen consumption and high operating pressures), longer residence times and, in particular, high catalyst costs due to high initial expense and rapid catalyst deactivation which usually necessitates disposal and hence monetary losses.

10 To improve proﬁtability and reduce waste and negative environmental impact, unsupported metal catalysts which can be used in such a manner as to achieve the same improved performance as the once-through disposable versions, but without having to dispose of the deactivated catalyst (contaminated with coke and its associated heteroatomic and metallic species), have been developed [30]. This improvement is only possible by recycling and reusing the catalyst multiple times, often requiring regeneration to regain lost activity and/or selectivity. One such catalyst, which has received a large amount of attention recently is Mo [1, 5, 18, 32].

The MoS2 active phase has been shown to be more active overall (for example in coke inhibition and hydrodesulphurisation, HDS) [33], more selective toward the desired middle distillate fuels range and produce less coke than the conventional

Fe-based alternatives as shown in Table 2.2. MoS2 may be introduced as dispersed particles produced ex-situ (in the form of micelles for instance) [26, 32] or formed in-situ by the addition of an oil- or water-soluble Mo compound (such as Mo naphthenate or ammonium heptamolybdate respectively) to the reaction medium [26, 32].

Table 2.2: Comparison of ﬂuidised bed and slurry-phase hydroconversion reactor performance.

Property Process LC-Fining 1,2 CANMET 2 M-coke 3 Reactor type Fluidised bed Slurry-phase Slurry-phase Feed type Athabasca bitumen CLVR 4 Vacuum residue Operating temperature (◦C) 425 - 450 440 - 460 5 400 - 454 Operating pressure (bar) 100 - 150 100 - 150 5 69 - 172 Catalyst type NiMo/Al2O3 Fe sulphate Mo naphthanate Catalyst loading (wt%) 2 - 13 6 1 - 2 0.01 - 0.02 7 Conversion (wt% 8) 66 86 > 95 Demetallation (%) 60 - 70 88 9 > 90 Desulphurisation (%) 65 80 9 Not reported Liquid product 10 (vol%) 65.2 88.7 99.5 Residue and solid product (vol%) 34.8 11.3 0.5 1 - Gray [3]. 2 - Nalithem et al. [34]. 3 - Speight [9], Bearden and Aldridge [33]. 4 - Cold Lake vacuum residue. 5 - Rana et al. [8]. 6 - Baussell et al. [35]. Loadings for various residue feeds, converted from 0.05 - 0.31 lb/bbl of feed (feed density of 1019 kg/m3). 7 - Converted from 100 - 200 ppm by mass. 8 - Material boiling above 524◦C. 9 - Furimsky [4]. 10 - Combined light and heavy naphtha and light and heavy gas oil.

11 Mo-based catalysts, prepared from reﬁned molybdenum metal or molybdenum oxide for water-soluble salts or oil-soluble complexes or mechanically ground

MoS2 (perhaps molybdenite ore), are, however, significantly more expensive than iron-based alternatives. With the price of molybdenum oxide (roasted molybdenite concentrate) at the time of writing being around USD 21,000 per tonne [36, 37] and iron ore (fines taken as a slight overestimate of the price of red clay or similar options) being only USD 130 per tonne [36]. Thus, despite improving performance and reducing waste, it is vital that Mo-based catalysts be efficiently recovered, regenerated and recycled if the processes in which they are utilised are to be economically viable and competitive [4, 5, 8, 25, 26, 38]. Recent research into the use of Mo-based catalysts has indicated that, depending on the precursor used, the active MoS2 phase may be recovered from a batch reactor system, together with the solid coke formed during the reaction, and this coke-catalyst agglomerate reused [1, 32, 38]. The performance observed from such recycled catalyst is, however, seen to decline after several recycles.

This decline in observed activity is due to the deactivation of the MoS2 in the reactor. Whilst a significant amount of research has been conducted to study the mechanisms of deactivation in supported metal catalysts (such as those utilised in fixed and fluidised hydroconversion reactors) [39–41], the deactivation of unsupported catalysts, particularly those utilised in hydroconversion systems, is not well understood [1]. Fundamental to the regeneration and recycling of Mo-based unsupported metal catalysts is an understanding of such deactivation mechanisms, which itself depends on an understanding of the mechanism of the active catalyst.

2.3 Catalyst Testing The development of an understanding of the catalytic mechanism for active hydroprocessing catalysts, and for that matter, the deactivation mechanism, is severely hindered by the complexity of the feedstock in which these catalysts operate [28].

2.3.1 Heavy Oil and Residue Oil Studies Crude oil is comprised of many thousands of different organic, inorganic, organometallic, aqueous and solid species making identiﬁcation and quantiﬁcation of in-

12 dividual species in the feed or product virtually impossible [10, 28]. This inability to determine the exact chemistry of the reaction has led to various techniques whereby the reactions (and the associated understanding and modeling activities) are described in terms of observed parameters (such as hydrogen consumption during reaction) or grouped measurements such as simulated distillation (SIMDIS) for boiling point distribution or elemental analysis of a recovered fraction, such as CHNS (carbon, hydrogen, nitrogen and sulphur) analysis. While useful for determining the effectiveness of a given catalyst in a reaction (in terms of coke suppression for instance), these grouped or lumped measurements do not provide the chemical species information required to reliably determine reaction mechanism affected by the catalyst [6, 42, 43]. To obtain the information required to determine how and why different catalysts work, it is necessary to greatly simplify the system through the use of model compounds [6, 42].

2.3.2 Model Compound Studies Studying catalyst activity in the presence of residue oil has both benefits and drawbacks. It is the actual feed in which the catalyst is designed to operate and, as such, studies using residue oil are extremely useful for tailoring catalysts, optimising process conditions and developing reactor systems. As indicated in Section 2.3.1, however, the complexity of such a feedstock makes it almost impossible to identify or quantify the myriad of species present. Furthermore, the physical and chemical properties of such heavy feeds (presented for various oils and residua in Table 2.1), and how they change during the reaction, can complicate studies in which they are used. One example of such a complication is that the metal species present in many residue oil feeds may themselves be catalytically active in the hydroconversion reaction [28]. An alternative to using residue oil is to substitute it with a suitable model compound. The use of model compounds (which may be introduced as a single species or as a mixtures of several chemicals) greatly simplifies the reaction system as all species of which the simulated feedstock is comprised are known and quantified. Unfortunately these model compounds may not actually be present in the original

13 feed and yet they are expected to provide accurate and applicable information regarding the reaction mechanism. Furthermore, whilst a model compound is ideal for studying the activity of an active catalyst (given that the clean feedstock is unlikely to cause the same degree of deactivation as residue oil), deactivation studies become more complicated. For such an experiment a catalyst must be deactivated in residue oil then its activity tested in the model compound. This means that the deactivated catalyst is being evaluated in an environment different from that in which it originally lost its activity, an environment which may affect how it functions (deactivation caused by coke deposition may be reversed, for instance, if the coke dissolves in the model compound). The selection of the model compounds and the reactors in which they are tested is thus not a trivial endeavor [42, 44].

Selection Model compound selection is often an optimization problem whereby the complexity of the feed is juxtaposed with its analytical simplicity. A more complex feed (not necessarily a mixture of more species but perhaps simply a larger molecule) may offer better representation of the feedstock than does a simpler molecule, but as the complexity of the feed, and hence its similarity to the residue oil, increases, so does the difficulty and complexity of chemical analyses and hence mechanistic evaluations. A large, polynuclear aromatic molecule with multiple alkyl branches and heteroatomic constituents may be an accurate representation of the asphaltenic fraction of residue oil [15, 45], but the range of products collected from the reaction of such a molecule (gases, liquids and even solids from condensation and precipitation reactions) would require the use of multiple analytical procedures for full identification and quantification. A smaller model compound may make analysis and interpretation easier, but such a species may not offer an accurate representation of the residue oil feedstock. The use of model compounds to simplify reactions to the point of mechanistic understanding is by no means a novel concept and has been conducted in the oil processing and coal liquefaction fields for many years and continues to see a great deal of active research [15, 27, 42, 44–56]. Some studies focused on specific reactions, such as the hydrodesulphurisation (HDS) of benzothiophene [50] or the

14 hydrodeoxygenation (HDO) of benzyl phenyl ether and bibenzyl ether [53, 56]. Some aimed to determine the mechanisms associated with specific structural components such as the aromatic rings of naphthalene, fluorene and pyrene [42]. Oth- ers sought to develop novel model compounds which accurately represent specific species or groups of species in a feed, such as 1-dodecylpyrene [44], cholestane- benzoquinoline [2], 1,3,6,8-tetrahexylpyrene [15] or larger “archipelago” compounds [14] for representing the asphaltenic fraction of residue oil. Yet more studied the interactions between different fractions, with both represented by model compounds, such as the hydroconversion of aromatic species in the presence of a hydrogen donor solvent (diphenylpropane in tetralin for instance) [11]. Of the many species proposed and tested over the years, diphenylmethane (or DPM, shown in Figure 2.1) has been widely used and accepted [27, 45, 46, 52– 55] as a model compound in both residue hydroconversion and coal liquefaction studies. There are two main reasons for this: thermal stability and molecular and reaction simplicity. Diphenylmethane may be viewed, in a simplified manner, as having only two types of functional groups: an aromatic ring and a proximal alkyl-aryl C-C bond where the methyl bridge connects to the ring. This means that a thermally initiated reaction would have to sever the proximal Calkyl-Caryl bond to form a benzyl and a phenyl radical (as illustrated in Figure 2.1), both unstable, high energy species [46].

This Calkyl-Caryl bond is very difﬁcult to break and the resultant thermal reactivity of DPM is low [45, 46, 48, 49]. Appreciably rapid thermal decomposition of DPM would thus require initiation by hydrogenation of or, more commonly believed, radical addition to one of the stable phenyl rings, destabilising the molecule and allowing the thermal reactions to proceed [15, 21, 42, 44, 46, 49]. The radical attack is thought to occur at the ipso position of the ring by either radical hydrogen addition or hydrocarbon radical addition. Due to this stability, DPM is considered to be a good kinetic representative of the much larger, refractory aromatic species (such as the asphaltenes) present in residue oil.

15 ∆G 718K = 218 kJ/mol r + Thermal Diphenylmethane Benzyl Phenyl radical radical

Figure 2.1: Thermal cracking of diphenylmethane

Reactors and Conditions With the extent of literature published, one can ﬁnd examples of virtually any reactor system being successfully used for heavy oil, residue oil and model compound studies. These range from the magnetically stirred 230 mL stainless steel system by Matsumura et al. [55] which was operated in batch and semi-batch (for hydrogen gas) mode to the 150 mL stainless steel system of Liu et al. [48] which was operated in fully continuous mode. Far smaller systems, often termed “micro- reactors”, are also quite popular, usually being operated in batch mode. Examples of such units include the 50 mL reactor by Sato et al. [54] which was rocked to mix and the unagitated system developed by Savage et al. [44] which required only 45 mg of reactant. Both of these reactors were constructed of stainless steel.

2.4 Micro-Reactors for Catalyst Testing The reactor systems indicated in Section 2.3.2 are not, however, all the same. “Micro-reactor” is used in the literature to refer to a variety of different reactors with varying designs and operating conditions. These are generally divided into two categories: ﬂow and batch, each having distinct advantages and disadvantages. This study is concerned solely with batch micro-reactors.

2.4.1 Advantages and Disadvantages of Micro-Reactors Micro-reactors beneﬁt from requiring less reagent and catalyst charge and producing less waste. This makes them ideal for rapid, cost-effective and environmentally conscious screening of catalysts and reaction conditions [57–60]. Reduced dimensions improve heat and mass transfer and allow for more accurate control of process conditions due to faster responses [57–61]. This improved condition control often results in improved conversion and selectivity [57, 62]. As the dimensions of the reaction chamber decrease, so the wall surface area to volume ratio (A:V)

16 increases, affording the opportunity for solid catalysts to be securely adhered to this surface instead of being separately supported on particles [60, 61, 63]. There are, unfortunately, also numerous disadvantages associated with scaling down a reaction system. The lower product volumes necessitate care in recovery, work-up and analysis and may make the recovery of certain materials (such as spent catalyst) difficult or impossible [60–62]. Control of the process conditions can be quite sensitive and even small fluctuations can propagate rapidly through a micro- reactor system [57–60]. The increased wall area to volume ratio may negatively influence the system by inhibiting mixing through frictional effects [60] and exert- ing a noticeable, and often unwanted, catalytic influence [59]. One of the biggest challenges associated with such systems in solid-catalysed reactions, however, is that most implementations can not handle either solid particles or the precipitation of material. Due to the small dimensions, solids rapidly foul the reaction system and necessitate unacceptably frequent shut down and cleaning [59, 61, 64]. Mixing in micro-reactors is another complicated matter [60] and whilst many studies have been conducted in the field, these have focused almost exclusively on micro-flow units operating gas-gas, gas-liquid or liquid-liquid phase systems.

2.4.2 Micro-Reactors in Hydroconversion Studies Despite the complications associated with mixing and solid material (introduced solid catalyst particles or precipitating products), micro-reactors see extensive use in applications where these factors may be considered to be quite extreme, specifically in the slurry-phase hydroconversion of vacuum residue. In such an application, for instance in catalyst screening or model compound testing [14], the reaction mixture is a gas-liquid-solid system with micrometer-scale unsupported catalyst particles suspended in the reaction liquid, as discussed in Section 2.2.3. As such, not only is thorough mixing necessary to ensure adequate contact between the gas and liquid-solid slurry, but also to ensure that the solid particles remain suspended in the liquid and that the resulting slurry is well-mixed. Many of these systems experience coking, and hence the formation of additional solid material, during reaction which may precipitate and remain suspended or deposit on the reactor walls [14, 15]. The high temperature, S-rich environment also results in the sulphidation

17 of the metal reactor itself, potentially turning the surface of exposed walls into catalytically active centers [21, 59, 60], obfuscating the results of thermal experiments or the effect of the catalyst added for study. Numerous researchers have developed ways to overcome these difficulties. Foremost is the almost exclusive use of batch [2, 14, 15, 44, 46, 54, 56, 65], rather than flow [53], systems. In this manner the fouling and plugging of flow paths by catalyst or coke is avoided. Mixing is somewhat more complicated by the reduced size of these units. Very few examples exist of internal agitation [63] (due partly to fabrication and operational complexities of size and partly to the increased wall friction effects indicated in Section 2.4.1) with most relying on an externally applied mixing regime. Examples of such engineering include spinning [65], rocking [54] or vertical shaking [2, 15] of the reactor, whilst some researchers are content with no mixing [44], a potentially feasible approach in very small systems due to reduced diffusion distances [59] and the slow settling of ultra-fine particles. Unlike the mixing studies conducted for micro-flow reactors, no literature could be found examining the effectiveness of various mixing techniques in micro-batch systems. The phenomenon of wall activation is, for the most part, overlooked in the slurry hydroconversion field of research. This factor, greatly exacerbated as the wall area:reaction volume (A:V) ratio increases with reduced reactor size, is only briefly indicated by a few researchers [21, 59, 66] and little is done to prevent or mitigate these effects. Some reactors are designed to include inert layers (such as the inner glass tube in the flow systems by Matsuhashi et al. [53] or Khorasheh and Gray [67], or the batch system by Alshareef et al. [2]) to separate the reaction liquid from the metal walls whilst others have resorted to extensive cleaning between reactions (such as the multi-stage acid etching by Savage et al. [44]) to help reduce inter-reaction effects or build-up. Despite knowledge of these wall effects, many researchers do not mention them and no studies are known which indicate the rate at which this catalytic influence develops or its extent.

2.5 Catalyst Activity and Deactivation A discussion of catalyst activity and deactivation requires a knowledge of the active phase and how it is presented in the reaction mixture. Below is presented such a

18 discussion, with a particular focus on the MoS2 of interest in this research, together with an overview of catalyst deactivation and regeneration processes.

2.5.1 Catalyst Selection Catalyst selection for a given hydroconversion system is a complex endeavour which examines all aspects of the system such as the properties of the feedstock, the reactor type and the operating conditions. A signiﬁcant amount of literature has been published on this topic, with the work below being but a brief overview.

Supported and Unsupported Catalysts Supported catalysts are common in oil processing. Such catalysts usually take the form of metal sulphide (Mo or W promoted by Ni or Co) crystallites on a porous γ- alumina, zeolite, silica, silica-alumina or carbon support material [4, 28, 41]. Such supported metal catalysts ease handling (due to larger particle sizes) whilst maintaining a high level of metal dispersion (as small crystallites over the large surface area of the porous support material) [4, 28]. The support itself may also play a role in the chemistry of the reaction (acid catalysed reactions by zeolite supports for example). Such supported catalysts are, however, prone to rapid deactivation by both physical blockage of the pores by coke and contaminant metal crystallites and direct poisoning of the active sites themselves [4, 8, 28, 41]. Unsupported catalysts, formed in-situ by the decomposition of oil- or water- soluble precursors added to the feed or introduced as finely divided solid particles [5, 17, 32, 68, 69], overcome some of these difficulties but pose their own challenges. The micron- to nanometer-sized particles from the decomposition of oil-soluble precursors [31], shown to be more active than water-soluble precursors or mechanically ground particles [5, 18, 31, 68], offer extremely high dispersion [5, 31] but are also extremely difficult to separate from the solid coke precipitates with which they are recovered from hydroconversion reactions. Recent studies have thus focused on recycling the coke-catalyst agglomerates formed during such a process in their entirety [1, 18, 32]. Initially, the lack of a porous support inhibits the deactivation of these catalysts as there are no inactive surfaces upon which coke can amass, and potentially occlude the active metal crystallites, and there are no

19 pores susceptible to blockage [4, 8, 28, 39, 70]. Once sufﬁcient coke has formed and agglomerates with the catalyst particles, however, this coke-catalyst agglomerate is, for all intents and purposes, a supported catalyst. Whether agglomerated or not, dispersed catalysts do still undergo poisoning and/or fouling due to the formation of contaminant metal crystallites on the active phase Bartholomew [71].

Catalyst Active Phase In the slurry-phase systems, unsupported metal sulphides are the most commonly used catalysts [5, 17]. Of these, MoS2 is the most effective, being introduced to the system as an oil-soluble precursor (such as Mo naphthanate or octoate) [5, 31]. An oil-soluble precursor, dissolved in the feedstock, decomposes at elevated temperatures, allowing the metallic species to react with H2S or other sulphur-containing species in the oil, to form active metal sulphide particles [5, 33, 72]. Metallic salts introduced to the system (ammonium heptamolybdate for instance) also decompose to an active metal sulphide, with these reactions being shown to proceed through various oxy-sulphide salts [5]. As such, given the severe conditions and high sulphur concentrations, metal sulphides are the obvious choice for a stable active phase in such reactions.

2.5.2 Molybdenum Disulphide

This work focused on the use of MoS2 as an unsupported hydroconversion catalyst. With the proven effectiveness of this catalyst, numerous works have been published reporting the properties, structure, active sites and theorised catalytic mechanisms of this material [5, 21, 29, 31, 38, 68, 69, 73–78]. Below is presented a brief overview of this literature.

Structure and Active Sites

The structure of MoS2 is similar to that of other transition metal sulphides, particularly WS2, and presents as a layered S-Mo-S crystal as depicted in Figure 2.2

[28, 29, 69, 74, 75]. It may be seen that the MoS2 consists of trigonal prisms of S coordinated to Mo, forming large sheet-like structures. These sheets asso- ciate with one another by weak van der Waal’s forces to form stacked crystallites

20 [28, 29, 69, 74, 75]. The size of the sheets and the height of the stacks is inﬂuenced by the conditions under which the MoS2 is synthesized [28]. Occurring naturally in large, ordered sheets as the mineral molybdenite [29, 31, 79], synthesized MoS2 often presents as highly bent, disordered sheets termed a “rag” structure [29].

Figure 2.2: Rendering of arbitrary 5-layer stack of MoS2 (created in Accelrys Materials Studio 4.4 with published unit cell data [80])

With the homogeneous sulphur basal planes of the MoS2 sheets considered chemically inert, it is the rim and edge atoms where coordinatively unsaturated sites and/or S anion vacancies exist which are considered to be the catalytically active sites [5, 28, 38, 75–77, 81] (although the role played by the former is a point of debate [29]). Other defects, such as distortions or inhomogeneities in the sulphur basal plane, are also active [29, 75, 76]. The sulphur anion vacancies at such active sites afford them Lewis acid character, allowing for the adsorption of molecules with unpaired electrons [40], and given the high density of such sites at the rim/edge or along basal plane features such as folds, double or higher vacancy points may occur [40]. It has also been proposed that -SH groups on Mo catalysts exhibit Brønsted acid characteristics and associated cracking activity [82]. The activity of an MoS2 catalyst may thus be improved by increasing the proportion of rim/edge sites and basal plane defects. This is achieved by reducing the both size

21 of the sheets and the stack height or by intentionally distorting the MoS2 sheets, for instance by controlling the synthesis conditions or through chemical exfolia- tion [29, 38]. The promotional effects of defects and distortions may explain why molybdenite, even when milled to the same particle size, exhibits reduced hydroconversion activity as compared to synthetic MoS2 [31, 68].

Catalytic Mechanism Before determining the inﬂuence which a catalyst may have in a hydroconversion reaction, it is necessary to understand what thermal processes (both radical chain reactions and thermal hydrogenolysis) may occur [83]. Depicted in Figure 2.3a, the thermolysis of DPM initiates with the homolytic cleavage of the Calkyl-Caryl bond to form benzyl and phenyl radicals. These radicals may propagate the reaction by either radical addition to or H* abstraction from other DPM molecules. As an example of this propagation, a benzyl radical may abstract hydrogen from the alkyl carbon of DPM (shown in Figure 2.3b) to form benzene. The DPM radical thus formed may continue to crack or the reaction may terminate, for instance by radical addition with a benzyl radical to form 1,1,1-(1-Ethanyl-2-ylidene)tris- benzene (ETB) as shown in Figure 2.3c. With excess hydrogen, the reaction may proceed via abstraction of H* from dissolved H2 as shown in Figure 2.3d for a phenyl radical. The resulting H* radical may then propagate the reaction by either abstracting H* from other species, similarly to phenyl in Figure 2.3b, to form H2, or through radical addition and subsequent C-C homolysis as shown in Figure 2.3e [83]. Thermal hydrogenolysis [83] occurs by direct C-C cleavage and hydrogen insertion as shown in Figure 2.3f.

Despite the extensive research published regarding the use of MoS2 as a hydroconversion catalyst, the mechanism is still relatively poorly understood, with many of those proposed, and indeed widely held and perpetuated, being considered at best conceptual understandings [29, 75] or openly challenged as partially or totally incorrect [21]. The role of the catalyst is generally presented as performing the heterolytic (to form Mo-H and S-H moieties) or homolytic (to form two S-H moieties) dissociation of hydrogen. These species either remain on the catalyst surface or desorb into the liquid phase [5, 21, 28, 29, 46, 75–77].

22 + Thermal (a)

+ +

(b)

(c)

+ HH + H

(d)

+ H +

(e)

+ HH +

(f)

Figure 2.3: Mechanisms for the thermal decomposition of diphenylmethane (DPM). (a) Initiation by Calkyl-Caryl thermolysis of DPM. (b) Propaga- tion by hydrogen abstraction from DPM by a phenyl radical. (c) Termi- nation by radical addition between a DPM radical and a benzyl radical. (d) Propagation by hydrogen abstraction from dissolved H2 by a phenyl radical. (e) Propagation by radical addition of H* to DPM. (f) Thermal hydrogenolysis to form benzene and toluene by H2 insertion to DPM.

23 Since the 1960’s, a mechanism proposed by Curran et al. [78] that this process forms a reactive hydrogen species, usually referred to as “H*” as shown in Figure 2.4a, has been widely held and supported [5, 29, 31, 46, 68, 75–77]. These activated hydrogen species are theorised to spill-over across the surface of the catalyst [5], and its support where applicable, and even back into the liquid phase [29], “capping” or “quenching” hydrocarbon free radicals formed through the thermal decomposition of the feed (for instance by radical addition as shown for a phenyl radical in Figure 2.4b). The overall mechanism, in the context of DPM hydroconversion, is illustrated in Figure 2.6a. This theory has proved popular as it explains many of the trends observed in residue hydroconversion reactions upon the addition of a catalyst: increased hydrogen consumption, reduced overcracking to gas and reduced condensation to solid coke. Additionally, such H* species are thought to promote conversion of hydrocarbon feedstocks. This is theorised to occur by two mechanisms. Firstly, the rapid stabilisation of thermolysis radicals inhibits condensation reactions, promoting the decomposition of larger hydrocarbons. Sec- ondly, much like hydrocarbon free radicals through the process of radical hydrogen transfer (RHT) [15, 21, 44, 49, 84], activated hydrogen may serve to promote the decomposition of otherwise thermally stable aromatic hydrocarbons through their addition to an aromatic ring and subsequent cracking [44, 49, 84–87], as shown for DPM in Figure 2.6b. Recent research [21, 29] has raised doubts as to the validity of the theory proposed by [78]. These studies, based mostly around the work of LaMarca et al. [88], propose that following an initial high activation energy thermal cracking step, decomposition of the feed occurs predominantly by radical chain reactions through a series of radical hydrogen transfer (addition or abstraction) and scission steps before terminating, by radical recombination or radical-to-oleﬁn addition, as stable hydrocarbon products [21, 88]. In the context of DPM hydroconversion, initiation would occur as per Figure 2.3a. Hydrogenation by the catalyst and continued thermolysis breaks these species into shorter hydrocarbon radicals, for instance phenyl radicals into C3 radicals per Figure 2.4c. Radical addition of these short chain radicals to the DPM followed by β-scission, per Figure 2.4d, decomposes the DPM feed into benzyl radicals and alkyl-benzene species. These benzyl radicals may continue to react or terminate as per one of the mechanism presented above. A

24 generalised mechanism for this reaction is presented in Figure 2.6c. The radical chain reaction theory of LaMarca et al. [88] suggests that the capping or quenching of radicals would be detrimental to the overall performance of the system as it would inhibit radical interaction with the feed and subsequent cracking. Unfor- tunately, many of the studies aimed at elucidating the mechanism of liquid-phase hydroconversion (often in the presence of a hydrogen donor or shuttle solvent) have met with mixed results [21], some supporting the mechanisms proposed by Curran et al. [78] and others those of LaMarca et al. [88].

HH 2H Catalytic (a)

+ H

(b)

+ 4.5 H2 2 Catalytic (c)

+ +

(d)

Figure 2.4: Mechanisms for the thermocatalytic decomposition of diphenylmethane. (a) Initiation by formation of activated hydrogen species on catalyst surface. (b) Termination by stabilisation of a phenyl radical by hydrogen radical addition. (c) Propagation by thermocatalytic cracking of a phenyl radical to propane radicals. (d) Propagation by radical addition of a propane radical to DPM followed by β-scission.

A ﬁnal mechanism of interested in such a catalytic system is catalytic hydrogenolysis [89–91]. By this mechanism, dissolved hydrogen would dissociatively adsorb, heterolytically or homolytically, on the surface of the catalyst as dictated by a Langmuir adsorption isotherm [92]. A DPM aromatic ring adsorbed in the vicinity of these hydrogen species is likely to undergo hydrogenation, as shown in Figure 2.5a, to produce saturated products such as the 2-Benzyl-1,3-

25 cyclohexadiene illustrated, or hydrogenolysis, as shown in Figure 2.5b, to produce a mixture of benzene and toluene. Given the adsorption, reaction and desorption steps required for these reactions, it is possible that such a system would be governed by Langmuir-Hinshelwood-Hougen-Watson (LHHW) kinetics [92–94] whereby one step would be rate limiting whilst the others could be considered in quasi-equilibrium.

H H

Catalyst surface (a)

H H

Catalyst surface (b)

Figure 2.5: Reactions occuring on the catalyst surface between adsorbed hydrogen and diphenylmethane. (a) Hydrogenolysis of adsorbed phenyl ring by dissociatively adsorbed hydrogen to form partially hydro. (b) label 2.

With this in mind, the state-of-the-art knowledge regarding the role of the catalyst in residue hydroconversion is that, using H2 dissolved in the liquid, the catalyst hydrogenates oleﬁns (which are reactive and themselves promote overcracking and coke formation reactions) to more stable saturates, hydrogenates poly-aromatic species to form hydrogen donor or shuttling compounds and exacts catalytic hydrogenolysis [21].

2.5.3 Processes of Deactivation Whilst understanding the mechanism of catalytic activity is important, development of a meaningful regeneration and recycle regime requires an understanding of how the catalyst deactivates, the chemical and morphological changes which re-

26 H2 2H Catalytic

+ + Thermal Capping or quenching

Continued cracking or condensation reactions (a)

+ H +

H2 2H Continued cracking or Catalytic condensation reactions (b)

CxHy CxHy + + CxHy

Continued cracking or condensation reactions (c)

Figure 2.6: Simpliﬁed literature mechanisms for the thermocatalytic decomposition of diphenylmethane. (a) Capping of thermal radicals (adapted from Curran et al. [78]). (b) DPM destabilisation by active hydrogen radicals (adapted from Wei et al. [46]). (c) DPM destabilisation by radical hydrogen transfer (adapted from Gray and McCaffrey [21]).

27 sult in the loss of catalytic activity [7]. The deactivation of heterogeneous catalysts has been a topic of research for many years and the varied processes accounting for this phenomenon are, generally, very well understood. The interpretation of these processes to a specific application, however, is more difficult. In few systems is this more the case than residue hydroconversion wherein the myriad of species in the feed and the complexity of the reaction networks makes determination of the physio-chemical deactivation mechanisms extremely complicated [7]. The deactivation of catalysts, supported or unsupported, in heavy oil or residue hydroconversion is due to the formation and deposition of coke and metals and the interaction of organometallic and heteroatomic species with the active sites [1, 4– 7, 10, 18, 28, 39–41, 43, 70, 71, 95–104]. The exact nature of the deposits and the influence of these deposits and other species on the system are dependent on many factors: the composition of the feed, operating conditions, type of catalyst support and type of active phase to name a few [41]. For supported catalysts, it is generally held that initial deactivation occurs by rapid coke deposition on the surface and in the pores of the support. This deposition reaches a pseudo-steady-state at which point more gradual metal deposition continues to steadily deactivate the catalyst by both plugging pores and physically occluding and/or poisoning active sites [10, 28, 95, 98]. Final and total deactivation occurs rapidly when coke and metal deposits constrict and block the catalyst pores, eliminating all contact between the active sites and the reaction mixture [28, 95, 97]. Few theories have been proposed for unsupported catalyst deactivation, but the most plausible explanation is that the precipitation of unreactive graphitic coke envelopes and encapsulates active metal particles, preventing their participation in the reaction [1, 18]. Extensive descriptions of the various deactivation mechanisms including fouling, poisoning, thermal degradation, solid-state reactions and vapour-phase degradation and mechanical degradation are provided in Section A.1.

2.5.4 Catalyst Regeneration Methodologies Catalyst deactivation may be controlled in three ways: prevention, mitigation and regeneration. Prevention often involves a physical change to the reaction system

28 (such as removal of feed contaminants, the installation of guard beds or changing process conditions to avoid deactivation altogether [40, 71, 96]). Mitigating deactivation involves temporarily changing operating parameters to overcome the effects of deactivation whilst not interrupting the system (for instance by continually adding fresh catalyst to compensate for deactivation [1] or steadily increasing the temperature of the reaction [40, 71, 103] as is common in residue hydroprocessing). Despite these options, one of the most direct and commonly implemented methodologies for extending catalyst life is the use of some form of regeneration step prior to reintroduction of a recovered catalyst to the reactor [40, 105]. Such techniques, generally incorporating some form of thermal and/or chemical treatment, are many and are presented in the literature with varying success in different applications. Descriptions of such treatments are provided in Section A.2.

2.6 Summary of Findings from the Literature As oil reﬁneries shift to heavier feedstocks, the efﬁcient processing of vacuum residue into valuable liquid fuels becomes ever more important. Whilst carbon rejection is a cheap and well understood technique, catalytic hydroconversion (with unsupported catalysts in slurry-phase reactors) is more effective. MoS2 is one of the best catalysts available for this application but it is expensive. Recycling the solid coke-catalyst agglomerate has been shown to prolong its use but the extended time-on-stream results in deactivation. A regeneration methodology is required for catalytic hydroconversion with unsupported MoS2 to be economically viable. The complexity of the residue oil feedstock presents a challenge in understanding the mechanism of this catalytic hydroconversion reaction. This mechanism is mostly unknown with published theories being disputed due to the lack of experimental evidence. Mechanisms for deactivation are limited to conceptual models. Simplifying the system through model compound studies could allow for deduction of the reaction and deactivation mechanisms and aid in the development of effective regeneration methodologies.

29 Chapter 3

Experimental

Fundamental to any meaningful experimental investigation is reproducibility, with a signiﬁcant part of this being a thorough understanding of the experimental apparatus, procedures and calculations conducted. As such, descriptions of these aspects must allow for other researchers to examine, critique and, if desired, accurately recreate the experiments reported. This section presents the experimental objectives of this investigation, the apparatus and analytical equipment used to pursue those objectives and the associated procedures and calculations.

3.1 Experimental Objectives and Programme The objectives of this study are the practical means whereby the key questions posed in Chapter 1 may be addressed and answered, a conversion of the more theoretical questions into speciﬁc, experimental goals. With these goals in mind, a detailed experimental programme may be established with clear expectations regarding which factors are to be examined in each experimental series, what data is to be obtained and the reasoning behind these decisions. As each of the key questions of this study yields multiple objectives, each of which may require some degree of explanation, they are discussed separately in Section 3.1.1 below, with the experiments resulting from these objectives being programmatically presented in Section 3.1.2.

30 3.1.1 Interpretation of Questions to Objectives

Can model compound experiments provide sufﬁcient information to understand the mechanism of the MoS2 catalyst? Presented as “Phase 1” of the experimental programme (Section 3.1.2), answering this question requires multiple model compounds to be selected and subjected to a series of screening experiments. These experiments will determine both the applicability of each model compound in terms of representing a residue hydroconversion reaction (by comparing observed conversion with published residue hydroconversion experiments under the same conditions) and the degree to which the analytical results from each reaction allow for a mechanistic understanding to be developed (that is, if the chemical analysis of the products is simple enough to allow for meaningful, contextual interpretation). As discussed in Section 2.3.2, numerous model compounds have been used in published hydroconversion studies, with diphenylmethane (DPM) being selected as the focus for this work. The reason for this choice is that it affords both the resistance to thermal cracking desired in the representation of heavy species in residue hydroconversion reactions (largely due to it only possessing stable proximal Calkyl-Caryl bonds and aromatic rings) and due to its small and predictable cracking mechanism (illustrated in Figure 3.1a). The simplicity of this species does, however, have a negative aspect in that regardless of which Calkyl-Caryl bond breaks, the products will be identical (as illustrated in Figure 3.1a). Furthermore, secondary cracking of these products, for instance of toluene (after stabilisation of a benzyl radical) to a methyl and phenyl radical, would be unlikely (in this case due to the instability of the C1 radical), meaning that an accurate determination of the rate of stabilisation of the primary cracking radicals may be difﬁcult. Two other model compounds were thus included in this study. Diphenylethane (DPE) and diphenylpropane (DPP), shown in Figures 3.1b and 3.1c respectively, are structurally similar to DPM but with two- and three-carbon n-alkyl linkages respectively. It was thought that increasing the alkyl bridge length would be a trade-off between increased reactivity (the Calkyl-Calkyl bonds being more susceptible to thermolysis than the Calkyl-Caryl bonds), and hence reduced applicability to residue hydroconversion, and increased analytical insight. The latter point arises as there exists a complex series of subsequent cracking, isomerisation and stabilisa-

31 tion reactions for the primary radicals of DPE or DPP thermolysis with the relative rates of these pathways providing valuable information into the reaction system. Additionally, to simulate the model compound representing only a low concentration of difﬁcult to crack species in a residue feed, dilution experiments were included in the screening studies. Decahydronaphthalene (decalin), illustrated in Figure 3.2, was selected as the solvent for these studies as it and similar species (such as tetralin) are commonly used for this purpose in published work. The saturated nature of decalin makes it an effective hydrogen donor and/or hydrogen shuttle in this system, simulating species in a residue feed which play similar roles. “Phase 1” allowed a model compound to be selected (from the DPM, DPE and DPP tested) and a decision made as to whether decalin dilution should be used.

Does the reaction environment affect catalytic performance? “Phase 2” of the experimental programme, following selection of a model compound and a decision regarding the use of a diluent, is to determine the effect of various operational aspects of the reaction environment on the observed reaction. In residue hydroprocessing, the reaction is often quantified in terms of broad observed properties (such as hydrogen uptake, changes in viscosity or simulated distillation curves) rather than detailed mechanistic parameters. Given the focus of this study on this latter point, however, such detail is of utmost importance in terms of both the model compound reaction and catalyst morphology. As such, many operational factors apart from the standard reaction temperature, reaction pressure, catalyst loading and so forth must be examined, with their impact on the reaction being quantified and, to as great a degree as possible, mitigated. One such factor is the control of the reaction temperature and the heat-up rate. With the thermal cracking range of 400 to 445◦C exemplified in Figure 3.3, it may be seen that a slow heat-up rate may un- necessarily expose the reaction mixture to elevated temperatures thereby resulting in unwanted and unquantified reactions which may be avoided by a system offering a faster, better controlled heat-up. Similarly, it is desirable to cool the system rapidly following the reaction to minimise continued reaction beyond the desired reaction time. Other equally important factors include the extent and impact of wall activation (whereby a stainless steel reactor wall may become catalytically active), the influence of mixing on unsupported catalyst morphology in a “clean” system

32 Either or

(a)

(1)

(1) +

(2) (2) 2 x

(b)

(1)

(1) +

(2) (2) +

(c)

Figure 3.1: Thermal cracking of diphenylmethane, diphenylethane and diphenylpropane to their primary products to illustrate potential for subsequent reactions. (a) DPM cracking to phenyl and benzyl radicals. (b) DPE cracking to phenyl, benzyl and ethylbenzene radicals. (c) DPP cracking to phenyl, benzyl, ethylbenzene and propylbenzene radicals.

Figure 3.2: Decahydronaphthalene structure. and the impact of hydrogen:reactant ratio in such a simplified reaction mixture. Given the complexities associated with the management of many of these aspects, Phase 2 commenced with the design of a novel micro-reactor system, based on the observations and findings from Phase 1, to allow for the quantification of, improved control over and selective elimination of unwanted influences.

33 445

400

300

200

Stirred reactor heat-up

Desired heat-up Reactortemperature (°C)

100 ~ 20 ~ min ~ 8min ~

00:00 00:15 00:30 00:45 01:00 01:15

Heating time (hh:mm)

Figure 3.3: Comparison of a heating proﬁle typical of the stirred batch reactor used in this study and a theoretical desired heating proﬁle.

What are the hydroconversion reaction mechanisms associated with fresh and deactivated MoS2 catalysts? With a model compound reaction system selected and undesirable effects quantiﬁed or eliminated, “Phase 3” of the experimental programme could commence wherein detailed reaction mechanisms could be proposed based on the experimental results from this model compound testing regime. These results and proposed mechanisms could be contrasted and compared with published hypotheses regarding the mechanism of hydroconversion and the role of the active catalyst in such a system (in particular, the widely held hydrogen radical capping theories depicted in Figure 2.6). With both the model compound testing regime and proposed active catalyst mechanism in place, MoS2 catalyst samples, deactivated in residue hydroconversion reactions, could be evaluated and the results compared to both the fresh MoS2 results and the residue hydroconversion results obtained for the same deactivated

34 catalyst samples (the residue deactivated samples and associated reaction data being obtained from Rezaei and Smith [1]). Furthermore, testing and residue result comparison of thermally treated deactivated MoS2 (prepared as part of a published study by Rezaei and Smith [1]) would afford additional insight into the process of unsupported MoS2 deactivation and allow for an informed discussion regarding possible mitigation and regeneration methodologies for such catalysts.

3.1.2 Experimental Programme As discussed in Section 3.1.1, the experimental programme of this study may be divided into three phases. The ﬁrst phase, model compound screening and selection, was performed in a commercially available stirred batch reactor with the results being used for the second phase, the design and development of a micro-reactor to offer a more accurate and controlled testing platform. The third phase, evaluation of active and deactivated catalysts and the deduction of reaction and deactivation mechanisms, was performed in this novel micro-reactor. Table 3.1 provides a brief overview of the experiments performed during this study, with a full listing of all experiments in Section B.1.

Phase 1 - Model Compound Evaluation Select model compounds and diluent Diphenylmethane, diphenylethane and diphenylpropane were selected as the model compounds for this study. Decahydronaphthalene was chosen as the diluent. A dilution of 3 wt% model compound in decalin, together with undiluted model compound, were selected for evaluation.

Select catalyst and loading The oil-soluble precursor molybdenum octoate was selected for use in this

study, forming the active MoS2 catalyst in-situ through thermal decomposi-

tion with CS2 (added as the sulphur source at three times the stoichiometri- cally required amount). Catalyst loadings of 0 ppm, 600 ppm and 1800 ppm Mo were chosen for evaluation (following the works by Rezaei et al. [32] and Rezaei et al. [18]).

35 Table 3.1: Summary of experimental programme.

Reaction temperature 1 Reaction time Catalyst loading Model compound Dilution 2 (◦C) (h) (ppm Mo) (wt%) 445 1 0 - 600 DPM, DPE, DPP 3 420 - 435 1 0 DPP Phase 1 445 0 - 8 0 - 600 DPM 3 - 100 415 - 445 1 1800 DPM 100

Reactor 3 Mixing speed Reaction time Catalyst loading Total feed loading 4 µ 36 (RPM) (h) (ppm Mo) ( L) 1 1800 400 - 500 Inclined SS 5 0 1 - 4 0 - 1800 400 Vertical SS 5 0 0 - 4 0 - 1800 400 0 0 - 4 0 - 1800 150 - 400 Glass insert

Phases 2 and 3 0 - 2250 1 0 - 1800 1 - 4 0 - 1800 150 2000 1 1800 6 1 3 - All 250 cm stirred batch reactor experiments were conducted at an initial reaction pressure of 13.79 MPa H2 and a mixer speed of 700 RPM. 2 - Decalin used as diluent for normal experiments. 3 - All experiments were conducted using undiluted DPM at a reaction ◦ 4 temperature of 445 C and an initial reaction pressure of 13.79 MPa H2. - The volume of mixed feed (model compound, catalyst and 5 6 CS2) pipetted into the reactor or insert. - 316 stainless steel. - Three residue hydroconversion coke-catalyst samples. Select appropriate reactor and reaction conditions A 250 cm3 stirred batch reactor, described in Section 3.2.2, was used for the preliminary screening experiments of this study. This reactor was operated at industrially applicable conditions of 415 - 445◦C and 13.79 MPa with a reaction time (at temperature) of 0 - 8 h and a stirrer speed of 700 RPM.

Conduct model compound screening experiments Following calibration of the analytical instruments (described in Section 3.3) used in this study, each model compound was evaluated at the aforementioned dilutions, catalyst loadings and reaction conditions. Blank experiments using the solvent (decalin) and anticipated major products (benzene and toluene) were also performed. To ensure an accurate interpretation of the data obtained, experiments were repeated at least three times (limited reagent and catalyst quantities notwithstanding).

Evaluate results The results from the screening experiments were examined alone for analytical simplicity (allowing for a detailed mechanistic understanding) and against comparable published residue hydroconversion data [18, 32] to determine the applicability of each model compound and dilution to such a system. The most appropriate model compound and dilution was selected and major sources of uncontrolled inﬂuence in this reaction system (such as heating rate or wall activity) were identiﬁed.

Phase 2 - Novel Reactor System Design and Testing Quantify undesirable inﬂuences Identifying the slow heat-up rate and catalytic wall activity as the major factors which contribute to the reaction (outside of the “normal” operating parameters of temperature, pressure, etc.), quantiﬁcation of their effects on DPM model compound conversion and product yield could be performed to determine the extent of their impact and hence the priority of their control or mitigation.

Develop novel reactor system

37 A novel micro-reactor system was developed with a specific focus on those undesirable and difficult to control/mitigate factors shown to impact the reaction. This design and commissioning proceeded through an evolution- like process and is detailed in Section B.3 with the final design presented in Section 3.2.2.

Evaluate reactor performance and refine methodology Operation of the micro-reactor was scrutinised by conducting a series of model compound experiments to correlate with those performed in the 250 cm3 stirred batch reactor. These experiments allowed for a more precise quantification of the various influential factors, including those unique to a “clean” system or a micro-reactor setup. These factors included: heat-up rate on catalyst active phase formation and reaction, rate of wall activation and catalytic influence, hydrogen diffusion rates and limitations and the impact of mixing on this process and the effect of mixing on unsupported catalyst morphology in a clean system. The results from these experiments allowed for the adjustment and improvement of the testing methodology, improving accuracy and reproducibility.

Phase 3 - Catalyst Study and Deactivation Investigation Conduct active catalyst experiments With the development of an accurate and reproducible reaction system wherein the activity of the catalyst could be isolated and studied without additional inﬂuences (the ﬁrst of its kind for such hydroconversion reactions), model compound experiments were conducted to gather data relating to the reac-

tion rate and product distribution for an active MoS2 catalyst.

Deduce active catalyst mechanism

Using the rate data and product distribution for the active MoS2 catalyst, the hydroconversion mechanisms proposed in the literature were scrutinised and modiﬁcations or changes made based on these experimental observations. To examine the wider applicability of this mechanism, it was applied to data from published residue hydroconversion studies and found to explain and predict the trends observed.

38 Conduct deactivated and heat treated catalyst experiments

Two spent MoS2 catalyst samples were then evaluated in the model compound system. Both samples were recovered in the form of a coke-catalyst agglomerate from residue hydroconversion studies [1]. The ﬁrst had been subjected to repeated recycles through the residue hydroconversion system whilst the second had been heat treated in an inert atmosphere after only a single reaction. Both samples were shown to be deactivated in residue hydroconversion experiments [1], affording minimal beneﬁt over catalyst-free comparisons.

Deduce mechanism of deactivation Comparison of model compound conversion and product distribution, juxtaposed with the published residue hydroconversion data [1, 18, 32], for fresh

MoS2 and the two deactivated MoS2 samples allowed for the hypothesisa- tion of a mechanism for the deactivation of this unsupported catalyst, independent of additional effects associated with parameters such as wall activity or slow heat-up rates. This information allowed for an informed discussion as to possible prevention and regeneration methodologies.

3.2 Experimental Apparatus and Supplies A list of the chemical species, their speciﬁcations and suppliers, used in this investigation follows together with a description of the experimental and analytical apparatus, their operation and a summary of the reaction conditions used.

3.2.1 Reaction and Analytical Supplies

Model Compounds and Diluent

Diphenylmethane (DPM, (C6H5)2CH2, Acros Organics, 99%), diphenyleth-ane

(DPE, (C6H5)2C2H4, Alfa Aesar, 98+%) and diphenylpropane (DPP, C6H5)2C3H6, Alfa Aesar, 98%) were used in this study, without further puriﬁcation, together with decahydronaphthalene (decalin, C10H18, Sigma-Aldrich, mixture of cis- and trans-decalin, 98%) as a solvent.

39 Catalyst Precursor

The active MoS2 phase was formed in-situ through the reaction of the oil-soluble precursor molybdenum octoate (C16H30MoO4, The Sheperd Chemical Company, molybdenum 2-ethylhexanoate in 2-ethylhexanoic acid, 15.5wt% Mo) and carbon disulphide (CS2, Sigma-Aldrich, ≥99.9%). The reaction occured rapidly during the heat-up period at temperatures as low as 415◦C (the lowest reaction temperature investigated in this study).

Reaction Gases

Ultra-high purity nitrogen (N2, Praxair, PP 4.8 [99.998%]) and hydrogen (H2, Prax- air, UHP 5.0 [99.999%]) were used to purge air from and pressurise the reactions respectively.

Analytical Standards and Gases Calibration of the gas chromatographs (GC, see Section 3.3) was performed using a certiﬁed gas mixture (Praxair). Gas chromatography-mass spectroscopy (GCMS) calibrations for liquid analyses were performed using standards prepared from the aforementioned model compounds and diluent together with benzene (C6H6, ® OmniSolv , 99.94%) and toluene (C7H8, Fisher Scientiﬁc, 99.8%) (see Appen- dices C.2 and C.1 for details). All GC and GCMS apparatus used ultra-high purity helium (He, Praxair, UHP 5.0 [99.999%]) for operation.

3.2.2 Reactors and Conditions Industrial slurry-phase reactors are implemented almost exclusively as continuous systems. On a laboratory scale, such systems are challenging due to increased reagent and catalyst consumption (which may be expensive or available in limited quantities), the complexities of operation (continuous feed mixing, product treatment, purging and recycling), the difficulty of maintaining stable operating conditions in a relatively small reaction volume (fluctuations increasing experimental uncertainty and hampering reproducibility) and safety concerns arising from the continuous supply of toxic and flammable species to a unit operating at high temperatures and pressures.

40 The alternatives are batch or semi-batch reactors. Batch units usually consume less reagent and catalyst, allow for easier control of reaction conditions and simplify feed introduction and product recovery. One drawback to the use of batch reactors in hydroconversion is the potential for hydrogen starvation. Whilst continuous and semi-batch systems supply a constant stream of hydrogen to replenish that consumed by the reaction, the hydrogen available to the reaction in a batch unit is limited. As discussed in Section 3.1.1 and indicated in Phase 2 of Section 3.1.2, two reaction systems were utilised during this investigation. Both of these were batch reactors, the ﬁrst being a commercially available stirred batch reactor and the second a custom designed and built batch micro-reactor. These units are described below with additional details (such as operating procedures and reactor development) provided in Appendices B.2 and B.3. The hydrogen:model compound ratio was varied in the micro-reactor system to examine the possibility of hydrogen starvation in batch operation. The ratios studied were compared to those calculated for an equivalent semi-batch system.

Batch Reactor Used for model compound screening in Phase 1, a typical reaction in this 250 cm3 stirred batch reactor began with the loading of 80 g of feed (model compound, diluent, CS2 and Mo octoate catalyst precursor). The system was purged with N2

(500 sccm) before being purged (900 sccm) and pressurised with H2 to 13.8 MPa. With mixing held at 700 RPM, the temperature was ramped to 445◦C and held for between 0 and 8 h reaction time. After cooling, gas, liquid and solid products were recovered for off-line analyses.

Description and Safety The system used in Phase 1 of the study was a stirred batch reactor supplied by the Parr Instrument Company (Parr). A schematic of this system is presented in Figure 3.4 with a photograph of how it was implemented in the laboratory in Figure 3.5 and additional details in Section B.2. This reactor system was designed to allow for operation in either semi-batch or batch mode. Details of features for semi-batch operation are provided in Section B.2

41 with only those components relating to batch operation, selected for this study as discussed above, being described here. The total internal volume of the 316 stainless steel reactor was 250 cm3, with liquid reagent loading masses and volumes, together with other pertinent operating parameters, being presented in Table 3.2. Either nitrogen (for purging air prior to heating) or hydrogen (for purging nitrogen and for pressurisation as a reagent) could be fed to the reactor, with the flow rates controlled by a Brooks Instrument 5850S mass flow controller. The reactor was heated by six 200 W heating rods positioned within the reactor walls, all operating in unison and monitored by three OMEGA Engineering Inc. (OMEGA) K-type thermocouples, two positioned inside the reactor (one as a backup to the other) and one measuring the wall temperature, to ensure accurate and even temperature distribution. Temperature monitoring and control was performed by the workstation. A tight-fitting ceramic fiber insulating jacket was secured over the reactor, itself covered by a layer of foil-backed fiberglass insulation, to further improve control and efficiency. A cooling water loop passed through the reactor to help speed cooling following a reaction. The gas outlet line was wrapped with OMEGA high temperature heating tape and covered with braided glass insulation. A Superior Electric variable transformer controlled the current going to the heating tape so as to maintain an exit line temperature of between 60 and 65◦C to minimise condensation during operation and shut down. The pressure of the reactor was monitored by both an analogue Ashcroft (welded, AISI 316 tube & socket) pressure gauge and an Ashcroft (A1906EP50) pressure transducer (for improved accuracy and data logging). Mixing of the reactor was achieved using a 316 stainless steel stirrer bar turned by a Parr magnetic drive monitored and controlled by the workstation. The workstation, running the CalGrafix software package (CAL Controls Ltd, v3.0.0), controlled the temperature and mixing speed whilst monitoring and recording the temperature, mixing speed and reactor pressure. The entire reactor system was housed in a vented plexiglass cage equipped with

Honeywell gas detectors (for H2 and H2S) to ensure that any gas leaks were quickly noted and safely contained and removed. Rubber sheeting covered the ﬂoor of the cage to contain spills and prevent slipping. All bulk chemicals, gas cylinders, control boxes and the operator workstation were located outside the cage. The

42 reactor was equipped with a pressure rupture disk (Fike Corporation) such that, in the event of over-pressurisation, excess gas would be safely discharged through a buffer vessel (to capture any entrained liquid) and vented. High temperature alarms, a back-up reactor thermocouple and the internal fail-safe mechanisms of the temperature controller helped avoid runaway temperatures.

Table 3.2: Operating conditions of the stirred batch reactor.

Condition Value Temperature (◦C) 415 - 445 Pressure (at temperature) (MPa) 13.79 Reaction time (h) 0 - 8 Mixer speed (RPM) 700 Catalyst loading 1 (ppm Mo) 0 600 1800 (g Mo octoate) 0 0.315 0.945 2 CS2 loading (g) 0.344 (cm3) 0.273 Dilution ratios (wt%) 3 100 Model compound loading (g) 2.7 80 (cm3) 2.7 80 Decalin loading (g) 77.7 - (cm3) 86.7 - 3 3 3 H2:model compound ratio (cm /cm ) 70.7 2.1 (g/g) 0.31 0.01 (mol/mol) 26.4 0.79 1 - Catalyst masses calculated per total mass rather than model compound mass. 2 - Constant (at 1800 ppm Mo level) regardless of catalyst loading to ensure consistent sulphur concentrations. 3 - Calculated based on reaction conditions with mass and molar ratios assuming DPM as model compound.

43 44

Figure 3.4: Process flow diagram of stirred slurry-phase batch hydroconversion reactor. 1 - Mass flow controller and mass flow meters connected to flow control box. 2 - Mixer connected via power control box to workstation. 3 - All six heating rods connected in parallel via distribution box to power control box. 4 - Pressure transducer connected to workstation. 5 - Double thermocouple connected via power control box (for high temperature auto-shutoff) to workstation. Wall thermocouple monitored separately. 6 - Heating tape connected to variable transformer. Figure 3.5: Stirred slurry-phase batch hydroconversion reaction system as implemented in laboratory (reactor unloaded and not in operating position).

45 Operation The operating procedure given below is a brief outline of one experimental run with full details provided in Section B.2.2. Beginning with a clean, dry, open reactor, the procedure was as follows:

1. Weigh model compound, decalin (diluent), CS2 and Mo octoate and load into reactor

2. Seal reactor

3. Purge with 500 sccm N2 for 1 min

4. Purge with 900 sccm H2 for 1 min

5. Pressurise to 13.8 MPa with H2

6. Heat system to desired reaction temperature

7. Maintain temperature for the desired reaction time

8. Shut off heating and allow system to cool

9. Depressurise system, collecting a gas sample if desired

10. Open the reactor and recover liquid and suspended solids

11. Clean reactor and internals with acetone

12. Analyse reaction products (see also Section 3.3)

(a) Gas samples may be analysed by GC as collected (b) Suspended solids may be recovered by vacuum ﬁltration, washed with acetone and dried before analysis (c) Liquid samples require dilution with decalin and internal standard addition prior to GCMS analysis

46 Micro-Reactor The micro-reactor used for Phase 3, reaction and deactivation mechanism studies, is described below. In a typical experiment, 150 - 400 µL of feed (DPM, CS2 and Mo octoate) were added to the glass insert (inner diameter of 4 mm and length of 250 mm) and lowered into the stainless steel reactor shell (inner diameter of 6 mm and length of 500 mm). The system was purged with N2 (pressurising to 700 kPa and venting, repeating this cycle three times) and H2 (same cycling) before pressurising to 13.8 MPa. Vortex mixing was begun at 2000 RPM and the temperature was ramped to 445◦C and held for 1 h reaction time. Gas analysis was done by in-line GC with liquid and solid products being recovered for off-line analyses.

Description and Safety The reactor system described below is the final product of a lengthy design and testing program fully detailed in Section B.3.1. Figure 3.6 shows a schematic of the micro-reactor system which was designed and constructed based on the results of experiments conducted in the batch reactor of Section 3.2.2 with a specific focus on improving operating parameter control and response and quantifying and/or mitigating the various factors found to have unwanted influences on the reaction (such as wall activity). Figure 3.7 shows how this design was implemented in the laboratory. The micro-reactor developed for this study comprised a removable glass insert (inner diameter of 4 mm, length of 250 mm, total volume of approximately 2.85 cm3) housed within a 316 stainless steel shell (inner diameter of 6 mm, length of 500 mm). The shell was positioned vertically, resting on a vortex mixer, within an 800 W Lindberg 55031 tubular furnace such that the reaction mixture within the insert was centered in the isothermal zone. The temperature of the reaction mixture was measured by a 1/16” OMEGA K-type thermocouple extending directly into the liquid and controlled by an OMEGA CN8201 temperature controller which was interfaced with a Dell Precision Workstation 690 PC (3.73 GHz, 36 GB RAM, Windows 8 64 bit) allowing control and logging through OMEGA CN8-SW Multi- Comm software package (v3.16.000). Note that the insert was positioned such that the mouth extended beyond the heated zone as this was essential to prevent loss of volatile or supercritical species from the insert during reaction.

47 The system was equipped to allow for the supply of either nitrogen (for purging air) or hydrogen (for pressurising for reaction), with needle valves being used to control the flows. The gas supply line attaching directly to the reactor head was “pig-tailed” to allow vibration from mixing to be absorbed without damage to the lines or fittings. Pressure in the system was continuously monitored by both an analogue ENFM USA Inc. pressure gauge and an OMEGA PX409-3.5KGUSB pressure transducer, the latter connected to the aforementioned computer and data logged using OMEGA TRH Control (v1.03.11.297). After a reaction, the system was depressurised through a series of valves and the sampling port of an in-line Shimadzu Scientific Instruments Inc. (Shimadzu) GC-14B gas chromatograph (see Section 3.3). A bubbler was used to visualise the depressurisation flow and keep it suitably low to prevent liquid carry-over from the reactor. Following depressurisation, the glass insert could be removed from the shell and the liquid and suspended solids recovered for analysis. One of the greatest challenges for a system of such a small size was how to achieve effective mixing. This was overcome through the use of a vortex mixer for an externally applied mixing effect. The reactor shell was attached to a custom mixer cup affixed to a Talboys 9456TAHDUSA advanced heavy-duty vortex mixer with the reactor head secured using spring restraints, which served to both absorb vibration and hold the reactor in position. In this manner the base of the steel shell assembly moved in a circular motion with the head as the pivot point. A series of experiments were conducted to verify the efficacy of this mixing technique as discussed in Section 4.2.3 using a Megaspeed MS70K high speed camera recording at 20,000 frames per second. Reaction product from a catalytic experiment (to ensure accurate fluid composition and particle sizes) was sealed within its glass insert using a cork and positioned inside a glass shell. This shell was the same diameter and length as its stainless steel counterpart, held in position about the same pivot point above the vortex mixer. Glass beads of the same diameter as in the stainless steel shell were used to position the insert at the same height. Visual mixing evaluations were performed both with and without the centrally located thermocouple. Despite the small size of this reaction system, compared to the stirred batch reactor described in Section 3.2.2, and hence the reduced amounts of liquid and

48 gas available to leak during any given run, a multitude of precautions were implemented to ensure safe operation. The assembly was installed inside a protective enclosure constructed of aluminium sheeting with a polycarbonate door. To reduce the risk of hydrogen ignition, following pressurisation, the hydrogen supply was shut off as part of the standard procedure, limiting the amount of this gas available to the system in the event of a leak. Furthermore, the enclosure was sufﬁciently large that even in the event of extraction failure and release of the hydrogen from the pressurised system, the air-hydrogen mix would still be below the lower explosive limit. The temperature control system was fail safe and equipped with two alarms with automatic shut-off. The pressure monitoring software also had an alarm to alert of over-pressurisation and a pressure release valve was installed (set to open at 17.24 MPa). To ensure stability of the entire system to mixer vibration, accidental jostling or seismic events, the enclosure and all equipment was secured to one another and/or to the counter. Table 3.3: Operating conditions and loadings utilised in micro-reactor.

Condition/Factor Value/Range Temperature (◦C) 415 - 445 Pressure (at temperature) (MPa) 13.79 Reaction time (h) 0 - 4 Mixer speed (RPM) 0 - 2500 Model compound loading 1 (mg) 151 402 (µL) 150 400 Catalyst loading (ppm Mo) 0 1800 0 1800 (mg Mo octoate) 0 1.77 0 4.73 2 CS2 loading (mg) 0.65 0.65 1.72 1.72 (µL) 0.51 0.51 1.36 1.36 3 3 3 H2:model compound ratio (cm /cm ) 19.9 6.9 (g/g) 0.09 0.03 (mol/mol) 7.70 2.65 1 - All micro-reactor experiments conducted using DPM as model compound. 2 - Constant (at 1800 ppm Mo level) regardless of catalyst loading to ensure consistent sulphur concentrations. 3 - Calculated based on reaction conditions and volume of glass insert only.

49 Figure 3.6: Process ﬂow diagram of batch slurry-phase hydroconversion micro-reactor. 1 - Thermocouple connected via OMEGA CN8201 controller to workstation for monitoring and control. 2 - Pressure transducer connected to workstation for logging.

50 Figure 3.7: Batch slurry-phase hydroconversion micro-reactor system as implemented in laboratory.

51 Operation Inherently similar in operation to the stirred batch reactor, use of the micro-reactor system had several distinct differences. A brief description of the procedure for a typical run is given below with full details provided in Section B.3.2. Beginning with a clean, dry reactor shell and glass insert, the procedure was as follows:

1. Weigh DPM, CS2 and Mo octoate and load into glass insert

2. Position insert within shell and seal the reactor

3. Position assembly on vortex mixer and secure using spring mounts

4. Purge with N2 (three 700 kPa - vent cycles)

5. Purge with H2 (three 700 kPa - vent cycles)

6. Pressurise to 13.8 MPa with H2

7. Heat system to 445◦C

8. Maintain temperature for the desired reaction time

9. Shut off heating and allow system to cool

10. Depressurise system, starting in-line GC analysis when system pressure drops to 70 kPa

11. Open the reactor, remove insert and liquid and suspended solids

12. Clean insert and thermocouple with acetone

13. Analyse reaction products (see also Section 3.3)

(a) Suspended solids may be recovered by settling or centrifugation to remove the liquid product followed by washing with acetone and drying (b) Liquid samples require dilution with decalin and internal standard addition prior to GCMS analysis

52 3.3 Analytical Equipment and Data Analysis The experimental results of this study were obtained from the analyses of the solid, liquid and gaseous reaction products and their comparison with the liquid feed mixture to determine such comparators as model compound conversion and product yields. Analysis of the solid phase (predominantly MoS2) was exclusively qualitative in this study (to identify which species were present, their relative concentrations and structures) whilst both gas and liquid phase analyses (performed by GC and GCMS) were both qualitative and quantitative. A brief description of the instruments utilised in these analyses is presented below. Additional details, operating procedures, examples of the data obtained from each, its interpretation and limitations, calibrations and sample calculations are provided in Appendix C. An full uncertainty and propagation analysis is provided in Section C.2.3.

3.3.1 Gas Product Analysis Although gas samples for both the stirred batch and micro-reactor systems were analysed by gas chromatography, the method of their collection and the instruments used differed due to the differences in the experimental setups. For the 250 cm3 stirred batch reactor, product gas samples were collected from the vent gas following a run using Alltech Tedlar® gas sampling bags. These samples were introduced into a Hewlett Packard (HP) 5890A GC, equipped with a Porapak® Q 80/100 mesh packed column and flame ionisation detector (FID), connected to a Hewlett Packard 3396 Series II integrator. The micro-reactor system was equipped with an in-line GC, simplifying the gas analysis by allowing for direct sampling of the product gases during reactor shutdown. The GC used in this setup was a Shimadzu GC-14B, equipped with an Agilent Technoligies Inc. (Agilent) HP-PLOT U column (19095P-UO4, ID 0.530 mm, length 30 m, film 20.00 µm) and an FID, connected to a Shimadzu C-R8A Chromatopac integrator Calibration of both GCs was performed using a certified hydrocarbon gas mixture (see Section 3.2.1 for gas mixture information).

53 3.3.2 Liquid Product Analysis Liquid products from both reactor systems were analysed on a Shimadzu GCMS- QP2010 gas chromatograph - mass spectrometer (GCMS) equipped with a Shi- madzu SHRXI-5MS column (220-94764-02, ID 0.25 mm, length 30 m, ﬁlm 0.25 mm) and AOC-20i autosampler (10 µL syringe). Due to the sensitivity of the system, it was necessary to dilute all liquid samples with decalin prior to analysis (to prevent overloading the column or saturating the detector). Each sample was prepared in two dilution ratios, a richer sample for analysis of lower concentration species (such as minor products) and a leaner ratio for higher concentration species (such as unreacted model compound). One of the model compounds not used in that particular reaction was added to each diluted sample to act as an internal standard. Numerous standard samples for each model compound, benzene and toluene were prepared and analysed for GCMS calibration.

3.3.3 Solid Product Analysis Given that no solid precipitation products were expected to form, the solid product analyses of this study served to: conﬁrm the formation of the MoS2 active phase from the liquid precursors, identify contaminants or unexpected solid products and determine the structure of the MoS2 and other solid particles. Solids recovery was by solid-liquid separation followed by washing with acetone and drying. Due to the nature of the formation of the solids (precipitating as nanometer-sized particles), size reduction for analysis was not necessary, with each solid sample simply needing to be appropriately mounted for analysis in each instrument. Compositional analysis of solid samples was performed by X-ray diffraction (XRD) and scanning electron microscopy with energy dispersive X-ray spectroscopy (SEM/EDX). A Bruker D8 Focus Bragg-Brentano was used for XRD analyses and a Hitachi S-2600N for the SEM/EDX analyses. Structural information relating to the recovered solids was obtained by transmission electron microscopy (TEM) and ﬁeld emission scanning electron microscopy (FESEM). TEM analysis was performed on an FEI Tecnai G2. FESEM was performed on a Hitachi S-4700.

54 Chapter 4

Experimental Results

This chapter presents the results of the experimental program shown in Table 3.1 (the stirred batch reactor experiments followed by the batch micro-reactor) together with brief descriptions of the results and indications of major trends and points of interest. All values reported are subject to an experimental uncertainty of ±4.7% (see Section C.2.3) unless otherwise indicated.

4.1 Stirred Batch Reactor

4.1.1 Model Compound Screening To determine which of the three model compounds selected for evaluation in this study was the most suitable for extensive testing, a series of screening experiments were conducted under industrially applicable conditions of temperature and pressure. Each model compound was to represent a low concentration species in the total feed and was hence diluted in decalin for these tests. The conversion results are presented in Table 4.1. An additional test was conducted using DPM wherein dilution was performed with benzene instead of decalin, the data obtained being used to determine the consequences of high model compound conversions producing a reaction mixture which would be supercritical under reaction conditions. Deﬁnitions and sample calculations for the conversion and yield comparators discussed are presented in Section C.2.3.

55 Conversion Both DPE and DPP were found to exhibit complete conversion, making them unsuitable for use in this study. DPM, however, presented conversions on a more mediocre level, making it more suitable. Changing the DPM diluent to benzene had only a marginal impact on the observed conversion.

Table 4.1: Conversion results obtained for diphenylmethane, diphenylethane and diphenylpropane screening experiments performed in the stirred ◦ batch reactor at 445 C, 13.8 MPa H2, 1 h, 600 ppm Mo, 700 RPM at 3 wt% in decalin.

Model compound Catalyst loading Conversion (ppm Mo) (wt%) DPM 0 32.6 600 35.7 DPM 1 600 39.0 DPE 0 99.6 600 99.6 DPP 0 99.6 600 98.5 1 - Benzene used as the solvent to study if supercritical phase has an inﬂuence on the reaction.

In an attempt to reduce DPP conversion to a level suitable for study, a series of experiments was conducted examining the thermal reaction at decreasing reaction temperatures. The results are presented in Figure 4.1. It may be seen that whilst reducing the temperature below approximately 430◦C does reduce the observed conversion, even temperatures as low as 420◦C result in only a 3 wt% decline in the conversion (within experimental uncertainty of the other values).

Product Distribution A summary of the major products from each of the screening experiments is provided in Table 4.2. The presence of the catalyst is seen to have a dramatic inﬂu- ence on the observed product distribution, shifting the reaction toward more hydrogenated and cracked products. It is clear from the DPP data in Table 4.2 that even though a reduction in the temperature does not have a signiﬁcant effect on the ob-

56 100

97 DPP conversion (wt%)

420 425 430 435 440 445

Reaction temperature (°C)

Figure 4.1: Conversion results obtained for diphenylpropane hydroconversion experiments performed in the stirred batch reactor at 420 - 445◦C, 13.8 MPa H2, 1 h, 600 ppm Mo, 700 RPM at 3 wt% in decalin. Curve is for illustration of trend only. served conversion, the catalyst appears to quickly lose activity as the temperature declines. This is evidenced by the product distribution shifting from the catalytically hydrogenated and cracked distribution at higher temperatures to consisting of only two species (from the primary cracking of DPP) at lower temperatures. Whilst this may suggest thermal cracking of DPP to be the rate-limiting step in this reaction, the catalyst only modifying the primary products, the excessively high conversion makes this claim impossible to substantiate without further study. These results also indicate that the catalyst acts predominantly in a hydrogenation role rather than to exact catalytic hydrogenolysis. The latter mechanism would allow the catalytic reactions to rapidly produce stable aromatic products, making the product more selective rather than the less selective distribution of hydrogenated, isomerised and cracked species observed. GC analyses determined that all of the screening experiments produced only minor quantities of gaseous products. The products from the 3 wt% dilution experiments could not be quantified due to their low concentrations. Table 4.3 presents the gas analyses from undiluted DPM experiments for quantification and identification of the gases formed. As may be seen, even after 6 h, very low concentrations

57 Table 4.2: Major products observed during model compound hydroconversion screening experiments of diphenylmethane (DPM), diphenylethane (DPE) and diphenylpropane (DPP) for different catalyst loadings and reaction temperatures in the stirred batch reactor at 13.8 MPa H2, 1 h, 700 RPM at 3 wt% in decalin.

Product Species Composition 1 Species Composition 1

Reactant (area%) (area%) 0 ppm Mo, 445◦C 600 ppm Mo, 445◦C Methylcyclohexane 39 Methylcyclohexane 24 Toluene 32 Toluene 19 1-Methylcyclohexene 29 Ethylcyclohexane 17

DPM 3-Methylheptane 16 Benzene 13 Cyclohexylmethylbenzene 10

0 ppm Mo, 445◦C 600 ppm Mo, 445◦C Toluene 43 Toluene 33 Methylcyclohexane 18 Methylcyclohexane 23 Ethylbenzene 12 Ethylbenzene 17 DPE 1,1-Diphenylethane 9 Ethylcyclohexane 10 1-Methylcyclohexene 9 1,1-Diphenylethane 9 Anthracene 9 Benzene 8

0 ppm Mo, 445◦C 600 ppm Mo, 445◦C Ethylbenzene 49 Ethylbenzene 67 Toluene 33 Toluene 18 Methylcyclohexane 10 Cyclohexylethylbenzene 15 1-Methylcyclohexene 8 DPP 600 ppm Mo, 430◦C 600 ppm Mo, 420◦C Ethylbenzene 78 Ethylbenzene 79 Toluene 22 Toluene 21 1 - Composition indicated is the percentage area from the GCMS chromatogram on a DPM-free basis (i.e. percentage of products formed) and limited to those species comprising >5%. of gaseous products were detected. For thermal reactions, little change was observed in the gas product composition with time. For catalytic systems, however, the gaseous products were observed to initially be suppressed to undetectable levels whilst after longer reaction times, larger species and higher concentrations were observed than for the equivalent thermal experiments. To conﬁrm that the majority of products remained in the liquid phase and

58 Table 4.3: Gaseous products observed during hydroconversion of undiluted ◦ diphenylmethane in the stirred batch reactor at 445 C, 13.8 MPa H2, 700 RPM.

Catalyst loading Reaction time Product composition (wt%) (ppm Mo) (h) Methane Ethane Propane 1 1.3 0.0 0.0 0 6 1.2 0.1 0.0 1 0.0 0.0 0.0 600 6 1.8 0.4 0.2 establish a quantiﬁcation for the mass balance, the loaded and recovered masses for eleven undiluted DPM experiments were compared and the mass balance was found to close to 98.9 ± 0.9 % (Section F.2).

4.1.2 Benzene, Toluene and Decalin Blanks To determine the stability of the major DPM decomposition products (benzene and toluene) and the diluent (decalin) under reaction conditions, blank tests of these species were conducted. Table 4.4 presents the feed and product compositions of the benzene and toluene tests together with their associated conversions. Whilst only trace amounts of gaseous products were detected for these tests, Tables 4.5 and 4.6 show the major liquid products found. Decalin blanks were observed to produce only trace gaseous and liquid products, with a conversion of approximately 3 wt% (within experimental uncertainty of zero). Given this value and that the feed was pure decalin, it was concluded that decalin is, alone, unreactive under the reaction conditions. The liquid products which were formed, quantiﬁed in Table 4.7, all eluted from the column in the 9.00 - 14.00 minute period. This is the portion of the chromatogram removed when analysing other products so as not to expose the GCMS ﬁlament and detector to high solvent concentrations. As such, the likelihood of confusion between decalin and model compound decomposition products was minimised. Of interest from this data is that both benzene and toluene show approximately the same conversion under the same reaction conditions and that the liquid products formed are not only the same species, but are produced in roughly the same amounts. It should be noted that the major products from benzene reaction are C7

59 and C8 species, as are those from toluene decomposition (with the exception of some benzene being formed). Products from decalin decomposition were seen to be cracking and dehydrogenation products.

Table 4.4: Benzene and toluene blank test conversions performed in the ◦ stirred batch reactor at 445 C, 13.8 MPa H2, 1 h, 600 ppm Mo, 700 RPM, 3 wt% in decalin.

Blank species Composition Conversion Feed Product (wt%) (wt%) (wt%) Benzene 3.3 2.9 11.0 Toluene 3.9 3.6 10.1

Table 4.5: Major liquid products detected for benzene blank tests performed ◦ in the stirred batch reactor at 445 C, 13.8 MPa H2, 1 h, 600 ppm Mo, 700 RPM, 3 wt% in decalin.

Species Composition 1 (area%) Methylcyclohexane 48 3-Methylheptane 36 Ethylcyclopentane 15 1 - Composition indicated is the percentage area from the GCMS chromatogram on a benzene-free basis (i.e. percentage of products formed) and limited to those species comprising >5%.

4.1.3 Diphenylmethane Studies With DPE and DPP found to be unsuitable for this study due to their excessively high conversions under reaction conditions, the remainder of this work focused on DPM.

Diluted Diphenylmethane Continuing from the DPM results obtained during the model compound screening experiments, with 0 and 600 ppm Mo and 3 wt% dilution in decalin, the effect of reaction time on the DPM hydroconversion reaction was examined, operating the system for between 0 and 8 h. The results are presented in Figure 4.2 for 0 and

60 Table 4.6: Major liquid products detected for toluene blank tests performed ◦ in the stirred batch reactor at 445 C, 13.8 MPa H2, 1 h, 600 ppm Mo, 700 RPM, 3 wt% in decalin.

Species Composition 1 (area%) Methylcyclohexane 41 3-Methylheptane 31 Benzene 19 Ethylcyclopentane 9 1 - Composition indicated is the percentage area from the GCMS chromatogram on a toluene-free basis (i.e. percentage of products formed) and limited to those species comprising >5%. Table 4.7: Major liquid products detected for decalin blank test performed in ◦ the stirred batch reactor at 445 C, 13.8 MPa H2, 1 h, 600 ppm Mo, 700 RPM.

Species Composition 1 (area%) Butylcyclohexane 41 1,2,3,4-tetrahydronaphthalene 2 36 1-Butylcyclohexene 23 1 - Composition indicated is the percentage area from the GCMS chromatogram on a decalin-free basis (i.e. percentage of products formed) and limited to those species comprising >5%. 2 - Commonly known as tetralin.

600 ppm Mo. There are two key features to note from this data. Firstly, even at 0 h (when the reactor is immediately cooled upon reaching reaction temperature), a DPM conversion of approximately 30 wt% is obtained. This is indicative of the reaction proceeding at temperatures below 445◦C, a fact shown by the DPP temperature dependency in Figure 4.1 and discussed for DPM below in the context of Figure 4.7. Secondly, there appears to be little difference between the conversions observed for 0 and 600 ppm Mo experiments upon reaching reaction temperature. After a short delay, however, the conversion for both Mo loadings begins the rise, the 600 ppm Mo system more rapidly than 0 ppm Mo. This results in a distinctly sigmoidal curve. It is the combination of this sigmoidal curvature and the non- zero conversion at 0 h which make ﬁrst and second order kinetic ﬁts unsuitable for

61 modeling these results. For completeness, these ﬁts are show in Figure E.1 with the kinetic constants provided in Table E.2. Langmuir-Hinshelwood-Hougen-Watson (LHHW) kinetics would also not result in the different trends (one sigmoidal, one not) observed. By such kinetics, the difference in the shape of the curves would indicate different kinetic expressions, indicative of different species adsorbing on the active sites. Both systems begin with the same concentration of reactants and Langmuir adsorption theory would thus predict similar surface coverage for both systems. As both the 0 and 600 ppm Mo systems possess active sites but in different numbers (for 0 ppm Mo only the FeS of the reactor walls and internals is present but with 600 ppm Mo that FeS is present in addition to the MoS2), were adsorption of the same species governing the LHHW expression, the 0 ppm Mo system would be expected to present with the same trend to that of the reaction with 600 ppm Mo, simply lower. Were LHHW kinetics applicable to this system, these results would suggest different species adsorbing on the active sites of FeS and MoS2, a point to be considered in later analyses.

100

25 DPM conversion (wt%)

0 2 4 6 8

Reaction time (h)

Figure 4.2: Conversion results obtained for diphenylmethane hydroconversion experiments performed in the stirred batch reactor at 0 - 600 ppm ◦ Mo, 445 C, 13.8 MPa H2, 0 - 8 h, 700 RPM at 3 wt% in decalin together with sigmoidal trend lines. Curves are for illustration of trends and are not kinetic ﬁts. ◦ 0 ppm Mo. 600 ppm Mo. As discussed in Section 2.5.2, the major products from the catalytic hydrocon-

62 version of DPM are expected to be benzene and toluene. This is particularly true if the primary role of the catalyst were to perform catalytic hydrogenolysis which, in DPM hydroconversion, has been shown [89] to result in very selective equimolar yields of benzene and toluene. As shown in Section 4.1.1, however, this was not always the case for this series of experiments. Numerous other cracking, isomerisation and condensation species were observed to form during the reaction. With the gas selectivity shown to be low (see Table 4.3) and with the additional liquid product species changing from reaction to reaction, the mechanisms for their formation being unclear, product analyses for these experiments were limited to benzene and toluene yields, presented in Figures 4.3 and 4.4 respectively. As may be seen, the 600 ppm Mo experiments begin with low yields of both benzene and toluene with these increasing steadily with an increase in conversion, the toluene yield exceeding that of benzene for all experiments, a trend contrary to both catalytic hydrogenolysis and the mechanism of Curran et al. [78]. The 0 ppm Mo experiments show higher levels of benzene and toluene than 600 ppm Mo, with these levels also increasing with conversion (following what appears to be an initial delay). Once more, toluene yield exceeds that of benzene. To better visualise the benzene:toluene (B:T) molar ratio, Figure 4.5 is provided. It is clear that under all conditions evaluated in this series, the toluene yield exceeds the benzene yield, increasing rapidly upon reaching reaction temperature before appearing to level out.

Undiluted Diphenylmethane In an attempt to isolate the cause of the sigmoidal results and unexpected liquid product species, the diluent (which may have served a role as a hydrogen shuttle) was removed from the system and catalyst concentrations up to 1800 ppm Mo were evaluated (in an attempt to overcome catalytic wall effects). The results from these undiluted experiments, which included a temperature dependency study for DPM mimicking that of DPP from the screening studies, are presented below. The conversion of undiluted DPM with reaction time is shown in Figure 4.6, with this data showing a strongly sigmoidal shape. Whilst the conversion at 0 h is lower than for the equivalent diluted experiments (approximately 3 wt% as

63 1.0 )

0.8 DPMreacted

0.6

0.4

0.2 Benzene yieldmolar (mol/mol

0.0

30 40 50 60 70 80 90

DPM conversion (wt%)

Figure 4.3: Benzene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the stirred batch reactor at 0 - 600 ◦ ppm Mo, 445 C, 13.8 MPa H2, 0 - 8 h, 700 RPM at 3 wt% in decalin. Curves are for illustration of trends only. ◦ 0 ppm Mo. 600 ppm Mo.

1.0 )

0.8 DPMreacted

0.6

0.4

0.2 Toluene molar yieldmolar Toluene (mol/mol

0.0

30 40 50 60 70 80 90

DPM conversion (wt%)

Figure 4.4: Toluene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the stirred batch reactor at 0 - 600 ◦ ppm Mo, 445 C, 13.8 MPa H2, 0 - 8 h, 700 RPM at 3 wt% in decalin. Curves are for illustration of trends only. ◦ 0 ppm Mo. 600 ppm Mo.

64 1.0

0.8

0.6

0.4

0.2 Benzene:toluene ratiomolar (mol:mol)

0.0

30 40 50 60 70 80 90

DPM conversion (wt%)

Figure 4.5: Benzene:toluene molar ratio obtained for diphenylmethane hydroconversion experiments performed in the stirred batch reactor at 0 ◦ - 600 ppm Mo, 445 C, 13.8 MPa H2, 0 - 8 h, 700 RPM at 3 wt% in decalin. Curves are for illustration of trends only. ◦ 0 ppm Mo. 600 ppm Mo.

65 compared to 30 wt%), it is still not zero, indicating that reactions occur before the system reaches the reaction temperature. The conversion after extended reaction times is also lower in the undiluted system, but seemingly only for the 600 ppm Mo experiments (approximately 40 wt% after 6 h with no dilution as compared to approximately 80 wt% after 6 h for the diluted system). It should also be noted that the 0 ppm Mo experiments, following an initial delay, show a rapid rise in DPM conversion to a level within experimental uncertainty of the 600 ppm Mo tests within the 6 h reaction time. As was the case for the diluted experiments, neither first nor second order kinetics provided satisfactory fits for the 600 ppm Mo results, but a first order equation was found to be a reasonable approximation for the 0 ppm Mo system (see Figure E.2 and Table E.7). Once more the change in the shape of the curve indicates that, if LHHW kinetics were applied, the equation would be different due to adsorption of different species on the FeS and MoS2 active sites, making comparison using such kinetics of little value. In the undiluted systems the conversion for the 600 ppm Mo series is observed to initially exceed that of 0 ppm Mo with the latter attaining the same levels after 6 h. Increasing catalyst concentration to 1800 ppm Mo was observed to increase the conversion whilst reducing the temperature (shown in Figure 4.7) resulted in a more rapid decline in conversion than was found for DPP (see Figure 4.1). Figure 4.7 may be interpreted in the context of the Arrhenius Law (as shown in Section F.3.3) to obtain Figure 4.8. From this plot the activation energy, Ea, for DPM hydroconversion with 1800 ppm Mo is found to be 154 ± 3 kJ/mol. Samples exemplifying the major liquid products observed in these reactions are presented in Table 4.8 for the different reaction times and temperatures with the structures of several of the species with less intuitive IUPAC names provided in Table 4.9 together with acronyms used in this study for clarity. Unlike the many products observed in the diluted model compound experiments (shown in Table 4.2), the undiluted experiments show fewer major species, many of which can be seen to contain either DPM or some hydrogenated form of it as a structural component. There remain, however, numerous additional species comprising fractions of a percent of the product which are not shown and yet which combine to form an appreciable portion of the total. Despite this progress in simplifying the reaction, with the major products in all

66 50

10 DPM conversion (wt%)

0 2 4 6

Reaction time (h)

Figure 4.6: Conversion results obtained for undiluted diphenylmethane hydroconversion experiments performed in the stirred batch reactor at 0 - ◦ 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 6 h, 700 RPM with sigmoidal trend lines. Error bars indicate standard deviation. Curves are for illustration of trends and are not kinetic ﬁts. ◦ 0 ppm Mo. 600 ppm Mo. △ 1800 ppm Mo.

10 DPM conversion (wt%)

415 430 445

Reaction temperature (°C)

Figure 4.7: Conversion results obtained for undiluted diphenylmethane hydroconversion experiments performed in the stirred batch reactor with ◦ 1800 ppm Mo, 415 - 445 C, 13.8 MPa H2, 1 h, 700 RPM.

67 -1.0

-1.2

-1.4

-1.6 ln(-ln(1-X))

-1.8

-2.0

-2.2

1.38 1.40 1.42 1.44 1.46

-1 3

1/T (K ) x10

Figure 4.8: Logarithmic conversion results obtained for undiluted diphenylmethane hydroconversion experiments performed in the stirred batch ◦ reactor with 1800 ppm Mo, 415 - 445 C, 13.8 MPa H2, 1 h, 700 RPM against inverse reaction temperatures for determination of the activation energy. Curve indicates ﬁt by linear regression.

68 Table 4.8: Major products observed during undiluted diphenylmethane hydroconversion experiments performed in the ◦ stirred batch reactor with 0 - 1800 ppm Mo, 415 - 445 C, 13.8 MPa H2, 0 - 6 h, 700 RPM.

Species Composition 1 Species Composition 1 Species Composition 1 (area%) (area%) (area%) 0 h, 445◦C 1 h, 445◦C 6 h, 445◦C Toluene 51 Benzene 47 Benzene 43 Benzene 49 Toluene 40 Toluene 35 Fluorene 5 MBP1,2,3 11 ETB 4 ETB 7 0 ppm Mo MBP2 3 Fluorene 2 o-Xylene 1

◦ ◦ ◦ 69 0 h, 445 C 1 h, 445 C 6 h, 445 C Toluene 50 Benzene 24 Benzene 41 Benzene 50 Toluene 19 Toluene 36 MBP2,3 17 ETB 11 HexF 16 MBP1,3 10 600 ppm Mo Fluorene 13 Fluorene 2

1 h, 415◦C 1 h, 430◦C 1 h, 445◦C Toluene 41 Toluene 47 Toluene 49 Benzene 37 Benzene 42 Benzene 45 CHMB 16 CHMB 9 CHMB 4 3-Methylheptane 3 3-Methylheptane 2 HexF 1 1800 ppm Mo HexF 3 3-Methylheptane 1 1 - Composition indicated is the percentage area from the GCMS chromatogram on a DPM-free basis (i.e. percentage of products formed) and limited to those species comprising >1%. Table 4.9: IUPAC names, structures and acronyms of several species observed in the diphenylmethane hydroconversion liquid products.

IUPAC name Acronym Structure

2-Methyl-1,1’-biphenyl MBP1

1-Methyl-3-(phenylmethyl)-benzene MBP2

1-Methyl-4-(phenylmethyl)-benzene MBP3

1-Ethyl-2-(1-phenylethyl)-benzene EPB

1,1’,1”-(1-Ethanyl-2-ylidene)tris-benzene ETB

4-Benzylbiphenyl BBP

2-Methyl-1,1,1-triphenyl-propane MTP

1,2,3,4,4a,9a-Hexahydroﬂuorene HexF

Cyclohexylmethylbenzene CHMB

70 tests being the benzene and toluene expected, some of the products appearing under one set of conditions are absent under another. To graphically examine the yields of these studies, it was thus necessary to group many of the lesser species to create lumps, one representing the cracking species (those with fewer carbons than DPM, excluding benzene and toluene) and another the isomerisation and condensation (isom. and cond.) species (those with carbon numbers equal to or higher than DPM, excluding cyclohexylmethylbenzene). One species of particular interest, and its yield hence remaining outside of a lump, is cyclohexylmethylbenzene (CHMB) which was observed as a major product under 1800 ppm Mo but as only a minor product for the other catalyst loadings. The molar yields for benzene, toluene and CHMB are presented in Figures 4.9, 4.10 and 4.11 respectively, with the mass yields for the lumped cracking and isomerisation/condensation products in Figures 4.12 and 4.13 respectively. From this data it may be seen that the benzene and toluene yields initially increase with conversion, appear to pass through a maximum and then begin to decrease again. There is significant scatter in the results but it is noted that increasing the catalyst concentration appears to reduce the benzene yield whilst increasing the toluene yield. This is better seen in Figures 4.14 and 4.15 wherein the trend of decreasing B:T ratio with increasing catalyst loading is clear. It is also evident that increasing the catalyst loading increases the yield of CHMB, suggesting this to be a catalytic product. This trend, combined with the results in Table 4.8, indicate the primary function of the catalyst to be hydrogenation rather than hydrogenolysis. Whilst benzene and toluene yields exhibit a maxima with increasing DPM conversion, the yields of the cracking and isom./cond. products steadily decline. It appears that the yield of cracking species may increase with increasing catalyst concentration but the extremely low yield values make definitive trends difficult to substantiate. The formation of isom./cond. products, however, represents major reaction pathways (which was clear from Table 4.8). For this lump it may be seen that while 0 and 600 ppm Mo show roughly the same yields, increasing the catalyst loading to 1800 ppm Mo clearly suppresses their formation. An additional benefit to undiluted model compound studies on this scale was that a measurable pressure change could be recorded during the reaction as shown in Figure 4.16. Such a significant change was not observed during diluted experi-

71 1.0 )

0.8 DPMreacted

0.6

0.4

0.2 Benzene yieldmolar (mol/mol

0.0

0 10 20 30 40 50

DPM conversion (wt%)

Figure 4.9: Benzene molar yield results obtained for undiluted diphenylmethane hydroconversion experiments performed in the stirred batch ◦ reactor at 0 - 1800 ppm Mo, 415 - 445 C, 13.8 MPa H2, 0 - 6 h, 700 RPM. Curve is for illustration of trend only. ◦ - 0 ppm Mo. - 600 ppm Mo. △ - 1800 ppm Mo.

1.0 )

0.8 DPMreacted

0.6

0.4

0.2 Toluene molar yieldmolar Toluene (mol/mol

0.0

0 10 20 30 40 50

DPM conversion (wt%)

Figure 4.10: Toluene molar yield results obtained for undiluted diphenylmethane hydroconversion experiments performed in the stirred batch ◦ reactor at 0 - 1800 ppm Mo, 415 - 445 C, 13.8 MPa H2, 0 - 6 h, 700 RPM. Curve is for illustration of trend only. - 0 ppm Mo. - 600 ppm Mo. △ - 1800 ppm Mo.

72 0.2 ) DPMreacted

0.1 CHMB yield (mol/mol

0.0

0 10 20 30 40 50

DPM conversion (wt%)

Figure 4.11: Cyclohexylmethylbenzene molar yield results obtained for undiluted diphenylmethane hydroconversion experiments performed in the stirred batch reactor at 0 - 1800 ppm Mo, 415 - 445◦C, 13.8 MPa H2, 0 - 6 h, 700 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. 600 ppm Mo. △ 1800 ppm Mo.

0.06

0.04 ) DPMreacted

0.02 (g/g Other cracking products mass yield

0.00

0 10 20 30 40 50

DPM conversion (wt%)

Figure 4.12: Mass yield of other cracking products (lumped) obtained for undiluted diphenylmethane hydroconversion experiments performed in the stirred batch reactor at 0 - 1800 ppm Mo, 415 - 445◦C, 13.8 MPa H2, 0 - 6 h, 700 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. 600 ppm Mo. △ 1800 ppm Mo.

73 0.4

0.3 ) DPM reacted DPM

0.2

0.1 mass yield (g/gmass yield Other isom. cond. and products 0.0

0 10 20 30 40 50

DPM conversion (wt%) Figure 4.13: Mass yield of isomerisation and condensation products (lumped) obtained for undiluted diphenylmethane hydroconversion experiments performed in the stirred batch reactor at 0 - 1800 ppm Mo, ◦ 415 - 445 C, 13.8 MPa H2, 0 - 6 h, 700 RPM. Curves are for illustration of trends only. ◦ - 0 ppm Mo. - 600 ppm Mo. Trend line for both 0 and 600 ppm Mo. △ 1800 ppm Mo.

1.25

1.00

0.75

0.50

0.25 Benzene:toluene ratiomolar (mol:mol)

0.00

0 10 20 30 40 50

DPM conversion (wt%)

Figure 4.14: Benzene:toluene molar ratio obtained for undiluted diphenylmethane hydroconversion experiments performed in the stirred batch ◦ reactor at 0 - 1800 ppm Mo, 415- 445 C, 13.8 MPa H2, 0 - 6 h, 700 RPM. Curve is for illustration of trend only. ◦ 0 ppm Mo. 600 ppm Mo. △ 1800 ppm Mo.

74 1.5

1.0

0.5 Benzene:toluene ratiomolar (mol:mol)

0.0

0 600 1200 1800

Catalyst loading (ppm Mo)

Figure 4.15: Benzene:toluene molar ratio obtained for undiluted diphenylmethane hydroconversion experiments performed in the stirred batch ◦ reactor at 0 - 1800 ppm Mo, 415- 445 C, 13.8 MPa H2, 0 - 6 h, 700 RPM compared with catalyst loading. Error bars indicate standard deviation. Curve is for illustration of trend only.

75 ments, presumably as a smaller quantity of model compound resulted in minimal gaseous exchange and undetectable pressure variations. From this data it may be seen that the pressure drop (a combination of the decline as hydrogen is consumed and the rise as gaseous products, shown to be minimal, are formed) has an inverse sigmoidal trend. After an initial delay, both 0 and 600 ppm Mo experiments exhibit a decline in the pressure (corresponding to a rapid uptake of hydrogen) followed by a leveling off. The 1800 ppm Mo data shows a greater pressure drop (indicative of more hydrogen being consumed and/or less gaseous products being formed) than the 0 and 600 ppm Mo experiments, but mimics the trend over the given range of conversions.

0.0

-0.5

-1.0

-1.5 Pressure change (MPa)

-2.0

0 10 20 30 40 50

DPM conversion (wt%)

Figure 4.16: Pressure change observed for undiluted diphenylmethane hydroconversion experiments performed in the stirred batch reactor at 0 ◦ - 1800 ppm Mo, 415- 445 C, 13.8 MPa H2, 0 - 6 h, 700 RPM. Curves are for illustration of trends only. ◦ - 0 ppm Mo. - 600 ppm Mo. Trend line for both 0 and 600 ppm Mo. △ 1800 ppm Mo. XRD and TEM analyses were performed on the solid material recovered from the undiluted DPM experiments. The objectives of these tests were to conﬁrm the formation of the desired MoS2 from the Mo octoate precursor and CS2, determine the dimensions (sheet size and stack height) of the MoS2 crystallites and to identify any other solid species formed during the reaction (such as FeS which would form on the walls of the reactor and slough off into the liquid).

76 Figures 4.17 and 4.18 show the XRD results for solids from 600 ppm and 1800 ppm Mo experiments respectively. As may be seen, the major crystalline species present in both samples is MoS2 with the 1800 ppm Mo sample showing a small amount of graphite and the 600 ppm Mo sample showing trace FeS. The apparent disappearance of FeS in the 1800 ppm Mo samples is believed to be due primarily to dilution in the recovered solid. The solid samples analysed were recovered from multiple experiments, conducted in a randomised order, and thoroughly mixed. It is proposed that the amount of FeS is the same in both samples but the lower amount of total MoS2 in the 600 ppm Mo experiment means the same mass of FeS from the reactor walls represents a higher concentration in this sample than the 1800 ppm Mo one and thus presents more prominent peaks. Using the Scherrer equation (the equation and associated peak analysis may be found in Section C.3.2) with the integral breadth broadening factor of the 002 basal peak, the average stack height of the 600 ppm Mo sample was calculated to be 2.31 nm, corresponding to approximately 3.75 sheets. This indicates that the majority of the crystallites were comprised of between three and four sheets. For the 1800 ppm Mo sample, the stack height was calculated to be 2.01 nm, approximately 3.26 sheets. Whilst this 15% difference may indicate slightly higher stacks in the case of 600 ppm Mo, this may simply be experimental variation given the numerous steps associated with the recovery, preparation and analysis of solid samples by XRD. TEM analyses for the same 600 ppm Mo and 1800 ppm Mo samples were conducted (presented in Figures 4.19 and 4.20 respectively) to examine the structure of the MoS2 crystallites, obtain a comparative value for the stack height and determine the sheet width. Figure 4.19 further indicates the inter-plate and d-spacing of the crystallites at 0.62 and 0.27 nm respectively. These compare well with the literature values for MoS2 of 0.6155 and 0.2738 nm respectively [106]. Both solid samples assumed the same “rag” structure of sheets and showed almost identical sheet width and stack height distributions (presented in Figures 4.21 and 4.22, respectively, for the 600 ppm Mo sample). From these analyses, the most common sheet width was observed to be 3 - 4 nm, with a most abundant stack height of two sheets.

77 5000

MoS {002}

4000

MoS

3000 78

2000 Counts(-)

1000

FeS

20 40 60 80

Angle, 2 (°)

Figure 4.17: X-ray diffractogram for solid material obtained from undiluted diphenylmethane hydroconversion exper- ◦ iments performed in the stirred batch reactor at 600 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 700 RPM. 4000

MoS {002}

MoS

3000

2000 79 Counts(-)

1000

Carbon (C, graphite)

20 40 60 80

Angle, 2 (°)

Figure 4.18: X-ray diffractogram for solid material obtained from undiluted diphenylmethane hydroconversion exper- ◦ iments performed in the stirred batch reactor at 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 700 RPM. 80

Figure 4.19: Transmission electron microscopy image for solid material obtained from a diphenylmethane test in the ◦ stirred batch reactor at 600 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 700 RPM showing the both the inter-plate spacing (d002 = 0.62 nm) and the d-spacing (d100 = 0.27 nm). Figure 4.20: Transmission electron microscopy image for solid material obtained from a undiluted diphenylmethane test in the stirred batch reac- ◦ tor at 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 700 RPM.

81 100

25 Number Number of sheets (counts)

0 1 2 3 4 5 6 7 8 9 1011121314

Sheet length (nm)

Figure 4.21: Sheet size distribution of MoS2 crystallites using transmission electron microscopy data for solid material obtained from a undiluted diphenylmethane test in the stirred batch reactor at 600 ppm Mo, ◦ 445 C, 13.8 MPa H2, 1 h, 700 RPM.

20 Number Number of stacks (counts)

1 2 3 4 5 6 7 8 9

Number of sheets in stack (counts)

Figure 4.22: Stack height distribution of MoS2 crystallites using transmission electron microscopy data for solid material obtained from a undiluted diphenylmethane test in the stirred batch reactor at 600 ppm Mo, ◦ 445 C, 13.8 MPa H2, 1 h, 700 RPM.

82 4.2 Batch Micro-reactor With the larger stirred batch reactor yielding unsatisfactory results (specifically in terms of non-zero conversions at 0 h, appreciable experimental uncertainties and variable product species), a micro-reactor system was designed to mitigate the effects of wall catalysis and slow heat-up rates, the two factors thought to contribute the most to experimental variations. The development of the micro-reactor is presented in detail in Section B.3.1 and progressed through three designs: an inclined stainless steel, a vertical stainless steel and the final glass insert system. The first two were constructed of seamless stainless steel tubing (3 mm inside diameter) with a central thermocouple and no mixing. The third used a glass insert (4 mm inside diameter) with a central thermocouple, positioned within a stainless steel shell (6 mm inside diameter) and agitated with a vortex mixer. All three micro-reactors were heated in a tubular furnace. DPM was used as the model compound with liquid loadings of 150 - 400 µL and catalyst concentrations of 0 - 1800 ppm Mo. The key comparators for these studies were the observed DPM conversion, benzene, toluene, CHMB, cracking and isom./cond. product yields (as discussed in Section 4.1.3 and defined in Section C.2.3) and the B:T molar ratio. Due to the small volume of model compound as compared to the total gas volume of the system, the pressure variations during reaction were not a reliable comparator. Whilst quantitative gas analysis by in-line GC showed only trace amount of gaseous hydrocarbon products, the changing qualitative “fingerprint” of these products was found to provide some insight into the reaction (this was only performed for the vertical stainless steel and glass insert systems).

4.2.1 Inclined Stainless Steel Micro-Reactor The ﬁrst functional incarnation of the micro-reactor was an inclined stainless steel unit. As shown in Table B.3, this system was used to study the effects of multiple factors including: wall activation, hydrogen:DPM ratio (with and without a catalyst), heat-up rate, catalyst loading and reaction time. A series of 21 experiments conducted with 1800 ppm Mo to study the activation of a fresh stainless steel wall. The results from these tests are presented in Figure 4.23. For multiple 1 h experiments, the DPM conversion was observed to

83 rise rapidly and then slow. To promote stabilisation of the reactor walls, experiments with longer reaction times were conducted. Wall activity was only observed to plateau after a cumulative reaction time of approximately 48 h. To ensure complete activation, the system was then left, loaded with liquid over a period of 72 h (one weekend). Subsequent 1800 ppm Mo tests confirmed that DPM conversion had stabilised but at a far higher level as seen in Figure 4.24. The significant influence of the wall activity was concerning.

4 h

3 h

1 h

2 h

5 DPM conversion (wt%)

0 5 10 15 20

Sequential experiment number (-)

Figure 4.23: Conversion results obtained for diphenylmethane hydroconversion experiments performed in the inclined micro-reactor at 1800 ppm ◦ Mo, 445 C, 13.8 MPa H2, 1 - 4 h reaction time (as indicated), 0 RPM to study wall activation. N - 1800 ppm Mo activation study. Recalling that the hydrogen:DPM ratio was a point of concern in the stirred batch system as per Table 3.2, a preliminary examination into this effect in an 1800 ppm Mo catalytic reaction was examined by increasing the DPM volume from 400 to 500 µL. All comparators were found to be within experimental uncertainty and, as such, the hydrogen:DPM ratio for this system is sufﬁcient so as not to interfere with the reaction mechanism (for instance through hydrogen starvation). More intensive experiments were conducted when the glass insert was introduced. To determine the inﬂuence of the faster heat-up rate (the micro-reactor able to achieve reaction temperature in 20 min as compared to 80 min for the stirred batch system), the results from rapid heat-up experiments were compared, in Table 4.10,

84 40

10 DPM conversion (wt%)

0 1 2 3 4

Reaction time (h)

Figure 4.24: Conversion results obtained for diphenylmethane hydroconversion experiments performed in the inclined micro-reactor at 1800 ppm ◦ Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM to compare the results obtained for an active and inactive wall. Error bars indicate standard deviation. △ - 1800 ppm Mo with active wall. N - 1800 ppm Mo activation study. with those where the rate was set to mimic that of the stirred batch reactor (note that these were obtained during the stabilising period of reactor wall activation). Whilst many of the values lie within experimental uncertainty, the faster ramp rate does appear to result in lower conversion, less benzene (and hence a reduced B:T ratio) and less CHMB. These results are believed to be a more accurate representation of the reaction occurring at 445◦C rather than in the tail end of the heat-up as was observed in the stirred batch reactor. Further corroboration of this is seen in Figure 4.25 (and subsequent datasets) where DPM conversion at 0 h is negligible. So as to compare the performance of this micro-reactor with that of the stirred batch reactor, to determine the inﬂuence of the catalyst and the kinetics of the reaction, a series of experiments was conducted varying the catalyst concentration and the reaction time. The DPM conversions, products yields and B:T ratios are presented below. It is noted that there is still a fairly large scatter associated with these results, disappointing with improved reproducibility being one of the aims of switching to a micro-reactor. DPM conversion data in Figure 4.25 shows the conversion at 0 h to be the

85 Table 4.10: Comparison of results for diphenylmethane hydroconversion obtained in the inclined micro-reactor at 1800 ppm Mo, 445◦C, 13.8 MPa H2, 1 h, 0 RPM to examine the effect of heating rate.

Ramp rate DPM conversion Product yield ◦ ( C/min) (wt%) (mol/molDPM reacted) Benzene Toluene CHMB 5.5 8.6 0.68 0.73 0.19 21.5 7.5 ± 1.1 0.63 ± 0.10 0.74 ± 0.10 0.05 ± 0.01

Product yield B:T ratio (g/gDPM reacted) (mol/mol) Cracking Isom./cond. 0.05 0.05 0.94 0.05 ± 0.02 0.07 ± 0.02 0.85 ± 0.06 desired 0 wt%, indicating that minimal reaction occurs during the rapid heat-up period. The curves shown are merely trend lines as ﬁrst and second order kinetic ﬁtting was not possible giving the complex shape of the data. The close correlations of these curves suggests that LHHW kinetics do not apply either for, unlike the difference in the trends observed in the stirred batch reactor (Figures 4.2 and 4.6) and attributed to different kinetic expressions resulting from different species adsorbing on the FeS and MoS2 active sites, the similarity in the trends of Figure 4.25 suggest comparable mechanisms at play (i.e. the same species adsorbing on both FeS and MoS2 active sites). Following an initial, rapid rise in conversion, the catalytic systems pass through a curious stabilisation period around 2 h before continuing to rise. This is in contrast with the clearly sigmoidal shaped curves from the stirred batch reactor. It is noted that whilst the conversion for 600 ppm Mo is comparable to that from the stirred batch reactor, neither 0 nor 1800 ppm Mo share this characteristic. 0 ppm Mo exhibits a smoothly increasing curve with a higher conversion than previously seen whilst 1800 ppm Mo is observed to have a DPM conversion lower than that of 600 ppm Mo below approximately 3 h reaction time. To examine the changes in the reaction mechanism that result in these trends, the products yields must be studied. From the benzene, toluene, CHMB, cracking and isom./cond. product yields in Figures 4.26 through 4.30, it may clearly be seen that the trends differ from those of the stirred batch reactor. Whilst all catalyst loadings show roughly the same benzene yield, 1800 ppm Mo shows a slightly suppressed toluene yield. The benzene:toluene ratio (shown in Figure 4.31) for

86 40

10 DPM conversion (wt%)

0 1 2 3 4

Reaction time (h)

Figure 4.25: Conversion results obtained for diphenylmethane hydroconversion experiments performed in the inclined micro-reactor at 0 - 1800 ◦ ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Error bars indicate standard deviation. Curve is for illustration of trend and is not a kinetic ﬁt. ◦ - 0 ppm Mo. - 600 ppm Mo. △ - 1800 ppm Mo. all three remains, however, very close to unity. The CHMB yield may be seen to increase with increasing catalyst loading, mimicking the trend seen in the stirred reactor but at higher levels. The yield of other isom./cond. products declines with increasing catalyst concentration as was noted in the stirred reactor, shadowing the trends over the conversion range although at elevated levels for 0 and 600 ppm Mo. Unlike the stirred system, however, the trends for cracking products show a distinct divergence of the 1800 ppm Mo system, producing more cracked species at higher conversion than either the 0 or 600 ppm Mo experiments, the latter two showing similar results to one another. Given the similarities between the results observed in both the stirred batch reactor and inclined micro-reactor between 0 and 600 ppm Mo, the 600 ppm Mo catalyst concentration was omitted from subsequent experiments. To help determine what mechanistic changes result in the observed differences in conversion and yield between the inclined micro-reactor and stirred batch systems, Table 4.11 shows the composition of several key liquid products. These samples cover both the catalyst loading and reaction time ranges. It is noted that the

87 )

1.0

DPM reacted DPM 0.8

0.6

0.4

0.2 Benzene molar yield (mol/mol molar yield Benzene 0.0

10 20 30 40

DPM conversion (wt%)

Figure 4.26: Benzene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the inclined micro-reactor at ◦ 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. 600 ppm Mo. △ 1800 ppm Mo. reaction product is now clean enough that all products above 0.25 area% may be shown without excessive complexity. It seems apparent that whilst the 0 ppm Mo and, to a lesser extent, 600 ppm Mo systems exhibit some degree of random reaction resulting in the formation of numerous side products, the 1800 ppm Mo reaction has access to a more streamlined pathway, producing only ﬁve products in roughly the same quantities regardless of DPM conversion.

4.2.2 Vertical Stainless Steel Micro-Reactor In preparation for the implementation of the glass insert micro-reactor, the inclined stainless steel unit was oriented vertically and the catalyst concentration (limited to 0 and 1800 ppm Mo) and reaction time experiments repeated. The change in orientation would reduce the gas-liquid interface area as well as the contact between the liquid and the active reactor and thermocouple walls (the thermocouple sheath itself is 316 stainless steel and subject to the same activation as the reactor), affecting the observed conversion and product yields. Figure 4.32 shows the DPM conversion for the series. The curves shown are

88 ) 1.0

0.8 DPM reacted DPM

0.6

0.4

0.2 Toluene molar yield (mol/mol molar yield Toluene

0.0

10 20 30 40

DPM conversion (wt%)

Figure 4.27: Toluene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the inclined micro-reactor at ◦ 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. 600 ppm Mo. △ 1800 ppm Mo.

0.2 ) DPM reacted DPM

0.1 CHMB molar yield (mol/mol molar yield CHMB

0.0

10 20 30 40

DPM conversion (wt%)

Figure 4.28: Cyclohexylmethylbenzene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the in- ◦ clined micro-reactor at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2,0- 4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. 600 ppm Mo. △ 1800 ppm Mo.

89 0.06

0.04 ) DPMreacted

(g/g 0.02 Other cracking products mass yield

0.00

10 20 30 40

DPM conversion (wt%)

Figure 4.29: Mass yield of other cracking products (lumped) obtained for diphenylmethane hydroconversion experiments performed in the in- ◦ clined micro-reactor at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2,0-4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. 600 ppm Mo. △ 1800 ppm Mo.

0.4

0.3 ) DPMreacted

0.2

0.1 mass yield (g/g Other isom. and cond. products

0.0

10 20 30 40

DPM conversion (wt%)

Figure 4.30: Mass yield of isomerisation and condensation products (lumped) obtained for diphenylmethane hydroconversion experiments performed in the inclined micro-reactor at 0 - 1800 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. 600 ppm Mo. △ 1800 ppm Mo.

90 1.25

1.00

0.75

0.50

0.25

0.00 Benzene:toluene molar ratio (mol:mol)Benzene:toluene

10 20 30 40

DPM conversion (wt%)

Figure 4.31: Benzene:toluene molar ratio obtained for diphenylmethane hydroconversion experiments performed in the inclined micro-reactor at ◦ 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. 600 ppm Mo. △ 1800 ppm Mo.

91 Table 4.11: Major products observed during diphenylmethane hydroconversion experiments performed in the inclined ◦ micro-reactor at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 - 4 h, 0 RPM (using acronyms from Table 4.9).

Species Composition 1 Species Composition 1 Species Composition 1 (area%) (area%) (area%) 1 h 2 h 4 h Toluene 48 Toluene 46 Toluene 45 Benzene 40 Benzene 41 Benzene 40 CHMB 6 CHMB 6 CHMB 7 HexF 4 HexF 3 MBP2,3 3 0 ppm Mo EPB 2 EPB 2 HexF 2 MTP 2 ETB 2 Fluorene 1

92 1 h 2 h 4 h Toluene 46 Toluene 47 Toluene 48 Benzene 41 Benzene 42 Benzene 41 CHMB 6 CHMB 7 CHMB 7 HexF 4 HexF 3 HexF 2 600 ppm Mo ETB 2 ETB 1 Fluorene 1 EPB 2 MBP3 1

1 h 2 h 4 h Toluene 41 Toluene 40 Toluene 42 Benzene 38 Benzene 38 Benzene 39 CHMB 17 CHMB 19 CHMB 16 HexF 2 HexF 2 HexF 2 1800 ppm Mo 3-Methylheptane 1 3-Methylheptane 1 3-Methylheptane 2 1 - Composition indicated is the percentage area from the GCMS chromatogram on a DPM-free basis (i.e. percentage of products formed) and limited to those species comprising >0.25%. first order fits with the kinetic coefficients presented in Table 4.12. The hard in- flection at 2 h observed in the inclined system has all but disappeared, with any variation remaining being encapsulated by experimental uncertainty. Curiously, the results of the 0 ppm Mo experiments now exceed those of the 1800 ppm Mo ones (although the kinetic coefficient of 0.082 h−1 for the 0 ppm Mo system is very close to the 0.088 h−1 observed in the stirred batch reactor). Once more, the application of LHHW kinetics would suggest similar kinetic expressions for both the 0 and 1800 ppm Mo systems, i.e. the same species adsorbing on the active sites of both FeS and MoS2, a conclusions which is in contradiction with the observation in the stirred batch reactor. It is thus clear that LHHW kinetics are not applicable to this catalytic system.

10 DPM conversion (wt%)

0 1 2 3 4

Reaction time (h)

Figure 4.32: Conversion results obtained for diphenylmethane hydroconversion experiments performed in the vertical stainless steel micro-reactor ◦ at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Error bars indicate standard deviation. Curves shown are ﬁrst order kinetic ﬁts as per Table 4.12. ◦ 0 ppm Mo. △ 1800 ppm Mo.

93 Table 4.12: Coefﬁcients for the kinetic models of diphenylmethane hydroconversion for data obtained in the vertical micro-reactor at 0 - 1800 ◦ ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM, depicted in Figure 4.32.

Catalyst loading Kinetic coefﬁcient 1 (ppm Mo) (h−1) 0 0.082 ± 0.002 1800 2 0.076 ± 0.001 1 - rDPM 0ppm = k0ppm.CDPM, rDPM 1800ppm = k1800ppm.CMo.CDPM. 2 - Reporting k1800ppm.CMo for comparison in context of the experiments.

The product yield data is presented in Figures 4.33 through 4.36. Interestingly in this system, no cracking products were detected for any of the experiments and the only isom./cond. product observed was fluorene, with an fluorene yield plot thus replacing the isom./cond. lump below. Despite the similarity in the DPM conversion data, the product composition for these experiments is surprising. The 1800 ppm Mo system now clearly exhibits lower benzene and toluene yields than its 0 ppm Mo counterpart and yet the B:T ratio (seen in Figure 4.37) remains almost the same for both and now slightly below unity (at approximately 0.95). The CHMB yield for 1800 ppm Mo, maintaining a similar range to the inclined reactor experiments, now appears to pass through a minimum as conversion increases. 0 ppm Mo shows a larger initial yield of CHMB than was observed in the inclined system and this yield is seen do decrease sharply with conversion. These trends (excluding the increase in the 1800 ppm Mo data) are reminiscent of the stirred reactor results and again suggest that catalytic hydrogenolysis in not a significant mechanism in this reaction system. If the catalyst were performing both hydrogenation and hydrogenolysis, one would anticipate saturated rings to be present in the product is appreciable quantities (the result of CHMB hydrogenolysis) which is not observed. Instead, both catalyst concentrations result in comparable yields of fluorene which decreases with increasing conversion. Figures 4.38 and 4.39 show typical gas product analyses for the 0 and 1800 ppm

Mo systems after reaction times of 1 and 4 h. Quantiﬁcation of C1 to C4 species is provided in Table 4.13. As was seen in the stirred batch reactor, the gaseous

94 1.0 )

0.8 DPMreacted

0.6

0.4

0.2 Benzene yieldmolar (mol/mol

0.0

0 10 20 30

DPM conversion (wt%)

Figure 4.33: Benzene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the vertical micro-reactor at ◦ 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. △ 1800 ppm Mo. products constitute a minimal mass toward the total products. The distribution of these species is, however, very interesting. The trend of the catalytic system to produce more and larger fragments is apparent. For the 0 ppm Mo systems, increasing the reaction time from 1 to 4 h results in an increase in the gas products formed, with the only apparent difference being a minor increase in C6 isomers. The 1800 ppm Mo systems, exhibit roughly the same species as the 0 ppm Mo ones, but the total amount and distribution differs. 1800 ppm Mo is noted to produce more of all gaseous species and that, with increasing reaction time, the amounts of these species increase. This includes notable increases in the C5 and C6 isomers not observed to such an extent in the 0 ppm Mo experiments. Despite the magnitude of the benzene and toluene peaks in these chromatograms, their composition in the gas phase remains less than 1 wt%. The larger benzene peak, compared to toluene, is due to its lower vapour pressure and more rapid rate of vaporisation. The low concentrations escaping to the gas phase were not thought to alter the liquid composition to any noticable extent.

95 1.0 )

0.8 DPMreacted

0.6

0.4

0.2 Toluene molar yieldmolar Toluene (mol/mol

0.0

0 10 20 30

DPM conversion (wt%)

Figure 4.34: Toluene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the vertical micro-reactor at ◦ 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. △ 1800 ppm Mo.

Table 4.13: Major gaseous products observed during diphenylmethane hydroconversion experiments performed in the vertical micro-reactor at ◦ 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2,1-4h,0RPM.

Conditions Composition (wt%) Methane Ethane Propane iso-Butane Butane 0 ppm Mo, 1 h 0.17 0.01 0.02 0.001 0.002 0 ppm Mo, 4 h 0.27 0.03 0.03 0.002 0.01 1800 ppm Mo, 1 h 0.18 0.05 0.05 0.001 0.03 1800 ppm Mo, 4 h 0.37 0.07 0.08 0.004 0.04

96 0.2 ) DPMreacted

0.1 CHMB yieldmolar (mol/mol

0.0

0 10 20 30

DPM conversion (wt%)

Figure 4.35: Cyclohexylmethylbenzene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the verti- ◦ cal micro-reactor at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2,0-4h,0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. △ 1800 ppm Mo.

0.2 ) DPMreacted

0.1 Fluorene molar yieldmolar Fluorene (mol/mol

0.0

0 10 20 30

DPM conversion (wt%)

Figure 4.36: Fluorene molar yield for diphenylmethane hydroconversion experiments performed in the vertical micro-reactor at 0 - 1800 ppm Mo, ◦ 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. △ 1800 ppm Mo.

97 1.25

1.00

0.75

0.50

0.25

0.00 Benzene:toluene molar ratio (mol:mol)Benzene:toluene

0 10 20 30

DPM conversion (wt%) Figure 4.37: Benzene:toluene molar ratio obtained for diphenylmethane hydroconversion experiments performed in the vertical micro-reactor at ◦ 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. △ 1800 ppm Mo.

98 4

Methane

Benzene 4

1.5 Ethane

Propane

C isomers

isoButane

Toluene 1.0

Butane

C isomers

0.5 Signal Signal intensity (-) x10

0.0

00:00 05:00 10:00 15:00 20:00 25:00 30:00

Retention time (mm:ss) (a)

5 4

1.5

1.0 Signal Signal intensity (-) x10

0.5

0.0

00:00 05:00 10:00 15:00 20:00 25:00 30:00

Retention time (mm:ss) (b)

Figure 4.38: Examples of gas chromatograms obtained for diphenylmethane hydroconversion experiments performed in the vertical micro-reactor ◦ at 0 ppm Mo, 445 C, 13.8 MPa H2, 1 - 4 h, 0 RPM. (a) 1 h reaction time. (b) 4 h reaction time.

99 5

3 4

1.5

1.0 Signal Signal intensity (-) x10

0.5

0.0

00:00 05:00 10:00 15:00 20:00 25:00 30:00

Retention time (mm:ss) (a)

5 4

1.5

1.0 Signal Signal intensity (-) x10

0.5

0.0

00:00 05:00 10:00 15:00 20:00 25:00 30:00

Retention time (mm:ss) (b)

Figure 4.39: Examples of gas chromatograms obtained for diphenylmethane hydroconversion experiments performed in the vertical micro-reactor ◦ at 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 - 4 h, 0 RPM. (a) 1 h reaction time. (b) 4 h reaction time.

100 4.2.3 Glass Insert Micro-Reactor Use of the glass insert micro-reactor allowed for yet more control over the factors inﬂuencing the reaction. Exposure to the catalytically active reactor wall was eliminated (although the active thermocouple wall was still in contact with the liquid), the rapid heat-up and cool-down rates allowed for better control over the temperatures to which the reaction mixture was exposed and implementation of an external vortex mixing system allowed for agitation.

Unmixed Comparison with Stainless Steel Micro-Reactor The first studies using this system were unmixed experiments for comparison with the vertical stainless steel micro-reactor. Figure 4.40 shows a direct comparison wherein it may be seen that the glass system exhibits a conversion roughly half that of the stainless steel reactor. It is noted that the experimental uncertainty is now lower and that the 1800 ppm Mo experiments exceed their 0 ppm Mo counterparts for all reaction times studied. Once more the trends may be roughly ap- proximated by first order kinetic models, the coefficients presented in Table 4.14. The closeness of the 0 and 1800 ppm Mo results are suspected to be due to the catalytic influence of the thermocouple walls (FeS) acting as additional catalytic centres. There appears to be little indication that the FeS and MoS2 form any manner of synergistic effect, an observation supported by published studies in this regard [107–109].

Table 4.14: Coefﬁcients for the kinetic models of diphenylmethane hydroconversion for data obtained in the glass insert micro-reactor at 0 - 1800 ◦ ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM, depicted in Figure 4.32.

Catalyst loading Kinetic coefﬁcient 1 (ppm Mo) (h−1) 0 0.036 ± 0.001 1800 2 0.042 ± 0.001 1 - rDPM 0ppm = k0ppm.CDPM, rDPM 1800ppm = k1800ppm.CMo.CDPM. 2 - Reporting k1800ppm.CMo for comparison in context of the experiments.

Figures 4.41 through 4.44 illustrate the yields for benzene, toluene, CHMB and the isom./cond. lump with the B:T ratio in Figure 4.45. Only trace amounts of other

101 30

10 DPM conversion (wt%)

0 1 2 3 4

Reaction time (h)

Figure 4.40: Comparison of conversion results obtained for diphenylmethane hydroconversion experiments performed in the vertical stainless steel and glass insert micro-reactors at 0 - 1800 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Error bars indicate standard deviation. Curves shown are first order kinetic fits as per Table 4.12 for the stainless steel reactor and Table 4.14 for the glass insert. • 0 ppm Mo, stainless steel. N 1800 ppm Mo, stainless steel. ◦ 0 ppm Mo, glass insert. △ 1800 ppm Mo, glass insert. cracking species were detected for either catalyst concentration for the reaction times investigated. This was also the case for CHMB in the 0ppm Mo experiments, wherein only trace concentrations were observed. Furthermore, fluorene was the major constituent of the isom./cond. lump except below 2 h reaction time when both fluorene and 4-benzylbiphenyl (BBP) were observed. As per the vertical micro-reactor, the benzene yield for 1800 ppm Mo is noted to be below that of the 0 ppm Mo but in the glass insert, the toluene yields for both catalyst concentrations quickly reach the same level and follow the same trend. The benzene:toluene ratio also differs with 0 ppm Mo showing a high initial ratio rapidly declining with increasing conversion to join that of the roughly constant 1800 ppm Mo at approximately 0.92. Figure 4.43 presents the CHMB yield for 1800 ppm Mo, only trace amounts detected for 0 ppm Mo. It may be seen that the yield remains almost constant for all conversions at a level below those observed for the stainless steel micro-reactor.

102 The isom./cond. lump products are again seen to follow a decreasing trend with increasing DPM conversion with 1800 ppm Mo below that of 0 ppm Mo.

1.0 )

0.8 DPMreacted

0.6

0.4

0.2 Benzene yieldmolar (mol/mol

0.0

0 5 10 15 20

DPM conversion (wt%)

Figure 4.41: Benzene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor ◦ at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. △ 1800 ppm Mo.

Visual Mixing Studies With no literature published on vortex mixing in micro-reactors, determination of a suitable mixing speed was a two step process. The first determined, visually, at what mixing speed and/or liquid volume a vortex was formed, the impact of the centrally-located thermocouple and whether or not particles were suspended by the motion. The second was a study of the impact of different mixing speeds (both with and without a full vortex) on the system performance. The visual study was performed using a high speed camera to film a glass mock-up of the micro-reactor system. Reaction product from an 1800 ppm Mo ◦ experiment (445 C, 13.8 MPa H2, 1 h, 0 RPM) was sealed in its insert and positioned inside the glass shell as described in Section 3.2.2. These observations were performed at ambient conditions (approximately 101.325 kPa and 20◦C) with Section 5.2.4 discussing how the fluid, and hence mixing, properties would change

103 1.0 )

0.8 DPMreacted

0.6

0.4

0.2 Toluene molar yieldmolar Toluene (mol/mol

0.0

0 5 10 15 20

DPM conversion (wt%)

Figure 4.42: Toluene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor ◦ at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. △ 1800 ppm Mo.

0.2 ) DPMreacted

0.1 CHMB yieldmolar (mol/mol

0.0

0 5 10 15 20

DPM conversion (wt%)

Figure 4.43: Cyclohexylmethylbenzene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the glass ◦ insert micro-reactor at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. △ 1800 ppm Mo.

104 0.2 ) DPMreacted

0.1 mass yield (g/g Other isom. and cond. products

0.0

0 5 10 15 20

DPM conversion (wt%)

Figure 4.44: Mass yield of isomerisation and condensation products for diphenylmethane hydroconversion experiments performed in the glass ◦ insert micro-reactor at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. △ 1800 ppm Mo.

1.25

1.00

0.75

0.50

0.25 Benzene:toluene ratiomolar (mol:mol)

0.00

0 5 10 15 20

DPM conversion (wt%)

Figure 4.45: Benzene:toluene molar ratio obtained for diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor ◦ at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. △ 1800 ppm Mo.

105 at reaction conditions. A problem quickly became evident as illustrated in Figure 4.46. With 400 µL of liquid, even the maximum mixing speed of 2500 RPM of the heavy duty vortex mixer was insufﬁcient to establish a vortex and agitate the full liquid volume, with the solid particles remaining settled. Further studies determined the height of the vortex to be a function of the mixer speed (this relationship being illustrated in Figure 4.47) and independent of liquid volume. It was thus necessary, to ensure agitation of all liquid in the insert and suspension of the solid particles, for the volume of liquid to be reduced to only 150 µL. With this volume, shown in Figure 4.48, solid suspension occurs at 2250 RPM and above. It may be seen, however, that a true vortex is not established with ﬂuid instead taking the form of a rotating concave wave for all speeds tested.

Figure 4.46: 2500 RPM mixing of 400 µL of reaction product obtained for diphenylmethane hydroconversion performed in the glass insert micro- ◦ reactor at 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM as seen in glass mock-up. This data represents an ideal system where the liquid movement is unobstructed.

106 30

10 "Vortex" height (mm)

1500 2000 2250 2500

Mixing speed (RPM)

Figure 4.47: Effect of mixing speed on “vortex” height as studied in glass mock-up using reaction product obtained for diphenylmethane hydroconversion performed in the glass insert micro-reactor at 1800 ppm ◦ Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM. Curve is for illustration of trend only. Data for 150 µL and 400 µL were indistinguishable.

When implemented in the micro-reactor, however, a thermocouple extends into the liquid of the insert. To determine what effect this obstruction would have on the mixing, the central thermocouple was duplicated in the glass mock-up, the results being presented in Figures 4.49 and 4.50 for 400 µL and 150 µL liquid loadings, respectively. It is apparent that the smooth wave motion is no longer present with the movement of the suspended thermocouple allowing it to act as a stirrer bar, agitating the liquid, suspending solids and entraining large amounts of gas. This is most effective for 150 µL where gas entrainment begins at approximately 2000 RPM and increases through 2250 RPM before the liquid wave reforms on the insert walls at 2500 RPM. In the 400 µL system even 2500 RPM is insufﬁcient to allow for the same extent of agitation to be achieved. One caveat to these visual mixing evaluations is that whilst they were performed using reaction product in an accurate mock-up, they were done at ambient conditions and not the 445◦C and 13.8 MPa of the reactor. As such, certain considerations must be made in their interpretation as discussed in Section 5.2.4.

107 108

(a) (b) (c) (d)

Figure 4.48: 1500 - 2500 RPM mixing of 150 µL of reaction product obtained for diphenylmethane hydroconversion ◦ performed in the glass insert micro-reactor at 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM as seen in glass mock-up. (a) 1500 RPM. (b) 2000 RPM. (c) 2250 RPM (lowest speed at which solid suspension was observed). (d) 2500 RPM. (a) (b)

Figure 4.49: 2000 - 2500 RPM mixing of 400 µL of reaction product obtained for diphenylmethane hydroconversion performed in the glass ◦ insert micro-reactor at 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM as seen in glass mock-up with central thermocouple. (a) 2000 RPM. (b) 2500 RPM.

109 110

(a) (b) (c) (d)

Figure 4.50: 1500 - 2500 RPM mixing of 150 µL of reaction product obtained for diphenylmethane hydroconversion ◦ performed in the glass insert micro-reactor at 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM as seen in glass mock-up with central thermocouple. (a) 1500 RPM. (b) 2000 RPM. (c) 2250 RPM. (d) 2500 RPM. Comparison of Liquid Loading Volumes The necessity of using a reduced volume required a comparative study to be performed. A lower liquid loading would change the hydrogen:DPM ratio and potentially affect the reaction mechanism. Table 4.15 shows the results from these experiments. As may be seen, reducing the liquid volume from 400 to 150 µL results in several key changes to the reaction, most notably increasing the DPM conversion (perhaps an indication of diffusional limitations in this unmixed system), increasing the yield of CHMB and decreasing the yield of isom./cond. products.

Table 4.15: Comparison of results for diphenylmethane hydroconversion obtained in the glass insert micro-reactor at 0 - 1800 ppm Mo, 445◦C, 13.8 MPa H2, 1 h, 0 RPM, 150 - 400 µL liquid to examine the effect of liquid loading volume.

Conditions DPM conversion Product yield (wt%) (mol/molDPM reacted) Benzene Toluene CHMB 0 ppm Mo, 150 µL 7.7 ± 0.5 0.92 ± 0.02 0.91 ± 0.02 0.03 ± 0.00 0 ppm Mo, 400 µL 3.3 ± 0.3 0.92 ± 0.03 0.78 ± 0.02 0.00 ± 0.00 1800 ppm Mo, 150 µL 11.1 ± 1.1 0.83 ± 0.01 0.86 ± 0.01 0.12 ± 0.05 1800 ppm Mo, 400 µL 5.7 ± 0.2 0.84 ± 0.02 0.89 ± 0.02 0.02 ± 0.03

Product yield B:T ratio (g/gDPM reacted) (mol/mol) Cracking Isom./cond. 0.00 0.04 ± 0.02 1.02 ± 0.00 0.00 0.15 ± 0.01 1.12 ± 0.06 0.00 0.00 0.97 ± 0.05 0.00 0.08 ± 0.02 0.94 ± 0.01

Thermocouple Wall Activity Before beginning mixing experiments, a series of tests was conducted to examine the wall activity of the central thermocouple. This series differed from the reactor wall activation study in Section 4.2.1 by conducting the tests without catalyst (so as to rule out activation by MoS2 deposition). As such, a series of 0 ppm Mo experiments were run after installation of a fresh thermocouple to determine the activation of the stainless steel sheath and its inﬂuence on the reaction mechanism. The DPM conversion for these sequential experiments is shown in Figure 4.51 wherein it may be seen that an initially low conversion rises rapidly with each subsequent

111 test, plateauing at a value comparable with that seen with the aged thermocouple (shown in Table 4.15). Figure 4.52 shows combined benzene and toluene yield data and Figure 4.53 the corresponding B:T ratio. It may be noted that without the activity of the thermocouple, the system tends to the formation of excess toluene but that the yields quickly equalise with the B:T ratio climbing toward unity as the thermocouple wall becomes more active. The CHMB yield in Figure 4.54 shows a roughly steady formation of this species with it being noted that its presence (not seen in previous glass insert experiments) being due to the reduced liquid loading as shown in Table 4.15. The isom./cond. lump product yield shown in Figure 4.55 is seen to increase with wall activity. This fraction comprises only ﬂuorene at lower conversions but includes BBP as the conversion increases with wall activity (notably these are the same products observed for the aged thermocouple). The species present and the trends observed suggest that the catalytically active thermocouple wall (predominantly FeS) performs a role comparable to that of the MoS2 catalyst being investigated, i.e. that it is foremost a hydrogenation catalyst.

2 DPM conversion (wt%)

0 2 4 6 8 10

Sequential experiment number (-)

Figure 4.51: Conversion results obtained for diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor at 0 ppm ◦ Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM to study thermocouple wall activation. Curve is for illustration of trend and is not a kinetic ﬁt.

112 1.25

1.00 )

0.75 DPMreacted

0.50 (mol/mol

0.25 Benzene and toluene yieldmolar

0.00

0 3 6 9

DPM conversion (wt%)

Figure 4.52: Benzene and toluene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the glass insert ◦ micro-reactor at 0 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM to study thermocouple wall activation. Curves are for illustration of trends only. • - Toluene. ◦ - Benzene.

1.25

1.00

0.75

0.50

0.25

Benzene:toluene ratiomolar (mol:mol) 0.00

0 3 6 9

DPM conversion (wt%)

Figure 4.53: Benzene: toluene molar ratio results obtained for diphenylmethane hydroconversion experiments performed in the glass insert micro- ◦ reactor at 0 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM to study thermocouple wall activation. Curve is for illustration of trend only. ◦ 0 ppm Mo. △ 1800 ppm Mo.

113 0.2 ) DPMreacted

0.1 CHMB yieldmolar (mol/mol

0.0

0 3 6 9

DPM conversion (wt%)

Figure 4.54: Cyclohexylmethylbenzene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the glass ◦ insert micro-reactor at 0 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM to study thermocouple wall activation. Curve is for illustration of trend only.

0.2 ) DPMreacted

0.1 mass yield (g/g Other isom. and cond. products

0.0

0 3 6 9

DPM conversion (wt%)

Figure 4.55: Mass yield of isomerisation and condensation products for diphenylmethane hydroconversion experiments performed in the glass ◦ insert micro-reactor at 0 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM to study thermocouple wall activation. Curve is for illustration of trend only. 114 Effect of Mixing Speed With the activity of the thermocouple wall stabilised, an investigation into the effect of mixing speed on the reaction could be performed. Figure 4.56 shows the DPM conversion over the range of mixing speeds investigated wherein a clear and unexpected trend may be seen. Increasing the mixing speed beyond 1500 RPM results in a rapid decline in the conversion for both the 0 and 1800 ppm Mo systems. The difference between these two datasets, representing the Mo catalyst activity, appears to show a maximum at 2000 RPM. Whilst the indicated decline at higher mixing speeds may be an experimental artifact, for the purposes of this study, 2000 RPM was identified as the optimum mixing speed and selected for more detailed study. Interestingly the DPM conversion observed for 0 ppm Mo at 2250 RPM is very similar to those values obtained for the system with an inactive thermocouple (see Figure 4.51). Figures 4.57 and 4.58 show the benzene and toluene yields, respectively. Note that the data is presented versus the DPM conversion rather than mixer speed as the non-linearity of conversion with mixing speed would make meaningful comparisons from such graphs difficult. Whilst the toluene yields appear to follow the same trend for both 0 and 1800 ppm Mo, the 0 ppm Mo system is noted to exhibit higher benzene yields (and hence higher B:T ratios as seen in Figure 4.59) than its 1800 ppm Mo counterpart. 1800 ppm Mo was seen to have higher CHMB yields than 0 ppm Mo and whilst 0 ppm Mo showed no cracking products, 1800 ppm Mo was observed to produce 3-methylheptane. This cracking product only appeared at 1500 RPM with a yield of 2.1 ± 0.2 mol/molDPM reacted. The only isom./cond. product observed was fluorene for 0 ppm Mo at 0 RPM with a yield of 0.04 ± 0.02 mol/molDPM reacted.

Analysis of Recovered Solids The maximum in the observed catalytic activity at 2000 RPM, together with visual evaluation of the product samples changing from ﬁne, suspended particles to large, glitter-like ﬂakes, prompted an investigation into the structure of the recovered solids by TEM, SEM and SEM-EDX. Samples selected for analysis covered both mixing speed and reaction time ranges but were limited to those experiments with

115 12

4 DPM conversion (wt%)

0 1500 2000 2250

Mixing speed (RPM)

Figure 4.56: Conversion results obtained for diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor at 0 - 1800 ◦ ppm Mo, 445 C, 13.8 MPa H2, 1 h, 0 - 2250 RPM. Curves are for illustration of trends and are not kinetic ﬁts. ◦ 0 ppm Mo. △ 1800 ppm Mo. X Difference between 0 and 1800 ppm Mo conversion.

1.0 )

0.8 DPMreacted

0.6

0.4

0.2 Benzene yieldmolar (mol/mol

0.0

0 4 8 12

DPM conversion (wt%)

Figure 4.57: Benzene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor ◦ at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. △ 1800 ppm Mo.

116 )

1.25

1.00 DPMreacted

0.75

0.50

0.25 Toluene molar yieldmolar Toluene (mol/mol 0.00

0 4 8 12

DPM conversion (wt%)

Figure 4.58: Toluene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor ◦ at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. △ 1800 ppm Mo.

1.25

1.00

0.75

0.50

0.25

Benzene:toluene ratiomolar (mol:mol) 0.00

0 4 8 12

DPM conversion (wt%)

Figure 4.59: Benzene:toluene molar ratio obtained for diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor ◦ at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. △ 1800 ppm Mo.

117 0.2 ) DPMreacted

0.1 CHMB yieldmolar (mol/mol

0.0

0 4 8 12

DPM conversion (wt%)

Figure 4.60: Cyclohexylmethylbenzene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the glass ◦ insert micro-reactor at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 0 RPM. Curves are for illustration of trends only. ◦ 0 ppm Mo. △ 1800 ppm Mo.

118 1800 ppm Mo (there were no solids from 0 ppm Mo experiments). Figures 4.61 and 4.62 show the TEM results for 0 and 2000 RPM samples, respectively. Whilst inter-plate and d-spacing measurements (shown in Figure 4.61b) confirm that the solid material is MoS2, the shape and size of the crystallites is seen to change dramatically with conditions. It is clear that the solid material from the 0 RPM experiments is almost identical to that from the stirred batch reactor (compare with Figures 4.19 and 4.20) in terms of stack height, sheet width and “rag” arrangement, the material present after 2000 RPM has undergone extreme changes. The short stacks of small sheets, observed for all reaction times at 0 RPM and initially at 2000 RPM, appear to agglomerate and fuse with prolonged reaction time at 2000 RPM to form large sheets up to 100 nm wide and stacks more than 20 sheets thick. Analysis by SEM-EDX is shown in Figure 4.63 with compositional results in Table 4.16. It may be seen that whilst the 0 RPM system is comprised almost entirely of S and Mo with a close correlation between these species (confirmed by their almost 1:1 quantification), at 2000 RPM an appreciable amount of Fe is found in the sample. This Fe, likely present as FeS as seen in Figure 4.17 and by the close correlation in Figure 4.63b, is thought to form by sulphidation of the stainless steel thermocouple wall (as evidence by the Ni in the solid) and slough off into the liquid due to the vigorous agitation of the system. Note that as these EDX analyses were performed without the use of an internal standard, it is difficult to obtain independent quantification results. The compositions are presented for comparative purposes only.

Table 4.16: Results for scanning electron microscopy with energy dispersive X-ray quantiﬁcation for solid material obtained from a diphenylmethane hydroconversion performed in the glass insert micro-reactor at 1800 ◦ ppm Mo, 445 C, 13.8 MPa H2, 4 h, 0 - 2000 RPM.

Mixing speed Composition (wt%) (RPM) S Mo Fe Ni 0 51 48 1 0 2000 53 27 18 2 To further examine the structure of the recovered solid material, FESEM was performed on solid samples representing a range of mixing speeds and reaction

119 times as shown in Figures 4.64 through 4.65. The solid material from the 0 RPM experiment presents as small, divided platelets collected loosely into larger structures. This material does not appear to change with increased reaction time. Mate- rial from reactions with 2000 RPM mixing, however, appear as larger plates fusing into tight sheets with increased reaction time. These observations compare well with TEM data in Figures 4.61 and 4.62. Of interest is that this fusing is apparent even when comparing the 0 h results shown in Figures 4.64e and 4.65e. This supports the theory that in both mixed and unmixed systems small particles of MoS2 precipitate separately during heat-up. In the mixed system, these particles then agglomerate and fuse. There is no evidence to support that MoS2 crystallite growth in the mixed system is due to initial MoS2 particles acting as nucleation centres. To study this crystal growth and simultaneously determine the extent to which it occurs at higher mixing speeds, two 2250 RPM experiments were conducted, the results presented in Figures 4.66 and 4.67. The mixing in one experiment began at the start of heating (as per the standard procedure), the other only once reaction temperature was achieved. After 1 h reaction time, very little difference is discernible between the two samples. To clarify the presentation of the plates and fused structures, one of the 2250 RPM samples was angled during FESEM analysis, the large, smooth sheets of agglomerated and fused crystallites being clearly visible in Figure 4.68.

Optimum Mixing Speed Evaluation With the optimum mixing speed selected (2000 RPM providing the maximum DPM conversion difference between 0 and 1800 ppm Mo), the reaction was studied over a range of reaction times both with and without agitation. The DPM conversion results are presented in Figure 4.69 wherein it may be seen that the trends for both the 0 and 2000 RPM are similar with 1800 ppm Mo showing a rapid increase followed by a more steady slope. For 0 ppm Mo, however, a clearly sigmoidal curvature was apparent. At 1 h the 1800 ppm Mo conversion results for 0 and 2000 RPM are almost identical but deviate with longer reaction times. This deviation is noted for all reaction times with 0 ppm Mo. The sigmoidal DPM conversion results once again resulted in poor kinetic ﬁts (see Figure E.3 and Table E.40).

120 (a)

(b)

Figure 4.61: Transmission electron microscopy images for solid material obtained from diphenylmethane hydroconversion performed in the glass ◦ insert micro-reactor at 1800 ppm Mo, 445 C, 13.8 MPa H2,0-4h,0 RPM. Images used to conﬁrm MoS2 formation and examine changes in crystallite dimensions (stack height and sheet width) with time in the absence of without mixing. (a) 0 h. (b) 4 h showing the both the inter-plate spacing (d002 = 0.62 nm) and the d-spacing (d100 = 0.27 nm).

121 (a)

(b)

Figure 4.62: Transmission electron microscopy images for solid material obtained from diphenylmethane hydroconversion performed in the glass ◦ insert micro-reactor at 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 2000 RPM. Images used to examine changes in crystallite dimensions (stack height and sheet width) with time when the reactor was mixed. (a) 0 h. (b) 4 h.

122 (a) 123

(b)

Figure 4.63: Scanning electron microscopy with energy dispersive X-ray images for solid material obtained from diphenylmethane hydroconversion performed in the glass insert micro-reactor at 1800 ppm Mo, 445◦C, 13.8 MPa H2, 4 h, 0 - 2000 RPM. Elemental maps inverted for clarity with dark spots indicating positive detection. (a) 0 RPM, 4 h. (b) 2000 RPM, 4 h. (a) (b)

(e) (f)

Figure 4.64: Field emission scanning electron microscopy images for solid material obtained from diphenylmethane hydroconversion performed in the glass insert micro-reactor at 1800 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 4 h, 0 RPM. Images used to examine particle agglomeration with time in the absence of mixing. (a) 4 h. (b) 4 h. (c) 4 h. (d) 4 h. (e)0h. (f)2h.

124 (a) (b)

(e) (f)

Figure 4.65: Field emission scanning electron microscopy images for solid material obtained from a diphenylmethane hydroconversion performed ◦ in the glass insert micro-reactor at 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 2000 RPM. Images used to examine particle agglomeration with time when the reactor was mixed. (a) 4 h. (b) 4 h. (c) 4 h. (d) 4 h. (e)0h. (f)2h.

125 (a) (b)

(c)

Figure 4.66: Field emission scanning electron microscopy images for solid material obtained from diphenylmethane hydroconversion performed ◦ in the glass insert micro-reactor at 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 2250 RPM (including during heat-up as per normal procedure). Images used to examine particle formation and agglomeration when the reactor was mixed during heat-up and reaction.

126 (a) (b)

(e)

Figure 4.67: Field emission scanning electron microscopy images for solid material obtained from diphenylmethane hydroconversion performed ◦ in the glass insert micro-reactor at 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 2250 RPM (excluding during heat-up). Images used to examine particle formation and agglomeration when the reactor was mixed only during reaction and not heat-up.

127 128

(a) (b)

Figure 4.68: Field emission scanning electron microscopy images for solid material, the sample angled for better interpretation of the structure, obtained from diphenylmethane hydroconversion performed in the glass insert micro- ◦ reactor at 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 2250 RPM (excluding during heat-up). Images used to examine the macro-structure of the particle agglomerates. 30

0 RPM

2000 RPM

10 DPM conversion (wt%)

0 1 2 3 4

Reaction time (h)

Figure 4.69: Conversion results obtained for diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor at 0 - 1800 ◦ ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 2000 RPM. Error bars indicate standard deviation. Curves are for illustration of trends and are not kinetic ﬁts. • 0 ppm Mo, 0 RPM. N 1800 ppm Mo, 0 RPM. ◦ 0 ppm Mo, 2000 RPM. △ 1800 ppm Mo, 2000 RPM.

129 The benzene and toluene yields are illustrated in Figures 4.70 and 4.71 with the corresponding B:T ratio in Figure 4.72. For both 0 and 2000 RPM a trend of lesser benzene yields for 1800 than 0 ppm Mo is observed with the 2000 RPM data perhaps declining slightly while the 0 RPM results do not. Both catalyst loadings show increasing toluene yields with increasing conversion, a trend mirrored in the 1800 ppm Mo 2000 RPM system. The mixed 0 ppm Mo system, however, shows a steady decline in toluene yield with increasing DPM conversion. The B:T trends appear similar over a given conversion range with 1800 ppm Mo falling below 0 ppm Mo. It is also noted that the 2000 RPM samples exhibit lower B:T vratios than their unmixed counterparts.

1.0 )

0.8 DPMreacted

0.6

0.4

0.2 Benzene yieldmolar (mol/mol

0.0

0 5 10 15 20 25

DPM conversion (wt%)

Figure 4.70: Benzene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor ◦ at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 - 2000 RPM, 150 µL liquid load. Curves are for illustration of trends only. • 0 ppm Mo, 0 RPM. N 1800 ppm Mo, 0 RPM. ◦ 0 ppm Mo, 2000 RPM. △ 1800 ppm Mo, 2000 RPM. The yield of CHMB is shown in Figure 4.73 from which a decrease in the this yield with increased mixing may be noted. Whilst mixing appears to suppress CHMB yield for 1800 ppm Mo, the effect is promotional for 0 ppm Mo. If the trends for 400 and 150 µL comparisons are recalled (see Table 4.15), an increase in the hydrogen:DPM ratio resulted in an increase in CHMB production. It seems

130 1.25 )

1.00 DPMreacted

0.75

0.50

0.25 Toluene molar yieldmolar Toluene (mol/mol 0.00

0 5 10 15 20 25

DPM conversion (wt%)

Figure 4.71: Toluene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor ◦ at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 - 2000 RPM, 150 µL liquid load. Curves are for illustration of trends only. • 0 ppm Mo, 0 RPM. N 1800 ppm Mo, 0 RPM. ◦ 0 ppm Mo, 2000 RPM. △ 1800 ppm Mo, 2000 RPM.

1.25

1.00

0.75

0.50

0.25

Benzene:toluene ratiomolar (mol:mol) 0.00

0 5 10 15 20 25

DPM conversion (wt%)

Figure 4.72: Benzene:toluene molar ratio obtained for diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor ◦ at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 - 2000 RPM, 150 µL liquid load. Curves are for illustration of trends only. • 0 ppm Mo, 0 RPM. N 1800 ppm Mo, 0 RPM. ◦ 0 ppm Mo, 2000 RPM. △ 1800 ppm Mo, 2000 RPM. 131 likely that mixing the 0 ppm Mo system allows for the dissolution of more hydrogen and hence the formation of more CHMB. Other cracking and isom./cond. species were few and limited to speciﬁc conditions. Only trace cracking species were observed for the 0 ppm Mo experiments but 1800 ppm Mo presented a near-constant yield of 3-methylheptane of approximately 0.022 mol/molDPM reacted for all reaction times. The yield of other isom./cond. species is presented in Figure 4.74 from which it may be seen that whilst such species appeared in both 0 RPM systems (1, 2 and 4 h for 0 ppm Mo but only 2 and 4 h for 1800 ppm Mo), only higher conversion 0 ppm Mo reactions showed any such species when stirred (only observed in 4 h experiments). The compositions of these isom./cond. lumps are shown in Table 4.17.

0.2 ) DPMreacted

0.1 CHMB yieldmolar (mol/mol

0.0

0 5 10 15 20 25

DPM conversion (wt%)

Figure 4.73: Cyclohexylmethylbenzene molar yield results obtained for diphenylmethane hydroconversion experiments performed in the glass ◦ insert micro-reactor at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 - 2000 RPM, 150 µL liquid load. Curves are for illustration of trends only. • 0 ppm Mo, 0 RPM. N 1800 ppm Mo, 0 RPM. ◦ 0 ppm Mo, 2000 RPM. △ 1800 ppm Mo, 2000 RPM.

Quantiﬁcation of the C1 to C4 gaseous products are presented in Table 4.18. Several trends may be noted from this data:

• Increasing reaction time increases the yield of all gaseous products, especially the larger species,

132 0.06 )

0.04 DPMreacted

0.02 mass yield (g/g Other isom. and cond. products

0.00

0 5 10 15 20 25

DPM conversion (wt%)

Figure 4.74: Mass yield of isomerisation and condensation products for diphenylmethane hydroconversion experiments performed in the glass ◦ insert micro-reactor at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2,0-4 h, 0 - 2000 RPM, 150 µL liquid load. Curve is for illustration of trend only. • 0 ppm Mo, 0 RPM. N - 1800 ppm Mo, 0 RPM. ◦ - 0 ppm Mo, 2000 RPM.

• Increasing the catalyst concentration increases the yield of all gaseous products, especially the larger species,

• Mixing decreases the yield of all gaseous products, especially the smaller species.

Spent Residue Hydroconversion Catalyst Evaluation The ﬁnal series of experiments involved evaluation of deactivated catalysts from residue hydroprocessing reactions. Three samples were tested, all consisting of coke-catalyst agglomerate recovered from residue hydroprocessing reactions performed by Rezaei et al. [32] in the same stirred reactor as this study but operated in ◦ semi-batch mode at 445 C, 13.8 MPa H2 (ﬂowing at 900 sccm), 1 h, 700 RPM using Cold Lake vacuum residue as the feed. In these reactions molybdenum chloride in reversed micelles was used as the MoS2 precursor, with the resulting catalyst shown to possess near identical properties and functionality to that derived from

133 Table 4.17: Major constituents observed in the isom./cond. product lump obtained during diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor at 0 - 1800 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 4 h, 0 - 2000 RPM, 150 µL liquid load (using acronyms from Table 4.9).

Catalyst loading Mixing speed Reaction time Species Composition 1 (ppm Mo) (RPM) (h) (area%) 1 Fluorene 100 Fluorene 51 2 ETB 49 0 Fluorene 39 0 4 HexF 39 ETB 22 HexF 69 2000 4 Fluorene 31 HexF 59 2 Fluorene 41 1800 0 HexF 60 4 Fluorene 40 1 - Composition indicated is the relative contribution of each species to the isom./cond. lump based on the percentage area from the GCMS chromatogram on a DPM-free basis (i.e. percentage of products formed).

134 Table 4.18: Major gaseous products observed during diphenylmethane hydroconversion experiments performed in the ◦ glass insert micro-reactor at 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 - 2000 RPM, 150 µL liquid load.

Catalyst loading Mixing speed Reaction time Composition (wt%) ×103 (ppm Mo) (RPM) (h) Methane Ethane Propane iso-Butane Butane

135 1 23.5 2.3 1.4 0.0 0.7 0 4 37.2 3.5 5.0 0.4 2.5 0 1 3.5 0.4 0.3 0.0 0.2 2000 4 16.8 2.1 2.9 0.0 0.0 1 37.9 7.3 5.9 0.5 4.8 0 4 47.6 9.8 12.8 1.3 9.4 1800 1 18.3 6.2 3.9 0.4 3.6 2000 4 14.7 3.5 3.9 0.5 3.3 Mo octoate. The recovered coke-catalyst agglomerate was washed with toluene, ground and dried prior to use. The three samples tested in this investigation are presented below, introduced to ensure 1800 ppm Mo was present in all reactions. ◦ Each was evaluated at 445 C, 13.8 MPa H2, 2000 RPM for 1 h with a 150 µL liquid load of undiluted DPM, these conditions being selected to obtain the best resolution of catalyst performance (the greatest difference between thermal and catalytic systems).

• “Fresh” coke-catalyst recovered after only a single residue hydroconversion experiment,

• Deactivated coke-catalyst recovered after ﬁve residue hydroconversion recovery and reuse cycles,

• “Fresh” coke catalyst thermally aged under 100 sccm He at 700◦C for 15 h [1].

Table 4.19 shows the results from these experiments together with coke yield and hydrogen conversion data (used to evaluate catalyst performance) from their use in residue hydroconversion reactions for comparison. In the context of residue hydroconversion, wherein a 0 ppm Mo experiment exhibits a coke yield of 21 wt% [32], both the heat treated and ﬁfth recycle catalysts are considered deactivated due to their elevated coke yields as compared with the fresh specimen. Hydrogen conversion is also seen to decline, although the change is not as dramatic, with the fresh catalyst exhibiting a greater hydrogen conversion than either the heat treated or recycled catalyst experiments (a 0 ppm Mo experiment yielding a hydrogen conversion of 13 % [32]). Evaluation in DPM hydroconversion, however, suggests the heat treated catalyst to be the most active (its DPM conversion and product yields on par with 1800 ppm Mo introduced as Mo octoate) with both the fresh coke-catalyst and the ﬁfth recycled species showing poor performance in terms of DPM conversion and a clearly change to the reaction mechanism per the product yields. Of interest was that only trace amounts of other cracking or isom./cond. species were present in the liquid products. For additional detail, the gas product analyses for these DPM hydroconversion tests are provided in Table 4.20. It is clear that whilst the fresh catalyst (both heat treated and not) show gas product yields

136 on par with previous 1800 ppm Mo experiments, the DPM hydroconversion with the ﬁfth recycle coke-catalyst agglomerate results in the formation of signiﬁcantly more gaseous products (although combined these still only account for less than 1 wt% of the gas).

137 Table 4.19: Results observed during diphenylmethane hydroconversion experiments performed in the glass insert ◦ micro-reactor at 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 2000 RPM, 150 µL liquid load using various coke- catalyst agglomerates from residue hydroconversion experiments.

Coke-catalyst Residue hydroconversion DPM hydroconversion

138 1 2 type Coke yield H2 conversion DPM conversion Product yield (mol/molDPM reacted) B:T ratio (wt%) (%) (wt%) Benzene Toluene CHMB (mol/mol) Fresh 2.9 18 0.5 0.65 1.28 0.0 0.51 Fresh, heat treated 3 11.5 15 11.1 0.74 0.81 0.21 0.92 Recycled 11.5 14 1.9 0.19 1.36 0.10 0.14 1 - Coke-catalyst agglomerate samples courtesy of Rezaei and Smith [1]. 2 - Coke yield expressed as wt% of total product. 3 - “Fresh” catalyst heat treated under 100 sccm He at 700◦C for 15 h. Table 4.20: Major gaseous products observed during diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor ◦ at 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 2000 RPM, 150 µL liquid load using various coke-catalyst agglomerates from residue hydroconversion experiments.

Catalyst type 1 Composition (wt%) ×103 (ppm Mo) Methane Ethane Propane iso-Butane Butane Fresh 13.2 1.9 1.3 0.0 1.6 Fresh, heat treated 3 15.7 1.5 1.1 0.0 1.1 Recycled 279 106 80 6 51 1 - Coke-catalyst agglomerate samples courtesy of Rezaei and Smith [1]. 2 - “Fresh” catalyst heat treated under 100 sccm He at 700◦C for 15 h.

139 Chapter 5

Discussion of Experimental Results

Presented below is an in-depth discussion of the results obtained during this study, their interpretation and implications. The results of additional experimental tests (not included in the main program of Table 3.1), data processing, thermodynamic analyses, etc. are presented, as needed, to provide additional detail. This discussion is organised according to the “Phases” of the work presented in Section 3.1.2.

5.1 Model Compound Evaluation

5.1.1 Model Compound Screening The conversion and product distribution results for the screening of the three model compounds selected for evaluation in this study (diphenylmethane [DPM], diphenylethane [DPE] and diphenylpropane [DPP]), diluted to 3 wt% in decahydronaphthalene (decalin), are presented in Section 4.1.1. It was determined that both DPE and DPP are extremely reactive under the operating conditions of 445◦C and 13.8 MPa

H2, achieving near complete conversion at both 0 and 600 ppm Mo (see Table 4.1). DPM showed lower conversions in the range of 30 - 40 wt%, making it more suitable for continued mechanistic studies in this work. Unfortunately, only DPE and DPP would allow for additional information to be

140 gathered by studying the secondary cracking of their alkyl branches. In an attempt to compromise between the desired operating temperature of 445◦C and a lower conversion, DPP was studied at temperatures as low as 420◦C (Figure 4.1). This temperature reduction saw a conversion decrease of only approximately 3 wt%. The product distributions from all three model compounds (presented in Table 4.2) were greatly informative. Many of the expected species were present: toluene from DPM, and toluene and ethylbenzene from DPE and DPP. Of note is the variety of species formed, with Table 4.2 showing only those present as >5 area% in the GCMS analyses. This suggests a complex reaction network, with the primary cracking radicals shown in Figure 3.1 undergoing signiﬁcant continued cracking, isomerisation and condensation reactions. As shown in Figure 2.6 for DPM, whether by the mechanisms proposed by Curran et al. [78], Wei et al. [46] or LaMarca et al.

[88], the addition of the MoS2 catalyst should serve to simplify the reaction product by providing lower energy pathways to stable species. What is observed, however, is that the addition of 600 ppm Mo increases product diversity and shifts the distribution toward more hydrogenated and cracked species (Table 4.2). Unfortunately the complexity of these product distributions makes clariﬁcation of the mechanisms almost impossible. From the DPM product compositions in Table 4.2, it is clear that the addition of MoS2 results in the formation of benzene, C2 substituents on rings and cyclohexylmethylbenzene (CHMB, being hydrogenated DPM). The benzene appears to support the mechanism of Curran et al. [78] (see Figure 2.6a) whereby the primary thermolysis radicals are rapidly stabilised. This theory does not, however, explain the increase in product diversity. The variety of product species and the C2 alkyl branches support the mechanism of LaMarca et al. [88] (see Figure 2.6c) whereby hydrocarbon radicals initiate cracking but this does not support the formation of benzene. The presence of hydrogenated rings and CHMB suggest that the catalyst has a strong hydrogenation role in the reaction which may be key to understanding the system. Even with this simple model compound it is apparent why so many contrasting mechanisms have been proposed and opposed in the literature [21, 78, 88] (see Section 2.5.2 for details). Comparing the conversion and product distribution results for DPM, DPE and DPP it is extremely difﬁcult to identify the role of the catalyst.

141 The DPP product compositions in Table 4.2 also afford some understanding regarding the thermal requirements of the reaction. It appears that below 445◦C, certainly below 430◦C, the rate of thermal decomposition has slowed with only the primary cracking products, ethylbenzene and toluene, being observed. At these temperatures it appears that the catalyst has reduced activity with none of the characteristic hydrogenation species being present. This seems counter-intuitive as the lower energy barrier for catalytic reactions should allow these to dominate at lower temperatures. This trend is due to the susceptibility of the DPP alkyl bridge to thermolysis, achieving >96% conversion even at 420◦C.

5.1.2 Diphenylmethane Studies

Diluted Diphenylmethane With DPM selected as the most appropriate of the three model compounds tested in this study, given its mid-range conversion under reaction conditions and simpliﬁed product spectrum, a more in-depth analysis was performed. These experiments were conducted at the same conditions as the screening experiments but with varying reaction time, the DPM conversion and product yields being studied in detail to determine reaction mechanisms. One helpful piece of information when analysing the decomposition of DPM, particularly when using the mechanisms proposed in literature [21, 29, 46, 78, 88] (see Figure 2.6) as a starting point, is the stability of the major species, namely the products (benzene and toluene) and the diluent (decalin), under reaction conditions. This information is provided in Tables 4.4 through 4.7. It is clear that both benzene and toluene, once formed, react under the given conditions to form various hydrogenation, cracking and isomerisation products as shown in Figure 5.1. The conversions of benzene and toluene are comparable, as are the products formed, but both these and decalin are less reactive than DPM. Whilst toluene hydrogenation to methylcyclohexane, isomerisation to ethylcyclopentane or cracking to 3- methylheptane or benzene all seem to be trivial mechanisms, the presence of these same product species in almost the same ratios in the benzene blank unexpectedly suggests a very similar mechanism (one which is able to produce C7 species from

142 a C6 reagent). The decalin blank shows different product species in minimal quantities and its decomposition is thus not responsible for the C7 species observed. It seems clear that this is a process of hydrogenation with minimal dehydrogenation (all products from benzene decomposition being saturated species) and that cracking, recombination and isomerisation are involved (benzene forms C7 species and toluene forms C6 species). It is theorised that toluene decomposes to benzene which is subsequently hydrogenated to cyclohexane. This cyclohexane rapidly cracks and recombines to form C7 alkanes which isomerise and cyclicise to the observed products (cyclicisation perhaps occurring during the abstraction of H*). The equivalent conversions of benzene and toluene decomposition and the presence of benzene in the toluene product suggest that the removal of C1 from toluene occurs faster than the hydrogenation of benzene. Thermodynamic simulations, presented in Figure 5.2, show the Gibbs free energy of reaction (∆Gr) of toluene to methane (an example for illustration) to be lower than that of benzene to methane, the former reaction hence occurring more readily than the latter. An alternative explanation would be that both benzene and toluene undergo hydrogenation, cracking, recombination and isomerisation with the formation of benzene from toluene being a separate process. Uncertainty in the exact mechanism aside, it appears that under reaction conditions, both benzene and toluene decompose at roughly equivalent rates but toluene does so to form benzene as a product. This would result in higher yields of benzene (i.e., higher benzene:toluene or B:T molar ratios) than may otherwise be expected. Returning to the DPM conversion data shown in Figure 4.2, it may be seen that both 0 and 600 ppm Mo experiments reach reaction temperature with a DPM conversion of approximately 30 wt%. This non-zero conversion at the start of the experiment is due to the slow heat-up of the reactor (as shown in Figure 3.3) which affords the thermal reaction an opportunity to begin prior to reaching operating temperature. The closeness in conversion between 0 and 600 ppm Mo upon reaching 445◦C suggests that minimal catalyst activity occurs during heat-up. This trend is supported by the DPP temperature data in Table 4.2 which showed a clear decline in catalytic products when the reaction temperature was lowered. The conversion results show clear sigmoidal trends. Common in enzymatic reactions (in Michaelis-Menten kinetics for instance) [112], sigmoidal curvature is

143 ethylcyclopentane methylcyclohexane 15 area% 48 area% 3-methylheptane 36 area% (a)

methylcyclohexane ethylcyclopentane 41 area% 3-methylheptane benzene 9 area% 31 area% 19 area% (b)

Figure 5.1: Major products from the thermocatalytic decomposition of benzene and toluene performed in the stirred batch reactor at 445◦C, 13.8 MPa H2, 1 h, 600 ppm Mo, 700 RPM, 3 wt% in decalin. Compositions indicate GCMS area percentage as shown in Tables 4.5 and 4.6. (a) Benzene. (b) Toluene.

718K ∆Gf = -338 kJ/molBenzene + 9 H2 6 CH4

718K ∆Gf = -405 kJ/molToluene + 10 H2 7 CH4

Figure 5.2: Thermodynamic simulations of benzene and toluene decomposition to methane simulated in Accelrys Materials Studio (v4.4) using DMol3 [110, 111] geometry optimisation and frequency calculations. Simulation details and parameters are provided in Section F.4. often indicative of cooperative systems. This may occur on the catalyst, whereby conversion at an active site is promoted by the adsorption of species on neighbouring sites, or in the reaction fluid, whereby the formation of specific species may promote the conversion of others. Such sigmoidal trends do not correspond with “standard” first or second order kinetic models, and attempts to fit such models

144 were unsuccessful (see Section E.1.2). Similarly Langmuir-Hinshelwood-Hougen- Watson kinetics were found to be unsuitable as changes in the conversion trends would suggest changes in the kinetic expression for the same catalyst, a conclusion which is contrary to the assumptions of this mechanism. Either or both of these cooperative effects may be at play in this system. On the catalyst surface, for example, dissolved H2 may need to adsorb and react in suf- ﬁcient quantities before DPM conversion can proceed at an appreciable rate (this applies to both the 0 and 600 ppm Mo systems as even 0 ppm Mo has access to the catalytic reactor walls). In the liquid phase, precursors may be necessary to initiate cracking reactions (as per the mechanisms for hydrogen and hydrocarbon radicals proposed by Curran et al. [78] and LaMarca et al. [88] in Figures 2.6b and 2.6c respectively) for a similar effect. Simulations (shown in Section F.7) performed in AspenTech Aspen Plus (v7.3) indicated H2 solubility in decalin to be approximately 25 mol% (0.4 wt%) and in DPM to be approximately 20 mol% (0.3 wt%) under reaction conditions (note that whilst the Peng-Robinson equation-of-state property method used in this simulation is not always accurate for the estimation of H2 solubilities in hydrocarbons, the results correspond well with experimental literature values [113–116] and are thus used for illustrative purposes). Given this high concentration of dissolved H2, its adsorption and reaction on the catalyst is unlikely to be responsible for the initial lag observed in the DPM conversion. It seems more probable that some hydrocarbon species must be responsible, the concentration being required to increase either on the catalyst surface or in the liquid to attain the observed increase in reaction rates. Given the increase in conversion with 600 ppm Mo, it is apparent that whatever the nature of this species and its mechanism, the catalyst plays a role in its formation and/or subsequent reaction. An additional point of interest in the DPM conversion results of Figure 4.2 is the plateauing of both the 0 and 600 ppm Mo conversion results after extended reaction times. This may be due simply to reduced DPM concentrations for the 600 ppm Mo, but a similar trend for 0 ppm Mo is observed at a lower conversion.

With H2 in vast excess and thus its limitation unlikely, this leveling off may indicate consumption of the rate-enhancing species either on the catalyst surface or in the liquid. The presence of decalin may also complicate the analysis. Hydrogenated species

145 have been shown to act as H shuttles (discussed in Section 2.3.2) and the decalin solvent may thus aid in the hydrogenation (and subsequent cracking and/or isomerisation) of the various species present. Following donation of H by decalin, re-hydrogenation may be promoted by the catalyst, hence the increase in rate with 600 ppm Mo. With the great excess of decalin in these diluted experiments, this contribution does not help to explain the leveling off observed in the DPM conversion. To better understand the mechanism, product yields must be examined. Given the great variety in other cracking, isomerisation and condensation products observed, product analyses for these diluted DPM experiments was limited to benzene and toluene, the major anticipated products. Together with these yields, the benzene:toluene (B:T) molar ratio served as a measure of continued reaction of the benzyl and phenyl radicals as per the mechanism of Figure 2.6a. Despite DPM conversions for 0 and 600 ppm Mo being the same upon reaching reaction temperature, their products are not (as shown in the product yield and ratio results in Figures 4.3 through 4.5). The 0 ppm Mo system exhibits greater yields of both benzene and toluene than does the 600 ppm Mo reaction with both series having B:T molar ratios of <1:1. An effective catalyst should either produce sufﬁcient H* to rapidly stabilise the benzyl and phenyl radicals with an ideal ratio of 1:1 [5, 29, 31, 46, 68, 75– 78] or perform catalytic hydrogenolysis to directly produce benzene and toluene in this ratio. As mentioned above, however, toluene reacts slightly more readily than benzene (and even forms benzene as it decomposes), resulting in a higher than expected B:T ratio. If we examine the benzyl and phenyl radicals, however, it is noted (as per the Gibbs free energy results shown in Figure 5.3) that phenyl radicals decompose to shorter radicals (methyl radicals in this example) more readily than do benzyl radicals (despite neither being spontaneous). This helps explain the trends observed. With 0 ppm Mo, the phenyl radicals decompose and the benzyl radicals remain to be stabilised to toluene, resulting in a lower B:T ratio. Increas- ing the catalyst concentration increases the rate of stabilisation and hence reduces the extent of phenyl radical decomposition, raising the B:T ratio. Whilst explaining the B:T ratios observed, this theory does not address some key trends.

146 • The 0 ppm Mo B:T ratio increases with increasing DPM conversion to a level comparable with the 600 ppm Mo reaction.

• The benzene and toluene yields decrease with increasing catalyst concentration (these should increase if the catalyst were promoting the stabilisation of the benzyl and phenyl radicals before they could decompose).

• The cracking and isomerisation product yields increase with increasing catalyst concentration per Table 4.2 (these should decrease by the same argument as above).

Although these few experiments do not provide enough data for a deﬁnitive mechanism to be presented, several conclusions may be drawn. The catalyst serves a hydrogenation role and, in doing so, promotes the formation of various hydrogenation, cracking and isomerisation products, thereby reducing the yield of benzene and toluene. As the concentration of these products increases, DPM conversion increases (perhaps per the mechanism proposed by LaMarca et al. [88] as shown in Figure 2.6c) and the rate of stabilisation of benzyl and phenyl radicals increases (for instance by radical addition or radical hydrogen transfer), increasing the yields of benzene and toluene and raising the B:T ratio. To clarify the DPM reaction mechanism, it is necessary to simplify the system by eliminating the H shuttle, decalin.

718K ∆Gr = 548 kJ/molPhenyl + 6.5 H2 6 CH3

718K ∆Gr = 750 kJ/molBenzyl + 7 H2 7 CH3

Figure 5.3: Thermodynamic simulations of phenyl and benzyl radical decomposition to methyl radicals simulated in Accelrys Materials Studio (v4.4) using DMol3 [110, 111] geometry optimisation and frequency calculations. Simulation details and parameters are provided in Section F.4.

147 Undiluted Diphenylmethane Figure 4.6 presents the conversion data for the undiluted DPM experiments. Apart from omitting the decalin diluent, the maximum reaction time was reduced to 6 h (a point where the plateau observed previously was well established) and 1800 ppm Mo catalyst loadings were included. Whilst the DPM maximum conversion was reduced to approximately 40 wt% at the reduced reaction time of 6 h, it should be noted that the conversion range under reaction conditions is equivalent to that of the diluted experiments. The undiluted DPM experiments reached reaction temperature with a conversion <5 wt% whilst the diluted DPM experiments, with maximum DPM conversions of approximately 80 wt% after 8 h, reach reaction temperature with approximately 30 wt% DPM conversion. The sigmoidal trends observed in the diluted system are still obvious but with a major change, starting at roughly the same DPM conversion upon reaching reaction temperature, the 0 and 600 ppm Mo systems diverge with 600 ppm Mo exceeding 0 ppm Mo after 1 h before re-converging to present comparable conversions after 6 h. 1800 ppm Mo exhibits the highest DPM conversion after 1 h. All 1800 ppm Mo experiments were conducted for 1 h but at different reaction temperatures to determine the impact of temperature on catalyst activity (as evidenced by changing product compositions, analogous to the DPP experiments discussed in Section 5.1.1). Unlike the diluted experiments, a first order kinetic fit now appears suitable for modeling of the 0 ppm Mo data (suggesting fewer other species affecting the reaction and perhaps a “cleaner” thermal product) whilst neither first nor second order models were found to accurately approximate the 600 ppm Mo results (the fits presented, for completeness, in Section E.1.3). XRD and TEM data (Figures 4.17 through 4.20) confirmed the formation of the desired MoS2 active phase. The average crystallite size was found to be less than 4 nm across (see the particle size distribution in Figure 4.21) and stacked to a thickness of less than three sheets (the stack height distribution provided in Figure 4.22). These narrow stacks would exhibit a high dispersion (the proportion of the active rim-edge atoms versus total atoms). Of interest in the XRD data are the FeS and C peaks present in the 600 and 1800 ppm Mo samples respectively. The FeS, which is also catalytically active in hydroconversion reactions [4, 8],

148 likely forms on the reactor walls and internals in the high temperature, sulphur- rich environment of the reaction and sloughs off into the liquid over time. This contaminant is certainly present in the 1800 ppm Mo sample too but given that the rate of FeS formation and sloughing is surely equivalent regardless of MoS2 loading (a great excess of CS2 being added to each reaction), the relative amount of FeS in the 1800 ppm Mo XRD sample would be less signiﬁcant. Another point of interest are the graphite peaks in the 1800 ppm Mo sample. Although small, these suggest the condensation of large polycyclic aromatic species, the formation of which from DPM would require extensive radical recombination. Returning to the DPM conversion data (Figure 4.6), another explanation for the change in the observed trend may be H starvation. In the diluted system, with 3 wt% model compound, the H2:DPM molar ratio was approximately 26 mol/mol

(Table 3.2). This, together with the H-donating decalin solvent, would make H2 supply a minor factor in the overall reaction. In the undiluted system, however, the volume of the reactor simply limits the amount of H2 present at the start of the reaction with a H2:DPM molar ratio of only approximately 0.8 mol/mol. This makes 100% conversion of DPM to benzene and toluene impossible as insufﬁcient

H2 is present for the reaction. This limiting factor may be responsible for the 600 ppm Mo DPM conversion leveling at the same point as the 0 ppm Mo. The lower conversions observed for 0 and 600 ppm Mo at 0 h in the undiluted DPM reactions, as compared to their diluted counterparts, may be due to the absence of decalin. From the diluted studies it was clear that DPM conversion resulted in high yields of hydrogenated products (which subsequently cracked and isomerised). Decalin, acting as a H shuttle [11, 21], would have promoted such hydrogenation reactions, allowing for thermolysis to occur readily even at lower temperatures. Without this H shuttle, the reaction is forced to hydrogenate species directly from dissolved H2 or await higher temperatures for thermolysis of the DPM and/or catalytic inﬂuence. To study this effect, the 1800 ppm Mo experiments were conducted at temperatures down to 415◦C as shown in Figure 4.7. While the true interest in this series was to compare product distributions with similar undiluted DPM experiments, an interesting segue is a comparison with the DPP temperature data. In such a comparison, the DPM conversion was observed to be far more temperature dependent with DPP conversion almost unchanged at

149 430◦C whilst DPM showed a loss of almost 30%. This is due to two effects: the DPP is more susceptible to thermolysis due to its three-member alkyl linkage, and the decalin acted as a H shuttle, promoting hydrogenation and hence conversion. Examining the products formed during these undiluted DPM hydroconversion experiments (Table 4.8) it is noted that the dominant products upon reaching reaction temperature are benzene and toluene. These species are present in roughly the same proportions for 0 and 600 ppm Mo and would suggest minimal catalyst activity below 445◦C but for the 1800 ppm Mo data. The data for this higher catalyst loading at lower temperatures shows clear evidence of hydrogenation and cracking products not seen in the other experiments and suggests that the catalyst, while proceeding at a slower rate at the lower temperatures, is still active. These reactions are also noted to be more selective with between two and six species present as >1% of the GCMS product area, likely due to decalin not participating as a H shuttle with the feasible reaction mechanisms thus limited. Of interest is that ﬂuorene is present in both the 0 and 600 ppm Mo samples and CHMB and 3-methylheptane are observed in the 1800 ppm Mo samples (recalling that 3-methylheptane was a product of both benzene and toluene decomposition as shown in Figure 5.1). The product yields for the undiluted DPM experiments are presented in Fig- ures 4.9 through 4.15. For benzene and toluene the high yields (corresponding with these being major products) are noted as is the maximum through which they pass at a DPM conversion of approximately 30 - 40 wt%. The benzene yield is seen to be suppressed at higher catalyst concentrations while the toluene yield is promoted. The B:T molar ratio shows an apparent relationship with catalyst concentration, declining as the amount of MoS2 increases, falling below 1:1 at 1800 ppm Mo. These trends suggest the catalyst to be preferentially inhibiting or con- suming benzene and forming toluene. These observations are inconsistent with the hydrogen activation mechanisms presented in Figures 2.6a and 2.6b [46, 78]. By these mechanisms, an increase in the catalyst concentration should affect an increase in the H* concentration. Per the mechanism of Curran et al. [78] ( 2.6a) this increased H* concentration would increase the rate of benzyl and phenyl radical capping, reducing the disparity in benzene and toluene yield with the B:T ratio tending toward 1:1. By the mechanism of Wei et al. [46] ( 2.6b), a higher H* concentration should promote benzene formation (the benzyl radical being susceptible

150 to continued cracking), resulting in an increase in the B:T ratio with increased catalyst concentration. These results do, however, appear to corroborate the mechanism of LaMarca et al. [88] (Figure 2.6c). With this mechanism requiring a supply of hydrocarbon radicals and with the catalyst clearly promoting DPM consumption, it is logical that the catalyst promotes the formation of the required hydrocarbon radicals. Unfortunately, the product distributions do not support this. Both 0 and 600 ppm Mo experiments show appreciable quantities of methyl- and ethyl-substituted DPM (the MBP and ETB species), but none of the required ipso-substituted species required for separation of the two rings. Despite these ipso-substituted species being formed as radicals (and hence susceptible to rapid cracking), some would be expected to stabilise and be observed in the product. Whilst this mechanism may play a role in the 0 and 600 ppm Mo experiments, the lack of either methyl- or ethyl-substituted species in any of the 1800 ppm Mo systems suggests that it is not promoted by the catalyst. The mechanism further fails to explain the presence of fluorene and/or hexahydroflourene (hexF) in many of the experiments. The presence of the numerous short-chain substituents does, however, suggest that hydrocarbon radicals are formed during the reaction and are added to the DPM rings. The apparent change in the reaction mechanism at a DPM conversion of 30 - 40 wt% is proposed to be due to the formation of a supercritical phase. Under reaction conditions, DPM is a liquid but benzene and toluene are not. Data from Afeefy et al. [117] indicate the critical temperature and pressure of DPM to be approximately 500◦C and 2.8 MPa, benzene to be 290◦C and 4.8 MPa and toluene to be 320◦C and 4.1 MPa. Figure 5.4 shows the proportion of each species in the vapour and the density of both the vapour and liquid phases with changes in DPM conversion. As DPM conversion increases, the density of the liquid phase declines (benzene and toluene being less dense than DPM) whilst that of the gas phase increases (benzene, toluene and DPM vapour being more dense than H2). Simulations indicate that at a DPM conversion of approximately 40 wt% the reaction mixture becomes a supercritical fluid. This promotes gas-liquid mass transfer

(the H2 and hydrocarbons forming a single phase) but reduces the effectiveness of the catalyst. Formation of the supercritical phase increases the reaction volume and hence decreases the overall catalyst concentration. Furthermore, a reduction in the density of the ﬂuid would result in less effective suspension of the solid catalyst

151 particles for a given mixer speed. The inﬂuence of these phase changes and their mass transfer effects on the reaction may be clearly seen in the benzene, toluene and isom./cond. product yields in Figures 4.9, 4.10 and 4.13 respectively. Both benzene and toluene yields decline upon formation of the supercritical phase whilst the yield of isom./cond. products increases, all three returning to levels similar to those observed at low DPM conversions. This is theorised to be due to the inability of the catalyst to effectively perform its role when the supercritical phase is formed (due to reduced solid-ﬂuid contact). Under these conditions, stabilisation of the benzyl and phenyl radicals (to toluene and benzene) formed by DPM decomposition occurs more slowly, with these radicals undergoing continued reaction to isom./cond. products as seen. For- mation of the supercritical phase is not believed to play a role in the plateauing of the DPM conversion results seen in Figure 4.6. The improved H2/hydrocarbon mixing is instead likely to promote DPM conversion as shown in the benzene diluent experiments presented in Table 4.1 (where operation under supercritical conditions was seen to promote DPM conversion by approximately 10%). To understand the mechanism, the yields of the lumped cracking and isom./cond. products must be considered together with that of a species prominent in the 1800 ppm Mo results, cyclohexylmethylbenzene (CHMB). One observation is the striking similarity between the yields for 0 and 600 ppm Mo, a trend most likely due to the activity of the reactor walls and internals. This complication aside, it is clear that whilst the catalyst suppresses the formation of isom./cond. products, consistent with the promotion of radical capping before they can react and recombine with one another, it promotes the formation of cracked products, inconsistent with such a theory. The catalyst was also seen to strongly promote the formation of CHMB, with the yield for 1800 ppm Mo greatly exceeding that of both 0 and 600 ppm Mo. Interestingly, the yields of cracking, isom./cond. and CHMB products decline with increasing conversion (while remaining in the liquid phase) as the yields of benzene and toluene rise. This suggests the initial reactions to be somewhat uncontrolled, proceeding by various slow mechanisms, and yet forming species which focus the mechanism to one which exacts the rapid production of benzene and toluene. This theory is supported by the sigmoidal conversion trends. An interesting theory to consider revolves around the identity of the isom./cond.

152 Benzene

Toluene

DPM

25 Proportion in

Supercritical vapour phase (%)

boundary

0 ) -3

0.3

Liquid

0.2

Vapour

0.1 Fluid densityFluid (g.cm

0 10 20 30 40

DPM conversion (wt%)

Figure 5.4: Simulated separation of reaction species to the vapour phase (the proportion of the total species in the system reporting to the vapour) and corresponding liquid and vapour densities with changing diphenylmethane conversion assuming equimolar benzene:toluene product. Per- formed in AspenTech Aspen Plus (v7.3). Simulation details and parameters are provided in Section F.5. species, namely fluorene and hexF. As shown in Figure 5.5, fluorene may be formed by the abstraction of H* from DPM, closing of the inter-ring bond (which stabilises the radical onto a tertiary carbon) and re-aromatisation by abstraction of a second H*. HexF may be formed in a similar manner from CHMB with the initial H* being abstracted from the saturated ring. It is also possible that fluorene, once formed, may be hydrogenated to hexF. Thus, while only small quantities of CHMB were detected in the 600 ppm Mo experiments, the presence of appreciable quantities of hexF betrays its participation in the reaction and suggests it to be extremely reactive. In the 0 and 600 ppm Mo systems, the radicals resulting from the abstraction of H* from DPM or CHMB may be stabilised by radical addition with hydrocarbon radicals present in the reaction mixture resulting in the methyl- and ethyl-substituted species observed. The 1800 ppm Mo system proceeds via

153 an additional step for, with ample catalyst, the DPM is hydrogenated to CHMB. H* may be more easily abstracted from the saturated ring of CHMB than from the DPM, minimising H* abstraction from the DPM and hence reducing the formation of ﬂuorene. Rapid re-hydrogenation, by the catalyst, of the CHMB radical formed by such a donation would also limit hexF formation (overall, hexF formation requires two H* be abstracted and its formation may hence be interrupted by re-hydrogenation of the intermediates). It is unlikely in the 1800 ppm Mo systems that H* would be abstracted from the DPM before regaining it from such species formed by the catalyst by the following logic. To regain H* before stabilisation by a hydrocarbon radical (to form MBP for instance which was not observed in the 1800 ppm Mo product) would suggest a very high concentration of such H* species on the catalyst surface or in the liquid. It would be easier for molecules requiring H* to react with those already formed by the catalyst rather than abstract them from DPM. Stabilisation of radicals by H* was shown in Section 5.1.2 to be unlikely (through an examination of benzene and toluene yields and ratios). The above proposed reactions were supported by thermodynamic simulations determining the Gibbs free energies of reaction as presented in Section F.4.

- H - H

H DPM Fluorene

+ 3H2 + 3H2

- H - H

CHMB H HexF

Figure 5.5: Proposed mechanism for the formation of fluorene and hexahydrofluorene from diphenylmethane. One possible reason for the abstraction of H* from DPM and CHMB, and the subsequent self-stabilisation of these molecules to fluorene and hexF respectively, may be the lack of H2 in the system. As discussed above, these undiluted DPM experiments could not achieve complete conversion due to H starvation. Examining the pressure measurement data, however, suggests that the 1800 ppm Mo system consumes more H more rapidly than either the 0 or 600 ppm Mo experiments. This not only supports the observation of the CHMB hydrogenation product but may in-

154 dicate that the systems do not reach a point where they would be inﬂuenced by H starvation (the rate of H2 transfer in the 1800 ppm Mo case is not being matched in the 0 and 600 ppm Mo reactions, indicating their rates of consumption in the liquid, and the subsequent concentration gradients for gas-liquid diffusion, are not as high and are not limiting the reaction). Based on these results, a new mechanism for the thermocatalytic hydroconversion of DPM may be proposed as shown in Figure 5.6. Here it is noted that the catalyst, rather than activating hydrogen as per the mechanisms of Figures 2.6a and 2.6b [46, 78], appears to hydrogenate the DPM to CHMB. This CHMB seems to serve the role of a H shuttle or perhaps as a precursor to the short chain hydrocarbon radicals, stabilising the benzyl and phenyl radicals to toluene and benzene and/or promoting DPM conversion as in the mechanism of 2.6c [88]. The DPM itself may act as a H donor to stabilise benzyl and phenyl radicals. Unfortunately, due to the slow heat-up rate of the reactor, the inﬂuence of the catalytically active wall, the possibility of H starvation effects and the supercritical phase formation, a more precise mechanism could not be substantiated from the data obtained in the 250 cm3 stirred batch reactor.

5.1.3 Summary of Model Compound Evaluation The data obtained from the stirred batch reactor was both extremely useful and enlightening. The DPM, DPE and DPP model compounds were evaluated, with DPE and DPP being found too reactive for use in this study. The decalin diluent, whilst not decomposing during the reaction, was shown to inﬂuence the conversion of the model compounds and their product distributions through its role as a H shuttle. Both benzene and toluene were seen to decompose, although to only a minor degree, and appeared to follow the same decomposition mechanism. In undiluted DPM, the catalyst precursors were conﬁrmed to form the desired

MoS2 active phase with a suitable particle size and distribution. Catalytic wall activity was conﬁrmed both through comparison of 0 and 600 ppm Mo data and the presence of FeS in the recovered solids. Mechanisms for the hydroconversion reaction currently proposed in the literature [21, 46, 78, 88] were seen to be inap-

155 H abstraction + +

Thermal Continued cracking, + isom., cond. reactions 2 2 + 3H Catalytic + 3H Catalytic 156 + + H abstraction 2 + zH Thermal Radical addition and cracking

C H + + CxHy x y Radical addition and Short chain radicals stabilisation

Figure 5.6: Proposed thermocatalytic decomposition mechanism of diphenylmethane from data gathered in the stirred ◦ batch reactor with 0 - 1800 ppm Mo, 415 - 445 C, 13.8 MPa H2, 0 - 6 h, 700 RPM. propriate and the beginnings of a new mechanism were proposed. Full details of this mechanism could not be reliably deduced due to various complicating factors of the reactor system including:

• Non-zero conversion upon reaching the reaction temperature due to slow heat-up rates (due itself to the size of the reactor),

• Complicated product compositions with various uncertain mechanistic routes due to either H starvation or the inﬂuence of the catalytically active walls,

• Obscuring of catalytic activity by the catalytic activity of the wall (speciﬁ- cally 0 ppm Mo versus 600 ppm Mo),

• An unacceptable level of experimental uncertainty (likely due to the above reasons),

• Achievement of a supercritical phase between 30 and 40 wt% conversion, beyond which the mechanism changes dramatically.

To further study this system it was necessary to accomplish several goals:

• Reduce reactor size for faster heat-up,

• Increase the H2:DPM ratio,

• Isolate the reaction from the reactor walls and internals,

• Reduce uncertainty / improve reproducibility,

• Avoid supercritical phase formation (most easily by remaining at lower conversions rather than changing operating temperature or pressure).

5.2 Novel Reactor System Design and Testing With data from the stirred batch reactor allowing DPM to be selected as a suitable model compound and for the development of an initial DPM hydroconversion mechanism, a novel micro-reactor system was designed to overcome some of the complicating factors and reﬁne the results.

157 5.2.1 Inclined Stainless Steel Micro-Reactor The inclined stainless steel micro-reactor allowed for data to be collected and analysed for preliminary comparison with that from the stirred batch system. This reactor was unmixed but its small size allowed for rapid heat-up (20 min compared to 80 min for the stirred batch system) and an improved H2:DPM molar ratio (2.7 compared to 0.8). Cool-down times showed slight improvements, the micro- reactor dropping to less than 400◦C in 60 s compared to 90 s for the stirred batch system. One downside of size reduction was an increase in the A:V ratio which is inversely proportional to the reactor diameter. With the stainless steel walls of the stirred batch reactor being shown to have a noticeable catalytic influence, and that with an A:V ratio of 1.3 cm2/cm3 (see Section F.8 for A:V diagrams and calculations), wall effects would be clear in the stainless steel micro-reactor with its A:V of 20.2 cm2/cm3, an increase of approximately 15.6 times. The first series of experiments conducted in this system was thus a study of the activation of the fresh stainless steel walls (of both the reactor and the thermocouple), the DPM conversion data for these tests being shown in Figures 4.23 and 4.24. These results indicated that for an 1800 ppm Mo system, the wall activity accounted for a conversion increase of 85±3% for all reaction times (except 0 h where both activating and stable systems showed negligible conversion). Using this result, a rough correlation could be determined whereby each unit of the A:V ratio increased the conversion of an 1800 ppm Mo system by approximately 4.2%. That is, each 1 cm2/cm3 corresponded to approximately 76 ppm Mo (an effective influence with the active phase present as FeS on the walls and internals of the reactor). The rapid heat-up rate of the micro-reactor was a success with all experiments in these systems shown to reach reaction temperature having undergone minimal DPM conversion. Comparison of this faster heating rate with one set to mimic that of the stirred batch reactor (DPM conversion and product yield results provided in Table 4.10) confirming this (although experimental uncertainty made only rough conclusions possible). In general it appeared that the slower heat-up rate did allow DPM to begin reacting before reaching reaction temperature. The products from this lower temperature reaction correspond well with those gathered in the stirred

158 batch reactor, with the results of the slower heat-up indicating a lower B:T ratio, a higher yield of CHMB and less isom./cond. products. With the walls of the reactor and thermocouple activated and stable, the experiments conducted in the stirred batch system were repeated in the inclined micro- reactor. The maximum reaction times were limited to 4 h so as to avoid supercritical phase change occurring at elevated DPM conversions. The DPM conversion results are shown in Figure 4.25. Comparison with the stirred batch reactor conversion data in Figure 4.6 shows the conversion achieved in the micro-reactor to be less than 10 wt% below that of the stirred batch reactor. Superficially this indicates a good correlation between the reaction systems, but when the far greater A:V ratio in the micro-reactor, and the greater conversion which should result, are considered, these lower conversion results reveal a more complicated system. Whilst both the 600 and 1800 ppm Mo systems show a rapid initial increase in DPM conversion with reaction time, exceeding that of the 0 ppm Mo reaction, all three catalyst concentrations were observed to converge after approximately 2 h and rise steadily thereafter. Furthermore, the 600 and 1800 ppm Mo systems appear almost identical with the 600 ppm Mo reaction possibly exceeding that of 1800 ppm Mo at some points during the reaction. To understand these trends, the product distributions must be considered (Fig- ures 4.26 through 4.31 and Table 4.11). The first clear trend is that the results of 0 and 600 ppm Mo are extremely similar in terms of benzene and toluene yields and the associated B:T ratio. This is certainly due to the high amounts of wall catalysis overwhelming the lower Mo loading. Many of the trends observed in the stirred batch reactor are repeated, with increased catalyst loading promoting CHMB and cracked species formation while inhibiting isom./cond. products. Curiously the toluene yield now declines with 1800 ppm Mo, a fact which pushes the B:T ratio for these experiments above 1:1. Despite the greater influence of wall catalysis, these reactions are markedly more selective, with even the most diverse mixture (0 ppm Mo after 4 h) showing only seven different species above 0.25 area% on the GCMS chromatogram. These well-defined product species, together with the increase in the yield of cracked products for 1800 ppm Mo above 20 wt% conversion may help explain the trends observed and add another piece to the mechanistic puzzle.

159 All of the reaction products now contain CHMB in appreciable quantities, confirming that this is a catalytic product resulting from the hydrogenation of DPM. All experiments also show hexF, the result of H* abstraction from CHMB and its subsequent stabilisation. Only small quantities of fluorene are detected and only for 0 and 600 ppm Mo after longer reaction times suggesting that H* abstraction from DPM is an undesirable alternative. This is supported by the reduced quantities of isom./cond. products in the 1800 ppm Mo experiments wherein sufficient CHMB is present to avert H* abstraction from DPM and its subsequent attack by hydrocarbon radicals to form MBP or ETB. The decline in the CHMB yield and the increase in the toluene yield with increasing conversion are related and suggest that CHMB decomposes during the reaction to form both toluene and a saturated hydrocarbon radical. Thermody- namic simulations, presented in Figure 5.7, support this mechanism with the ∆Gr associated with CHMB decomposition to benzyl radicals being favoured over that to phenyl radicals.

718K ∆Gr = 281 kJ/molPhenyl +

∆G 718K = 163 kJ/mol r Benzyl +

Figure 5.7: Comparison of Gibbs free energies of reaction for the thermolysis of cyclohexylmethylbenzene simulated in Accelrys Materials Studio (v4.4) using DMol3 [110, 111] geometry optimisation and frequency calculations. Simulation details and parameters are provided in Section F.4.

The fate of the cyclohexyl radical is unclear at this point. It does not appear in the product as cyclohexane and must thus decompose to short chain hydrocarbon radicals. The increased concentration of these radicals would, however, force additional H* abstraction from both DPM and CHMB (which may stabilise to ﬂuorene or hexF respectively) or radical addition to the DPM (to form MBP or ETB), resulting in an increase in the isom./cond. product yield. This is counter to the trend observed. It thus seems that these hydrocarbon radicals instead interact preferentially with other species in the reaction mixture, a mechanism which requires more

160 data to clarify. With the highest yields of CHMB of the three series, 1800 ppm Mo would be expected to show the highest toluene yield too. This is not the case with it instead showing the lowest yield. This trend is connected with the 3-methylheptane observed in all 1800 ppm Mo samples. At such elevated catalyst concentrations (1800 ppm Mo combined with the wall catalysis) and CHMB yields, two pathways are possible. Toluene may be hydrogenated and crack to 3-methylheptane but this would produce various other products not observed in these experiments (see the toluene blank test product composition results in Table 4.6). Alternatively, the less favoured CHMB to phenyl radical pathway may represent a signiﬁcant route of decomposition. Each mole of DPM decomposing via this pathway would not only produce a methyl-cyclohexane species (which would rapidly undergo thermolysis to 3-methylheptane) but it would simultaneously deprive the system of a mole of toluene whilst maintaining the production of benzene (hence the relatively unchanged benzene yields for different catalyst concentrations). One ﬁnal trend yet to be explained in the inclined micro-reactor is the convergence of the 0, 600 and 1800 ppm Mo DPM conversion results above 2 h reaction time. This trend is believed to be linked to the cracking product yields which are seen to be irregular for 600 ppm Mo and increase sharply above 2 h for 1800 ppm Mo. This micro-reactor was designed to allow volatile species (such as shorter cracked products) to leave the liquid phase, condense in the cold zone and run back. At an angle of only 30◦, however, the droplets required to form before running back would be quite large (note that the A:V ratio does not change as the bulk liquid volume declines as may be seen in the calculations provided in Section F.8). As suggested by the mechanism of LaMarca et al. [88] ( 2.6c), the cracking products leaving the liquid phase may promote decomposition of the DPM. The conversion convergence and the fate of the cyclohexyl radical now become clear. In- creased catalyst loads promote the formation of CHMB and its decomposition to benzyl and cyclohexyl radicals. These cyclohexyl radicals decompose into short chain hydrocarbon radicals and themselves promote the decomposition of DPM. This explains the more rapid initial rise in conversion for 600 and 1800 ppm Mo experiments. The short chain hydrocarbons do, however, leave the liquid phase and begin condensing elsewhere in the reactor, slowing the rate of consumption of

161 DPM and speeding the formation of more short chain radicals. When the droplets do eventually return to the bulk, the concentration of such cracking species is observed to rise again (most sharply in the case of 1800 ppm Mo samples as these include the 3-methylheptane from benzyl radical hydrogenation and cracking). Figure 5.8 summarises the proposed mechanism thus far together with the results from additional thermodynamic simulations. From the Gibbs free energies of reaction shown it is apparent that catalytic hydrogenation of DPM to CHMB and its subsequent thermal cracking is comparable to the thermal cracking of DPM. By this mechanism, the catalyst serves to hydrogenate DPM to form CHMB which undergoes thermolysis to either benzyl and cyclohexyl radicals or, less favourably, phenyl and methyl-cyclohexyl radicals (see Figure 5.7). In the former case, benzyl radicals stabilise to toluene (for instance by H* abstraction) while the cyclohexyl radicals crack to short chain hydrocarbons which promote the decomposition of DPM (although whether by DPM attack or product stabilisation is unclear). In the latter case, phenyl radicals stabilise to benzene while methyl-cyclohexyl radicals crack to 3-methylheptane (and possibly further to short chain hydrocarbon radicals as per cyclohexyl radicals).

5.2.2 Vertical Stainless Steel Micro-Reactor With the orientation of the inclined stainless steel micro-reactor appearing to affect operation, and to gather data for comparison with the glass insert micro-reactor, the stainless steel reactor was operated in a vertical orientation. This system used the same reactor body and thermocouple (with the activated walls) as the inclined unit. This change in orientation was shown to impact the gas-liquid transfer area (as seen in Section F.8.1, vertical orientation of the reactor halved the surface area from that of the 30◦ inclined orientation from 0.234 to 0.117 cm2) and condensed liquid run-back rate but not the A:V ratio (see calculations in Section F.8). From Figure 4.32 it is apparent that a change in the orientation affected a change in the DPM conversion. Beneficially, the DPM conversion after 4 h is now below the supercritical level. Interestingly, the thermal and catalytic systems now lie within experimental uncertainty of one another (the 1800 ppm Mo, surprisingly, being lower on average) and may be represented by first order reaction fits. The

162 718K ∆Gr = 218 kJ/molDPM H abstraction + +

Thermal + Continued cracking, isom., cond. reactions 2 2 + 3H Catalytic + 3H

Catalytic 718K ∆Gr = 56 kJ/molDPM

+ + H abstraction

∆G 718K = 163 kJ/mol 163 r CHMB Thermal

+ Radical addition and cracking 2 + zH Thermal

CxHy + + CxHy Radical addition and Short chain radicals stabilisation

Figure 5.8: Proposed thermocatalytic decomposition mechanism of diphenylmethane from data gathered in the inclined ◦ micro-reactor with 0 - 1800 ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM with simulations performed in Accelrys Materials Studio (v4.4) using DMol3 [110, 111] geometry optimisation and frequency calculations. Simulation details and parameters are provided in Section F.4. coefficient of this fit for 0 ppm Mo was seen to be comparable to that of the 0 ppm Mo experiments of the stirred batch reactor (0.082 h−1 and 0.088 h−1 respectively) and suggests the reason for the similarity of these trends to be H2 limitations. This reasoning fits well with the observed trends if rapid dissolution of H2 into but slow diffusion through the liquid phase are assumed. The inclined system, with a larger surface area, reaches reaction temperature with more dissolved H2 than does the vertical system. This accounts for the more rapid initial increase in DPM conversion in the inclined system and, as dissolved H2 is consumed near the gas-liquid interface, little can penetrate deeper into the liquid with the reaction being seen to slow at increased reaction times. The lesser gas-liquid surface area of the vertical system means less H2 is available upon reaching reaction temperature (this mimics the stirred reactor system where H2 was limited by bulk supply rather than surface area). The initial rate is thus lower than the inclined system and stays steady for all reaction times as H2 is consumed near the gas-liquid interface and is unable to diffuse deeper. In the vertical system with 1800 ppm Mo, additional consumption of

H2 by the catalyst to form hydrogen-rich products (such as CHMB) limits the supply even further, resulting in the slight decrease observed when compared to the 0 ppm Mo experiments. This argument relies on two assumptions: the dissolution of

H2 into the liquid is fast (resulting in the supply near the gas-liquid interface upon reaching reaction temperature) and the diffusion of H2 through the liquid is much slower (resulting in the confinement of H2 close to the gas-liquid interface and the relationship of the hydrogen concentration with the gas-liquid surface area). The former assumption is often made in the literature when working with a pure gas which is highly soluble in the liquid [116, 118–120], such as H2 has been shown to be in DPM in this study (recall that under reaction conditions, H2 can dissolve in DPM to approximately 20 mol%). In studies using batch systems, this rapid dissolution often makes quantification of the gas-side mass transfer coefficient (kGa) extremely difficult due to the complexities of measuring the rapid drop in gas pressure [119]. To support the use of this assumption in this work, experiments were conducted whereby the pressure of the micro-reactor, loaded with DPM and at 20◦C, was rapidly increased and monitored with time (the pressure being sampled to within 7 kPa [1 psi] every 1 s by the pressure transducer). The results are shown in Figure 5.9. Simulations into H2 solubility (simulation and

164 calculation details provided in Section F.7) showed an increase in solubility with temperature (although not the expected trend, this is supported by published studies [113, 114, 121, 122]), so these tests at 20◦C could be considered under-estimations of the rate of dissolution. The pressure test data showed that following pressurisation to 12.76 MPa, the gas pressure dropped by more than 82 kPa within 40 s.

This pressure change, by ideal gas assumptions, would result in a H2 concentration of approximately 14.0 mol% in the liquid (calculations detailed in Section F.7.2). Whilst additional factors may exaggerate this value (such as spring-back of the transducer mechanism following the rapid pressurisation or adsorption of H2 on metal surfaces), it is clear that the H2 does rapidly dissolve in the DPM (possibly saturating the liquid near the gas-liquid interface) and the assumption is valid.

12.76

12.74

12.72

12.70 Hydrogen pressure (MPa)

12.68

0 20 40 60

Time from maximum pressure (s)

Figure 5.9: Pressure drop after initial, rapid pressurisation with H2 in vertical micro-reactor loaded with 150 µL diphenylmethane at 20◦C to study the H2 dissolution rate. The latter assumption, that diffusion through the liquid is significantly slower than this dissolution, required that the H2 concentration along the vertical axis be modeled as a function of both time and liquid depth. Described in detail in Section F.7, this required estimation of the liquid-side mass transfer coefficient −9 2 (kLa) using the modified Wilke-Chang correlation [123] (1.09×10 m /s in DPM under reaction conditions), data from molecular simulations performed in Accelrys Materials Studio (v4.4) and Fick’s second law of diffusion to produce the concen-

165 tration proﬁles presented in Figure 5.10. It may be seen that even after 60 min, without consumption due to reaction, H2 penetration by diffusion is limited, even at a depth of only 1 cm (recall that the depth of 400 µL of liquid in the micro- reactor is 3.42 cm). This supports the assumption of liquid-side diffusion being slower than dissolution from the gas, thereby justifying the DPM conversion inter- pretations above. ) 3

1.5

Saturation concentration = 1.376 kmol/m

1.0

t = 15 min

t = 30 min

t = 45 min

t = 60 min 0.5

0.0

t = 15 min

0.10

t = 30 min

t = 45 min

t = 60 min

0.05

0.00 Hydrogen concentration in diphenylmethane (kmol/m

0.0 0.5 1.0

Distance from gas-liquid interface (cm)

Figure 5.10: Concentration proﬁles of H2 diffusing from saturated gas-liquid interface into diphenylmethane at 445◦C and 13.8 MPa as per Fick’s second law of diffusion using data from AspenTech Aspen Plus (v7.3) solubility simulations1 and the modiﬁed Wilke-Chang correlation2 [123] (with data from Accelrys Materials Studio simulations3).

1,2 - Simulation details, parameters and calculations provided in Section F.7. 3 - v4.4 using DMol3 [110, 111] geometry optimisation and frequency calculations as shown in Section F.4.

Whilst the DPM conversion data indicated the dependence of the system on H2 diffusion rates through the liquid, the product distributions in the vertical stainless steel reactor (Figures 4.33 through 4.37) provided additional detail regarding the mechanism of the reaction. Although the B:T molar ratio for both 0 and 1800 ppm Mo experiments remained roughly equivalent and constant (slightly below 1:1) for all DPM conversions, the 1800 ppm Mo system exhibited lower yields of both benzene and

166 toluene than were observed for 0 ppm Mo. This suggests the primary route for benzene and toluene formation to be the same for both systems. Both systems showed an increase in benzene and toluene yield with increasing DPM conversion. The reduced toluene yield with 1800 ppm Mo was previously attributed to catalytic hydrogenation of DPM to CHMB and its subsequent decomposition to phenyl and methyl-cyclohexyl radicals. This pathway was characterised by 3-methylheptane in the reaction product (the result of methyl-cyclohexyl thermolysis). In the vertical reactor, however, only trace amounts of other cracking products were detected in either the 0 or 1800 ppm Mo systems, this despite the CHMB yields being in the same order of magnitude as those of the inclined reactor. Furthermore, the only isom./cond. product appearing in measurable quantities was ﬂuorene, with hexF being noticeably absent from the analyses. The ﬂuorene yield trends for both 0 and 1800 ppm Mo were seen to be equivalent, suggesting the formation of this species to be independent of catalytic activity. This is a plausible assumption given that its formation likely occurs via the abstraction of H* from DPM by hydrocarbon radicals and that the ∆Gr for such single-step abstraction by benzyl and phenyl radicals is only slightly higher than with a short chain radical (such as a propyl radical) acting as an intermediate (as seen in Figure 5.11).

718K ∆Gr = 68 kJ/molDPM + +

718K ∆Gr = 56 kJ/molDPM + +

718K ∆Gr = 12 kJ/molPropane + +

Figure 5.11: Comparison of Gibbs free energies of reaction for the abstraction of H* from diphenylmethane by benzyl and propyl radicals simulated in Accelrys Materials Studio (v4.4) using DMol3 [110, 111] geometry optimisation and frequency calculations. Simulation details and parameters are provided in Section F.4.

Despite quantiﬁcation of the gaseous products indicating their representation of only a minimal mass product stream (gas product compositions shown in Table 4.13), their composition is indicative of the liquid reaction occurring. As seen in the

167 stirred batch reactor, the addition of MoS2 catalyst not only increases the yield of gaseous products, but also shifts the spectrum toward larger species (particularly those in the C5 and C6 range as shown in Figures 4.38 and 4.39). This, together with the higher yield of CHMB in the 1800 ppm Mo experiments, supports the theory of CHMB decomposition to short chain hydrocarbon radicals. It is proposed that CHMB forms rapidly at low conversions (catalysed by both the active walls and MoS2, hence the slight increase in the 1800 ppm Mo system) and begins to decompose as previously discussed. The limited H2 available in the liquid limits the amount of CHMB formed and as decomposition progresses, so this initial concentration declines. The yields of CHMB eventually tend toward the pseudo-stable levels observed in the inclined system as the reaction shifts to relying on H2 as it dissolves into the liquid rather than the initially higher amount accumulated during heat-up. The short chain hydrocarbon radical products from this CHMB decomposition serve to stabilise the benzyl and phenyl radicals from DPM thermolysis. Whether this occurs by radical addition (for instance a phenyl radical and a methyl radical to form toluene), radical hydrogen transfer (such as from propane to a phenyl radical to form benzene and a propyl radical) or by radical disproportionation (a propyl radical and a phenyl radical reacting to form benzene and propene, the latter being re-hydrogenated by the catalyst), examples depicted in Figure 5.12, is unclear. Several such mechanisms may be responsible and some may even work in series (propane, for instance, proceeding through a chain reaction of the second, third and forth mechanisms before terminating by radical addition). As these radicals form and are stabilised, they may vaporise and leave the liquid mixture but, due to the orientation of the reactor, rapidly condense and return. For this reason, the concentration of short chain hydrocarbons in the liquid remains higher than was observed in the inclined system (where larger droplets had to form and run back into the bulk reaction mixture). This consistently higher concentration of short chain hydrocarbons and their radicals has three key effects. Firstly, the stabilisation of the benzyl and phenyl radicals from DPM is promoted, hence the higher yields of benzene and toluene than in the inclined reactor. Secondly, this supply of stabilising species reduces the extent to which H* abstraction from DPM and CHMB occurs, decreasing ﬂuorene yields and all but eliminating hexF formation (ﬂuorene production having a lower

168 overall ∆G f than hexF), a trend also observed in the inclined system. Finally, the higher concentration of these short chain species in the liquid promotes the reverse recombination reaction of their formation from CHMB. This limits the extent to which CHMB decomposition occurs and favours that reaction with the lower ∆Gr, speciﬁcally the formation of benzyl and cyclohexyl radicals. This accounts for the negligible yield of 3-methylheptane (the methyl-cyclohexyl radical is an un- favoured product), the B:T ratio below 1:1 (toluene is preferentially formed, albeit slightly) and the lower benzene and toluene yields in the 1800 ppm Mo experiments (a higher yield of CHMB means a lower yield of benzene and toluene).

718K ∆Gr = -311 kJ/molPhenyl + CH3

718K ∆Gr = -77 kJ/molPhenyl + +

718K ∆Gr = -274 kJ/molPhenyl + +

718K ∆Gr = -45 kJ/molPropene + H2

Figure 5.12: Comparison of Gibbs free energies of reaction for various phenyl radical stabilisation mechanisms as simulated in Accelrys Ma- terials Studio (v4.4) using DMol3 [110, 111] geometry optimisation and frequency calculations. Simulation details and parameters are provided in Section F.4.

5.2.3 Unmixed Glass Insert Micro-Reactor Progressing from the vertical stainless steel micro-reactor experiments to the use of the glass insert micro-reactor allowed for a determination of the inﬂuence of the active reactor and thermocouple walls. The glass insert micro-reactor, using the same activated thermocouple as the vertical stainless steel unit, was shown to have an A:V ratio of 4.0 cm2/cm3, down from the 20.2 cm2/cm3 of the stainless steel micro-reactor (see Section F.8) but still higher than the stirred batch reactor at 1.3 cm2/cm3. As is clear from the DPM conversion and product yield results in Figures 4.40 through 4.45, this decrease in the A:V ratio had a notable impact on

169 both the observed conversion and product distribution. Presented in Figure 4.40, the DPM conversion using the glass insert was observed to be adequately modeled by first order kinetics. A comparison of the kinetic coefficients between the glass insert micro-reactor and the vertical stainless steel micro-reactor (Tables 4.14 and 4.12 respectively) indicated a rate decline of 53±3% in the 0 ppm Mo reactions and 42±2% in the 1800 ppm Mo reactions. This difference in changes between the two micro-reactors accounts for the 1800 ppm Mo DPM conversion data now lying above that of the 0 ppm Mo system. Unlike the stainless steel systems, these trends no longer lie within experimental uncertainty. The reason for the decline in conversion is due to the decline in the A:V ratio. Reduced contact with activated stainless steel surfaces reduces the effective catalyst concentration from these sources from approximately 1500 ppm Mo in the stainless steel micro-reactor to approximately 300 ppm Mo with a glass insert (at 76 ppm Mo per cm2/cm3 as shown in Section 5.2.1). It should be noted that the change in DPM conversion with changing A:V ratio is not linear and suggests that even small quantities of catalyst are able to promote the decomposition of DPM. With this greater difference in the effective catalyst loading between the 0 and 1800 ppm Mo experiments, the product yields begin to reveal a more accurate mechanism. Foremost is that only trace quantities of CHMB were detected in the 0 ppm Mo products, a predictable result based on the deductions of this being a catalytic hydrogenation product. The benzene and toluene yields and the B:T ratio now all exhibit differences between 0 and 1800 ppm Mo experiments. At low DPM conversions, the 1800 ppm Mo system rapidly hydrogenates DPM to CHMB which subsequently decomposes to benzyl and cyclohexyl radicals. This not only promotes the formation of toluene (as evidenced by an initially higher toluene yield for 1800 ppm Mo), but simultaneously reduces the formation of benzene. It is proposed that the benzene yield is suppressed due to a double effect. The first is a reduction in the thermolysis of DPM to benzyl and phenyl radicals due to its hydrogenation to CHMB. The second is the formation of short chain hydrocarbon radicals from this CHMB which may stabilise the phenyl radicals to alkyl-substituted species which may then crack to benzyl radicals and stabilise to toluene. As the reaction progresses, such short chain hydrocarbon radicals are formed in sufficiently large quantities in the 0 ppm Mo system too and the toluene yield for that system

170 rises. The benzene yield for 1800 ppm Mo stays suppressed as the higher catalyst concentration continues to produce CHMB. This theory is supported by the isom./cond. product yield which indicates a greater suppression of these reactions in the 1800 ppm Mo experiments, a trend explained in Section 5.2.2. The lack of signiﬁcant cracking products in both the glass insert micro-reactor and the vertical stainless steel micro-reactor support the theory proposed surrounding the mechanisms of Figure 5.12. By this theory, the short chain hydrocarbon radicals (formed predominantly from CHMB and hence promoted by the catalyst) perpetuate a radical chain reaction, stabilising other radicals in the system and hence limiting their own concentration.

5.2.4 Mixed Glass Insert Micro-Reactor With an understanding that the mechanism is limited by the liquid-side diffusion of H2 (the dissolution being comparatively fast), agitation of the system to increase gas-liquid contact and H2 distribution through the liquid was deemed necessary to understand the mechanism without H starvation limitations and to ensure solid suspension (a factor which would be crucial when working with coke-catalyst agglomerates). With the vortex mixing of a slurry-phase micro-reactor having no precedence in the literature, determination of the nature of the mixing and quantiﬁcation of its effectiveness were conducted in two ways. The ﬁrst was by visual inspection of the reaction mixture at different mixing speeds. The second, by analysis of the impact of mixing on the DPM hydroconversion reaction.

Visual Mixing Studies A “conventional” stirred tank is designed with the ratio of liquid depth, H, to vessel width, D, of approximately 1.0 [119, 120]. Bafﬂes may be used with the standard width of D/10. The impeller of such a system would be positioned H/3 from the bottom of the vessel [120]. By these standards, the glass insert of the micro-reactor used in this study would certainly be “unconventional” [119] with H/D = 62.5. Deﬁnitions aside, the complexity of designing and operating such a long, narrow impeller were found prohibitive. Albal et al. [119] studied the use of various mix-

171 ers in conventional and unconventional vessels and determined that, with impeller mixing, the mixing time for gas-liquid and liquid-liquid systems was reduced in the unconventional arrangement, likely due surface breakage and “surface entrainment” occurring at lower mixing speeds [119, 120]. Albal et al. [119] also noted improved homogeneity and dispersion for gas-liquid-solid systems. Whilst not directly applicable to this study where no impeller was employed, these results were encouraging. Of interest is that the “critical stirrer speed” for surface breakage [119] and solids suspension [119, 124] was determined by visual studies using glass vessels. A similar methodology was employed in this study using a glass mock-up of the micro-reactor system and a sample of reaction product as described in Section 4.2.3. Vortices in liquids are well known and understood [125] and may be viewed as rotating masses of fluid with minimal axial or radial mixing, properties which usually make them indicative of poor mixing. This, however, is usually only true in larger vessels with powered mixing where the velocity of a particle increases proportionally to its distance from the central axis, a rotational or rigid-body vortex. In a confined system, such as the glass insert, friction with the walls would inhibit particle motion away from the central axis, resulting in a velocity gradient and formation of an irrotational vortex. The velocity gradients in such a vortex would create sheer forces between inner and outer layers of fluid, moving at different rotational velocities about the central axis, and promote liquid-liquid and liquid- solid mixing. Formation of the vortex core would also increase gas-liquid exchange by increasing the transfer area. Whilst this was the theory supporting the development of a vortex-mixed reactor, the analysis of mixing in true vortices in this study is moot. As may be seen from the high speed camera stills in Figures 4.46 through 4.50, vortex formation does not occur in this system, at least not under ambient conditions. Instead it appears that the cohesive forces of the liquid mixture exceed the centrifugal forces of the vortex mixer and result in the formation of a rotating wave. Whilst the radial depth of this wave was found to remain roughly constant at 1.7 mm, the height of the wave/“vortex” was almost linearly related to mixer speed (over the range investigated). At the maximum sustainable mixer speed of 2500 RPM, the “vortex” height was approximately 28.5 mm, insufficient to agitate the liquid and suspend

172 the solid particles. Reducing the liquid volume to 150 µL allowed the entire liquid volume to be drawn up into the wave at 2500 RPM but 2000 RPM was found to be the critical speed at which particle suspension occurred. Implementation of the glass insert was not, however, simply liquid in a tube but rather included the central thermocouple. Studying the effect of this thermocouple on mixing was extremely enlightening with it appearing to act as a stirrer bar. With 400 µL of liquid a mixing speed of 2500 RPM was required for surface entrainment whereas with 150 µL, entrainment began at 2000 RPM, appeared to peak at 2250 RPM but at 2500 RPM had been replaced by the familiar wave of liquid. Although simply visual confirmation of mixing, these results suggested a significant degree of gas-liquid-solid agitation. The use of actual reaction product for these visual mixing studies ensured a representative solid and liquid properties (density, viscosity, particle size, etc.). These tests were, however, conducted at ambient conditions (approximately 20◦C and 101.325 kPa). It was thus necessary to determine how the physical properties of the mixture, and hence mixing dynamics, would change at reaction conditions (445◦C and 13.8 MPa) and with changing DPM conversion. With the solid density and particle size assumed to be constant, interest was focused on the liquid with the density, viscosity and surface tension being of particular interest [119, 120]. The results of physical property simulations performed in AspenTech Aspen Plus are presented in Table 5.1. The density, viscosity and surface tension are all seen to decrease when DPM is raised to operating conditions and further decrease with increasing DPM conversion. Changes in these properties all influence the efficacy of mixing. Albal et al. [119] showed increases in kLa with decreasing viscosity and surface tension. These trends were linked to surface entrainment (which promotes gas-liquid contact) which occurs more readily in less viscous fluids with lower surface tensions. The reduced viscosity and density also reduces mixer power requirements or, as is this case in this study, where the power remains fixed, increases the extent of agitation. This is clear if the Reynolds number (Re), shown in Equation 5.1, is examined. The kinematic viscosity, ν = µ/ρ, of DPM at ambient conditions is approximately 80 times larger than at reaction conditions. With Re inversely proportional to ν, this means that with all other factors being equal, the changes in density and viscosity would result in an approximately 80

173 fold increase in the Reynolds number, indicating significantly more turbulence and agitation of the fluid. An additional effect of the density changes is that, with a fixed mass of fluid, the volume will increase by approximately 2.5 fold when raised to reaction conditions and continue to do so as the reaction progresses.

ρ.v.DH v.DH Re = µ = ν (5.1) It should, however, be noted that the characteristic dimension for the glass insert micro-reactor is only 0.003 m (3 mm) and an evaluation of liquid agitation using Re is likely not appropriate as adhesive interactions between the liquid and the vessel walls will impose a laminar regime even at high mixing speeds. More applicable to mixing analyses in a system of this size would be the Schmidt (Sc) and Sherwood (Sh) numbers shown in Equations 5.2 and 5.3 respectively. Sc provides a quantification of the ratio between momentum diffusivity and mass diffusivity whilst Sh quantifies the ratio between convective mass transfer and diffusive mass transfer. Given the small dimensions of the micro-reactor, it was anticipated that the role of molecular diffusion in overall mixing of the system would be significantly greater than in the larger stirred batch reactor (i.e. smaller Sc and Sh numbers). This would imply that convective mixing would have only a minor effect on the reaction at all but the highest mixing speeds.

ν Sc = (5.2) D

K .L Sh = Sh Sh (5.3) D With detailed modeling of the mixing within this reactor beyond the scope of this project, the most efﬁcient means of optimising the mixer speed was to conduct experiments to evaluate its inﬂuence on DPM conversion and product distribution as presented below.

Effect of Liquid Volume Prior to studying the effect of mixer speed on the reaction, it was necessary to determine how reduction of the loaded liquid volume from 400 to 150 µL (visually

174 found to be necessary for agitation of the entire liquid volume) would affect the system in terms of DPM conversion and product distribution. This volume reduction would have no impact on the A:V ratio (see Section F.8) but would increase the H2:DPM ratio by a factor of approximately 2.6 and, perhaps more importantly, would reduce the liquid depth in the vertically oriented glass insert from 3.4 cm to only 1.3 cm. This would increase the proportion of liquid reaction mixture which is accessible to H2 diffusing down through the unmixed liquid from the gas-liquid interface, resulting in a more hydrogen-rich reaction environment. The results from this comparison are shown in Table 4.15. It is clear that hydrogen plays a key role in both DPM conversion and product distribution, with this data lending additional conﬁrmation to the mechanisms proposed above. The DPM conversion at 150 µL is seen to be approximately double that of 400 µL for both 0 and 1800 ppm Mo experiments. The benzene and toluene yields and B:T ratio remained mostly unchanged for the 1800 ppm Mo system, whilst with 0 ppm Mo an increase in the toluene yield was accompanied by no change in the benzene yield (with a subsequent decline in the B:T ratio). This indicates that the formation of toluene is favoured with greater hydrogen supply as is the formation of CHMB which was seen to increase signiﬁcantly with a reduction in the liquid volume. These trends support the proposed mechanism of DPM hydrogenation to CHMB with its subsequent thermolysis to benzyl and cyclohexyl radicals, the benzyl radical being stabilised to form toluene. Additional consumption of DPM to CHMB with 1800 ppm Mo manifest as the lower benzene and toluene yields observed in those experiments. Whilst no increase in the yield of other cracking products in the liquid phase was observed, the isom./cond. product yield was found to decline with a decrease in liquid volume. Once more this supports the proposed theory whereby the cyclohexyl radicals decompose to short chain hydrocarbon radicals which participate in radical chain reactions to stabilise other radical species in the mixture.

Thermocouple Wall Activation

With the sulphided stainless steel thermocouple sheath being the only non-MoS2 catalyst present in the glass insert 0 ppm Mo experiments, an investigation was

175 conducted to observe a reaction as close to truly thermal as possible by installing a fresh thermocouple and monitoring the DPM conversion and product distribution over sequential 0 ppm Mo experiments. The DPM conversion and product yield results are presented in Figures 4.51 through 4.55. Similar to the trend observed in the stainless steel micro-reactor wall activation experiments (discussed in Section 5.2.1), the DPM conversion rises rapidly with each experiment as the thermocouple sheath becomes sulphided and catalytically active. From an initial conversion of <2 wt%, the system rises to approximately 8 wt% over the course of six experiments before plateauing. This more rapid stabilisation in wall activity is believed to be due to the smaller A:V ratio (4.0 cm2/cm3 with the glass insert micro-reactor versus 20.2 cm2/cm3 with the stainless steel micro-reactor) resulting in a greater excess of CS2 and hence more extensive sulphidation with each successive experiment. With only a single datapoint at the lowest DPM conversion, the best representation of a truly thermal reaction, a discussion of trends is difﬁcult. It does, however, appear that the initiation steps of DPM decomposition are strongly supported by the catalytic hydrogenation of DPM to CHMB. This CHMB then decomposes to benzyl radicals (the stabilisation of which manifest as elevated toluene yields and a suppressed B:T ratio) and short chain hydrocarbon radicals. These short chain species then promote the decomposition of DPM by either destabilisation by radical attack or more rapid stabilisation of thermolysis products, thereby inhibiting reverse recombination reactions. Ev- idence for both mechanisms is observed as the amount of catalyst grows and the production and consumption of CHMB increases. The higher concentrations of short chain species appear to be able to stabilise phenyl radicals from DPM thermolysis, raising the benzene yield, but the increase in the isom./cond. product yield supports the mechanism of their destabilisation of DPM (by H* abstraction or addition, such as shown in the mechanism of Figure 2.6c).

Effect of Mixing With the unmixed glass insert micro-reactor experiments completed and a rough guide for the desirable mixing speed determined by visual observations, it was necessary to evaluate the reaction over a range of mixing speeds to determine

176 Table 5.1: Physical property simulation results for for diphenylmethane hydroconversion mixtures at both ambient conditions and under reaction conditions for interpretation of visual mixing results. Performed in AspenTech Aspen Plus (v7.3).

Conditions Properties 1 177 DPM conversion Temperature Pressure Density, ρ Dynamic viscosity, µ Surface tension, γ (wt%) (◦C) (MPa) (g/cm3) (Pa.s) x105 (N/m) x103 0 20 0.1 1.01 300 38 0 445 13.8 0.39 1.6 1.1 19 445 13.8 0.28 1.0 0.4 Simulation details and parameters are provided in Section F.6. 1 - Simulations were only conducted to a DPM conversion of 19.4 wt% as above this value surface tension calculations did not converge. the optimal speed for the ﬁnal active and deactivated catalyst evaluation studies. The DPM conversion and product yield results of these tests are presented in Fig- ures 4.56 through 4.60. Increased mixing speeds, corresponding to improved gas-liquid-solid mass transfer (greater gas-liquid interface, better liquid-liquid mixing and solid particle suspension), would be anticipated to result in an increase in DPM conversion, leveling as the reaction transitions from mass transfer limitations to kinetic restrictions. The observed trends do not support this theory with both 0 and 1800 ppm Mo showing minimal change up to 1500 RPM followed by a decline in DPM conversion with a further increase in mixing speed. These trends support the use of Sc and Sh to quantify mixing as it appears that very high mixing speeds are required for convective mixing to present to any signiﬁcant degree. The product distributions, however, are supported by the proposed mechanism. The 1800 ppm Mo experiments show lower benzene but higher CHMB yields than do the 0 ppm Mo experiments. This is indicative of DPM hydrogenation to CHMB (and preferentially to toluene) rather than the thermolysis of DPM to equimolar benzyl and phenyl radicals. This does not, however, explain why the 1800 ppm Mo toluene yield trend follows smoothly from that of 0 ppm Mo. The logical reasoning would be that the rate limiting step is that of CHMB decomposition which now progresses at the same speed in both systems. This rate results in the highest toluene and lowest benzene yields observed thus far, minimal isom./cond. products and yet comparatively low CHMB yields. This maximised rate of CHMB decomposition is likely due to improved mixing whereby homogenisation of the liquid allows the short chain hydrocarbon radicals (formed, for instance, from cylohexyl radical thermolysis) to rapidly attack or stabilise other species, including regulating the reversible decomposition reaction and hence toluene formation by this pathway. The higher concentration of CHMB also allows, as is the case for 1800 ppm Mo at 1500 RPM, decomposition to phenyl and methylcyclohexyl radicals and hence the observed formation of the 3-methylheptane. Regarding the mechanism whereby the short chain radicals promote DPM conversion, the mechanism of LaMarca et al. [88] presented in Figure 2.6c seems unlikely. Were this the reaction occurring, the concentration of short chain radicals would be depleted as they are consumed through addition reactions. This would

178 result in more rapid consumption of CHMB as, by Le Chatelier’s principle [126], the reaction is driven forward by the consumption of the products. This would result in higher toluene yields (due to higher concentrations of benzyl radicals) in the 1800 ppm Mo system rather than the contiguous trend observed between it and 0 ppm Mo. Despite mixing being theorised to promote CHMB thermolysis, the DPM conversion is observed to decline at higher mixing speeds. This phenomenon is proposed to be due to two effects, one involving the hydrodynamics of mixing and the other catalyst morphology. As observed in the high speed camera stills in Figure 4.50, mixing speeds above 1500 RPM result in extensive surface entrainment. This reduces the contact between the liquid and active thermocouple surface, reducing the extent to which it may catalytically promote the reaction. It is for this reason that the DPM conversion and product distribution results for 0 ppm Mo at different mixing speeds so closely mimic those of the thermocouple activation tests. As the effective catalyst loading in 1800 ppm Mo experiments is a combination of unsupported MoS2 and the thermocouple wall, a decrease in liquid contact with the wall decreases the effective catalyst loading. As such, the difference between the 0 and 1800 ppm Mo data in Figure 4.56 should be a representation of the MoS2 activity without thermocouple wall catalysis. This activity should increase with increasing mixer speed as gas-liquid-solid mass transfer is promoted. This trend is not observed. Following a rise in DPM conversion with mixer speed, reaching a maximum at approximately 2000 RPM, the conversion of the 1800 ppm Mo experiments declines as mixing speed continues to increase. Clues to this secondary decline came from visual inspection of the reaction product. Whereas the product from unmixed 1800 ppm Mo experiments presented as ﬁne particles clouding and colouring the liquid, those from higher mixing speeds appeared as reﬂective platelets (much like glitter) in colourless liquid. These observations led to an extensive analysis of the recovered solid material by XRD, SEM-EDX and FESEM with the results presented in Figures 4.61 through 4.68.

Analogous to the conﬁrmation of MoS2 formation in the stirred batch reactor, inter- sheet and d-spacing measurements from TEM images again conﬁrmed that the Mo octoate precursor decomposed in-situ to form the desired active phase. Results

179 from 0 h experiments further indicating that this had been achieved before reaching reaction conditions. SEM-EDX results further support this, with Table 4.16 indicating the solid recovered from 0 RPM experiments to be comprised almost entirely of Mo and S with Figure 4.63a showing the atoms of these two elements to occur in close physical proximity to one another (supporting that they are present as MoSx). Samples from 2000 RPM experiments, however, were observed to contain both Fe and Ni together with the Mo and S. The Fe and Ni were likely present as sulphides, the former showing close physical correlation with S (Figure 4.63b), similar to those seen in the XRD results from the stirred batch reactor presented in Figure 4.17. These would have formed on the thermocouple sheath (316 stainless steel containing 16.5 - 18.5 wt% Cr, 10.5 - 13.5 wt% Ni, 2.0 - 2.5 wt% Mo and balance Fe [127]) in the sulphur-rich reaction environment and sloughed off into the liquid due to the extreme agitation.

The structure of these MoS2 crystallites, however, differ significantly from those observed in previous experiments. In the 0 RPM systems FESEM and TEM imagery shows the narrow (3 - 4 nm), short (less than five layers) stacks of MoS2 sheets arranged in a rag structure to form loose collections of small particles. Both the dimensions and structures of these particles are comparable to those observed in the stirred batch reactor. These structures are observed at 0 h (Figure 4.64e), indicating that they form and agglomerate during heat-up, through to 4 h (Figure 4.64c) with little change apparent to their dimensions with reaction time. With a mixing speed of 2000 RPM, however, the morphology of the catalyst changes significantly. At 0 h (Figure 4.65e), TEM imagery indicates growth of the crystallites in terms of both width and stack height compared to 0 RPM after the same time (Figure 4.62a versus 4.61a), the particles clearly beginning to agglomerate into ordered structures, forming larger flat sheets as seen by FESEM (Figure 4.65e). This continued through 2 h (Figure 4.65f) with the structures after 4 h shown by FESEM and TEM (Figures 4.65c and 4.62b respectively) to comprise stacks of more than ten sheets, some in excess of 100 nm wide. Angling the sample during FESEM analysis re- vealed the striking images of Figure 4.68 wherein stacks of MoS2 sheets tightly agglomerate and fuse together to form much larger structures. These increases in stack height and sheet width together with the obstruction of some crystallites by others after agglomeration reduce the number of active rim/edge atoms accessible

180 to the reaction liquid for a given catalyst loading. This results in a decline in the “dispersion” and reduces the overall efficacy of that catalyst. It is for this reason that the conversion with 1800 ppm Mo declines at higher mixing speeds. The formation of larger crystallites may be thought to occur by one of two mechanisms, nucleation or agglomeration/fusion. In the 0 RPM system, with Mo octoate distributed through the liquid feed, small MoS2 crystallites precipitate in relative isolation, depleting the local Mo octoate concentration as they do so (depicted in Figure 5.13a). With agitation, MoS2 crystallites precipitating in pockets of slightly higher local concentration may act as nucleation centers. As these particles formed, they deplete the local Mo octoate concentration but, with agitation, are quickly exposed to additional dissolved precursor (shown in Figure 5.13b). Additional precipitation may then occur on these first crystallites, forming larger sheets than observed in the unmixed system. By agglomeration/fusion, the precipitation of MoS2 crystallites occurs with sufficient speed that they form in isolation regardless of mixing. Once formed in the 0 RPM system, these particles settle slowly and, for the most part, never make contact with one another due to liquid layers being maintained between individual particles (Figure 5.13c). In an agitated system, the particles are forced into contact (see Figure 5.13d), agglomerating as they do so due to the same van der Waal’s forces that hold the sheets of each stack together. Once in such close proximity the electronically defective edge atoms of two separate sheets are able to fuse together to form a larger, lower energy sheet (which happens to be less catalytically active). Which of these processes occurs in the system may be deduced from the FE- SEM and TEM results, particularly the images in Figures 4.66 and 4.67. These figures show the solid material recovered after 1 h for a system which was heated without mixing, the 2250 RPM mixing only begun upon reaching reaction temperature, and one which was heated with mixing, as is the standard procedure. As may be seen, the results show very little difference in catalyst morphology, suggesting that small crystallites form rapidly in isolation before being brought into contact with one another. This is further supported by the TEM results which, for 2000 RPM in Figure 4.62, show that what begin as small crystallites upon reaching reaction temperature have agglomerated and fused after 4 h. This “morphological deactivation” (the reduction in the ratio of active rim-edge

181 (a) (b)

Figure 5.13: Conceptual diagram of MoS2 crystallite precipitation and movement in glass insert micro-reactor and the effect of mixing. (a) Un- mixed MoS2 precipitation with local precursor depletion. (b) Mixed MoS2 precipitation with initial particles moving to area of higher precursor concentration. (c) Unmixed MoS2 settling showing particle separation by liquid and inhibited agglomeration. (d) Mixed MoS2 crystallites showing collisions due to mixing.

sites to the bulk MoS2 with increasing sheet size) does not appear to occur in the stirred batch reactor where, even with mixing and 8 h reaction time, no such effects were observed. Two theories may be proposed for this lack of particle growth in the stirred batch system. The ﬁrst is postulated to be due to the “cleanliness” of the glass insert micro- reactor. Close contact between the MoS2 crystallites would be necessary for the

182 observed agglomeration and fusion to larger particles to occur. In the glass insert micro-reactor, the MoS2 crystallites are limited as to what solid particles are available for collision in the liquid. This results in virtually all collisions being

MoS2-MoS2, resulting in agglomeration into large, pure MoS2 structures as observed. In the stirred batch reactor, previous residue hydroconversion experiments have resulted in the deposition of unreactive graphitic species on the walls of the reactor and internals. Particles of these species slowly slough off into the liquid during reaction (as seen in the XRD analysis for 1800 ppm Mo in this system in

Figure 4.18), reducing the probability of MoS2-MoS2 collisions and resulting in agglomerates containing both MoS2 and various other solid species. Such “collision obstruction particles” likely inhibit MoS2 agglomeration in residue hydroconversion reactions too (be these solids introduced with the feed or coke forming during reaction), ruling MoS2 morphological deactivation out as a suspect in such systems. This theory seems unlikely though as a significant concentration of obstructing particles would be necessary to significantly influence the purity, and hence fusion, of the MoS2 agglomerates. The second theory relates to the nature of mixing in each of the reactors. In the stirred batch reactor, the liquid is mixed as a large, moving bulk with the MoS2 particle concentration remaining roughly equivalent throughout. This reduces the contact between MoS2 particles and hence their agglomeration and fusion. In the glass insert micro-reactor, however, the mixing results in the formation of thin films as seen in the high speed camera stills in Figure 4.50. These films could result in elevated local concentrations of MoS2 particles, promoting agglomeration and fusion.

5.2.5 Summary of Micro-Reactor System Design and Testing Data obtained from the studies conducted in the inclined and vertical stainless steel micro-reactors were useful in determining crucial aspects of the reaction mechanism despite these not being the ﬁnal micro-reactor design. From these unmixed systems with high A:V ratios, additional details into the catalytic mechanism could be deduced, often the result of the very limitations that made these systems ultimately unsuitable for studying the deactivated coke-catalyst agglomerates (such

183 as H starvation). This included a rough “effective catalyst” measure for activated stainless steel walls whereby every unit of A:V ratio represented approximately 76 ppm Mo. The results provided conﬁrmation of CHMB as a catalytic hydrogenation product. Decomposition of CHMB to benzyl and cyclohexyl radicals was seen to be reversible and dictated by the concentration of short chain hydrocarbon radicals formed from cyclohexyl radical thermolysis. These short chain hydrocarbon radicals served to stabilise benzyl and phenyl radicals from DPM thermolysis in a radical chain reaction mechanism. This stabilisation was found to be a necessary step to inhibit the recombination of benzyl and phenyl radicals to DPM. In the absence of any catalyst to form CHMB and hence the stabilisation radicals, DPM conversion is minimal. High concentrations of short chain hydrocarbon radicals, however, were found to abstract H* from DPM and CHMB, resulting in the formation of ﬂuorene and hexF respectively.

H2 was found to be a limiting reagent in unmixed systems due to the depth of liquid and slow rate of diffusion with reduced liquid loading volumes (150 µL instead of 400 µL) promoting the overall conversion of DPM through a promotion of CHMB formation. A reduced liquid loading volume was also found to be necessary to ensure adequate agitation of the system. Externally applied vortex mixing was found to be very effective. It did not, as anticipated, establish a vortex in the liquid but instead a circulating concave wave. In the presence of the central thermocouple, this wave was disrupted with extensive surface entrainment and particle suspension being visually observed at ambient conditions. Simulations indicated that gas-liquid-solid mixing efﬁciency would improve under reaction conditions and as conversion increases. Agitation was found to have two unexpected effects. Firstly, surface entrainment reduced liquid contact with the thermocouple, reducing the A:V ratio and hence DPM conversion for both the 0 and 1800 ppm Mo experiments. Secondly, increased contact between MoS2 crystallites in such a “clean” system resulted in extensive agglomeration and fusion, increasing the stack height and sheet width in a process termed “morphological deactivation”. The optimum mixing speed for a catalytic reaction, reducing liquid contact with the thermocouple whilst avoiding the worst of the morphological deactivation, was determined to be 2000 RPM.

184 Experiments using this optimum mixing speed could now be used to reﬁne and further develop the DPM hydroconversion mechanism proposed thus far (presented in Figure 5.8), determining the inﬂuence of the active MoS2 catalyst. The study could then progress to the deactivated coke-catalyst agglomerate from residue hydroconversion with MoS2 and hence propose a deactivation mechanism.

5.3 Catalyst Study and Deactivation Investigation

5.3.1 Active MoS2 With a proposed mechanism for DPM hydroconversion supported by experiments in multiple reactors and under various conditions, the culminating series of active catalyst experiments could be conducted in the glass insert micro-reactor at the optimal mixing speed. The DPM conversion and product composition and yield results from these experiments, comparing 0 and 2000 RPM data, are presented in Figures 4.69 through 4.74 and Tables 4.17 and 4.18. It is clear that, as initially observed in the stirred batch reactor data of Figure 4.2, the DPM conversion data matches neither ﬁrst nor second order kinetics and is instead a sigmoidal trend. This trend is easily explained by the proposed mechanism. CHMB is required for rapid DPM conversion with the short chain hydrocarbon radicals from CHMB thermolysis stabilising the benzyl and phenyl radicals from DPM thermolysis, thereby driving that reaction forward. CHMB is a catalytic product which is formed easily with either 0 or 1800 ppm Mo in the unmixed systems with good contact between the liquid and active thermocouple wall. In the mixed systems, CHMB forms more rapidly with 1800 ppm Mo than 0 ppm Mo.

This is because the former still possesses catalyst in the form of suspended MoS2 particles, whilst reduced contact with the active thermocouple wall signiﬁcantly reduces the catalytic effect in the latter. CHMB, once formed, decomposes rapidly before establishing a stable concentration regulated by the short chain hydrocarbon radicals. This is observed as a rapid rise in the 0 ppm Mo DPM conversion to attain levels comparable with those achieved with 1800 ppm Mo. These trends are supported by the CHMB yields which indicate 0 ppm Mo rising from negligible levels at 2000 RPM before leveling. CHMB for 1800 ppm Mo shows a rapid ini-

185 tial decline with 2000 RPM, which corresponds to the CHMB being more quickly consumed with improved mixing, before also leveling. CHMB yields are lower for both 0 and 1800 ppm Mo at 0 RPM as the lack of mixing inhibits the extent to which the short chain hydrocarbon radicals can affect the reverse recombination reaction. Benzene and toluene yields and the B:T ratio also follow anticipated trends. The toluene yields for all four experimental series rapidly converge to almost identical values. Beginning at comparable levels, the benzene yields for 0 and 1800 ppm Mo experiments at 2000 RPM begin to decline from their 0 RPM counterparts, with 1800 ppm Mo remaining slightly below 0 ppm Mo in both cases. These yields result in the slightly lower B:T ratios observed for the mixed experiments with 1800 ppm Mo being marginally below 0 ppm Mo. The reasons for these changes are due to the source of the benzene and toluene. Whilst benzene forms predominantly from DPM thermolysis, toluene forms from both this reaction and from the decomposition of CHMB. The 1800 ppm Mo systems are able to produce CHMB, and hence toluene, more rapidly during the initial stages of the reaction. As the reaction continues, CHMB production in the 0 ppm Mo mixed experiment increases along with toluene yields. As hydrogenation of DPM to CHMB increases, so DPM thermolysis to benzyl and phenyl radicals (and ultimately toluene and benzene) decreases. This does not inﬂuence toluene formation but does have a negative impact on benzene yield which is seen to decline in the mixed experiments as improved gas-liquid-solid mass transfer allows improved CHMB production and consumption. These changes in the benzene yield result in the decline in the B:T ratio observed for the 2000 RPM systems with the value below 1:1 indicative of more toluene than benzene being formed. Examining the yields and identities of the isom./cond. products also lends support to the proposed mechanism. In the unmixed systems, the reverse recombination reaction of CHMB is not inhibited to as great an extent as in the mixed systems (due to poor contact between the short chain hydrocarbon radicals required to recombine). This allows these short chain radicals to abstract H* from DPM and CHMB to form the various species observed (ﬂuorene and hexF from DPM and CHMB radical stabilisation respectively and ETB from stabilisation of radical DPM by a benzyl radical [in higher concentrations due to CHMB thermolysis]). In

186 the 2000 RPM systems, such isom./cond. species only present after long reaction times. This corresponds with gradually rising toluene levels and suggests the benzyl radicals from CHMB thermolysis, required for the reverse reaction, are slowly stabilised before recombination. This leaves a gradually rising concentration of rogue short chain hydrocarbon radicals. Gaseous product analyses for these experiments support these mechanisms. In unmixed experiments, short chain hydrocarbon radicals stabilise (for instance by H* abstraction from DPM) and escape to the gas phase prior to continued reaction. Poor gas-liquid contacting in these systems results in minimal return of these species to the liquid phase. An increase in the catalyst loading is seen to increase the amount of these gases. The dominant increase occurs after only 1 h, indicative of the source of these species being CHMB decomposition. The higher yields of larger hydrocarbons for 1800 ppm Mo experiments is likely an indication of the mechanism of the radical chain reaction for, in such systems, the catalyst is able to hydrogenate unsaturated fragments, stabilising them and allowing them to escape to the gas phase before undergoing thermolysis. In mixed systems, stabilised hydrocarbon fragments which might otherwise escape to the gas phase are rapidly exposed to other radicals and may undergo continued reaction, reducing their ap- pearance in the gaseous product. One curious trend is the decrease in gaseous products with reaction time observed for 1800 ppm Mo at 2000 RPM. This is believed to be due to the extreme surface entrainment allowing stable hydrocarbon species to be re-dissolved from the gas phase and continue reacting. The ﬁnal thermocatalytic DPM hydroconversion mechanism as it occurs in the mixed, glass insert micro-reactor may thus be presented as per Figure 5.14

5.3.2 Deactivated Coke-MoS2 Agglomerate

From the active MoS2 studies above, the optimal mixing speed and reaction time for evaluating the deactivated coke-catalyst (coke-MoS2) agglomerates from residue hydroconversion reactions would be 2000 RPM and 1 h. This combination was shown to offer the greatest difference between 0 and 1800 ppm Mo results and should be best for determining the coke-catalyst activity. The results for these tests and comparative results for residue hydroconversion tests are shown in Table 4.19.

187 Benzyl radical Phenyl radical

+ +

Toluene Benzene

Radical chain reaction stabilisation Thermal

DPM + CxHy

Fluorene 2 2 + 3H + 3H Catalytic Catalytic

CHMB + CxHy

HexF

H abstraction at high Thermal Cyclohexyl radical concentrations radical

Benzyl radical 2 + zH Thermal

CxHy Short chain radicals

Figure 5.14: Proposed thermocatalytic decomposition mechanism of di- ◦ phenylmethane at 445 C, 13.8 MPa H2.

In the residue hydroconversion reactions in which the coke-catalyst samples examined in this study were generated (performed by Rezaei and Smith [1]), two of the quantiﬁers of catalyst activity were coke yield and H2 conversion. An active catalyst exhibited high H2 conversions and low coke yields whilst a deactivated catalyst showed lower H2 conversions and increased coke yields. Per the hydrogen activation mechanism, widely held in the literature to govern such systems [5, 29, 31, 46, 68, 75–78], these trends may be explained as the catalyst activating dissolved H2 to H* which then caps the radicals formed by the thermolysis of species such as asphaltenes before they can undergo recombination reactions and precipitate as coke. Despite a coke yield (wt%) increase of almost 400% between

188 the active and deactivated catalysts (more than 700% as compared to 0 ppm Mo), however, the difference in H2 conversion is far smaller. The active catalyst consumes only between 20% and 30% more H2 than the deactivated specimens and less than 40% more than a 0 ppm Mo reaction. This indicates that H2 is being incorporated into the reaction mixture even without the addition of MoS2 catalyst but that it is not having a significant influence on the condensation reactions. Before examining the DPM hydroconversion results, it is thus necessary to understand the residue hydroconversion reaction in light of the mechanism developed in this study. For clarity, the asphaltene model compound bibenzyl-cholestane (BBCh), studied by Alshareef et al. [2], is used as an example. Of specific interest is why in 0 ppm Mo experiments only 40% less H2 conversion occurs as compared to fresh MoS2 and yet the coke yield is more than 700% higher. It is theorised that even in the 0 ppm Mo reactions, hydrogenation of such a polyaromatic feed molecule is promoted by other catalytically active species (such as FeS on the reactor walls and internals or other metal sulphides formed from metals in the residue feed). It is proposed that whilst these low catalyst concentrations are sufficient to promote partial hydrogenation of the aromatic residue (resulting in an appreciable H2 conversion), this hydrogenation does not occur to an extent required to significantly increase the decomposition of these species to short chain hydrocarbon radicals. Such radicals would still be formed as the alkyl substituents and some of the saturated rings of the BBCh undergo thermolysis. This would leave the partially hydrogenated polycyclic core of the molecule behind. Such cracking reactions, and the stabilisation of the resultant radicals, requires the addition of hydrogen which may be made available by species acting as H shuttles (much like the role performed by decalin in the stirred batch reactor) or, as seen in the undiluted DPM experiments, through the abstraction of H* from the remaining portions of the molecule, the partially saturated core. Now radicalised, the core may stabilise internally to form a more aromatic molecule or may condense with other radicals (resulting in the elevated coke yields observed). The short chain hydrocarbons formed at the end of the process may escape as gases.

Addition of MoS2 does not reduce the cracking of the alkyl substituents, with these still forming short chain hydrocarbon radicals. Instead, the catalyst serves to promote additional hydrogenation of the aromatic core of the BBCh (although not

189 significantly greater, it does appear sufficient). This reduced aromaticity allows extensive portions of the BBCh to reversibly decompose to short chain hydrocarbon radicals, establishing a mechanism similar to that observed for CHMB decomposition. Together with the hydrogenation ability of the catalyst, these radicals may participate in chain reactions to stabilise other radicals in the system rather than abstracting H* from the BBCh cores or simply escaping as gases. Hydrogenation of the BBCh molecule would also allow it to better function as an H shuttle and, if H* is abstracted from such a species and it does condense, its higher hydrogen content may avoid precipitation as solid coke, allowing it to remain in the liquid phase and continue rapidly reacting (likely cracking once more). Upon deactivation of the coke-catalyst in residue hydroconversion, whether by heat treating or repeated recycling, the reactions appear to reach similar coke yield levels which are less than that of a 0 ppm Mo reaction. It is possible that this is an indication that one of the hydrogenation roles of the catalyst has been inhibited, it is either no longer able to hydrogenate the asphaltenic species or it is no longer able to hydrogenate short chain hydrocarbons and thereby promote the radical chain stabilisation reactions. Specific reasons aside, it is clear that both the heat treated and recycled coke-catalyst agglomerates have lost a significant amount of their coke suppression activity and may be considered deactivated.

Figure 5.15: Bibenzyl-cholestane, an asphaltene model compound proposed by Alshareef et al. [2]. Evaluation of the coke-catalyst samples by DPM hydroconversion in the glass insert micro-reactor suggest activity trends which do not agree with those of residue hydroconversion. From the DPM conversion results it appears that the heat treated

190 catalyst retains full activity (its conversion equivalent to active MoS2) whilst both the fresh and recycled catalyst are deactivated, with DPM conversion below that even attributed to the thermocouple wall activity. The benzene, toluene and CHMB yields were found to correlate well with the trends observed for active MoS2 at different conversion levels and suggest that rather than the active sites changing (by poisoning for instance), there are simply fewer of them available for reaction. Ex- amination of the gas products from these reactions indicated that whilst the fresh and heat treated coke-catalyst agglomerates produced roughly equivalent amounts of gaseous hydrocarbons, similar to those of active MoS2 under the same conditions, the recycled agglomerate exhibited a signiﬁcant increase in these species. Of interest is that whilst the active catalyst showed a trend toward larger gaseous hydrocarbons, the products from these reactions greatly favour methane. Rezaei and Smith [1] showed that a portion of the coke introduced with the coke-catalyst agglomerate re-dissolves and reacts. Table 5.2 shows the solubility of each of the coke-catalyst agglomerates in DPM (the methodology for these tests presented in Section F.9). Note that these tests were conducted at relatively low temperatures for ease of handling and to avoid any possibility of thermal reactions. As may be seen, the coke solubility increases with temperature and is thus predicted to be even greater at the operating temperature of 445◦C. It is also clear that the heat treating removed all soluble coke whilst recycling the catalyst has resulted in the generation of large amounts of soluble coke. Interestingly, despite the solubility of this coke, no other cracking or isom./cond. species were detected in the liquid products. The DPM hydroconversion results may thus be explained. The coke of the coke-catalyst agglomerates dissolves in the DPM model compound and reacts. This coke is comprised of large polycyclic species with numerous alkyl substituents which hydrogenate more easily than DPM and hence out-compete the model compound for catalyst active sites. This includes those sites on the thermocouple wall, hence the DPM conversions fall below previous 0 ppm Mo experiments. A similar competition phenomenon was demonstrated by Yu et al. [128] with Fe-S/ZSM-5 in di(1-naphthyl)methane (DNM) and DPM. The researchers observed that whilst, separately, both DNM and DPM react under the given conditions, in a mixture of the two species, only DNM undergoes reaction. In the case of the heat treated

191 Table 5.2: Solubility of coke-catalyst agglomerates, from residue hydroconversion, in diphenylmethane at 20 - 100◦C.

Coke-catalyst DPM temperature Amount dissolved type 1 (◦C) (wt%) 20 3.2±0.3 Fresh 100 3.8±0.3 20 0.0 Fresh, heat treated 2 100 0.0 20 1.0±0.1 Recycled 100 25.8±2.1 1 - Coke-catalyst agglomerate samples courtesy of Rezaei and Smith [1]. 2 - “Fresh” catalyst heat treated under 100 sccm He at 700◦C for 15 h. coke-catalyst sample, the soluble/reactive coke has been removed and the DPM experiences no competition for the catalyst sites, hence conversion and product distribution much alike that of active MoS2. In the case of the recycled agglomerate, multiple reactions with decreasing catalyst activity have resulted in large amounts of coke being deposited. The alkyl branches of this coke react readily to form short chain gaseous hydrocarbons but little decomposition of the polycylic coke structures occurs, hence the lack of other liquid species.

5.3.3 Mechanism of MoS2 Deactivation in Residue Hydroconversion These observations for the activity of various coke-catalyst agglomerates in residue and model compound hydroconversion reactions may be used to propose a deactivation mechanism for unsupported MoS2 in residue hydroconversion reactors. Catalyst deactivation in residue hydroconversion is an observed phenomenon indicative of a decline in the interaction between the active sites of the catalyst and the various liquid species. This may be due to various processes (see Section A.1), some which alter the chemistry or morphology of the active sites and some which simply obscure them. The results from this study indicate that the active sites have not been altered (for instance by poisoning), with DPM model compound evaluation of the heat treated coke-catalyst agglomerate possessing activity comparable to fresh unsupported MoS2. That the same coke-catalyst exhibits reduced residue hydroconver-

192 sion and yet unchanged DPM hydroconversion suggests physical restrictions to active site access. This theory is further supported per the discussion of Section 5.3.2 whereby catalysts deactivated by heat treatment or multiple recycles exhibit near identical coke yields and H2 conversions which are distinct from the thermal results. This implies that not only is the mechanism of deactivation in both catalysts similar (perhaps the same) but that these deactivated catalysts are still able to promote some hydrogenation pathways (proposed to be the hydrogenation of short chain hydrocarbons in support of the radical stabilisation chain reactions). The reasoning for the fresh coke-catalyst agglomerate being active in residue hydroconversion and yet exhibiting minimal DPM conversion activity is due to the preferential reactivity of soluble coke introduced with the agglomerate. This coke out-competes the DPM for both catalyst active sites, including those of the thermocouple walls, (for hydrogenation and subsequent thermolysis) and for short chain hydrocarbon radicals (for H* abstraction and product stabilisation).

The process whereby unsupported MoS2 crystallites are deactivated by this theory is depicted in Figure 5.16. During the ﬁrst residue hydroconversion reaction, the precursor reacts to form nanometer-sized MoS2 crystallites (Figure 5.16a). These unsupported catalyst particles are able to freely access all species in the liquid, from asphaltenes to short chain hydrocarbon radicals. Coke inevitably forms during the course of this reaction (Figure 5.16b) and begins agglomerating with the catalyst particles. This coke is not homogeneous and is often categorised as either “soft coke” or “hard coke” (described in Section A.1.1). Soft coke, most of which is soluble and reactive under reaction conditions but precipitates during cooling and catalyst recovery, acts as a precursor for hard coke. Hard coke is solid and unreactive under reaction conditions. As the coke-catalyst agglomerate is recycled, not only is the hard, unreactive coke from the previous experiment re-introduced to the system but the soft coke in the agglomerate speeds the formation of additional hard coke. These, together with the soft and hard cokes which form as they did in the ﬁrst experiment, contribute to a growing proportion of coke in the coke-catalyst agglomerate (Figure 5.16f). Heat treating of the coke-catalyst agglomerate serves to vaporise or react the majority of soft coke. This simultaneously reduces the amount of soluble coke in the agglomerate (as shown in Table 5.2) whilst increasing the amount of hard coke present [1] (Figure 5.16d). This theory is not novel

193 and was most recently supported by the work of Rezaei and Smith [1], with the conceptual model presented in Figure A.1e. That deactivation mechanism, however, suggested that the MoS2 crystallites become completely occluded in the hard coke, unable to interact with the system. The present study proposes that the hard coke instead forms a porous carbonaceous structure upon which the MoS2 crystallites become supported. As the coke- catalyst agglomerate is recycled, so the extent of this carbonaceous support grows to incorporate all of the MoS2 crystallites (Figure 5.16f). A similar effect occurs during heat treating wherein the soft coke is forced to react to form hard coke (Figure 5.16d), expanding the support structure in a manner nearly identical to that formed by repeated recycles (hence the similarities in coke yield and H2 conversion for the heat treated and recycled agglomerates in Table 4.19). These hard coke structures do not, however, occlude the MoS2 crystallites with molecules as large as DPM able to enter the pores and react uninhibited. The pores are, however, suf- ﬁciently small so as to exclude the large asphaltenes from reaching the active sites.

Ultimate deactivation of the now supported MoS2 catalyst would likely occur in a manner akin to that of other supported catalysts, pore fouling and catastrophic loss of surface area and active site access (depicted in Figures A.1c and A.1d).

The recovery of soft coke together with the hard coke supported MoS2 masks the catalyst activity in the fresh and recycled coke-catalyst agglomerates. Shown in Figures 5.16c and 5.16g, the soft coke redissolves in the DPM hydroconversion reactions and out-competes DPM, masking the catalyst activity. As all soft coke is removed from the heat treated agglomerate, this masking does not occur when testing that sample in DPM hydroconversion (Figure 5.16e). This insight into the formation of a porous carbonaceous support of hard coke which physically excludes large species from reaching the otherwise active MoS2 crystallites presents numerous possibilities for catalyst regeneration. As per the full occlusion models, inhibition of hard coke formation is key to limiting deactivation. This may best be achieved by removing the soft coke from the coke-catalyst agglomerate, for instance by solvent extraction, prior to recycling. With hard coke forming in each reaction regardless of soft coke reintroduction, entrapment of the

MoS2 in the porous support is inevitable. This does not, however, necessitate removal of the hard coke. A process as simple as mechanical crushing may be suf-

194 ﬁcient to destroy the porous structure and re-expose the active sites for access by larger species. Alternatively some manner of chemical etching may serve to en- large the pores or otherwise destabilise the support.

Figure 5.16: Proposed conceptual mechanism for the deactivation of unsupported MoS2 in residue hydroconversion reactions.

5.3.4 Summary of MoS2 Activity and Deactivation With catalytic wall effects minimised, the mechanism for the thermocatalytic hydroconversion of DPM in the presence of unsupported MoS2 was determined. This reaction relies on the catalytic hydrogenation of the DPM to CHMB and its subsequent reversible thermal decomposition to short chain hydrocarbon radicals. These radicals stabilise the products of DPM thermolysis, driving that reaction forward, in a radical stabilisation chain reaction which is promoted by the catalyst hydro- genating unsaturated intermediates. Evaluation of various coke-catalyst agglomerates from residue hydroconver-

195 sion reactions indicated the DPM model compound reaction to be a “worst case” measurement technique with soft coke (soluble in DPM under reaction conditions) out-competing the DPM. Despite this additional factor, the deactivation of unsupported MoS2 in residue hydroconversion by encapsulation by hard, unreactive coke as proposed in the literature [1] was conﬁrmed. The morphological features of this carbonaceous obstruction were, however, determined to be different. As opposed to solid coke fragments completely enclosing MoS2 crystallites, it was found that the hard coke instead forms a porous structure upon which the otherwise unchanged crystallites are supported. The pore size of this support is sufﬁciently large to allow unrestricted access by short chain hydrocarbons or DPM, but all but prevents access by larger species such as asphaltenes. It was proposed that catalyst deactivation may be inhibited by removal of soft coke prior to recycling of the coke-catalyst agglomerate but that the inevitable formation of the hard coke support required some form of chemical or mechanical destruction to re-expose the active MoS2 crystallites.

196 Chapter 6

Conclusions

This study began with three key questions in support of two hypotheses (Chapter 1). After extensive experimental investigations using both a commercially available stirred slurry-phase batch reactor and novel slurry-phase micro-reactors developed specifically for this study, these key questions were answered and the hypotheses supported. It was determined that sufficient information can be obtained from model compound experiments to accurately determine the mechanism of unsupported MoS2- catalysed hydroconversion. Diphenylmethane (DPM) was found to be a suitable model compound, its simplicity of structure allowing for the analytical resolution required to elucidate details of the reaction mechanism. DPM was also shown to be more difficult to crack than many of the species comprising “soft coke” (the carbonaceous deposits which re-dissolve and react when re-introduced to the reactor) in the coke-catalyst agglomerate, making it a chemically simplistic model compound for the representation of the more stable (to both thermal and thermocatalytic reactions) species in residue feedstocks. Numerous factors beyond the superficial reaction conditions of temperature, pressure, reaction time, etc. were found to have significant contributions toward the observed feed conversion and product distribution. These included: the presence of a hydrogen shuttle such as decalin (which promoted DPM conversion), reactor heat-up rate (slower rates allowing thermal reactions to proceed before the catalyst was fully active) and reactor and internal component construction mate-

197 rials (316 stainless steel being shown to sulphide and become catalytically active, contributing to DPM conversion and product distribution). A novel mixed slurry- phase micro-reactor with glass insert was designed and shown to suitably overcome many of these effects whilst introducing one of its own. The vortex mixing regime implemented in this clean system was shown to result in signiﬁcant MoS2 agglomeration and crystallite growth, reducing DPM conversion.

Using DPM as a model compound, a new mechanism of MoS2-catalysed hydroconversion was proposed from data from the various reactor systems used in this study. It was determined that the various mechanisms proposed in literature, most purporting the role of the catalyst to be the splitting of dissolved H2 to an activated H* species, are inaccurate. This study proposes that the unsupported MoS2 instead hydrogenates the DPM to form cyclohexylmethylbenzene (CHMB). This CHMB then undergoes thermolysis to form short chain hydrocarbon radicals which stabilise other radicals in the system by a mechanism of chain radical stabilisation, a process supported by the hydrogenation ability of the catalyst. This mechanism was used to successfully explain trends from residue hydroconversion studies with unsupported MoS2 in the literature. Using this model compound testing methodology, the activity of coke-catalyst agglomerates, deactivated to varying extents in residue hydroconversion, was evaluated. The mechanism of deactivation was proposed to be due to the formation of a porous carbonaceous structure by “hard coke” (insoluble, unreactive, graphitic deposits) with the MoS2 crystallites becoming supported on this structure. The pore size was sufﬁciently small to physically exclude access by larger species in residue hydroconversion (such as asphaltenes) but large enough to allow uninhibited access by DPM. Methods for overcoming this deactivation were proposed as either solvent extraction of “soft coke” between coke-catalyst recycles (removing hard coke precursors rather than re-introducing them to the reaction) or destroying the porous support by either mechanical size reduction or chemical etching.

198 Summary Summarising in the context of the original hypotheses, a mechanistic understanding of diphenylmethane hydroconversion was developed for active and deactivated unsupported MoS2 catalysts. This mechanism allowed for the deduction of the process of deactivation of such catalysts in residue hydroprocessing reactions. The deactivation was proposed to be due to the morphology of the coke, which precipitates during residue hydroconversion reactions and agglomerates with the MoS2 crystallites with this coke changing with continued recycling of the coke-catalyst agglomerate recovered after the reaction. The hardening of the coke results in the formation of an insoluble, porous carbonaceous support for the catalyst particles and subsequent deactivation in residue hydroconversion by physical exclusion of larger species.

Recommendations for Future Work This work focused extensively on the use of diphenylmethane as a residue hydroconversion model compound. The simplicity this afforded in terms of analysis and interpretation allowed for the development of reaction and deactivation mechanisms for MoS2. It is recommended that these studies be extended by using larger model compounds which better represent residue feedstocks (such as bibenzyl- cholestane [2]) and the proposed mechanisms evaluated in those reactions. The results of this work have highlighted the importance of the first 1 h of reaction time on the formation of species crucial to the conversion of the feed. Future work should focus on this time period to determine the initial reactions occurring when larger model compounds or actual residue feeds are processed, the steps involved, activation energies and rates. If a rate limiting step could be identified in the generation of the seemingly essential short chain hydrocarbon radicals, it may be possible to improve the overall efficiency of the process by overcoming or cir- cumventing this delay. Further to this, work utilising radical scavengers and radical generators would serve to provide further details regarding the reaction mechanism in more complex systems. The use of such species in residue hydroconversion should be investigated. In such a system, radicals scavengers may suppress coke formation by stabilising larger hydrocarbon radicals which may otherwise undergo

199 condensation reactions. Radical generators may promote residue conversion by destabilising feedstock species, promoting their thermolysis. CHMB, for instance, may be added to a residue feedstock to generate short chain hydrocarbon radicals, its decomposition monitored by observing toluene formation.

With insights into MoS2 deactivation in residue hydroconversion, it is further recommended that work be conducted with a focus on determining the effectiveness of soft coke removal by solvent extraction and physical or chemical size reduction of hard coke. These may be used as separate processes or in conjunction, such as the former to prolong catalyst life before use of the latter. It is possible that the complex structure of the recycled coke-catalyst agglomerate obscures other forms of deactivation (such as poisoning) when evaluated by DPM hydroconversion. It is recommended that similar recycled coke-catalyst samples by heat treated to remove soft coke and tested by DPM hydroconversion to determine if any changes have occurred to the MoS2 active sites during prolonged reaction times. The glass insert micro-reactor developed for this study is ideal for the proposed model compound evaluations as the rapid heat-up and cool-down allow for better control of reaction conditions (time and temperature) than can be achieved in larger systems. Work should also be performed to determine if this technology can be used for residue hydroconversion. Suited for the rapid screening of catalysts and catalyst combinations, this micro-reactor technology is not only applicable to residue hydroconversion catalysts (quickly identifying promising catalysts before testing with a residue feedstock) but may be used for a wide variety of heterogeneous catalyst testing applications. To isolate the reaction mixture from the catalytically active thermocouple sheath, it is recommended that work into repositioning the thermocouple (perhaps from below into a slot in the bottom of the insert) or coating it with a stable, inert material (such as with gold) be conducted.

MoS2 crystallite size change with mixing speed and time may be used to prepare crystallites of speciﬁc sizes for study or use in different applications. The mechanisms governing this phenomenon require further study. This may include computational ﬂuid dynamics (to better understand the mixing of the system), particle tracking and particle Reynolds number calculations (to determine collision frequencies and particle motion) and TEM analyses of solids recovered after varying reaction times (to determine the mechanism and rate of crystallite growth).

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217 Appendices

218 Appendix A

Catalyst Deactivation and Regeneration

A.1 Processes of Catalyst Deactivation In essence, the chemical and morphological changes of catalyst deactivation all result in a reduction in the successful contact between the reagent species and the active sites. This may, for instance, be due to the physical obstruction of the active sites (such as by fouling by coke deposits) or a reduction in the number of active sites available for reaction (some, perhaps, being lost due to poisoning). Numerous papers and textbooks have been published on the topic, with the information below being but a brief overview of the main types of catalyst deactivation in the context of supported and unsupported hydroconversion catalysts.

A.1.1 Fouling Deactivation by fouling is common to many hydrocarbon processing operations but is particularly severe in heavy oil hydroprocessing where supported catalysts are used [16, 40, 129]. As illustrated in Figure A.1 for both supported and unsupported catalysts, fouling involves physically blocking the reacting species from reaching the active sites, thereby inhibiting the reaction and resulting in observed deactivation of the catalyst.

219 In a system as complex as heavy oil, even the chemistry of the deposits and the mechanism of their formation is a topic of intense research. Classiﬁed as either metal or carbonaceous deposits, the latter usually termed “coke”, fouling by such species is the major cause of supported and unsupported catalyst deactivation [1, 7, 40, 70, 71, 130]. Metallic deposits, usually in the form of sulphides, are prominent in heavy oil hydroprocessing due to the high concentrations of metals (notably Fe, V and Ni) present as organometallic compounds (usually a class known as porphyrins) in the feed [41, 43, 96]. Hydroconversion reactions involving such compounds result in the release of the metal atoms and their sulphidation to and deposition as metal sulphide crystallites. Whilst it has been indicated that some of these species (such as FeS) exhibit catalytic activity in and of themselves [98], their presence usually serves to inhibit the activity of the catalyst added to the system. The mechanism for the formation of carbonaceous coke deposits is currently a topic of great interest and, whilst it is agreed to involve a complex network of reactions and precipitative phase changes, the complexity and diversity of heavy oil feedstocks have led to many theories being proposed. Thus far, however, such complications have made determining anything resembling an exact mechanism impossible [10, 28, 95, 97, 131]. It is currently widely held that coke deposits, rich in heteroatomic species such as S and N, form from a group of compounds in the oil known as asphaltenes [2, 10, 11, 14, 28, 132, 133]. Whilst the exact chemistry and structure of these large polycyclic aromatic species, usually the largest in an oil feed, and their many and varied substituents and side chains remains, for the most part, a mystery, the general mechanism of their decomposition to coke is accepted in the literature. Brieﬂy, this mechanism involves two aspects. Firstly, the asphaltenic species undergo some degree of thermal cracking to form large, polycyclic aromatic radicals which may then react with one another in condensation reactions and precipitate to form solid carbonaceous deposits [2, 14, 19, 20, 134]. Secondly, the solubility of the asphaltenic species in the oil feed is dependent on the chemistry of the bulk liquid to such an extent that, upon cracking and hydrogenation of the oil, its solvent properties change and the solubility of the asphaltenes is reduced. This phase-separation process allows asphaltenic species to precipitate [19, 20, 134]. Further complicating the study of these deposits is that species sol-

220 uble and reactive under reaction conditions may precipitate as solid coke when the product is recovered, obscuring understanding of what solid species are actually present during reaction. Even solid species formed during the reaction are not stable end products. The chemistry of such coke deposits continues to change as the reaction proceeds or as recovered coke-catalyst samples are recycled [1, 18, 40, 43]. This process is often referred to as coke “aging” and has been described as the coke moving from a young, soft, partially soluble, reactive deposit to an older, harder, insoluble, graphitic deposit [97]. As the chemistry of the coke changes, so does its morphology and its effect on the system. More important than the amount of coke formed is the location of the deposits in determining the impact on the deactivation of the catalyst [5, 10, 28, 39, 70, 95, 97, 131]. It has been shown [70, 130] that coke is suppressed in the vicinity of the active catalytic sites, presumably due to the rapid capping or quenching of the thermal radicals in that vicinity, inhibiting the degree to which these deposits may directly coat the active sites. This trend does not, however, extend to metal deposition as this process is actually enhanced in the vicinity of the catalyst (due to hydrodemetallation of the organometallic precursors) [70, 97]. It has also been shown that different types of metals (for instance V as opposed to Ni) preferentially deposit at different locations in a supported catalyst pellet [7, 96, 103]. As shown in Figure A.1, deposition of either carbonaceous or metallic species can quickly interfere with catalyst-reagent contact. Figure A.1a shows how deposition of coke on the catalyst support surrounding a catalyst particle may build up, imposing diffusional restrictions and eventually enclosing and completely de- activating that catalyst particle [4, 7, 8, 28, 39, 95, 97]. Figure A.1b shows how metal deposits, which unlike coke may form directly adjacent to or even on the catalyst, may block access to some or all active sites [7, 28, 95]. The proximity of these deposits to the active metal crystallites may also allow a degree of solid-state reaction and subsequent deactivation (due to a change in the chemistry or morphology of the active phase) as discussed in Section A.1.3. Figures A.1c and A.1d provide examples of pore blocking with uniform deposition and pore-mouth plugging respectively. Which process takes precedence during a reaction is dependent on the chemistry of the feed, reaction conditions and pore dimensions [10]. Both

221 processes serve to increase diffusional restrictions through the porous system, inhibiting access to active sites beyond. Advanced models of interconnected porous systems and the impact of deposits and blockages on these networks have been conducted [39] to better understand the deactivation trends associated with them. Figure A.1e shows the deactivation of an unsupported catalyst by coke deposition [1]. By this mechanism, the catalyst particles are recovered together with both the coke which was solid under reaction conditions (hard coke) and that which pre- cipitated during cooling (soft coke) [1, 18]. When recycled, some of the soft coke dissolves and reacts whilst some remains and ages. This aged coke, together with the original hard coke and any more which deposits, encapsulates and deactivates the catalyst particles within [1]. Interestingly, the hard coke agglomerate which captures the unsupported catalyst particles essentially creates a porous carbon support. It is unclear what promotional or inhibitory effects this imparts to the catalyst and no literature could be found dealing with this aspect of the catalytic system. For instance, dual-function catalyst-support chemistry may promote the reaction but fouling processes typi- cally associated with “normal” supported catalysts may become a major route to deactivation.

A.1.2 Poisoning The second major source of hydroprocessing catalyst deactivation is poisoning [16, 40, 41, 95]. Whereas fouling involves the physical blockage of reactants from active sites by various types of deposition, poisoning occurs when species (usually contaminants but the competitive adsorption of reactants, reaction intermediates or products have a similar effect [129]) chemisorb onto the active sites of the catalyst, occupying them and either competing with or completely preventing the desired reactions [40, 71]. Generally the “strength” of poison adsorption is relative to the adsorption of the desired species (as this is the competitive process influencing the observed activity) but some poisons may have such great affinity for the active sites that they affect not only the sites on which they adsorb, but may have electronic (influencing the chemistry of neighbouring sites) and even structural (altering the surface of the catalyst) effects [71]. Furthermore, the unwanted occupation

222 (a) (b)

Figure A.1: Catalyst deactivation due to fouling by carbonaceous or metallic deposits. (a) Deposits on the surface of a supported catalyst surround and enclose catalyst particles. Adapted from Richardson et al. [130]. (b) Crystallites of contaminant metals deposit on or adjacent to catalyst particles, blocking active sites. (c) Pores are uniformly fouled, gradually increasing diffusional resistances. Adapted from Menon [97]. (d) Pore mouth plugging rapidly inhibits diffusion and completely seals pores. Adapted from Menon [97]. (e) Unsupported catalyst agglomeration with soft and hard cokes and the effect of coke aging upon continued reaction or recycle. Adapted from Rezaei and Smith [1]. of active sites may obstruct other reactions by preventing adsorbed reactants from contacting one another or by inhibiting surface diffusion of adsorbed species [71]. Apart from the strength of competition of a poison (sometimes termed the “toxicity” [71]), poisons may be classiﬁed in two general manners: the selectivity of their poisoning and the reversibility of their effects. Poisons may be classiﬁed as either selective, anti-selective or non-selective,

223 portrayed in Figure A.2 by relating the normalised activity of a catalyst against the normalised concentration of the poison [71, 129]. Figure A.2a shows how a selective poison adsorbs preferentially onto the most active sites, resulting in a rapid and severe loss in activity even at very low concentrations. An anti-selective poison on the other hand, Figure A.2b, adsorbs preferentially onto the least active sites, whilst a non-selective poison adsorbs without preference onto all sites as shown in Figure A.2c, with such poisons being characterised by the proportionality between catalyst deactivation and poison concentration.

1.0

(c)

(b)

0.5

(a) (rate(t) / rate(t = 0)) = rate(t / (rate(t) Normalisedcatalyst activity

0.0

0.0 0.5 1.0

Normalised poison concentration

(C (t) / C (t = 0))

Poison Poison

Figure A.2: Selectivity of poisoning behaviour. Adapted from Bartholomew [71]. (a) Selective poisoning preferentially occupying most active sites. (b) Non-selective poisoning showing proportionality to poison concentration. (c) Anti-selective poisoning whereby least active sites are preferentially occupied. Further to selectivity, the effect of an adsorbed poison may be reversible, irreversible or quasi-reversible as illustrated in Figure A.3 wherein the normalised activity of a catalyst in a ﬂow system is examined with time following the intro-

224 duction and subsequent removal of a poison from the feed stream [40]. A reversible poisoning effect, shown in Figure A.3a, indicates that the poison is only weakly adsorbed (recall that the toxicity of a poison is relative to the desired reaction) and desorbs rapidly and (perhaps almost) completely under reaction conditions once the poison is removed from the feed. Figure A.3b shows how the effects of irreversible poisoning cannot be overcome by simply removing the offending species from the feed. In such a situation, the loss of activity is permanent under reaction conditions. Increasing the severity of the reaction, for instance by increasing the temperature, may aid in the regeneration of catalyst activity but it is often unclear if this is due to desorption of the poison or to its decomposition under more severe conditions [40]. Quasi-reversible poisoning, Figure A.3c, is essentially a combination of the previous two effects. In such a situation, some of the poisoning molecules may desorb at an appreciable rate, resulting in the regeneration of catalyst activity whereas some molecules may be desorb at a far slower rate or not at all. These differences may be due to multiple poisons being present in the feed or to the variation in catalytic site activity whereby poisons may desorb more readily from weaker sites.

Poison stop Poison start Poison stop Poison start Poison start Poison stop

1.0 1.0 1.0

0.5 0.5 0.5

0.0 0.0 0.0 Normalisedcatalyst activity (rate(t)/rate(t=0)) Normalisedcatalyst activity (rate(t)/rate(t=0)) Normalisedcatalyst activity (rate(t)/rate(t=0))

0.0 0.5 1.0 0.0 0.5 1.0 0.0 0.5 1.0

Arbitrary reaction time (-) Arbitrary reaction time (-) Arbitrary reaction time (-)

(a) (b) (c)

Figure A.3: Reversibility of poisoning under reaction conditions. Adapted from Furimsky and Massoth [40]. (a) Reversible poisoning allows activity to be regained. (b) Irreversible poisoning shows permanent loss of activity. (c) Quasi-reversible poisoning shows some activity restoration but some permanent deactivation.

225 In the context of heavy oil hydroconversion, the complex composition of the feed includes a large variety of poisons. Foremost among these are the asphaltenes [40, 71], Ni and V organometallics [40, 71, 99, 102, 103] (usually porphyrins [43, 96, 135]), heteroatomic compounds (with nitrogenous species being particularly problematic) [9, 40–42, 71, 72, 76, 102, 136] and the Ni and V metal crystallite deposits themselves [9, 41, 43, 71, 103, 103, 137, 138]. Although the speciﬁc poisoning processes are difﬁcult to study with an oil feed, due to the many simultaneous poisoning and deposition reactions, numerous model compound studies have been conducted to try to determine which species are most responsible for deactivation, the nature of these interactions and how they may be modeled for better understanding [42, 43, 102, 139, 140].

In the case of MoS2, the coordinatively unsaturated rim/edge sites, and the Lewis and Brønsted acid characteristics thereof, allow for interaction with and strong adsorbtion of numerous species, including the plethora of poisons in the oil feed [40]. The difficulty in understanding the chemistry of asphaltenes and the large organometallics in the oil means that, for the most part, they have been identified as poisons by virtue of their presence on the active sites of a spent catalyst despite the precise mechanisms of their adsorption and their chemical and morphological impacts being unknown [28, 40, 43, 71, 100, 102]. For many of the poisons it is not even clear that they remain intact on the surface, with some studies suggesting that they may undergo simple surface reactions (such as polymerisation) which contribute to their toxicity and make desorption more difficult than expected [40]. For asphaltenes and porphyrins, the adsorption is thought to be a donor-acceptor relationship with the delocalised π system of the poison being the electron donor and the Lewis or Brønsted acid of the catalyst being the acceptor [40, 135]. Het- eroatomic species have been indicated to adsorb by interaction of unpaired electrons (for instance in nitrogenous compounds) or aromatic π systems with Lewis acid sites or through the action of Brønsted acid protons to form positively charged surface species [28, 40, 71, 102]. Whilst the aforementioned poisoning mechanisms leave the active sites mostly unchanged [40], poisoning by metal crystallites deposited on or in the near vicinity of the active catalyst crystallites proceeds through more of a solid-solid interaction (see Section A.1.3) to form an alloy, altering

226 the structure of the catalyst and its active sites [71, 100, 103] in such a manner as to render them inactive in the desired reactions.

A.1.3 Others Whilst fouling and poisoning are thought to be by far the dominant causes of catalyst deactivation in heavy oil hydroconversion (and have hence received the majority of research interest), numerous other processes exist which, given that they may inﬂuence different systems (particularly model compound reactions) to a greater or lesser extent, are brieﬂy presented below.

Thermal Degradation High temperature operations (or prolonged lower temperature service [40]) have the potential to deactivate the catalyst through various chemical and morphological changes, processes that may be grouped as sintering, solid-solid reactions or vapour-phase degradation, all resulting in the loss of catalytically active sites [71]. Sintering [40, 71, 129, 141–143], almost exclusively referring to supported metal catalysts, involves the loss of catalyst surface area, and hence active sites, due to the growth of the catalyst crystallites or the collapse of the porous support. At elevated temperatures, metal crystallites may agglomerate as: entire crystallites migrate across the surface of the support (colliding and coalescing into larger particles), metal atoms detach from crystallites and migrate across the support to accumulate on a larger crystallite, or metal atoms actually vapourise and redeposit in different locations and conﬁgurations as illustrated in Figure A.4. Being temperature dependent processes, the sintering of supported catalysts is often described by the Huttig¨ and Tamman temperatures, being 0.3 × TMelting and 0.5 × TMelting ◦ ◦ (437 C and 729 C for MoS2) respectively [129]. The Huttig¨ temperature is the point at which atoms from surface defects begin to become mobile whilst the Tam- man temperature is when bulk atoms become mobile [129]. These mechanisms of movement may occur individually or simultaneously, are kinetically slow and are generally irreversible. Unfortunately, a lack of fundamental understanding of these processes and reliable kinetic data limit the understanding of sintering [71].

227 (a) (b) (c)

Figure A.4: Sintering of a supported metal catalyst. Adapted from Bartholomew [71]. (a) Migration and agglomeration of metal crystallites. (b) Migration of metal atoms separated from large crystallites and agglomerates. (c) Vapourisation and re-deposition of metal.

Solid-state Reactions Although similar to poisoning and often a high temperature process, solid-state deactivation is distinguished in that it results in the chemical conversion of the catalytically active phase into a less active or inactive phase [71, 144]. This process is most often associated with multi-component catalysts, such as Ni-promoted Mo, but may play a role in a system operating with unsupported MoS2 as high surface area catalysts with high energy defect structures are particularly prone to this form of deactivation [145]. Solid-state reactions may involve the active catalyst phase reacting with the support material (for instance the removal of promotional K in an

Al2O3 supported Fe/K catalyst through its reaction with the support to form KAlO2 [71]), with species in the feed (such as active Fe/K carbides in Fischer-Tropsch processing deactivation to inactive carbides [146, 147]), through incomplete redox reactions (such as the deactivation of MoO3 to MoO2 in the partial oxidation of propene [145]) or with solid deposits in their vicinity (such as the formation of various alloys between MoS2 and Ni or V crystallites [40, 71, 100, 103]). In all cases, the catalyst remains a solid but is converted from an active phase or form to an inactive version. Depending on the nature and extent of the change, such deactivation may be completely reversible or permanent. It should be noted, however, that many of these solid-state reactions are exacerbated in typical regeneration procedures (such as heating in an oxygen rich atmosphere) and, as such, a thorough understanding of the deactivation process is required before implementing such a process [40].

228 Vapour-phase Degradation Vapour-phase deactivation may occur by one of two routes: direct vapourisation and volatile compound formation [71] as illustrated in Figure A.5. Direct vapourisation only occurs at extremely high temperatures (in excess of 500◦C) when the active phase itself vapourises and either redeposits in a less optimal conﬁguration or location or is carried out of the reactor altogether. A lesser understood process that has a similar result, volatilisation of the active phase may occur over a wide range of conditions and whilst the properties of the volatile carbonyls, oxides or sulphides thus formed are well known, the mechanism, rate and extent of their formation under reaction conditions and the inﬂuence on overall catalyst performance is not [71].

Figure A.5: Vapour-phase degradation through volatilisation of a metallic crystallite active phase. Adapted from Bartholomew and Farrauto [144].

Mechanical Degradation Whilst mechanical degradation often refers to size reduction of supported catalysts, which is not strictly a deactivation process aside from involving the entrainment of ﬁnes out of the reactor, it may also encompass particle growth in the context of unsupported catalysts, a process which will deactivate a catalyst in much the same way as sintering. Supported catalysts usually take the form of a powder, compacted or held together by a binder, and formed into larger particles such as granules, spheres, ex- trudates or pellets [71]. These larger structures are generally much weaker than the powder particles of which they are comprised and are subject to various physical

229 size reduction processes during both handling and reaction such as crushing (due to the application of a heavy load during transport or in, for instance, a fixed bed) or attrition (movement during transport or collisions with other particles or reactor walls in fluidised or slurry reactors). Size reduction during reaction leads to the formation of fines which either negatively influence the physical operation of the reaction system (such as by causing blockage of a fixed bed and increased pressure drop) or result in the unintended entrainment of catalyst particles in the product (and hence the removal of catalyst from the system) [71].

In relation to unsupported catalysts, such as the MoS2 of interest in this study, particle size reduction would be beneﬁcial as it would serve to promote activity by increasing the number of active rim/edge sites and other defects. In such a system, mechanical degradation manifests as an opposite effect, particle growth. Whether by agglomeration into larger clusters (which inhibit access to inner sites by reactants) or true crystal growth (with an increase in the size of the MoS2 sheets reducing the proportion of rim/edge to total atoms), increase in the size of the

MoS2 particles will result in a decrease in the observed activity of the catalyst.

While the intentional growth of MoS2 (and similar structures such as MoSe2) has been reported, for instance by chemical vapor deposition [148, 149] to obtain ordered sheets for electrochemical applications, no reports have been published regarding the growth of MoS2 structures in hydroconversion reactions. A study by

Ratnasamy et al. [150], preparing MoS2 by the decomposition of MoS3, indicated that the product crystal size and arrangement could be controlled by reaction conditions, with crystal growth being attributed to successive stacking of fused layers and migration of -SH groups. Such crystal growth of a hydroconversion catalyst would likely result in irreversible deactivation as separation of the larger crystallites would likely be almost impossible.

A.2 Catalyst Regeneration Processes There exist two main methodologies for the regeneration of a deactivated hydroconversion process, thermal or chemical.

230 Thermal Regeneration These techniques involve subjecting the spent catalyst sample to elevated temperatures, under a specific atmosphere for a period of time, with the goal generally being to gasify the coke and thereby expose any occluded or obscured active surface. The most common form of this technique is performed under an oxidising atmosphere, with any carbonaceous deposits essentially being burnt off the surface of the catalyst [99, 105, 151] and any volatile species vaporising due to the elevated temperature. This technique is by far the most commonly implemented in the hydroconversion industry, seeing applications in the treatment of both fluidised and fixed bed catalysts. Thermal oxidative treatment does, however, have significant drawbacks. It is ineffective in the removal of non-volatile oxidation-resistant species in the coke (such as the metal crystallites known to form during hydroconversion of heavy oils), it provokes various thermal reactions in the coke (such as condensation and polymerisation) which may even worsen catalyst performance and there exists the possibility of causing thermal degradation of the catalyst itself (such as sintering or volatiliation) [96, 99, 152]. This lattermost point is particularly problematic when processing spent supported metal catalysts as the elevated temperatures may cause sintering of the active metal crystallites or degradation of the support (due, for instance, to the migration of species into and out of the lattice) [153]. It must be noted, that unsupported catalysts, when recovered with coke from heavy oil hydroconversion reactions, are essentially supported and are thus susceptible to the same degradation effects.

Chemical Regeneration Regeneration of a spent catalyst by chemical treatment is not a commonly implemented technique in industrial hydroconversion processes [153] although its numerous variations in application do lend it a degree of ﬂexibility not seen in thermal treatment. This is because chemical treatment of a spent catalyst may be accomplished as either a more physical process (such as chemically etching a layer of coke from the catalyst surface) or as solvent extraction operation [100, 130, 152]. It is solvent extraction which has been shown to offer the most opportunity

231 for the regeneration of spent slurry-phase hydroconversion catalysts for, due to solvent-solute affinity interactions, certain solvents may be able to remove not only carbonaceous deposits from a spent catalyst, but also many other species which cannot be removed by thermal treatment (being perhaps resistant to oxidation or prone to condensation reactions at higher temperatures) [105]. Furthermore, solvents may be selected in such a manner as to remove only specific target species from the coke deposits [152]. This approach not only allows for a degree of analytical information regarding the deposits to be obtained (in terms of the amount of the target species in the coke), but affords the opportunity for the development of a regeneration technique which removes only those species whose interactions with the active phase negatively impact upon catalyst performance. This technique is, however, not without problems. Given the wide range of species present in the coke (including various metals) it is possible that the solvent may react, especially if the extraction is performed at slightly elevated temperatures, with unpredictable results. The solvents required to achieve the desired extraction may be extremely expensive and not economically feasible on an industrial scale without high performance solvent recovery units. Many of the solvents shown to be of use in spent hydroconversion catalyst regeneration (such as acetone, quinoline, tetrahydrofuran, methylene chloride or pyridine [105, 133, 152, 153]) are extremely toxic or flammable, making them dangerous to work with and neces- sitating the use of solvent recovery units to ensure minimal environmental impact. Furthermore, some solvents have been shown to react with certain species in the catalyst (most often adsorbing onto the coke or support) [40, 133], aggravating the problem of deactivation.

232 Appendix B

Experimental Apparatus and Procedures

Referred to extensively in Chapter 3, this appendix contains detailed information regarding the experimental programme and the experimental and analytical apparatus used in this study.

B.1 Detailed Experimental Programme Tables B.1 through B.1 list all experiments conducted and a study of this table will present the reader with a thorough road map of what tests were performed. Due to the complexity of this programme, these tables refer to each series of experiments by an alphanumeric code. These codes, listed together with explanations of the experimental sets they represent, are implemented solely for clarity and legibility and are neither used nor referenced elsewhere in this work.

a.1 - Thermocatalytic stability of major products (benzene and toluene) and diluent (decalin).

a.2 - Diluted model compound (DPM, DPE, DPP) screening for suitability.

a.3 - Diluted DPM in benzene to determine the effect of a supercritical phase.

a.4 - Diluted DPP at reduced reaction temperatures, attempting to lower conversion.

233 a.5 - Diluted DPM with extended reaction times, determining rate of reaction.

a.6 - Undiluted DPM with extended reaction times.

a.7 - Undiluted DPM with increased catalyst loading and reduced reaction temperatures, attempting to compensate for wall catalysis and evaluate catalyst and wall temperature dependency respectively.

b.1 - Series of 21, testing wall activation.

b.2 - Hydrogen:model compound ratio.

b.3 - Heating rates, 20 min versus stirred batch reactor’s 80 min.

b.4 - Catalyst concentration and reaction times (inclined).

b.5 - Catalyst concentration and reaction times (vertical).

b.6 - Catalyst concentration, hydrogen:model compound ratio and reaction times.

b.7 - Series of 8, testing TC wall activation.

b.8 - Catalyst loading and mixing speed with reduced liquid volume.

b.9 - Reaction time at optimal mixing speed.

b.10 - 1st recycle residue-deactivated coke.

b.11 - 5th recycle residue-deactivated coke.

b.12 - Heat treated 1st recycle residue-deactivated coke.

B.2 Batch Reactor Speciﬁcations and Operation Further to the information presented in Section 3.2.2, herewith follow additional details describing the stirred batch reactor system, including those features utilised during semi-batch operation, a full illustration of the unit and detailed batch mode operating procedures. As semi-batch operation was not a part of this study, however, operating instructions for this mode are not provided.

234 Table B.1: Experimental programme as performed in the stirred batch reactor (part 1 of 2).

Reaction temperature Reaction time Catalyst loading Model compound Dilution 1 Code 2,3 (◦C) (h) (ppm Mo) (wt%) 600 Benzene 3 600 Toluene 3 a.1 600 Decalin - 0 3 600 DPM 445 1 600 3 (benzene - a.3) 235 0 a.2 DPE 3 600 0 DPP 3 600 435 430 1 0 DPP 3 a.4 425 420 1 - Reported as wt% of model compound in decalin unless otherwise indicated. 2 - All experiments were conducted at an initial reaction pressure of 13.79 MPa H2 and a mixer speed of 700 RPM. 3 - Wherever possible, experiments were repeated at least three times to aid in the quantiﬁcation of experimental uncertainty. Table B.2: Experimental programme as performed in the stirred batch reactor (part 2 of 2).

Reaction temperature Reaction time Catalyst loading Model compound Dilution 1 Code 2,3 (◦C) (h) (ppm Mo) (wt%) 0 0.5 1 1.5 600 2 4 DPM 3 a.5 6 8 1 445 236 4 0 6 8 0 1 0 6 DPM 100 a.6 0 1 600 6 445 430 1 1800 DPM 100 a.7 415 1 - Reported as wt% of model compound in decalin unless otherwise indicated. 2 - All experiments were conducted at an initial reaction pressure of 13.79 MPa H2 and a mixer speed of 700 RPM. 3 - Wherever possible, experiments were repeated at least three times to aid in the quantiﬁcation of experimental uncertainty. Table B.3: Experimental programme as performed in the micro-reactor (part 1 of 4).

Reactor Mixing speed Reaction time Catalyst loading Total feed loading 1 Code 2,3 (RPM) (h) (ppm Mo) (µL) 1800 400 b.1 400 1 1800 b.2 500 1800 400 b.3 1 2 0 3 237 4 Inclined SS 4 0 1 2 600 400 3 b.4 4 1 2 1800 3 4 1 2 - The volume of mixed feed (model compound, catalyst and CS2) pipetted into the reactor or insert. - All experiments were ◦ 3 conducted using undiluted DPM at a reaction temperature of 445 C and an initial reaction pressure of 13.79 MPa H2. - Wherever possible, experiments were repeated at least three times to aid in the quantiﬁcation of experimental uncertainty. 4 - 316 stainless steel. Table B.4: Experimental programme as performed in the micro-reactor (part 2 of 4).

Reactor Mixing speed Reaction time Catalyst loading Total feed loading 1 Code 2,3 (RPM) (h) (ppm Mo) (µL) 0 1 2 0 3 238 4 Vertical SS 4 0 400 b.5 0 1 2 1800 3 4 1 2 - The volume of mixed feed (model compound, catalyst and CS2) pipetted into the reactor or insert. - All experiments were ◦ 3 conducted using undiluted DPM at a reaction temperature of 445 C and an initial reaction pressure of 13.79 MPa H2. - Wherever possible, experiments were repeated at least three times to aid in the quantiﬁcation of experimental uncertainty. 4 - 316 stainless steel. Table B.5: Experimental programme as performed in the micro-reactor (part 3 of 4).

Reactor Mixing speed Reaction time Catalyst loading Total feed loading 1 Code 2,3 (RPM) (h) (ppm Mo) (µL) 0 1 2 0 3 4 400 0 1 2 1800 3 239 4 Glass insert 0 b.6 0 1 2 0 3 4 150 0 1 2 1800 3 4 1 2 - The volume of mixed feed (model compound, catalyst and CS2) pipetted into the reactor or insert. - All experiments were ◦ 3 conducted using undiluted DPM at a reaction temperature of 445 C and an initial reaction pressure of 13.79 MPa H2. - Wherever possible, experiments were repeated at least three times to aid in the quantiﬁcation of experimental uncertainty. Table B.6: Experimental programme as performed in the micro-reactor (part 4 of 4).

Reactor Mixing speed Reaction time Catalyst loading Total feed loading 1 Code 2,3 (RPM) (h) (ppm Mo) (µL) 0 1 0 150 b.7 0 1500 2000 1 0 2250 0 b.8 150 1500 1 1800 2000 2250 240 Glass insert 1 (continued) 2 0 3 4 150 b.9 1 2000 2 1800 3 4 1800 b.10 1 1800 150 b.11 1800 b.12 1 2 - The volume of mixed feed (model compound, catalyst and CS2) pipetted into the reactor or insert. - All experiments were ◦ 3 conducted using undiluted DPM at a reaction temperature of 445 C and an initial reaction pressure of 13.79 MPa H2. - Wherever possible, experiments were repeated at least three times to aid in the quantification of experimental uncertainty. B.2.1 Description and Specifications The details below are presented as information further to the succinct description of the unit given in Section 3.2.2, with a more complete illustration of the apparatus in Figure B.1 and photographs in Figures B.2 through B.5. The Parr reactor system used in this study and shown in Figure B.2 was designed for high temperature, high pressure operation in both batch and semi-batch mode. The heart of this apparatus is the reactor itself, pictured in Figure B.3. Con- structed of 316 stainless steel, the reactor has an internal volume of 250 cm3 and is able to withstand continuous operating temperatures in excess of 600◦C and 34.47 MPa. The operating volume of the reactor is 240 cm3 with the gas inlet line, stirrer bar, thermowell (housing the dual thermocouples) and cooling water loop (all constructed of 316 stainless steel and shown in Figure B.4) occupying the other 10 cm3. The total volume of the system during semi-batch operation comprises an additional 92 cm3 in the gas lines leading to and from the reactor and 300 cm3 in the condensers before the back-pressure regulator (BPR). During batch operation, valves immediately prior to and after the reactor are shut and the additional gas line volume is reduced to approximately 10 cm3. The reactor walls themselves housed the heating elements, six evenly positioned 200 W Parr heating rods. These heating rods were connected, in parallel, to the heater power distribution box, itself connected via a Parr 4875 Power Controller and Parr 4857 Process Controller to the workstation. In this manner, the software running on the workstation, using temperature information from the reactor thermocouples, was able to control all of the heating elements simultaneously to ensure even heat distribution. The temperature of the reaction mixture was measured by two OMEGA K-type thermocouples located inside the thermowell (see Figure B.4, the thermowell being necessary to prevent any damage or corrosion to the thermocouples. For safety, both thermocouples were monitored continuously during operation and whilst failure of one would not interrupt an ongoing experiment, built-in systems would prevent the unit from being started again if the faulty thermocouple were not replaced. To prevent excessive heat loss (and the poor control and safety hazards which would thus have resulted), the reactor was surrounded in a custom- fitted ceramic fiber insulation jacket, laced on using braided steel cable. As this

241 jacket did not cover the reactor head, an additional layer of insulation, foil-backed fiberglass insulation secured using braided steel cables, was installed. An OMEGA K-type thermocouple, connected to an OMEGA DPi8C temperature monitor, was positioned between the ceramic fiber jacket and the reactor wall. Monitoring this temperature during start up helped ensure that the reactor wall never heated up in great excess of the contents, a scenario which would have resulted in hotspots and unpredictable reactions occurring on the wall. Given the large mass of the reactor, a cooling system was necessary to render control of the reaction time. Towards this end, a cooling water loop passed through the reaction mixture with both ball and needle valves allowing for control of the cooling water flow rate during shutdown procedures. Heating and cooling was not limited to the reactor itself. A length of the outlet line, from the reactor to the condensers, was heated with OMEGA high temperature heating tape and covered with braided glass insulation. The current to this line was controlled using a Superior Electric 3PN116C variable transformer (120 V, 10 A) so as to maintain a temperature, as measured during operation with a separate temperature probe at multiple points along its length, of between 60 and 65◦C. This was necessary to minimise the condensation of products in this line during both operation and shut down. Mixing of the reactor was achieved using a hybrid anchor/pitched-blade 316 stainless steel stirrer, shown in Figure B.4, turned by a Parr magnetic drive monitored and controlled by the workstation via a Parr 4875 Power Controller and 4857 Process Controller. Three instruments worked to monitor the pressure of the system. The reactor pressure was monitored by an analogue Ashcroft (welded, AISI 316 tube & socket) pressure gauge and an Ashcroft (A1906EP50) pressure transducer. The former being a “hands on” check of the latter which itself was connected to the workstation for monitoring and logging. Another analogue pressure gauge, a Duro United In- struments unit, provided further monitoring by measuring the pressure of the gas outlet line just prior to the condensers. Whilst for batch operation the reactor operating pressure was dictated by a pre-determined initial pressure (see Section B.2.2), semi-batch operation employed a back-pressure regulator positioned between the condensers and the scrubber to control the pressure of the reactor whilst allow

242 gas flow to proceed uninterrupted. Given the severe operating conditions of this unit, securely sealing the reactor required the use of a special clamp (shown in Figure B.4) to hold the reactor head and body together. With a Parr Grafoil® gasket (a high temperature flexible graphite gasket) in place, the closing latch served to hold the clamp in place while the eight securing bolts were tightened to a torque of 54.22 Nm (corresponding to 40 ft/lb, this value is specified by the gasket manufacturer for the desired operation at 13.8 MPa). As with any multi-bolt system, the tightening pattern was important and a standard criss-cross technique, illustrated in Figure B.5, was used to torque the bolts sequentially to 10, 20, 30, 35 and 40 ft/lb. Failure to follow this pattern could result in misalignment of the reactor head and body, damage to the gasket or gasket seat faces and severe leaks at high pressure. Inlet gas flow was limited to one gas at a time, flowing through a Brooks Instru- ment (Brooks) 5850S mass flow controller (1000 sccm, calibrated for H2) which was connected to a Brooks control box (Read Out & Control Electronics 0154). Two gases were connected to this unit, nitrogen (for purging the system before and after reaction) and hydrogen (for purging nitrogen from the system and as a reagent). Outlet flow rates, essential during semi-batch operation and useful during batch (to measure the flow rate whilst depressurising during shutdown), were monitored by a Brooks 5860S mass flow meter (1000 sccm, calibrated for H2) connected to the same Brooks control box indicated above. After exiting the reactor, product gases passed through two 150 cm3 stainless steel condensers and a 1000 cm3 stainless steel scrubber. The purpose of the condensers was to ensure that all hydrocarbon products above C4 were captured and collected rather than condensing elsewhere in the lines, in the BPR or in gas sample bags. During semi-batch operation, when hot gaseous product continuously flows through the system, this is an important feature. During batch operations, however, gaseous product only leaves the reactor after cooling to room temperature and initial experiments conducted using the condensers to cool this product gas during depressurisation resulted in no liquid product recovery. As such, condenser use was omitted for the remaining batch experiments. Gas exiting the condensers and flowing through the BPR for pressure regulation, was passed through a scrubber to remove excess H2S (with a 1M NaOH solution) prior to venting. During semi- batch operation this was necessary for quantification of the total H2S formed, gas

243 samples could be collected from the total product and, as such, the scrubber was unnecessary and hence unused. The workstation used with this apparatus (3 GHz AMD Athlon, 2.0 GB RAM, Windows 7 32 bit), running CalGraﬁx (CAL Controls Ltd, v3.0.0), controlled the temperature and mixing speed whilst monitoring and recording the temperature, mixing speed and reactor pressure. Given the hazardous nature of the materials used and the conditions implemented, numerous safety features and protocols were in place as described in Section B.2.3.

B.2.2 Operating Procedure Prior to operation of the unit, two important calculations need to be performed, determining the loading masses and the initial H2 pressure. Presented in Equation B.1, the loading mass calculation equations are applicable to both the stirred batch reactor and micro-reactor systems, holding over a range of masses, dilution rates and catalyst concentrations. As an example, below the loading masses required for 1800 ppm Mo in a reaction of 3 wt% DPM in decalin are calculated. First some of the required values are deﬁned.

Reactor charge mass (g) = mLTotal = 80

Model compound purity (%) = Cmodel.pur = 99

Model compound dilution (wt%) = Cmodel.dil = 3

Mo catalyst loading (ppm) = CMo.ppm = 1800

Mo in precursor (wt%) = CMo.pre = 15.3

CS2 stoichiometric ratio (mol : mol) = Qmol.CS2.Mo = 3 3 DPM density at room conditions (g/cm ) = ρDPM = 1.006 3 Decalin density at room conditions (g/cm ) = ρDec = 0.896 3 CS2 density at room conditions (g/cm ) = ρCS2 = 1.261

Mo molar mass (g/mol) = MrMo = 95.96

CS2 molar mass (g/mol) = MrCS2 = 76.14 The loading masses for species in this example (DPM, decalin, Mo octoate and CS2 being mDPM, mDec, mMo.pre and mCS2 respectively) are then calculated to achieve the desired total mass. The contribution of the Mo octoate and CS2

244 245

Figure B.1: Process flow diagram of stirred slurry-phase batch hydroconversion reactor. 1 - Mass flow controller and mass flow meters connected to flow control box. 2 - Mixer connected via power control box to workstation. 3 - All six heating rods connected in parallel via distribution box to power control box. 4 - Pressure transducer connected to workstation. 5 - Double thermocouple connected via power control box (for high temperature auto-shutoff) to workstation. Wall thermocouple monitored separately. 6 - Heating tape connected to variable transformer. Figure B.2: Stirred slurry-phase batch hydroconversion reaction system as implemented in laboratory.

246 Figure B.3: Close-up of stirred reactor body.

247 Figure B.4: Close-up of bottom of stirred reactor head and internal components.

Figure B.5: Correct order for tightening of reactor clamp bolts. 248 allowed the mixture to be slightly heavier than the speciﬁed mLTotal, with this buffer compensating for liquid left on the mixing vessel walls. C m = m × model.dil = 2.40 g DPM LTotal 100 mDec = mLTotal − mDPM = 77.60 g −6 mMo = mLTotal ×CMo.ppm × 10 = 0.14 g mMo nMo = = 1.50 mmol (B.1) MrMo 100 mMo.pre = mMo × = 0.94 g CMo.pre nCS2 = Qmol.CS2.Mo × nMo = 4.50 mmol

mCS2 = nCS2 × MrCS2 = 0.34 g Determination of an initial hydrogen pressure (the pressure to which the reactor is charged at room temperature such that, after heating, the system is at the desired 13.79 MPa) was slightly less formulaic. The assumption of an ideal gas and the implementation of Equation B.2, proved woefully inaccurate. This was due largely to the solubility of hydrogen in both decalin and the various model compounds. Whilst this complication could be overcome by solubility and thermal expansion modeling (extensions of those studies presented in Section F.7, a simpler trial-and- error approach was used instead with a suitable initial pressure in the stirred batch reactor being found to be approximately 6.76 MPa compared to 5.31 MPa using ideal gas calculations (the additional gas dissolving in the reaction mixture as the temperature increases). P.V = n.R.T

P1.V1 P2.V2 = (B.2) T1 T2 P2.T1 P1 = T2 With loading masses and initial hydrogen pressure known, reactor operation may begin. This procedure assumes the reactor is open and that the reactor and internals are clean and dry. If the reactor is closed, begin with Step 4e and if it needs to be cleaned, begin with Step 4q. Always wear appropriate personal protective equipment (see Section B.2.3) before entering the lab or working with

249 the chemicals or equipment.

1. Overview and evaluation

(a) Ensure cage extraction is functioning (by observing the “tell”) and that the laboratory is negatively pressurised (by checking the differential barometer near the door) (b) Ensure sufficient chemicals are on hand to mix feed (c) Ensure sufficient hydrogen and nitrogen pressure for the experiment (d) Turn on the gas detectors and ensure they do not need routine mainte- nance or calibration (e) Replace rupture disk or reactor thermocouple if either failed in previous run (if unsure of previous usage, a failed thermocouple will be indicated when testing heating equipment in Step 1l whilst the only way to check the rupture disk is to remove it [not recommended unless known to need replacing] or attempt to pressurise the system) (f) Get graphite gasket ready, ensure it is smooth and not cracked or damaged (g) Test cooling water flow by opening outlet valve and slowly opening inlet flow valve (h) Turn off the flow but leave the outlet valve open (this ensures that steam formed during heating can escape safely) (i) Install mixer blade by holding the top exposed head of the magnetic drive and firmly screwing it in place (do NOT overtighten) (j) Power on the control boxes, start the workstation and run the CalGrafix software program (select “Add new instrument” and add “Device 1” on com 3) (k) Start external temperature indicator (for wall temperature) (l) Ensure all three thermocouples (two in reactor and one wall) are functioning correctly

250 (m) Ensure power box power is off and connect all heating rods to the power box, checking the cables of each for damage (especially where the cable connects to the rod) (n) Using a multimeter, test the resistance of the heater assembly (this should be 48 - 49 Ω) (o) Disconnect the cables from the power box (p) Check that the gasket ridge on both the reactor body and head are clean and undamaged

2. Feed preparation and loading

(a) Weigh empty reactor body, ensuring that the heater cables are on the balance too (a Mettler Toledo SB12001 was used) (b) Weigh a 250 cm3 or larger beaker (again, a Mettler Toledo SB12001 was used)

(c) Weigh Mo octoate, model compound, decalin (if used) and CS2 in beaker and mix thoroughly (it is best to weigh in the ascending order of volatility indicated to minimise variation due to evaporation) (d) Pour mixed feed into clean, dry reactor (e) Position reactor in holder on apparatus, feeding power cables through the base (be careful not to damage them) (f) Re-weigh beaker to calculate mass of mixed feed loaded (g) Position the graphite gasket on the reactor body (h) Have the two halves of the reactor clamp, with securing bolts unscrewed most of the way, nearby (i) Caution, this step is tricky! Hold the reactor in position (with the engraved “5000” mark facing the front) with one hand and use the other to position the clamp and close the latches (j) Seal the reactor by tightening the bolts to a torque of 40 ft/lb as indicated in Figure B.5 (k) Loosely position the insulation jacket

251 (l) Position the wall thermocouple beneath the insulation jacket (line up the tip with the “5000 psig” mark engraved on the side of the reactor body as this places it directly above one of the heating rods) (m) Secure insulation jacket in position using steel cable (n) Position ﬁberglass insulation and secure using steel cable

3. Startup and running

(a) Ensure the cooling water exit valve is open (b) Turn on the exit line heating power (c) Ensure both condensers and the scrubber are empty (d) Start mixer at desired speed (700 RPM for this study) (e) Open three-way valve to nitrogen (f) Check valves through system to ensure flow goes to vent (g) Open nitrogen cylinder and adjust MFC to approximately 700 sccm (h) Allow to purge for 1 min (i) Stop flow and close nitrogen cylinder (j) Repeat for hydrogen at 900 sccm for 1 min (k) With hydrogen flowing, close exit valve and pressurise the reactor to the desired initial pressure (l) Shut off the flow, close the hydrogen cylinder and turn the three way valve to nitrogen (m) Using a PerkinElmer electronic leak detector (N9306089), inspect all fittings for leaks, paying particular attention around the reactor head- body connection i. If leaks are found on tube fittings, attempt to tighten them. If tightening fails, shutdown the unit and replace the fitting(s) ii. If the head-body connection leaks, check torque and, if necessary shutdown the unit, inspect the gasket face, replace the gasket and try again

252 (n) Begin logging and start heating program. This heating program was a simple ramp from room temperature (approximately 20◦C) at 5◦C/min to the desired reaction temperature (usually 445◦C, taking 80 min) (o) To ensure that the wall does not heat excessively, when the wall temperature reaches 340◦C: i. Pause the heating program and manually hold the temperature stable (by turning the heating on and off, “PARK” and “PID”) ii. When the reaction mixture closes to within 10◦C of the wall, continue with the heating program (p) Allow reactor to reach operating temperature and begin timing when it does so (q) Test ﬁttings for leaks at operating temperature and pressure, referring to Steps 3(m)i and 3(m)ii (r) Maintain temperature for the desired duration, allowing the pressure to drop or rise as hydrogen is consumed or products are formed

4. Shutdown and cleaning

(a) 1 min prior to the desired duration, shut off the heater (set it to “PARK”), turn off the heater power box and disconnect the heater rod cables (b) Very slowly open cooling water ﬂow valve (c) Remove insulation using suitable gloves (d) Decrease stirring to approximately 250 RPM (e) Allow system to cool to ambient temperature (f) Monitoring the MFM readout, slowly open vent valve to depressurise the system at around 500 sccm (too rapid a depressurisation may result in liquid entrainment and loss) (g) If a gas sample is desired: i. Attach a vacuum pump to a suitable gas sampling bag and extract any residual gases ii. Close the valve on the bag and disconnect from vacuum pump

253 iii. Loosely connect the gas bag to the gas sampling line iv. Allow the system to vent for approximately 5 min to clear out the condensers and scrubber vessels

v. WARNING! The next step allows H2 and H2S to escape into the cage, use extreme caution! vi. Slowly open the gas sampling valve and allow gas to escape through the loose fitting to purge air from the line for approximately 5s vii. Tighten the fitting and open the gas bag valve to collect the sample viii. When done, close the valve on the gas bag and sampling line and remove the gas bag (h) Allow the system to vent to atmospheric pressure (i) Purge remaining reaction gases from the system with nitrogen as per Steps 3e through 3e (j) Shut off the mixer and control equipment (k) Loosen the securing bolts on the clamp (l) Holding the reactor body, carefully remove the clamp and lower the reactor (m) Carefully peel off the graphite gasket, letting as little as possible fall into the liquid (n) Leave the reactor below the internals for 1 min to allow liquid to drip from the internals (o) Re-weigh the reactor body to determine the final mass (p) Carefully pour the liquid and suspended solids from the reactor into a beaker (from which it can be decanted into storage containers) (q) Clean reactor and internals i. Use acetone (Fisher Scientific, 99.7%) to rinse remaining liquids and loose solids from reactor body and internals into a suitable organic waste container ii. Use paper towel to wipe off remaining solid material

254 iii. Remove the stirrer bar for cleaning iv. If some deposits cannot be removed with paper towel, use a brass wire brush to gently clean those places and wipe with acetone and paper towels v. Leave cleaned parts in fumehood to dry vi. Using a plastic scraper, very carefully remove the remnants of the graphite gasket from the sealing faces, paying particular attention to the ridges (scratching this sruface could prevent the reactor from sealing)

5. Product workup and analysis (described in detail in Appendix C for each instrument)

(a) Gas samples may be analysed by offline GC as collected in the gas bag (b) Suspended solids may be recovered by vacuum filtration using a suitable membrane filter, washed with acetone and dried before analysis (c) Liquid samples require dilution with decalin and internal standard addition prior to GCMS analysis

B.2.3 Safety Considerations

Personal Protective Equipment In addition to standard PPE (safety glasses, lab coat, solvent resistant gloves, closed shoes and long pants), working with volatile materials and toxic gases (speciﬁcally

H2S) requires that an organic vapour respirator be worn when inside the cage, loading or unloading the reactor or working with the products. Furthermore, as cleaning the reactor can cause solid deposits to become airborne, it is necessary that the respirator also ﬁlter out particulate matter. A 3M 6900 full face respirator with 3M organic vapour / P100 cartridges was used in this study.

Equipment Safety Features First and foremost, this system was operated as an attended unit, i.e. there were always at least two operators present during the reaction to ensure that rapid, ap-

255 propriate responses could be made should a problem arise. The major concerns for the operation of this unit were fire, explosion and toxic material spillage. Fire and explosion, linked in that one would likely lead to the other when working with pressurised hydrogen and flammable liquids, were mitigated through control of leaks and ignition sources. The entire reactor system was housed in a vented plexiglass cage (183 cm wide by 127 cm deep by 200 cm high) equipped with Honeywell GasPoint II detectors (for H2 and H2S) to ensure that any gas leaks were quickly noted and safely contained and removed. With batch operation, versus semi-batch or continuous, there is an inherent safety feature in the limitation of the amount of material, hydrogen gas of major concern in this study, which could be released from the system. As such, the enclosure could be a reasonable size and yet ensure that, should the extraction system fail and the entire volume of the batch reaction system (approximately 34700 cm3 at STP) leak out, the hydrogen concentration within the enclosure would be approximately 0.75 vol%, below the lower explosive limit (LEL) of 4.0 vol%. A small, lightweight strip of plastic, known as a “tell”, was positioned near the mouth of the extraction line to allow for quick confirmation that extraction was operational. The enclosure itself was located inside a negatively pressurised laboratory (with a pressure differential of 20 Pa as measured by a differential barometer across the door) to ensure that should any gaseous or volatile material leak or spill, it would be contained for safety and easier response and disposal. In this manner, the impact of ignition sources outside of this apparatus could be reduced. The reactor was equipped with a pressure rupture disk (Fike Corporation, P ST FS, 34.47 MPa [5000 psi] burst pressure) such that, in the event of over-pressurisation, excess gas would be safely discharged in a controlled and contained manner, passing through a 1000 cm3 stainless steel buffer vessel (to capture any entrained liquid) and vented rather than damaging the reactor or other fittings and escaping into the cage. The powerful heating elements comprising an intrinsic part of the reactor assembly, also represented possible ignition sources. As such, multiple safeguards were in place to prevent overheating. As mentioned above, the reactor temperature was monitored by two thermocouples with mechanisms in place to ensure that both be functioning when conducting an experiment. Both the Parr 4875 Power Controller and the Cal- Grafix software provided high temperature alarms, to alert users when pre-defined

256 limits were approached, and automatically shutoff when they were passed. Fur- thermore the Parr controller had a ground fault circuit interrupter (GFCI) to ensure that, should one of the heating elements short circuit, power would be cut to the heating system within one thirtieth of a second to reduce the chance of a spark. Rubber sheeting was used to cover the ﬂoor of the cage to contain any spills, aiding in quick cleanup, and to prevent slipping due to unseen spills or condensation (perhaps from the cooling water line or the ice bath). All bulk chemicals, gas cylinders, the control boxes and the operator workstation were located outside the cage to minimise the impact and extent of any problems arising inside the cage. Emergency shut-off switches for all of the apparatus were located near the door of the laboratory.

B.3 Micro-Reactor Design, Development and Operation The micro-reactor system used in this study is presented below including details of the various design and development stages, a full illustration of the ﬁnal unit and detailed operating procedures.

B.3.1 Design and Development Following the extensive experimentation (summarised in Tables B.1 and B.2) performed in the stirred batch reactor, it was deemed necessary to redesign the system, speciﬁcally to convert to a micro-reactor unit. There were ﬁve main reasons for this decision:

• To mitigate catalytic wall activity, due to the formation of active metal sulphides (particularly FeS) on the stainless steel reactor walls during reac-

tion, which was shown to be obscuring analysis of the MoS2 catalytic effect and preventing an accurate determination of the mechanism (described in Section 4.2.1)

• To improve the temperature control by shortening the heat-up and cool- down times (illustrated in Figure 3.3) so as to reduce the extent of unquan- tiﬁable reaction occurring outside the desired temperature range (shown in Section 4.1.3)

257 • To increase the hydrogen:model compound ratio beyond the levels in the stirred batch system

• To reduce the consumption of model compound species and hence both the cost and environmental impact of this study (only very small quantities were required for analysis, the remainder being disposed of)

• To reduce the amount (and hence the time, monetary investment and environmental impact of preparation) of deactivated coke-catalyst required for comparative analyses

The final micro-reactor design was the culmination of an evolutionary development with four distinct stages. These stages are presented in Figure B.6 wherein it may be seen that the system changed in terms of material of construction, size, shape and orientation. Each is described below, together with a simplified diagram of its implementation, the rationale behind its design features and, where applicable, the drawbacks which lead to its supersedure. Note that all tubing and fittings used in these designs were standard imperial sizes with imperial specifications. For consistency and clarity within this section and for future works, these values have been reported as such rather than converting to SI units.

Figure B.6: Comparison of micro-reactor designs.

(a) - Design 1 - 1/4” stainless steel, horizontal with upturned ends, ﬂow-through, unmixed. (b) - Design 2 - 3/8” stainless steel, horizontal with glass boat insert, ﬂow-through, unmixed. (c) - Design 3 - 1/4” stainless steel, inclined and vertical, capped, unmixed. (d) - Design 4 - 6 mm glass insert within 3/8” stainless steel shell, vertical, capped, mixed.

258 Design 1 Illustrated in Figure B.7, this preliminary micro-reactor design utilised seamless 1/4” 316 stainless steel tubing with a wall thickness of 0.049”. At 500◦C, this tubing was rated for operation up to 46.9 MPa, well above the 13.8 MPa reaction pressure. The remainder of the tubing was seamless 1/8” 316 stainless steel with a wall thickness of 0.028”, rated to 53.1 MPa at 500◦C. Either end of the reactor body was equipped with a VCR ﬁtting which, together with disposable silver-coated 316 stainless steel gaskets, allowed the entire reactor to be removed for loading, unloading and cleaning multiple times without concern for the integrity of the seal. The inclined end design was implemented to contain the liquid reaction mixture during operation. With a total length of 50 cm, both ends the reactor (approximately 7.5 cm on each) were bent along the same axis to an angle of 20◦. In this manner, the horizontal section of the reactor body remained within the heated zone of the 800 W Lindberg 55031 tubular furnace (with a 90-280 V solid state relay controlling power to the heater elements) whilst the inclined sections extended beyond the quartz wool insulation at either end and remained cool during operation. A small quantity of liquid was introduced beyond the bend and allowed to pool in the horizontal section. It was postulated that if this pool were shallow enough, the increase in the liquid surface area may negate the need for mixing to promote gas-liquid exchange (although solid-liquid mixing would still be a limiting factor). Any volatile species which formed during the reaction would condense in the upturned ends and run back down into the reaction mixture. The gaseous volume of the reactor could be easily purged by ﬂowing nitrogen or hydrogen across the surface of the liquid. Following a reaction, the product mixture could simply be poured from the reactor into a suitable storage vessel. This desired operation is depicted in Figures B.8a and B.8b. The pressure of this unit was monitored by both an analogue ENFM USA Inc. pressure gauge and an OMEGA PX409-3.5KGUSB pressure transducer, the latter connected to a Dell Precision Workstation 690 PC (3.73 GHz, 36 GB RAM, Win- dows 8 64 bit) and data logged using the OMEGA TRH Control (v1.03.11.297) software package. The pressure gauge, transducer and pressure relief valve (set to open at 17.2 MPa to prevent over-pressurisation) were mounted vertically above

259 the main gas line such that, in the event of liquid back-flow into the gas system during purging or routine depressurisation, these components would be protected from contamination and damage. A 1/16” OMEGA K-type thermocouple extended past the bend to the center of the reactor to monitor and control the temperature via an OMEGA CN8201 temperature controller (a single input, single output, AC relay model with RS485 serial connection) which was interfaced with the aforementioned computer, allowing control and logging through the OMEGA CN8-SW Multi-Comm software package (v3.16.000). Note that as the TRH Control and Multi-Comm software packages were incompatible with Windows 8, they were executed in Windows XP SP3 running under an Oracle VM Virtualbox virtual machine. The entire assembly was located within a vented enclosure constructed of 1/4” LEXAN MR10 polycarbonate (measuring 80 cm wide, 45 cm deep and 43 cm high) to ensure that any gas leaks or explosions were suitably contained. Due to the very specific nature of this design, the valve assembly for the unit could be kept relatively simple. The use of two water-filled bubblers, for visualis- ing gas flow rather than unnecessary measurement with a mass flow meter, allowed for various purging and depressurisation schemes to be implemented. This not only saved materials, but helped reduce the risk of user-error accidents. Unfortunately, this design was unsuccessful. It was determined that the surface tension of the liquid reaction mixture was too great as compared to the liquid-wall adhesive forces and, as such, the liquid formed a “plug” in the reactor (as shown Figure B.8c). This factor was fatal to the design. Unable to purge air from the reactor, the system could not be operated. Any attempt to flow gas through the reactor simply pushed the plug of liquid out. Furthermore, if the gas on either side could be removed of air by nitrogen, and subsequently nitrogen by hydrogen, it would be necessary to maintain equal pressures during both pressurisation and depressurisation to keep the plug from moving. This reactor was soon replaced by design 2.

Design 2 Wishing to maintain the ﬂow-through purging system of design 1 but introducing the liquid in a more reliable manner, the implementation of design 2 saw a glass

260 261

Figure B.7: Process ﬂow diagram of micro-reactor design 1. 1 - Thermocouple connected via OMEGA CN8201 controller to workstation for monitoring and control. 2 - Pressure transducer connected to workstation for logging. (a)

(b)

(c)

Figure B.8: Liquid behaviour in micro-reactor design 1. (a) Desired behaviour where liquid forms shallow pool or film on reactor wall, allowing for through-flow of gases. (b) Desired behaviour where any volatile or supercritical species condense outside of heated zone and return to reaction mixture. (b) Observed behaviour where liquid forms a slug due to strong cohesive forces, preventing gas flow.

262 boat within a steel shell as illustrated in Figure B.9. This design saw the replacement of the inclined-end 1/4” tubing of design 1 with straight seamless 3/8” 316 stainless steel tubing (with a wall thickness of 0.065” and operating pressure of 40.7 MPa at 500◦C) with VCR fittings on either end, located horizontally within the furnace with quartz wool insulation preventing excessive heat loss as before. The remainder of the system remained relatively unchanged in terms of temperature and pressure monitoring and control, bubblers and safety features. Within the 3/8” shell, however, the liquid was now held within 5 cm long glass boat with a semi-circular cross section (outside diameter of 10 mm, 1 mm wall thickness), positioned within the reactor’s isothermal zone. This isothermal zone, roughly midpoint along the furnace length, was determined by filling the tube with 1 mm glass beads, heating it to 445◦C (as measured by a thermocouple located midpoint in the beads) and slowly withdrawing the thermocouple whilst recording the temperature. The use of a glass insert within a shell was necessary in this situation for multiple reasons: it would be extremely difficult to use a standalone glass reactor under such extreme temperatures and pressures, the different thermal expansion coefficients of glass and steel would make a coating unsuitable, alternative metal coatings may be expensive or lack the required durability (such as thin gold coating) and alternative reactor materials may require activity tests (novel alloys for instance) or be cost prohibitive (such a custom sapphire reactors). This rather simple boat design overcame these concerns and offered multiple improvements. The liquid was now not in contact with a steel surface, eliminating the possibility of wall catalysis. The glass boat was small enough and light enough to be weighed on an analytical balance, allowing accurate mass balance calculations to be performed. The glass boats were cheap and easy to produce and could be reused or disposed of if damaged or contaminated without financial or severe environmental concerns. The size of the glass boats could be changed to easily alter the amount and surface area of liquid loaded. Alternative materials could even be used for the boat to study specific wall activation effects. Unfortunately, this design was also determined to be unacceptable as shown in Figure B.10. Firstly, the thermocouple now lay beside the glass boat rather than directly in the reaction liquid. Whilst the liquid volume was small and would hence rapidly equilibrate to the surrounding temperature, this was an additional unknown

263 which had been introduced. Secondly, the fatal ﬂaw of this design, was the inability of the glass boat to contain the reaction liquid under operating conditions. As the entire boat was in the heated zone volatile products (including any supercritical species, see Figure 5.4 for more details) were free to leave the boat and diffuse throughout the shell. This resulted in condensation during the reaction at the mouths of the shell where it exited from the heated zone and, upon cooling of the system, liquid was found to condense along the entire internal surface of the shell. Insufﬁcient liquid remained in the boat and the liquid in the shell could not be recovered.

Design 3 With the previous two designs being flawed, design 3 was created from a few steps back. A 25 cm long seamless 1/4” stainless steel reactor, capped at one end (no flow-through for purging) and inclined within the tubular furnace as illustrated in Figure B.11. Quartz wool was again used as insulation for both ends of the furnace, minimising heat losses and improving temperature control. Using a simplified version of the valve assemblies from designs 1 and 2 (reduced to only a single bubbler), the same pressure and temperature control apparatus (the thermocouple again introduced directly into the reaction liquid) and with the addition of the Shi- madzu GC14B on the gas exit line for ease of analysis, design 3 was a success. A small amount of liquid, introduced to the reactor, was found to pool in the lower end of the tube. Under operating conditions, volatile species could not pass beyond the heated zone and the incline of the reactor ensured that, upon condensation, they returned to the liquid mixture as shown in Figure B.12.The experiments conducted in this apparatus are presented in Table B.3. One drawback of the capped design was that flow-through purging was no longer possible. To overcome this, purging of air by nitrogen and nitrogen by hydrogen was accomplished by sequential pressurisation and depressurisation steps. It was determined that cycling the system to 800 kPa three times would be sufficient to reduce oxygen to acceptable levels. As shown in Equation B.3, assuming an atmospheric oxygen composition of φ ′′ O2.atm = 21vol%, by Dalton’s Law, its initial partial pressure in the system, PO2, would be 21.3 kPa. After the first pressure cycle (to PGTotal = 800kPa), its compo-

264 265

Figure B.9: Process ﬂow diagram of micro-reactor design 2. 1 - Thermocouple connected via OMEGA CN8201 controller to workstation for monitoring and control. 2 - Pressure transducer connected to workstation for logging. Figure B.10: Thermocouple location and ﬂuid movement during operation of micro-reactor design 2.

266 sition would reduce to 2.7 vol% and after three cycles that would be down to 0.04 vol%.

φO2.atm PO2 = .101.325 = 21.3kPa 100 (B.3) φ ′′ PO2 O2 = = 2.7vol% PGTotal The length of this reactor and the angle of incline were not arbitrary. As indicated in Table 3.2, the hydrogen:model compound ratio was noted to be lower than desired and hydrogen starvation was a cause for concern. This micro-reactor was designed so as to allow for the hydrogen:model compound ratio to be greatly increased and varied by changing the amount of liquid loaded. To perform these design calculations, equivalent semi-batch residue hydroconversion conditions [18] were used to determine the amount of hydrogen supplied to such a system and these were converted to equivalent batch volumes as shown below. The data for the residue hydroconversion reaction [18] is as follows:

Reactor charge mass (g) = mresid = 80 3 Residue density (g/cm ) = ρresid = 1.0402

Reaction pressure (psig) = Prxn = 2000 ◦ Reaction temperature ( C) = Trxn = 445

Reaction time (min) = trxn = 60 3 Reactor volume (cm ) = VGLTotal = 250

Hydrogen ﬂow (sccm) = VH2.ﬂow = 900

This information allowed for the calculation of the hydrogen required, at STP, for such a semi-batch operation as per Equation B.4. mresid 3 Vresid = = 77 cm ρresid 3 VGTotal = Vrxr −Vresid = 173 cm (B.4) 3 VGTotal.flow = VH2.flow.trxn = 54000 cm 3 VGTotal+flow = VGTotal +VGTotal.flow = 54173 cm

The desired volumetric (easiest for design) hydrogen:liquid ratios (Qvol.gas.liq), may thus be calculated.

267 PSTP.Vgas.STP.Trxn 3 VGTotal+flow = = 1025 cm TSTP.Prxn (B.5) ′ VGTotal+flow 3 3 QG.L = = 13.3 cm /cm Vresid As such, a reactor would need a total volume approximately 14 times larger than the liquid charge volume to have sufficient hydrogen to be equivalent to a semi-batch residue hydroconversion system (compared to only 3 times the volume for a batch system, per Table 3.2). As a compromise between the ideal semi-batch ratio, the undesirable batch ratio and the practicality of reactor size (in terms of fitting an inclined reactor with the tubular furnace), a final reactor length of 25 cm with a liquid loading of 0.4 cm3 was selected for this design. This afforded a ′ ′ QG.L value of approximately 7.5, corresponding to a QG.L of 2.9, and allowed the lower end of the reactor to be positioned in the isothermal zone of the furnace, as shown in Figure B.11. Approximately 2.5 cm of the reactor extended beyond the heated zone. This provided an area for volatile species to condense and protected the VCR fittings from damage which may result from both repeated heating cycles and repeat opening. Unfortunately, despite its success as a micro-reactor, the simplicity of this unit meant that it had inherent problems and did not achieve the desired goals of this development program. Not only was the reaction mixture still exposed to catalytically active walls (as shown in Section 4.2.1) but the design of the system prevented any form of mixing or agitation from being implemented. As such, the development and testing of design 4 began even before the experiments in design 3 were complete. As a move toward design 4, a second version of design 3 (shown in Figure B.13) saw the horizontal system turned vertically with an associated migration to a new and improved enclosure (discussed in detail with design 4). Additional experiments (listed in Table B.4) were conducted to study the effect of the vertical orientation (which would influence, in a small way, the area of the gas-liquid interface and the rate of condensed species return to the bulk liquid [illustrated in Figure B.14) on the reaction system before design 4 was implemented.

268 269

Figure B.11: Process ﬂow diagram of micro-reactor design 3, inclined implementation. 1 - Thermocouple connected via OMEGA CN8201 controller to workstation for monitoring and control. 2 - Pressure transducer connected to workstation for logging. Figure B.12: Fluid movement during the inclined operation of micro-reactor design 2.

270 Figure B.13: Process ﬂow diagram of micro-reactor design 3, vertical implementation. 1 - Thermocouple connected via OMEGA CN8201 controller to workstation for monitoring and control. 2 - Pressure transducer connected to workstation for logging.

271 Figure B.14: Fluid movement during the vertical operation of micro-reactor design 2.

272 Design 4 Design 4 saw the culmination of lessons learned from previous reactors into a targeted design which met all of the requirements desired for the completion of this study. The condensation and run-back concepts from design 1 combined with the glass insert ideas from design 2 and the scaling and orientation from design 3 were amalgamated and enhanced to produce the vertically-oriented, glass insert, vortex-mixed slurry-phase hydroconversion micro-reactor shown in Figures B.15 and B.16. This system isolated the reaction mixture from the metal walls, kept species which were volatile under reaction conditions conﬁned within the insert, allowed for rapid heat-up and cool-down for a more responsive temperature control, offered the ability to quickly unload and reload the system for quick turnaround and was able to mix the reaction medium. At the core of the micro-reactor unit was the removable glass insert, a tube 25 cm long with an outer diameter of 6 mm and a wall thickness of 1 mm, a total volume of 2.85 cm3 and sealed at one end to form a narrow “test tube” shape (the dimensions allowing for a range of hydrogen:model compound ratios to be tested depending on the liquid loading as indicated in Table 3.3 and discussed for design 3). During operation, this insert was positioned vertically within a 3/8” 316 stainless steel tube with a wall thickness of 0.065” and an operating pressure at 500◦C of 40.0 MPa. This tube, with its lower end capped, acted as a shell, containing the high pressure of the system and hence allowing the insert within to operate under the desired conditions without structural failure. The top of the shell had a VCR connection, allowing the insert to be removed and replaced with ease and ensuring a reliable seal through the use of disposable silver-coated 316 stainless steel gaskets. The shell-insert assembly was positioned within the vertically oriented 800 W Lindberg 55031 tubular furnace (with a 90-280 V solid state relay controlling power to the heater elements), itself bolted to a custom made aluminium base for spacing (the bottom of the shell projected out of the furnace and other components were to be located below), safety (a manufacturer-recommended minimum gap was required around the furnace for safe operation) and for stability. Quartz wool was used as a soft, removable insulation around the top of the shell to reduce thermal

273 drafts which would negatively affect temperature control. Similarly to design 2, the isothermal zone for the vertical orientation was determined by pre-heating the shell, filled with glass beads, to 445◦C and slowly withdrawing the thermocouple whilst recording the temperature. To maintain the reaction mixture within the insert in the isothermal zone of the furnace, glass beads 1 mm in diameter were used as a non-compressible (to ensure settling, and hence movement of the insert, did not occur over time) spacer material below the insert and the assembly was positioned within the furnace such that approximately 25 mm of the insert extended beyond the heated zone. This was important as it ensured that material volatilising under reaction conditions could condense and run back down rather than escape from the insert. The thermocouple for temperature monitoring and control was a 1/16” OMEGA K-type which extended into the reaction mixture from above to provide accurate monitoring and control via the OMEGA CN8201 temperature controller which was interfaced with the aforementioned Dell Precision Workstation 690 PC (3.73 GHz, 36 GB RAM, Windows 8 64 bit) running the OMEGA CN8-SW Multi- Comm software package. Again, as the TRH Control and Multi-Comm software packages were incompatible with Windows 8, they were executed in Windows XP SP3 running under an Oracle VM Virtualbox virtual machine. Given the small volume of the reaction mixture, the numerous thermal barriers between it and the heating elements (element - gas layer - shell wall - gas layer - insert wall - mixture) and that the control thermocouple was located within the mixture itself, the applicability of positioning the insert precisely in the previously determined “isothermal zone” was questionable. In keeping with a simplified, easy to operate design, the valve system remained mostly unchanged from design 3, allowing for the supply of either nitrogen or hydrogen gas with shut-off and needle valves being used to control the flows. This system differed from design 3, however, in that the gas line connecting to the reactor was formed into a “pig-tail” (seen in Figure B.17, a spiral of tubing allowing movement or vibration to be dampened thereby minimising damage to the lines and fittings. This feature was necessary to prevent damage caused by the mixing system described below. All gas lines leading to and from the shell-insert assembly were kept to 1/16” or 1/8” to minimise the volume of gas outside of the reaction zone, resulting in a total system volume (insert, shell, supply lines) of only 17.5 cm3. The

274 analogue ENFM USA Inc. pressure gauge and OMEGA PX409-3.5KGUSB pressure transducer, the latter connected to the aforementioned computer, remained for continuous monitoring of the system. After a reaction, the system was depressurised through shut-off and metering valves. The flow passed through the sampling port of an in-line Shimadzu GC-14B gas chromatograph before entering a water-filled bubbler, which was used to visualise the depressurisation flow and keep it suitably low to prevent liquid carry-over from the reactor, before being directed to vent. The dimensions of the reactor insert, specifically having an internal diameter of 4 mm but being 250 mm long, made mixing a particular challenge. Traditional internal mixing systems (such as stirrer bars) could not be used for not only would they necessitate the relocation (or perhaps incorporation into the mixer itself) of the thermocouple, but keeping the long, narrow shaft of such a mixer centered, to prevent collision with and damage to the insert, would be extremely difficult. For externally applied mixing, neither shaking nor inversion were feasible either, the former making continuous connection to the unit difficult (for pressure monitoring) and both the former and latter risking spillage of the reaction mixture from the insert. Instead an externally applied vortex mixing force was employed. The base of the shell was secured to a custom designed, high density PVC adapter which was connected to a Talboys 9456TAHDUSA advanced heavy-duty vortex mixer. The adapter allowed the hexagonal cap of the shell to fit snugly in the concave cone of the mixer, reducing vibration and wear. The vortex mixer selected for this study was rated for continuous operation between 300 and 2500 RPM with a load of 2.5 lb and allowed for precise control and display of the mixing speed and mixing time. Whilst such mixing speeds may seem excessive compared to normal internal mixers, the comparison is not direct. Whilst an internal mixer such as a stirrer bar imparts almost all of its energy to the liquid being mixed, in this system it was necessary to establish a full vortex to ensure suitable mixing of the entire liquid load, an endeavour wherein the centrifugal forces of rotation need to overcome the centripetal-like forces of cohesion and surface tension to force the liquid to the sides of the insert. To study these effects and determine a suitable mixing speed, a series of experiments were conducted using a Megaspeed MS70K high speed camera as shown in Section 4.2.3. These visual experiments

275 were conducted using a glass mock-up of the shell-insert assembly. An actual insert with reactor product sealed after a reaction was used placed within a glass tube with the same dimensions as the steel shell. This glass version of the system was then used to observe the dimensions of the vortex, the motion of the thermocouple during mixing and the suspension of solids within the liquid. Another complication of such high mixing speeds and the movement of the entire reactor assembly to exact this mixing was the vibration it caused. Vortex mixers are prone to “walking” whereby they slowly move across a surface due to their vibration. This would be unacceptable when such walking would move the shell off-center and possibly contact the furnace walls. To prevent this, small metal brackets were secured around the base of the mixer to confine it. Vibration of the reactor assembly was also problematic. “Hard holds”, whereby a clamp was used to physically restrain the top of the assembly, were quickly ruled out as the vibration was transferred to the clamp and any connected structures. The final design used a “soft hold”, a spring mechanism to hold the assembly down, hold it centered and absorb the majority of the vibration without transferring it to any other structures. One connection which was not protected by this system was the thermocouple which, as seen in Figure B.17, connected to the head of the reactor. Instead, the thermocouple shaft was protected by gently bending it around and firmly securing its connector to the reactor. In this manner, the vibration was transferred to the long, flexible connection wire which would be undamaged by the motion. Together with these improvements, design 4 also featured a number of improved safety features and protocols as discussed in Section B.3.3.

B.3.2 Operating Procedure Whilst operation of the micro-reactor is very similar to that of the stirred batch system described in Section B.2.2, there are some important differences. Before loading and operation may begin, it is, as was with the stirred reactor, necessary to determine both the feed masses to be mixed and the initial hydrogen pressure required to give the desired reaction pressure. Calculation of the former is performed in the same manner for both systems, and is described in detail in Section B.2.2.The

276 Figure B.15: Process ﬂow diagram of design 4, a novel slurry-phase batch hydroconversion micro-reactor. 1 - Thermocouple connected via OMEGA CN8201 controller to workstation for monitoring and control. 2 - Pressure transducer connected to workstation for logging.

277 Figure B.16: Novel slurry-phase batch hydroconversion micro-reactor as implemented in the laboratory.

278 Figure B.17: Detail of novel slurry-phase batch hydroconversion micro- reactor head assembly and connections.

279 latter was, once again, most easily determined by trial and error experiments, with a suitable starting pressure found to be approximately 12.0 MPa. This system may be operated as a single-run system or as a higher throughput multi-run whereby multiple inserts are loaded with mixed feed and each insert, having been removed from the reactor after reaction, is capped for later processing while the next is loaded and the reactor restarted. For completeness, this procedure covers the latter scenario starting with a clean and dry shell, with the head off and separate from the mixer, and sufﬁcient inserts for the desired number of runs. If the reactor is closed, begin with Step 4c, and if the shell or insert is dirty, with Steps 4m or 4u respectively.

1. Overview and evaluation

(a) Ensure enclosure extraction is functioning (by observing the “tell”) and that the laboratory is negatively pressurised (by checking the differential barometer near the door) (b) Ensure sufficient chemicals are on hand to mix the required amount of feed (c) Ensure sufficient hydrogen and nitrogen pressure for the experiment (d) If gas analyses will be performed, ensure the GC integrator has enough paper (e) Get a new stainless steel gasket ready (avoid touching or contaminating the surface of the gasket as dirt or dust may damage both the gasket and the VCR fittings) (f) Turn on the main power (ensure the emergency shut-off switch has not been triggered) (g) The vortex mixer will spin several times as it performs a self-test, attend to any error messages it may display (h) Power on the temperature controller (not the furnace) and start the workstation (i) Under the Windows XP virtual machine, run the TRH Control and Multi-Comm software packages

280 (j) Ensure that both the thermocouple and pressure transducer are functional and providing accurate readings (k) Start the GC14B and link it to the integrator (this is done by pressing Command - Shift - 5 - 1 - Enter) (l) Ignite the FID ﬂame, set the ﬂows and ensure the correct program is loaded (see Section C.1) (m) Allow the GC time to stabilise to its starting temperature

2. Feed preparation and loading

(a) Zeroa 2 cm3 sealable vial on the microbalance (a Sartorius ME5 was used in this study)

(b) Weigh the Mo octoate, model compound, and CS2 into the vial (it is best to weigh in the ascending order of volatility indicated to minimise

variation due to evaporation and it also helps to cool the CS2 in a re- frigerator beforehand) (c) Seal the vial tightly and shake to mix (d) Load the insert i. Zero a 100 cm3 measuring cylinder on the analytical balance (an AND GR-200 was used in this study). This will be used to hold the insert vertical during weighing and loading ii. Remove the stopper from the ﬁrst insert and weigh it iii. Shake the vial of mixture to ensure that it is thoroughly mixed and, using a glass pipette or analytical syringe (an SGE eVol XR was used in this study), carefully add the desired volume of mixture to the insert iv. Gently tap the insert to settle the mixture in the bottom v. Re-weigh the insert to determine the initial mass vi. Replace the insert stopper to minimise evaporation or accidental spillage and keep the insert vertical until use vii. Promptly rinse the pipette or analytical syringe with acetone (Fi- sher Scientiﬁc, 99.7%)

281 (e) Carefully lower the loaded insert into the shell and ensure that the level mark on the insert lines up with the top of the shell. This is very important. These marks are engraved on each insert to indicate the correct insertion depth. Failure to position the insert correctly can cause the thermocouple to bottom out and break the insert (if it is too high) or for the liquid to be outside the isothermal zone, the thermocouple to not make contact with the liquid, and vaporous species to escape and condense outside the insert during reaction (if it is too low) (f) Use a small piece of cork or other compressible, temperature-resistant material to prevent the insert from spinning during mixing (g) Use a lint-free cloth to wipe off the VCR sealing faces and gasket to be used (h) Position the gasket around the insert and remove the insert stopper (i) With the furnace door open (otherwise it will not reach), lower the reactor head in position (with the thermocouple entering the insert) and loosely tighten the nut (j) Position the reactor on the mixer and lock it into place (k) Tighten the reactor head nut (l) Use the spring restraint to secure the assembly in position, ensuring that it is vertical and not angled within the furnace

3. Startup and running

(a) Begin logging temperature and pressure (in the event of a problem, data immediately prior to its onset would be invaluable) (b) Purge the system with nitrogen i. Ensure that the reactor valve is open and the vent valves are closed ii. Using the gate and needle valves, slowly pressurise the reactor with nitrogen to 700 kPa iii. Shut off the nitrogen ﬂow iv. If the system cannot hold pressure, there is a leak, refer to Step 3f

282 v. Using the vent needle valve, carefully vent the system, observing the bubbler to ensure a slow, steady flow vi. Repeat twice more for a total of three purges (c) Purge the system with hydrogen i. Repeat Step 3b using hydrogen ii. During one of the purge cycles, when the system is at 700 kPa, use a PerkinElmer electronic leak detector (N9306089) to inspect all fittings for leaks, paying particular attention around the VCR fitting (there are two small holes in the nut for leak testing) iii. If a leak is found, refer to Step 3f (d) Pressurise the reactor with hydrogen to the desired initial pressure (e) Leak test all fittings once more. For safety, even minute leaks must be attended to promptly. An additional incentive is that, given the small volume of this system, small leaks can depressurise it very quickly and ruin the experiment (f) If a leak is found: i. Using a product such as Snoop® Liquid Leak Detector when working with nitrogen or the more sensitive PerkinElmer electronic leak detector when using hydrogen, identify the location(s) of the leak(s) ii. Attempt to tighten any leaking fitting (tube fittings or the VCR gasket fitting) iii. If the leak persists, shutdown the unit per Steps 3(f)vi and 3(f)vii iv. Replace any offending tube fitting(s) v. If the VCR fitting is leaking, inspect the sealing beads and gasket for damage. If beads are worn or scratched, replace the fitting. If these are undamaged, replace the gasket and try again. If the problem persists, replace the fitting vi. If the leak occurred before heating, re-purge and pressurise from Step 3b

283 vii. If the leak occurred after heating, the reactor must, unfortunately, be reloaded. Skip to Step 4b (g) Close the furnace door and position quartz wool snugly around the top of the reactor. The bottom is left open as the rotation of the reactor during mixing would dislodge any insulation here (h) Perform a ﬁnal sweep to ensure no steps have been missed (i) Turn on the furnace (j) Start the mixer and ensure it is running smoothly (k) Close and latch the enclosure door (l) Begin the heating program. This heating program was a simple ramp from room temperature (approximately 20◦C) at 18◦C/min to the desired reaction temperature of 445◦C, taking approximately 23 min (m) Allow reactor to reach operating temperature and begin timing when it does so (n) Test ﬁttings for leaks at operating temperature and pressure. If a leak is found, immediately shutdown the unit and refer to Step 3f (o) Maintain temperature for the desired duration, allowing the pressure to drop or rise as hydrogen is consumed or products are formed (p) If another run is to be conducted immediately after the current one, load and seal the next insert as per Step 2d and store it vertically until needed while the current run is underway (q) If a run was conducted prior to this one, process the insert as per Step 4p while the current run is underway

4. Shutdown and cleaning

(a) 30 s prior to the desired duration, shut off the heater (set Multi-Comm temperature to 0◦C) and turn off the furnace (not the controller) (b) When the duration has expired, shut off the mixer, carefully open the furnace door and remove the insulation using suitable gloves or tongs

284 (c) Allow system to cool to ambient temperature (if desired, a fan may be used to speed cooling after the reactor drops below 400◦C (d) Slowly open the vent valve and, monitoring the bubbler, depressurise the system at a slow, steady ﬂow (rapid depressurisation may result in liquid entrainment) (e) If a gas sample is desired, depressurise the system to 70 kPa and press Start on the GC (this ensures that hydrogen in the feed lines, not par- taking in the reaction, has been expelled and gas from inside the reactor is passing through the sample loop) (f) Allow the system to vent to atmospheric pressure (g) Purge remaining reaction gases from the system with nitrogen as per Step 3b (h) Stop data logging (i) Remove spring restraint, loosen reactor head nut and remove assembly from mixer (j) Disconnect reactor head from shell, gently tapping the TC against the sides of the insert as it comes out to remove as much liquid as possible from the sides of the TC (k) Carefully slide the insert out of the shell and stopper it (l) Important! Note if there is any liquid on the outside of the insert. If so, some escaped during loading or reaction and the shell must be cleaned. If not, skip to Step 4n (m) If the shell is suspected of being contaminated: i. Carefully pour out the glass beads into a container. If liquid escaped into the shell, some beads may be stuck in the bottom; rinse them out with acetone ii. Rinse the beads several times with acetone iii. Rinse reactor body with acetone and clean with paper towels and pipe cleaners iv. Leave beads and shell in drying oven at 100◦C until dry (usually 30 - 60 mins)

285 (n) Clean the thermocouple with acetone and paper towel and allow to air dry (o) If running another insert, return to Step 2e (p) Zero a 100 cm3 measuring cylinder on the analytical balance (q) Briefly remove the stopper from the insert and weigh it to determine the final mass (r) Using a suitable syringe (a 1 cm3 Popper glass tuberculin syringe with 20 gauge, 30.48 cm (12”) 316 stainless steel needle was used in this study), extract the liquid and suspended solids from the insert and store in a suitable container (pouring the product out is difficult and not recommended as the liquid often does not easily drip out and solids remain in the insert) (s) Important! Do not use, for extraction or storage, any plastic-containing equipment. The small, sometimes trace, amounts of benzene and toluene present in the product will be quickly absorbed, rendering any analytical results useless (t) Promptly rinse the syringe and needle with acetone (u) Clean the insert using acetone and a pipe cleaner. Rinse the stopper with acetone. Stubborn solid deposits in the insert may be removed by filling the insert with acetone and placing in a sonic bath (a Cole- Parmer Ultrasonic Cleaner, 08895-12, 100 W, 42 kHz was used in this study) for 60 seconds (v) Place the insert and stopper in a drying oven at 100◦C until dry (usually 30 - 60 mins) (w) Remove the insert from the oven and loosely stopper (to keep out dust and other contaminants)

5. Product workup and analysis (described in detail in Appendix C for each instrument)

(a) Gas samples are analysed during depressurisation by inline GC

286 (b) Suspended solids may be recovered by allowing them to settle (this may be sped up using a centrifuge) and carefully removing the liquid by syringe or pipette. The solids may then be dried in a vacuum oven or the settle-remove liquid process may be repeated with acetone to wash off unwanted liquids (c) Liquid samples require dilution with decalin and internal standard addition prior to GCMS analysis

B.3.3 Safety Considerations

Personal Protective Equipment Whilst the amounts of chemicals used for any given run in the micro-reactor is smaller than in the stirred batch reactor, the safety concerns relating to contamination and poisoning remain the same. As such, in addition to standard PPE (safety glasses, lab coat, solvent resistant gloves, closed shoes and long pants), an organic vapour respirator is to be worn when loading or unloading the reactor, working inside the enclosure or working with the products. As there are numerous sources of particulate matter in this study (such as from the furnace wall insulation), it is recommended that the respirator also ﬁlter out particulate matter. A 3M 6900 full face respirator with 3M organic vapour / P100 cartridges was used in this study.

Equipment Safety Features The primary safety feature of this apparatus was that it was operated as an attended unit, with at least two operators present during any reaction to ensure that rapid, appropriate responses could be made should a problem arise. Despite the small size of this reaction system, compared to the stirred batch reactor described in Section 3.2.2, this assembly too was installed inside a protective enclosure. As an improvement to the Plexiglass structure, this case was constructed of 1/16” aluminium sheeting with a 1/4” LEXAN MR10 polycarbonate (SABIC Polymershapes) door, was vented through an extraction attachment located at the top of the enclosure. The main gas vent line (exiting the water bubbler) was directed to this extraction line rather than into the enclosure space. A small,

287 lightweight strip of plastic, known as a “tell”, was positioned near the mouth of the extractor to allow for quick confirmation of operation. As with the stirred batch reactor, operating in batch mode and shutting off the gas supplies as part of the standard procedure limited the amount of hydrogen available to leak. Furthermore, this enclosure (measuring 72 cm wide, 66 cm deep and 106 cm high) was sufficiently large so as the ensure that, should the extraction system fail and the entire pressurised contents of the reactor escape during operation (approximately 2400 cm3 at STP), the hydrogen concentration within the enclosure would be 0.5 vol%, below the LEL of 4.0 vol%. This enclosure was located within a negatively pressurised lab (with a pressure differential of 20 Pa as measured across the door) to ensure that should any gases or volatile materials escape, they were confined for safety and easier response and disposal. The door of the enclosure had two locking latches to keep it securely in position even during an explosion. The temperature control system was fail safe (immediately shutting off should the thermocouple fail or the temperature change rapidly and significantly [an indication that the thermocouple has moved]) and equipped with two alarms. One alarm warned when a preset temperature was approached and the second shut off the heating when it was passed. The pressure monitoring software also had an alarm to alert of over-pressurisation with a pressure relief valve (set to 17.24 MPa) opening to relieve excess pressure (venting through the bubbler as a visual indication). To ensure stability of the entire system to mixer vibration, accidental jostling or seismic events, the enclosure and all equipment was secured to one another and/or to the counter. Two emergency shut-off switches were installed to ensure that in the event of a leak or other event, power to the system could be immediately cut to limit to possibility of ignition. The first was on the side of the enclosure itself and the second near the door of the lab.

288 Appendix C

Analytical Apparatus, Procedures and Data Analysis

C.1 Gas Product Analysis

C.1.1 Analytical Equipment and Procedures Although gas samples for both the stirred batch and micro-reactor systems and were analysed by gas chromatography, the method of their collection and the instruments used differed due to the differences in the experimental setups. Product gas analysis for the stirred batch reactor was performed by collecting an approximately 1000 cm3 sample, in an Alltech Tedlar® gas sampling bag, of the vent gas during reactor depressurisation. This sample was introduced to a Hewlett Packard (HP) 5890A GC, pictured in Figure C.1. The HP5890A, equipped with a Porapak® Q 80/100 mesh packed column and a ﬂame ionisation detector (FID), was connected to a Hewlett Packard 3396 Series II integrator for data collection and printout. It is important to note that, despite being rated for low permeability, the gas sampling bags are still permeable to some gases, and that this permeability is different for different gases. Hydrogen in particular is rated to be able to diffuse through the gas bag at 150 cm3/(100”2.24h.atm) compared to, for instance, nitrogen at 0.25. It was thus important that gas samples be analysed as soon after

289 collection as possible to minimise the extent of hydrocarbon concentration due to hydrogen loss. During this study, samples were analysed within 4 h of collection. The injector port flow rate for this instrument was set to 50 kPa and the detector to 60 kPa. The temperature ramping program is shown in Table C.1. With the gas bag connected to the sampling port of the GC, a steady pressure (by gently squeezing the bag) was applied and flow through the sample loop confirmed using the water bubbler. This served to flush out any air or previous sample and load the current sample for analysis. After flowing the sample for 10 seconds, analysis was begun which simultaneous triggered the pneumatic injection valves to introduce the sample onto the column, started the temperature program and began data collection on the integrator. Following analysis, the integrator provided a printout of the results.

Table C.1: Temperature program for Hewlett Packard 5890A gas chromatograph analyses.

Specification Value Initial temperature (◦C) 35 Hold time (min) 10 Ramp rate (◦C/min) 20 Final temperature (◦C) 120 Hold time (min) 14.75 Gas sampling and analysis in the micro-reactor system was greatly simplified. Equipped with an in-line GC, the vent gas during reactor depressurisation flowed through the sample loop of the GC on its way to the bubbler and to vent, allowing thorough, repeatable, reliable flushing and loading of the loop. The GC used in this setup was a Shimadzu GC-14B, pictured in Figure C.2, connected to a Shi- madzu C-R8A Chromatopac integrator, and equipped with two distinct analytical paths. The first utilised an Agilent Technoligies Inc. (Agilent) HP-PLOT U column (19095P-UO4, ID 0.530 mm, length 30 m, film 20.00 µm), shown as Column 1 in Figure C.3, and an FID, the second an Alltech Carbosphere® 80/100 packed column (5682PC, OD 1/8”, length 6’, 316 stainless steel), Column 2 in Figure C.3, with a thermal conductivity detector (TCD). Of the three Vici multiport injection valves visible in Figure C.2, the top two introduce sample to the FID and TCD loop respectively whilst the bottom valve allows a separate reference gas to sent to the TCD. One drawback to the use of an FID, as opposed to TCD, is that an FID is

290 Figure C.1: Hewlett Packard 5890A gas chromatograph used for ofﬂine gas sample analyses. only able to detect hydrocarbon products and not unreacted hydrogen or species such as H2S. This is because the signal detected by an FID is the current generated between the electrodes due to the pyrolysis of hydrocarbons to cations and electrons. In general, however, an FID is more sensitive than a TCD to hydrocarbon products and, given that the products of interest in this investigation were hydrocarbons, only the FID was used in this study. The air pressure to the FID was set at 60 kPa, hydrogen to the FID at 60 kPa, helium injector split at 40 kPa and helium to the column at 40 kPa. The temperature program is presented in Table C.2. One complication in the analysis of gas samples from the micro-reactor was that not all of the gas vented from the system partook in the reaction. Through trial and error (by analysing gas samples at different points during depressurisation) it was determined that the majority of unparticipating hydrogen gas could be removed by allowing the system to drop down to approximately 70 kPa before injecting the sample. Once at this target pressure, beginning the GC analysis simultaneously in- jected the sample onto the column, began the temperature program and started data collection on the integrator. Once the analysis was complete, the integrator pro-

291 vided a printout of the data gathered. Additionally the data was collected from the integrator for digital backup and further analysis using Shimadzu’s Chromatopac Data Archive Utility (v3.0).

Table C.2: Temperature program for Shimadzu GC-14B gas chromatograph analyses.

Speciﬁcation Value Initial temperature (◦C) 35 Hold time (min) 3 Ramp rate (◦C/min) 20 Final temperature (◦C) 170 Hold time (min) 20.3

C.1.2 Calibration, Data Acquisition, Analysis and Interpretation Figure C.4 show two chromatograms obtained from the GC-14B, with those from the HP5890A taking the same form (but the data not available digitally). As may be seen, the data obtained from such an instrument is of the form of retention time (x- axis) versus response (y-axis) which, together with suitable calibration data, allow for identification and quantification respectively. The response signal is proportional to the electrical charge (usually in the range of mV) detected by the FID. As this signal is generated by the cations and electrons formed during pyrolysis of hydrocarbons, the more carbon in a molecule (i.e. the larger the hydrocarbon), the greater the signal per molecule. The magnitude of this signal is analysed in terms of peak height, shape and area. The height and shape of a peak are determined by the nature of the sample separation and species elution from the column. The area of each peak is used together with a calibrated response factor, discussed below, to quantify each species. For a given area, i.e. a fixed amount of species, the height of the corresponding peak will change as the shape of the peak changes. A tall, narrow Gaussian peak is ideal and indicates very sharp, specific separation. A broader peak would be shorter and, whilst quantification may not be an issue for a single peak, broad peaks risk overlapping one another (an indication of multiple species eluting over the same retention time) and hence complicating the quantification of each. Additionally, the peak shape may be used to analyse such operating parameters as column loading whereby an ex-

292 Figure C.2: Shimadzu GC-14B gas chromatograph used for in-line gas analyses. cessive amount of sample saturates the stationary phase of the column and results in a “shark-fin” shaped peak. Calibration of both GCs was performed using a certified hydrocarbon gas mixture (Praxair) with a composition shown in Table C.3. Identification of peaks is performed by injecting a sample of known gas composition and observing the retention time of each species for, with all other conditions unchanged (column type and specifications, flow rates, etc.), the retention time of a given species will remain approximately the same between samples. So for instance by knowing the composition of the calibration gas samples in Figure C.4a

293 Figure C.3: Interior plumbing of Shimadzu GC-14B gas chromatograph used for in-line gas analyses.

(both GCs were calibrated using a certified hydrocarbon gas mixture discussed in Section 3.2.1), one may determine the retention time of each of those species, as indicated. These known retention times may be applied to unknown samples to identify key species, such as in Figure C.4b. The identities of species not included in the calibration mixture may be inferred from their retention time relation to known species or, if their identification is crucial, a sample may be analysed by GCMS (described for liquids in Section 3.3.2). Given the very low concentrations of the uncalibrated species in the gas phase, the “guilty by association” method was deemed suitable for this study. Quantification of species in GC analyses is also performed by analysing the calibration gas sample (one sample will suffice, but ideally a range of samples should be fit and a calibration curve developed). From the data presented in Figure C.4a for the calibration gas, the response factors for each species (e.g. RgcMeth) may

294 Table C.3: Gravimetric composition of certiﬁed calibration gas, supplied by Praxair, for use with gas chromatographs.

Component Composition (wt%) Ethylene 2 isoButane 2 Butane 4 Propane 6 Carbon dioxide 8 Ethane 8 Methane 10 Nitrogen 10 Hydrogen 25 Carbon monoxide Balance be determined, these being the ratio of GC area recorded (e.g. AgcMeth) to known concentration. Equation C.1 shows this calculation for methane as an example with the data and results for this and propane given in Table C.4 for comparison. With these values, the concentration of methane and propane in C.4b may be calculated as shown in Equation C.2 and presented in Table C.4. It should be noted that the chromatogram presented in Figure C.4a was acquired with a detector range of 100, whilst that in Figure C.4b used a detector range of 10. As such a factor 10 applies to the response factors when applied as seen in Equation C.2. From this the sensitivity of the FID may be seen whereby even the relatively large, sharp peak of propane corresponds to only 0.26 wt% in the gas phase. With all other conditions unchanged, these response factors remain constant for a given species between samples and a full list of these values for each GC are provided in Section D.1.

CMeth.cal 10 −6 RgcMeth = = = 4.22 × 10 wt%/counts (C.1) Agcmeth.cal 2370131 4.22 × 10−6 C = Rgc .Agc = .295335 = 0.125wt% (C.2) Meth.samp Meth Meth.samp 10 Knowing that the reaction began with only high purity hydrogen in the gas phase, any hydrocarbon products detected must be the gaseous products of reaction and could thus be identiﬁed, quantiﬁed and interpreted in the context of proposed reaction mechanisms.

295 Table C.4: Area, concentration and response factor data for gas chromatograph calibration of methane and propane.

Species Speciﬁcation Value GC peak area, AgcMeth.cal (counts) 2370131 Calibration concentration, CMeth.cal (wt%) 10 −6 Methane Response factor, RgcMeth (wt%/counts) 4.22 × 10 Sample peak area, AgcMeth.samp (counts) 295335 Sample concentration, CMeth.samp (wt%) 1.25 GC peak area, AgcProp.cal (counts) 3728070 Calibration concentration, CProp.cal (wt%) 6 Propane Response factor, Rgc (wt%/counts) 1.61 × 10−6 Sample peak area, AgcProp.samp (counts) 158971 Sample concentration, CProp.samp (wt%) 0.26 C.2 Liquid Product Analysis

C.2.1 Analytical Equipment and Procedures Whether from the stirred batch reactor or micro-reactor, all liquid products for this study were prepared in the same manner and analysed on the same instrument. This lent a necessary continuity to the results from the two systems with the liquid composition analyses being arguably the most important. These analyses were performed on an Shimadzu GCMS-QP2010 gas chromatograph - mass spectrometer (GCMS), pictured in Figure C.6, equipped with a Shimadzu SHRXI-5MS column (220-94764-02, ID 0.25 mm, length 30 m, ﬁlm 0.25 mm) and an AOC- 20i autosampler with 10 µL syringe, allowing for extremely accurate qualitative and quantitative information to be obtained. Shimadzu’s GCMSsolution (v2.50) software package was used for monitoring, control and data analysis (data collection, peak identiﬁcation and integration). Due to the complexity of this system, a more extensive description of operating parameters is required than for the GCs of Section C.1 as presented in Figure C.5. The autosampler for this unit was equipped with a twelve sample rack, allowing samples to be prepared and loaded for analysis while unattended. The handling and preparation of these samples was a delicate procedure due to the low concentrations and volatilities of some of the products. Following recovery from the reactor and separation of the solid material (discussed

296 24

Methane

Ethane

Ethylene 4

Propane

isoButane

Butane

Signalx10 intensity (-) Residual toluene

00:00 05:00 25:00 30:00

Retention time (mm:ss)

(a)

Benzene 7

Toluene

Methane 4

Ethane

Propane

isoButane

C isomers

5 or Butene

Butane

1 C isomers

6 C isomer Signalx10 intensity (-)

00:00 05:00 10:00 15:00 20:00 25:00 30:00

Retention time (mm:ss)

(b)

Figure C.4: Examples of gas chromatograms obtained from a Shimadzu GC- 14B with Agilent HP-PLOT U column and FID for qualitative and quantitative analysis. (a) Calibration gas (with detector range of 100). (b) Gas sample obtained from a diphenylmethane test in the micro- ◦ reactor at 445 C, 13.8 MPa H2, 4 h, 1800 ppm Mo, 0 RPM (with detector range of 10).

297 in Section 3.3.3), all liquid products were sealed in glass sample vials and stored in a freezer. This step served to minimise the evaporation of volatile species and ensured that the samples being analysed were accurate representations of the reactor product. Due to the sensitivity of the system, being configured for trace sample analyses, even with a split ratio of 100 both column overloading and detector saturation were possible. Column overloading in the GC occurs when more sample is introduced than can interact with the stationary phase. This results in poor peak shapes (such as “shark fins”), resulting in inaccurate area measurements, and poor separation (possibly with overlapping of peaks) which at best makes deconvolution necessary and at worst completely obscures some products. Detector saturation occurs when the concentration of species exiting the GC and being ionised by the MS filament exceeds safe and reliable detection limits of the MS detector. Such exposure not only damages and shortens the usable life of the filament but can damage the detector too. For this reason the detector will automatically shut off when certain levels are reached, protecting itself at the expense of the data not collected. To avoid these scenarios, it was necessary to dilute all liquid samples with decalin prior to analysis and control the MS program to only collect data for the retention times of interest. Each sample was prepared in two dilution ratios, a richer sample for analysis of lower concentration species (such as minor products) and a leaner sample for higher concentration species (such as unreacted model compound). During all analyses, the MS program was set to omit the retention times corresponding to the diluent peaks (in this study that was 9.00 to 14.00 min for decalin) and, for richer samples, those times corresponding to the model compound peaks too. One of the model compounds not used in that particular reaction was added to each diluted sample to act as an internal standard (for instance using DPE as an internal standard in DPM experiments). Calibration preparation and results are discussed in Section C.2 and essentially provide a relation between the ratio of target species and internal standard concentrations (φcal.wt) with the ratio of their areas from GCMS analyses (φcal.A). The dilution ratios and internal standard addi- tions were incorporated into the calibrations for each species for consistency and to reduce the likelihood of calculational errors. The aim of diluting each sample was to reduce the target species concentration

298 (CTarg) to approximately 0.5 wt% to avoid column overloading or detector saturation. The main complication was that, before analysis, the model compound conversion, and hence the concentrations of each species, was unknown. This was overcome by assuming a mediocre conversion for all experiments (20 wt% was found to work well), having a broad range of calibrated concentration ratios (φCal ranging from 0.0 to 2.0 wt%/wt% for model compounds and 0.0 to 4.0 wt%/wt% for products) and aiming for a diluted sample φTarg in the middle, i.e. 1.0. In this manner, the calibration held even for a 100% deviation in the assumed conversion (handling model compound conversions from 0 wt% to 40 wt%) such that sample repreparation due to results lying beyond the calibrated range was minimised. Products from this theoretical 20 wt% conversion sample were assumed to contain 10 wt% each of benzene and toluene, with the large calibration range allowing for the same extent of flexibility in these concentrations. The presence of benzene and toluene added a layer of complexity to these preparation procedures as these species were found to interact with plastic lab equipment, affecting the outcome of quantification analyses in unpredictable ways as discussed in Section C.2.3. Equation C.3 gives an example of dilution mass calculations for DPM. Some initial values are first defined, with the total sample mass (mT) being the amount required to fill the insert of a GC sample vial.

Total sample mass (mg) = mLTotal = 200

Estimated DPM conversion (wt%) = XDPM.est = 20

Estimated DPM concentration (wt%) = CDPM.est = 80

Estimated benzene concentration (wt%) = CBenz.est = 10

Estimated toluene concentration (wt%) = CTol.est = 10

Target species concentration (wt%) = CTarg = 0.5

Target concentration ratio (wt%/wt%) = φTarg = 1.0 3 DPM density at room conditions (g/cm ) = ρDPM = 1.006 3 Decalin density at room conditions (g/cm ) = ρDec = 0.896 3 Benzene density at room conditions (g/cm ) = ρBenz = 0.877 3 Toluene density at room conditions (g/cm ) = ρTol = 0.867

299 C ρ = DPM.est .ρ + Prod.est 100 DPM C C Benz.est .ρ + Tol.est .ρ = 0.9792 g/cm3 100 Benz 100 Tol CTarg m = .m = 1.000 mg DPM 100 T mDPM mProd.added = = 1.250 mg CDPM.est/100 (C.3) mProd.added VProd.added = = 1.28 µL ρProd.est mDPM mDPE = = 1.000 mg φTarg mDec = mT − mDPM − mDPE = 198.000 mg mDec VDec = = 220.98 µL ρDec

It may be noted that in Equation C.3, the volumes of the product being diluted and the decalin were calculated. The reasoning being that these were liquid at room temperature (the DPE internal standard being solid) and far easier to measure as such. When samples were being diluted, an SGE eVol XR analytical syringe was used for liquid measurement and the actual mass of each component was determined on a Sartorius ME5 microbalance to one thousandth of a milligram to ensure accuracy and repeatability. To ensure that no decomposition products were being obscured by the decalin and lost when the GCMS ﬁlament and detector were shut off during solvent elution between 9.00 and 14.00 min, a sample of DPM reaction product was analysed using n-pentane instead of decalin as the solvent (removing instead the initial 2.00 min for solvent elution). The chromatogram obtained is shown in Figure C.7 from which it may be seen that no products elute during the time period of interest. The few small peaks present in the decalin solvent cut time are residual decalin peaks from a previous injection. Note that it was not appropriate to use n-C5 as the solvent for all quantiﬁcation analyses as its high volatility introduced unacceptable experimental error to the results.

300 C.2.2 Data Acquisition, Analysis and Interpretation Similarly to data obtained from a GC, depicted in Figure C.4, the chromatogram from a GCMS analysis also takes the form of retention time versus response as exemplified in Figure C.8. Whilst identification of the peaks may be performed in the same manner as for a GC (by analysing known species and recording their retention times for comparison), a GCMS also records ion fragment spectra. These spectra, usually presented as the dimensionless mass:charge ratio (m/z) versus relative intensity of the signal, may be compared to a library of known species and all compounds in the reaction product, regardless of whether their retention times have been specifically determined or not, may be reliably identified. Figure C.9 gives an example of such a comparison with the experimental spectrum for benzene shown in Figure C.9a and that from the library in Figure C.9b. Quantification of each species was accomplished through the use of internal standard calibration curves. Detailed in Section C.2.3 for toluene, with data for the remaining species in Section D.1, standards for each species of interest were prepared over a selected concentration range with an internal standard added to each sample. Analysis of these samples allowed for the construction of a calibration curve relating the species:standard area ratio to the respective concentration ratio (such as that of toluene shown in Figure C.10). Not every species from the cracking reactions could be acquired as a pure compound for such calibrations, however, and as such some minor compounds were quantified by estimation using the calibration of a species of similar composition, structure and mass (for instance fluorene

[C13H10], at approximately 166 g/mol, by DPM [C13H12], at approximately 168 g/mol). In this manner the GCMS allowed for very accurate information to be obtained for each liquid sample, yielding both the identity and concentration of even those species comprising fractions of a percent by mass. As such, not only could the major observed reaction pathways be determined, but lesser reactions and reactive intermediates could be identiﬁed, quantiﬁed and added to the proposed reaction mechanisms, providing additional clarity and understanding.

301 302

(a) (b)

(c)

Figure C.5: Operating parameters for Shimadzu GCMS-QP2010 gas chromatography - mass spectroscopy analyses of liquid products. (a) AOC-20i autosampler operation. (b) Gas chromatograph temperature and ﬂow program details. (c) Mass spectrometer settings and program. Figure C.6: Shimadzu GCMS-QP2010 gas chromatograph - mass spectrometer used for liquid analyses.

303 25

Benzene

DPM

Toluene

15 5

Decalin solvent cut

(9.00 - 14.00 min)

3 304

Residual cis- and

2 trans-decalin from Signalx10 intensity (-)

previous injection

0 5 10 15 20

Retention time (min)

Figure C.7: Chromatogram for diphenylmethane hydroconversion product using n-pentane as the sample dilution solvent to identify any peaks eluting during the decalin solvent cut time. Experiment conducted using undiluted ◦ diphenylmethane in the stirred batch reactor at 445 C, 13.8 MPa H2, 1800 ppm Mo, 1 h, 700 RPM. 100 DPM

DPE

Benzene

50 5

Toluene

Isomerisation and

condensation

products

5 Signalx10 intensity (-)

Other cracking

products

2 4 6 8 14 16 18 20 22

Retention time (min)

Figure C.8: Example of a chromatogram obtained from gas chromatography - mass spectroscopy analysis for liquid obtained from a diphenylmeth- ◦ ane test in the stirred batch reactor at 445 C, 13.8 MPa H2, 6 h, 600 ppm Mo. Note the diphenylethane used as an internal standard.

305 100

Benzene with m/z = 78

10 Relativeintensity (-)

50 55 60 65 70 75 80

Mass:charge ratio, m/z (-)

(a)

100

10 Relativeintensity (-)

50 55 60 65 70 75 80

Mass:charge ratio, m/z (-)

(b)

Figure C.9: Identiﬁcation of benzene by comparison of sample and known ion fragment spectra. Note that m/z ¡ 50 is not displayed as peaks due to species such as O2 (m/z = 32), N2 (m/z = 28), etc. result in false positives. (a) Sample ion fragment spectrum from liquid obtained from a diphenylmethane test in the stirred batch reactor at 445◦C, 13.8 MPa H2, 6 h, 600 ppm Mo. (b) Ion fragment spectrum for benzene from the built-in NIST05 library of Shimadzu’s GCMSsolution v2.50 software package.

306 C.2.3 Calibration, Analysis and Experimental Uncertainty With all of the reagents, except hydrogen, being in the liquid phase at the beginning of each experiment and almost all of the products being in the liquid phase at the end, liquid analysis by GCMS was an integral part in calculating such comparators as model compound conversion, product selectivity and yield and for determining nuances in the product composition crucial to elucidating the reaction mechanism. As such, particular care and attention was paid to the accurate calibration of the GCMS for the major species of interest in this study: DPM, DPE, DPP, benzene and toluene. The results from these calibrations are presented in Section D.2. Whilst DPM, benzene and toluene (the major species with which most of this study was concerned) were calibrated using DPE as an internal standard, the extremely high conversions observed for DPE and DPP (see Section 4.1.1) made it more convenient to instead use a direct response factor calibration (as described in Section C.1.2 for gas samples) for these two species. As seen in the toluene example below, however, such calibrations were almost as accurate as those using an internal standard, provided additional precautions were taken. As an example of calibration preparation and data analysis (including uncertainty analysis and propagation), the calibration for toluene is provided below (with complete data in Section D.2.2). Toluene was selected as this example because it, and benzene to a lesser extent, complicate procedures due to their high volatilities and interactions with plastics. Through a discussion of these effects, the use of various materials and procedures may be explained and the “worst case” uncertainty calculated (this being for a volatile species prone to additional complications). With the major products from all three model compounds being benzene and toluene (as discussed in Chapter 3), there existed a clear difference in the volatility of products and reagents. It was thus necessary to take steps to ensure the integrity of the samples for, if a sample were exposed and species allowed to escape, the products would do so preferentially, resulting in a bias toward observed lower conversion of the model compound (as products escape, the unreacted model compound is concentrated in the sample). For this reason, all samples were tightly sealed and stored in a freezer immediately after recovery and only removed when being diluted for analysis.

307 An additional complication associated with benzene and toluene products is their interaction with plastics; most apparent for toluene, benzene exhibits this behaviour to a lesser extent. When in contact with certain plastics, including many of those utilised in laboratory equipment and supplies (such as polypropylene), toluene will be absorbed into the plastic. The rate and extent of this absorption varies with the type of plastic, area in contact, plastic thickness, toluene concentration, etc. rendering prediction of variations caused by this phenomenon virtually impossible. To further obscure the situation, this absorption process is not apparent from calibration data. When measuring pure toluene for calibration preparation, all plastic surfaces appear to rapidly become saturated, with the weighed and analysed amounts producing accurate calibration curves. When dealing with reaction products, however, the low concentrations of toluene meant that a relatively large percentage was absorbed from the sample. With toluene being the most susceptible of the species in such liquids, this absorption phenomenon most commonly manifested as a very low toluene concentration, leading to the misinterpretation of results. To overcome this problem, all plastic items were removed from handling and preparation procedures. This included pipettes, pipette tips, vial inserts and vial seals, being replaced with polytetrafluoroethylene (PTFE, also known as Teflon®.), which did not appear to exhibit this interaction, glass or stainless steel. For this reason, all product samples were stored in glass vials with PTFE seals, glass vial inserts were used, liquids were recovered with glass syringes and stainless steel needles and all volumetric measurements were conducted with an SGE eVol XR with syringes made of glass, stainless steel and PTFE. Sure of reliable sample preparation procedures, a series of calibration standards could be prepared, using the dilution methods presented in Equation C.3, for a range of DPE/species concentration ratios (φCal). When analysed in the GCMS, φ φ ′ these Cal values could be plotted against their respective area ratios ( Cal) to determine the internal standard calibration curve as presented in Figure C.10. This curve was obtained by linear regression using the weighted least-square method in OriginLab Corporation’s OriginPro (v8.6.0) software package which works by minimising the sum of squares of the residual values (the differences between the fitted curve and the experimental results). The internal calibration of toluene with DPE is thus defined per Equation C.4, with an R2 of 0.996, without forcing the

308 origin point and per Equation C.5, with an R2 of 0.999, forcing the fit through the origin. It is thus clear that, despite the general rule being to not force the origin point in such calibration curves, the accuracy of this data allows for a slight improvement of the fit with the removal of the intercept degree of freedom offering a slight decrease in the uncertainty in the slope from approximately ±1.5% to less than ±0.9%. This improvement is echoed by the increase in the R2 value when the curve is forced through the origin. The R2 value reported is the coefficient of determination is calculated as R2 = SSresid 1 − , where SSresid is the sum of squares of the residuals and SStotal is the SStotal total sum of squares (the difference between the experimental and the average). φ φ ′ φ ′ Cal = aint. Cal + bint =(1.93 ± 0.02). Cal − (0.01 ± 0.01) (C.4) φ φ ′ φ ′ Cal = aint.0. Cal =(1.91 ± 0.01). Cal (C.5)

4.5

Linear fit through origin

3.0

1.5 Concentrationtoluene/DPE ratio, (wt%/wt%)

0.0

0.00 1.25 2.50

GCMS area ratio, toluene/DPE (-/-)

Figure C.10: Gas chromatograph - mass spectrometer calibration curve for toluene using DPE internal standard. Figure C.11 shows the same calibration data as Equation C.4 but without the internal standard. What is notable here is the presence of three distinct trends corresponding to samples analysed on three different days. Each trend is well deﬁned and has a low uncertainty in and of itself, but compared to one another, these curves are clearly not compatible. The reasoning for these daily samples to present different response factors and yet, when considered with their internal standard to show

309 negligible differences, was established to be due to minor day-to-day ﬂuctuations in laboratory conditions. When working with such precision, environmental factors such as laboratory temperature, humidity or even the time of day or weather outside (which may inﬂuence the temperature based on the amount of sunlight entering through the windows) can affect volumetric and mass measurements and results in the variations observed. With an internal standard, such variations affect both the species of interest (i.e. the toluene) and the standard equally, essentially canceling one another out.

2.5

2.0

1.5

1.0

Tolueneconcentration (wt%) 0.5

0.0

0 5 10 15

T oluene GCMS area (-) x10

Figure C.11: Gas chromatograph - mass spectrometer calibration curve for toluene without an internal standard.

−7 2 × - Day 1 - CTol.samp = aext1.0.AgcTol.samp =(2.18 ± 0.01) × 10 .AgcTol.samp, R = 1.000. −7 2 ◦ - Day 2 - CTol.samp = aext2.0.AgcTol.samp =(1.90 ± 0.01) × 10 .AgcTol.samp, R = 1.000. −7 2 △ - Day 3 - CTol.samp = aext3.0.AgcTol.samp =(1.50 ± 0.02) × 10 .AgcTol.samp, R = 0.999.

To evaluate the quality of these calibration curves and quantify the experimental error associated with both sample preparation and analysis, six separate toluene samples of known concentration were prepared and each analysed three times. The results from these preparations are presented in Table C.5. As with all datasets obtained in this study, the experimental values were evaluated for outliers [154], as shown in Equation C.6 for the toluene concentrations. Here the ﬁrst and third quartiles are, Q1 being the value below which 25% of the measurements

310 lie and Q3 being the value above which 25% lie, are used to determine the interquartile range, IQ. In this study both minor and major outliers were determined, the former being potentially anomalous readings requiring additional evaluation and the latter being those readings disregarded as abnormal for some reason or other (sample contamination, power surges during analysis, human error, etc.). Minor outliers were those measurements lying outside the “inner fence”, the boundaries deﬁned by Q1 − 1.5 × IQ < x < Q3 + 1.5 × IQ, but still within the “outer fence”, Q1−3×IQ < x < Q3+3×IQ. Major outliers lay beyond the outer fence. As may be seen from Equation C.6 and Table C.5, none of the readings may be classiﬁed as even minor outliers.

Equation C.7 shows the calculation of the mean (C¯Tol.samp), standard deviation

(sTol.samp) of the toluene concentration data (assuming that this data represents a sample rather than the full population). The average uncertainty (¯s) associated with the combined volumetric and mass measurements (based on both the toluene and DPE samples shown) may thus be calculated as being ±0.004 wt% or ±0.745%.

Table C.5: Toluene and diphenylethane concentrations in known samples for preparation uncertainty determination.

Toluene DPE Sample (wt%) (wt%) 1 0.485 0.501 2 0.487 0.494 3 0.490 0.498 4 0.490 0.491 5 0.489 0.502 6 0.494 0.499 C¯ 0.489 0.497 s 0.003 0.004 s¯ 0.004

311 Q1 = 0.487 Q3 = 0.490 IQ = 0.003 1.5 × IQ = 0.0045 (C.6) 3 × IQ = 0.0090 Inner fence = [0.483,0.495] Outer fence = [0.478,0.499]

n ∑Ci C¯ = i=1 Tol.samp n 6 ∑Ctol.samp.i = i=1 6 = 0.489 wt%

∑n ¯ (C.7) i=1 Ci −C v n u ! sTol.samp = u u (n − 1) u t ∑6 ¯ i=1 (CTol.samp.i −CTol.samp) v u n ! = u u 5 u = t0.003 wt% Table C.6 displays the area results for the three analyses for each of the ﬁrst two samples, each datapoint representing the toluene peak area detected by the GCMS for each injection. These datasets may be analysed in the same manner as the concentrations in Equation C.7, to obtain the mean and standard deviation results indicated. Continuing these calculations for all eighteen injections, the overall average uncertainty associated with the GCMS areas may be determined to be ±20,503 or ±0.788%. ¯ ¯ With this AgcTol.samp, and corresponding AgcDPE.samp, data for the ﬁrst tolu-

312 Table C.6: GCMS areas for multiple injections of toluene samples of known composition for analysis uncertainty determination.

Sample Analysis number GCMS area, Agc Agc¯ s s¯ (-) (-) (-) (-) (-) 1 2,586,457 1 2 2,575,317 2,587,289 12,408 3 2,600,092 11,325 1 2,586,146 2 2 2,598,341 2,587,493 10,241 3 2,577,992 ene sample, Equations C.8 and C.9 use the internal standard and response factor calibrations (from Figures C.10 and C.11 respectively) to calculate the toluene concentration in the original sample. Table C.7 presents such results for all six of the samples together with the overall mean and standard deviation values. For all six samples, the uncertainty associated with repeated preparations and use of the internal standard calibration is thus approximately ±0.003 wt% or ±0.55%, and for response factor calibration is ±0.003 wt% or ±0.60%. φ ′ CTol.samp.int = aint.0. Cal.CDPE.samp ¯ AgcTol.samp = aint.0. ¯ .CDPE.samp AgcDPE.samp 2,587,289 (C.8) = 1.91. .0.501 4,951,566 = 1.91 × 0.523 × 0.501 = 0.501 ± 0.003 wt%

Ctol.samp.ext = aext2.0.AgcTol.samp = 1.90 × 10−7.2,587,289 (C.9) = 0.491 ± 0.003 wt% Propagation of uncertainty from the calibrations, sample preparation measurements, GCMS analyses and repeated sampling may be performed to determine the overall uncertainty associated with the calculated concentrations per Equation C.10 for the general function z = x.y using standard deviations. This calculation is illustrated for the ﬁrst toluene sample using internal standard and response fac-

313 Table C.7: GCMS areas and associated concentrations for repeated preparation uncertainty determination.

¯ Sample AgcTol.samp CTol.samp.int CTol.samp.ext (-) (wt%) (wt%) 1 2,587,289 0.501 0.491 2 2,587,493 0.493 0.491 3 2,599,888 0.502 0.493 4 2,602,980 0.499 0.494 5 2,606,359 0.495 0.495 6 2,619,912 0.504 0.497 C¯ - 0.499 0.494 s¯ - 0.003 0.003 tor calibrations in Equation C.11 and Equation C.12 respectively. Note that in Equation C.11, ∆φ is calculated using the uncertainties from the toluene and DPE

GCMS areas as per Equation C.13 and ∆CDPE and CDPE are from the preparations presented in Table C.5 and CTol is the result obtained from the calibration as shown in Table C.7. It may thus be seen that using the internal standard calibration results gives a slightly higher uncertainty of ±0.005 wt%, corresponding to ±1.0%, as compared with response factor calibration with an uncertainty of ±0.004 wt%, or ±0.8 %.

∆z ∆x 2 ∆y 2 = + z x y s (C.10) ∆x 2 ∆y 2 ∆z = z. + x y s φ ′ CTol.samp.int = aint.0. Cal.CDPE.samp

∆a 2 ∆φ ′ 2 ∆C 2 ∆C = C . + + DPE Tol Tol a φ ′ C s DPE 0.01 2 0.003 2 0.004 2 (C.11) = 0.501 × + + 1.91 0.523 0.501 s = 0.501 × (0.005)2 +(0.005)2 +(0.007)2 = 0.005 wt%q

314 CTol.samp.ext = aext2.0.AgcTol.samp

∆a 2 ∆Agc 2 ∆C = C × + Tol Tol a Agc s 2 0.01 × 10−7 12,408 2 (C.12) = 0.491 × + s 1.90 × 10−7 2,587,289 = 0.491 × (0.006)2 +(0.005)2 = 0.004 wt%q

Agc φ ′ = Tol AgcDPE ∆Agc 2 ∆Agc 2 ∆φ ′ = φ ′. Tol + DPE Agc Agc s Tol DPE 12,408 2 2,195 2 (C.13) = 0.523. + 2,587,289 4,951,566 s = 0.523. (0.0048)2 +(0.0004)2 = 0.003 q Applying these uncertainties to each of the concentrations used to calculate C¯ in Table C.7, as per Equation C.7, and including the contribution froms ¯ in the same table, the ﬁnal uncertainties are determined as shown in Equation C.14. It may thus be seen that any toluene concentration calculated in this study using the internal standard calibration curve of Figure C.10 is subject to an uncertainty of ±4.7% whilst using the Day 2 data from Figure C.11 yields 3.0%.

315 6 ∆C 2 s¯ 2 ∆C¯ = C¯ × ∑ Tol.i + Tol Tol Tol C C¯ si=1 Tol.i Tol 0.003 2 ∆C¯ = 0.499 × 0.0005 + Tol.samp.int 0.499 s = 0.024 wt% (C.14)

0.003 2 ∆C¯ = 0.494 × 0.0002 + Tol.samp.ext 0.494 s = 0.015 wt%

The true comparator, however, is how well these concentrations, obtained using each calibration method, relate to the prepared toluene concentration, indicated in Table C.5, of 0.489±0.003 wt%. It may thus be seen that, for these samples, the internal standard calibration identiﬁed a value approximately 2.0% higher whilst the response factor result was only 0.9% higher. Both of these lie well within their respective experimental uncertainties of ±4.7% and ±3.0% respectively. From these values it may appear that the response factor technique is more accurate, resulting in a lower overall uncertainty and better estimation of the “true” concentration. From Figures C.10 and C.11 it should, however, it should be noted that whilst this same internal calibration curve (and associated uncertainty) holds for the full range of investigation, the response factor calibration data used for the above calculations is but one of three that could have been selected (combining would have increased

∆aext and hence ∆Ctol.samp.ext). So whilst the response factor methodology is comparatively accurate, overall, the internal standard methodology is preferred despite slightly elevated uncertainties in some situations. The experimental uncertainty of ±4.7% calculated for toluene was assumed to be a “worst case” quantiﬁcation as toluene is a volatile species and was seen to interact with some laboratory equipment, both factors which would negatively impact on reproducibility. This quantiﬁcation was thus extended to the interpretation of the DPM, DPE, DPP and benzene results.

316 Reaction Comparator Calculation In this study three main comparators are used for evaluating each reaction, conversion, yield and benzene:toluene molar ratio. With a sample suitably recovered, diluted and analysed by GCMS as discussed above to obtain the mass composition of each species, the comparators may be calculated with relative simplicity. Using data from 0 ppm Mo, 2 h, 0 RPM in the glass insert micro-reactor, the reaction began with undiluted DPM and product contained 84.7 wt% DPM, 8.5 wt% toluene and 7.2 wt% benzene (totaling slightly above 100 wt% due to experimental uncertainty. The fractional DPM conversion, XDPM, may be calculated as ′ shown in Equation C.15. The mass and molar yields of toluene, YTol and YTol, are shown in Equation C.16 and Equation C.17 respectively. These yields are the ratio of desired product (in this case toluene) to total reactant consumed (DPM) as deﬁned by Fogler [112] and Sinnott [155]. Note that the “reaction yield” is used rather than the “plant yield” [155] as physical losses in this study are minimal. The B:T molar ratio, BT, is calculated per Equation C.18.

CDPM.in −CDPM.out XDPM = CDPM.in c .V − c .V = DPM.in LTotal DPM.out LTotal cDPM.in.VLTotal c − c = DPM.in DPM.out (C.15) cDPM.in 1.00 − 0.847 = 1.00 = 0.153

′ mTol YTol = mDPM.in − mDPM.out c .V = Tol LTotal cDPM.in.VLTotal − cDPM.out.VLTotal cTol = (C.16) cDPM.in − cDPM.out 0.085 = 1.00 − 0.847 = 0.56

317 nTol YTol = nDPM.in − nDPM.out ′ MrDPM = YTol. MrTol 168.23 (C.17) = 0.56 × 92.14 = 1.01

n BT = Benz nTol m .Mr = Benz Tol mTol.MrBenz c .V .Mr = Benz LTotal Tol cTol.VLTotal.MrBenz c .Mr (C.18) = Benz Tol cTol.MrBenz 0.072 × 92.14 = 0.085 × 78.11 = 1.00

C.3 Solid Product Analysis

C.3.1 Analytical Equipment and Procedures Given that no solid precipitation products were expected to form, the solid product analyses of this study served to: conﬁrm the formation of the desired MoS2 active phase from the liquid precursors, identify any contaminants or unexpected solid products and determine the structure and particle size of the MoS2 and other solid species. The recovery of solid material was, in most cases, a trivial solid-liquid separation followed by washing and drying. For both the stirred batch and micro- reactor units, the solid catalyst particles formed from the in-situ reaction of the oil-soluble catalyst precursor, molybdenum octoate, with sulphur compounds (either the CS2 added to the system or sulphides formed from its reaction with other species). It was surmised that the solubility of the precursor and its even distribution in the reaction mixture meant that formation of the MoS2, and the precipitation of particles thereof, occurred uniformly throughout the liquid. Two alternative pro-

318 cesses were examined. The first was the initial precipitation of particles in isolated zones of slightly higher concentration, on solid contaminant particles in the mixture (such as dust) or on defects on the reactor internals. These particles would them move through the reaction mixture acting as nucleation sites whereupon additional material could crystallise. The second theory was that the particles formed in isolation, however, once formed, these particles could agglomerate by electro- static forces to form larger crystallites (through either growth of the sheets or the addition of stack layers). In all situations, the solid species were recovered from the reactors as particles suspended in the liquid product. Being a straightforward solid-liquid separation, numerous techniques could be employed to isolate both the solid and liquid phases. In this study, the larger volume of stirred batch product made vacuum filtration an efficient option whilst the far smaller volume of micro-reactor product necessitated a step-wise settling and liquid removal process. For the stirred-batch product, the solid-liquid mixture was agitated to suspend all of the particles and quickly introduced to the funnel of a vacuum filtration system equipped with a filter membrane (with a pore size of 0.22 µm). With all of the liquid drawn through the filter, the system was stopped and the liquor recovered for analysis. The solids were then washed using three 100 cm3 volumes of acetone to remove residual liquid reaction products before being placed in a drying oven at 100◦C for 12 h to dry. For micro-reactor products, the solid particles were allowed to settle for 12 h with the liquid then pipetted off for later analysis. The solid slurry remaining was then washed with three 25 cm3 volumes of acetone, allowing the solids to settle and removing the liquid as before between washes, to remove residual liquid reaction products. The solids were them spread on a clean glass plate and placed in a drying oven at 100◦C for 12 h to dry. For the solid analysis techniques used in this study, size reduction was a standard preparation procedure with samples requiring grinding and/or sonication to both reduce them into a handleable form and suspend them as necessary. Due to the nature of the formation of the MoS2 particles (precipitating as nanometer-scale particles and clusters), however, size reduction for analysis was not necessary, with each solid sample needing simply to be appropriately mounted for analysis.

319 Compositional analysis of solid samples was performed by X-ray diffraction (XRD) and scanning electron microscopy with energy dispersive X-ray spectroscopy (SEM/EDX). Bruker D8 Focus Bragg-Brentano diffractometer, pictured in Figure C.12, operated at 10-80◦ 2θ, with a step size of 0.0386◦ 2θ using CoKα radiation, an Fe monochromator foil, 0.6mm (0.3◦) divergence slit, incident- and diffracted-beam Soller slits and a LynxEye detector was used for the XRD analyses. The long finefocus Co X-ray tube of the instrument was operated at 35 kV and 40 mA, with a take-off angle of 6◦. Data was analysed using the International Cen- tre for Diffraction Data PDF-4+ database. SEM/EDX analyses were performed on a Hitachi S-2600N variable pressure scanning electron microscope (Hitachi High- Technologies, Tokyo, Japan), pictured in Figure C.13, with a Gresham Scientific Instruments (now SGX Sensortech (MA) Ltd) Antares 10 (SiLi) detector at 1,500 times magnification for 240 s with beam energy of 30 keV and spot size of 70. Data was acquired with and analysed by Quartz Imaging System’s XOne (v8.0) software package. Structural information relating to the recovered solids was obtained by transmission electron microscopy (TEM) and field emission scanning electron microscopy (FESEM). TEM analysis was performed on an FEI Tecnai G2 (FEI Company Eindhoven, Netherlands), Figure C.14, operated at 200 kV with a LaB6 filament. FESEM was performed on a Hitachi S-4700 field emission scanning electron microscope (Hitachi High-Technologies, Tokyo, Japan), shown in Figure C.15, operated at up to 5 kV and magnifications up to 250,000 times.

C.3.2 Calibration, Data Acquisition, Analysis and Interpretation XRD allows for both the identiﬁcation of the solid species present in a sample and an estimation of their crystallite size. The diffractogram obtained from an XRD analysis (such as that in Figure C.16) plots an angle (2θ, corresponding to the angle of the incident X-ray beam and detector with regards to the sample) against counts (a measure of magnitude of the diffracted signal), whereby the incident X-Ray beam is diffracted through the crystal lattice and the returning radiation measured and recorded. The angle of diffraction is characteristic of crystals of a given species and as such, by comparison with a database of known materials, the

320 Figure C.12: Bruker D8 Focus Bragg-Brentano diffractometer used for XRD analyses. Reproduced with permission from Mati Raudsepp, Direc- tor of Electron-Microbeam/X-Ray Diffraction Facility, Department of Earth, Ocean and Atmospheric Sciences, UBC. element corresponding to each peak may be identified. Quantification of relative phases is possible with XRD data, but was not performed in this study (due to the minimal presence of contaminant material). Aside from quantification, the dimensions of the peaks may be used to determine the average crystallite size, p, using the Scherrer equation [156] shown in Equations C.19 and C.19 for the basal peak of MoS2 (where K is the Scherrer constant, λ is the X-ray wavelength

321 Figure C.13: Hitachi S-2600N variable pressure scanning electron microscope used for SEM/EDX analyses. Copyright UBC Bioimaging Fa- cility. Reproduced with permission from Garnet Martens, Research Manager, UBC Bioimaging Facility.

[Kα.ave = 0.17902 nm], b is the peak broadening factor and θ is the Bragg angle of the peak). This basal peak corresponds to X-rays diffracted from the 002 plane of the crystal lattice and was extracted from the Figure C.16 dataset and analysed using OriginPro, the results being presented in Figure C.17. From this analysis, two values could be obtained for the peak broadening factor, the “full width at half maximum” (FWHM) and the integral breadth (the integrated peak area divided by its intensity, often indicated as β rather than b). The value of the Scherrer constant depends on which of these broadening factors is used, the shape of the crystallites and the crystallite size distribution [156]. K = 0.76 was used as an estimation for the 002 peak [157], with the resultant sizes obtained being 2.15 nm and 2.01 nm using the FWHM and integral breadth broadening factors respectively. These dimensions correspond to the stack height of the MoS2 crystallites. Knowing the average inter-plate spacing to be 6.16 A˚ [106], these stacks may be determined to be between and three and four layers thick. The preparation of samples for such an XRD analysis involves making a slurry with a small amount of the washed, dried recovered solids in ethanol. This slurry

322 Figure C.14: FEI Tecnai G2 used for TEM analyses. Copyright UBC Bioimaging Facility. Reproduced with permission from Garnet Martens, Research Manager, UBC Bioimaging Facility. is used to coat the surface of a glass sample plate. When dried, the coated plate is loaded into a sample holder and into the instrument for analysis.

323 Figure C.15: Hitachi S-4700 ﬁeld emission scanning electron microscope used for FESEM analyses. Copyright UBC Bioimaging Facility. Re- produced with permission from Garnet Martens, Research Manager, UBC Bioimaging Facility.

K.λ p = b.cosθ K = 0.76 λ = 0.17902 nm (C.19) 2θ π θ = × = 0.142 rad 2 180π b = 3.67 × = 0.064 rad 180 p = 2.15 nm

324 4000

MoS {002}

MoS

3000

2000 Counts(-)

1000

Carbon (C, graphite)

20 40 60 80

Angle, 2 (°)

Figure C.16: Example of an X-ray diffractogram for solid material obtained from a diphenylmethane test in the stirred batch reactor at 1800 ppm ◦ Mo, 445 C, 13.8 MPa H2, 1 h, 700 RPM.

K.λ p = β.cosθ K = 0.76 λ = 0.17902 nm (C.20) θ = 0.142 rad Area π β = × = 0.069 rad Height 180 p = 2.01 nm TEM analysis allows for high resolution, high magnification images to be obtained (such as that shown in Figure C.18). Such an instrument focuses a high energy beam of electrons at a thin sample. This electron beam interacts with the sample, with the electrons being absorbed, transmitted through or undergoing some form of wave interaction before being detected by an imaging device (such as a flu- orescent screen for quick visualisation or a camera for image capture). From the images produced, the stacked plate structure of MoS2 may be confirmed, the length of the plates may be measured and the number of plates in each stack counted, al-

325 2500

Height = 2246

FW HM = 3.67°

1250

Area = 8822 Counts(-)

Peak location = 16.25°

10 15 20 25

Angle, 2

Figure C.17: MoS2 basal peak extracted and analysed from X-ray diffractogram for solid material obtained from a diphenylmethane test in the ◦ stirred batch reactor at 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 700 RPM. lowing for both particle size and stack height distributions to be determined. This process may be performed using image processing software or manually, as shown in Figure C.19, but measuring and counting each sheet in a given TEM image. The information obtained may be related to the rim-edge activity model discussed in Section 2.5.2 to help explain observed activity and/or selectivity trends. Addition- ally, the wave interaction patterns observed on TEM images may be measured to obtain the d-spacings, the distances between the planes in the atomic lattice (both between plates (d002) and between atoms within each plate, d100), and hence identify the material being studied through comparison with known crystal structures [106]. Sample preparation for TEM involves sonication (a Cole-Parmer Ultrasonic

326 Cleaner, 08895-12, 100 W, 42 kHz being used in this study) of a small amount of solid sample in ethanol to form a lightly coloured but translucent suspension. A single drop of this suspension is introduced to a copper TEM grid and allowed to air dry. These grids may then be loaded into the TEM for analysis. It is important not to “over-sonicate” the suspension as this may break apart the MoS2 structures and obfuscate analysis.

Figure C.18: Example of a transmission electron microscopy image, for solid material obtained from a diphenylmethane test in the stirred batch re- ◦ actor at 600 ppm Mo, 445 C, 13.8 MPa H2, 1 h, 700 RPM, showing the both the inter-plate spacing (d002 = 0.62 nm) and the d-spacing (d100 = 0.27 nm). SEM/EDX analysis can be used to determine three pieces of information, it: gives a fairly low magniﬁcation SEM image to allow for a rough structure to be observed, determines the elemental composition of the sample being observed (as the particles are so small, this is considered to be the bulk composition) and provides a graphical indication of where the various elements were detected (usually presented as a coloured overlay of the SEM image). An example of SEM/EDX results is presented in Figure C.20 to illustrate the strength of this technique in its simplicity. The instrument is able to generate such a compositional image by sweeping the EDX beam across the sample surface, analysing the composition in small, isolated regions. Unlike XRD where the signal detected is the diffracted

327 Figure C.19: Transmission electron microscopy image, for solid material obtained from a diphenylmethane test in the stirred batch reactor at 600 ◦ ppm Mo, 445 C, 13.8 MPa H2, 1 h, 700 RPM, showing manual iden- tiﬁcation of sheets for size and stack height distribution analysis. incident beam, in EDX the signal results from electrons moving between orbital shells. As the energy of the beam excites and expels electrons from the inner shells of atoms, electrons from outer shells drop down to ﬁll these “electron holes” in the orbitals, emitting excess energy as X-rays in the process. The amount of energy emitted is characteristic of the element and the orbitals the electrons are moving between (e.g. Kα or Kβ). The spectrum in C.20b shows the intensity (measured

328 in counts) against the energy of the X-ray beam (in keV). The graphical representation of this compositional analysis, as shown in Figure C.20a, allows for the close correlation between the Mo and S to be easily seen whilst the resulting spectrum in Figure C.20b, showing counts (again a measure of magnitude) versus X-ray beam energy (in eV), allows for relative quantiﬁcation of all detected elements to be determined (true quantiﬁcation is possible through the use of standards but was not performed in this study). As may be seen, the major elements detected in the region of sample analysed are Al, S, Fe and Mo.

The S and Mo are due to the MoS2 catalyst, the Al is from the sample mount (the X-rays penetrating through the sample to the support itself) and the Fe is present as FeS which was picked up from the reactor walls. This technique, together with the XRD analyses, allow for even trace amounts of contaminant crystallite species to be identified. Note that both rely on crystal diffraction and hence are unable to detect amorphous carbon deposits, although graphitic species may be identified. To prepare a sample for SEM/EDX, a small square of adhesive carbon tape is cut to fit the top of an instrument-specific aluminium sample “stub”. The film is removed from one side of the tape which is then stuck to the top of the stub and firmly rubbed to adhere. The remaining side of film is removed to expose the adhesive surface and a small amount of the solid sample gently pressed onto the tape to mount it. Excess sample is tapped off. This stub may then be mounted in the instrument for analysis. FESEM was used for visual inspection of solids recovered from many of the micro-reactor experiments, such as that portrayed in Figure C.21. This technique allowed for extremely high resolution, high magnification images to be obtained for examination of the surface structure of the solid material (something completely invisible in the penetrating TEM images). Both SEM and FESEM function by scanning across the surface of the sample with a focused electron beam. The electrons in this beam interact with the atoms on the surface resulting in various sig- nals which may be detected (secondary electrons, back-scattered electrons, X-rays, etc.) with the instrument used in this study, as with most SEM units, detecting the secondary electrons (these being electrons emitted from the atoms after excitation by the electron beam). The difference between SEM and FESEM is the electron source with the former employing a thermionic emitter (electrical current heats a

329 filament to produce the electron beam) and the latter a field emitter (the electron beam is generated by exposing the filament to a suitable potential difference field). The surface view obtained from the FESEM allowed for additional details to be observed in terms of particle dimensions, structure and interactions, lending further information for use in the explanation of observed phenomena. This technique is purely a visual observation of the surface and did not yield any information relating to composition or internal structure. Preparation of samples for FESEM analysis was identical to that of SEM/EDX and, indeed, the sample stubs used in this study were interchangeable, allowing the same samples to be analysed on both instruments.

330 (a)

9.0

7.5

S (K) 4 331

Mo (K )

Fe (K )

Al (K )

Mo (K ) Fe (K ) Signalx10 intensity (-)

0 5 10 15 20

Detected X-ray energy (keV)

(b)

Figure C.20: Example of results from scanning electron microscopy with energy dispersive X-ray spectrum for solid ◦ material obtained from a diphenylmethane test in the micro-reactor at 445 C, 13.8 MPa H2, 4 h, 1800 ppm Mo, 0 RPM. (a) Microscopy image with S and Mo elemental overlay. (b) Spectrum for elemental identification and relative quantification. Figure C.21: Example of field emission scanning electron microscopy image for solid material obtained from a diphenylmethane test in the micro- ◦ reactor at 445 C, 13.8 MPa H2, 1 h, 1800 ppm Mo, 2250 RPM.

332 Appendix D

Calibration Data

D.1 Gas Chromatograph

D.1.1 HP5980A Calibration Results Table D.1 shows the response factor results from analysis of the calibration gas (Table C.3) in the HP5980A GC.

Table D.1: Response factors for calibration of the HP5980A gas chromatograph.

Species Known composition GC area, Agc Response factor, Rgc (wt%) (counts) (wt%/counts) Methane 10 7,425,950 1.35 × 10−6 Ethane 8 10,907,189 7.33 × 10−7 Propane 6 11,453,962 5.24 × 10−7 Butane 4 9,826,615 4.07 × 10−7 iso-Butane 2 4,976,133 4.02 × 10−7 Ethylene 2 2,816,035 7.10 × 10−7

D.1.2 Shimadzu GC-14B Calibration Results Table D.2 shows the response factor results from analysis of the calibration gas (Table C.3) in the Shimadzu GC-14B gas chromatograph.

333 Table D.2: Response factors for calibration of the Shimadzu GC-14B gas chromatograph.

Species Known composition GC area, Agc Response factor, Rgc (wt%) (counts) (wt%/counts) Methane 10 2,370,131 4.22 × 10−6 Ethane 8 4,034,054 1.98 × 10−6 Propane 6 3,728,070 1.61 × 10−6 Butane 4 3,377,234 1.18 × 10−6 iso-Butane 2 1,401,906 1.43 × 10−6 Ethylene 2 503,226 3.97 × 10−6

334 D.2 GCMS-QP2010 Gas Chromatograph - Mass Spectrometer Liquid Calibration Results

D.2.1 Benzene

4 )

1 Benzene/DPEconcentration ( ratio

0.0 0.5 1.0 1.5 2.0

Benzene/DPE GCMS area ratio ( ')

Figure D.1: Calibration plot for benzene in Shimadzu GCMS-QP2010 with diphenylethane as internal standard.

335 Table D.3: Calibration data for benzene in Shimadzu GCMS-QP2010 with diphenylethane as internal standard.

Standard Sample preparation GCMS areas Ratios for calibration Mixed (mg) Conc. (wt%) (counts) (Benzene/DPE) DPE Decalin Benzene Benzene DPE Benzene DPE Conc. ratio, φ Area ratio, φ ′ 1 3.63 700.185 0.888 0.126 0.515 838,357 5,298,538 0.245 0.158 2 3.557 699.317 1.862 0.264 0.505 1,753,212 5,168,002 0.523 0.339 3 3.528 688.211 2.811 0.405 0.508 2,642,137 5,159,592 0.797 0.512

336 4 3.569 695.314 3.77 0.537 0.508 3,411,393 5,205,174 1.056 0.655 5 3.594 694.434 4.625 0.658 0.511 4,139,088 5,259,052 1.287 0.787 6 3.633 691.105 5.597 0.799 0.519 4,894,086 5,242,716 1.541 0.934 7 3.415 691.304 6.585 0.939 0.487 5,585,137 4,965,357 1.928 1.125 8 3.446 690.186 7.565 1.079 0.491 6,197,846 4,974,188 2.195 1.246 9 0.969 195.242 2.127 1.072 0.489 7,953,868 4,782,252 2.195 1.281 10 1.057 195.338 2.656 1.334 0.531 9,672,426 5,251,979 2.513 1.418 11 0.957 193.59 3.168 1.602 0.484 11,201,302 4,789,779 3.310 1.801 12 0.964 194.277 3.78 1.899 0.484 12,362,103 4,615,135 3.921 2.063 13 1.002 192.652 4.33 2.187 0.506 13,632,215 5,037,153 4.321 2.084 D.2.2 Toluene

4 )

1 Toluene/DPEconcentration ( ratio

0.0 0.5 1.0 1.5 2.0

T oluene/DPE GCMS area ratio ( ')

Figure D.2: Calibration plot for toluene in Shimadzu GCMS-QP2010 with diphenylethane as internal standard.

337 Table D.4: Calibration data for toluene in Shimadzu GCMS-QP2010 with diphenylethane as internal standard.

Standard Sample preparation GCMS areas Ratios for calibration Mixed (mg) Conc. (wt%) (counts) (Toluene/DPE) DPE Decalin Toluene Toluene DPE Toluene DPE Conc. ratio, φ Area ratio, φ ′ 1 3.523 700.057 0.524 0.074 0.500 404,610 5,051,556 0.149 0.080 2 3.478 697.319 0.873 0.124 0.496 678,327 5,075,010 0.251 0.134 3 3.528 696.135 1.784 0.254 0.503 1,384,823 5,135,113 0.506 0.270 4 3.445 694.970 2.684 0.383 0.491 2,067,536 4,982,662 0.779 0.415 338 5 3.442 692.220 3.600 0.515 0.492 2,787,794 5,004,987 1.046 0.557 6 3.495 689.656 4.440 0.636 0.501 3,419,018 5,177,636 1.270 0.660 7 3.524 691.697 5.310 0.758 0.503 4,014,245 5,154,818 1.507 0.779 8 3.552 690.995 6.069 0.866 0.507 4,505,836 5,199,258 1.709 0.867 9 3.526 689.843 6.987 0.998 0.503 5,172,385 5,165,962 1.982 1.001 10 1.050 196.061 1.985 0.997 0.527 7,102,727 5,193,991 1.890 0.971 11 0.959 195.014 2.567 1.293 0.483 8,784,976 4,716,196 2.677 1.323 12 0.977 195.277 3.042 1.526 0.490 10,158,147 4,798,618 3.114 1.503 13 0.988 194.322 3.529 1.775 0.497 11,769,648 4,865,703 3.572 1.717 14 0.960 192.974 4.113 2.077 0.485 13,486,925 4,732,121 4.284 2.024 D.2.3 Diphenylmethane

2 )

1 DPM/DPEconcentration ( ratio

0.0 0.5 1.0 1.5 2.0

DPM/DPE GCMS area ratio ( ')

Figure D.3: Calibration plot for diphenylmethane in Shimadzu GCMS- QP2010 with diphenylethane as internal standard.

339 Table D.5: Calibration data for diphenylmethane in Shimadzu GCMS-QP2010 with diphenylethane as internal standard.

Standard DPE premade Sample preparation GCMS areas Ratios for calibration Mixed (mg) Conc. (wt%) Mixed (mg) Conc. (wt%) (counts) (Toluene/DPE) DPE Decalin DPE DPE premade Decalin Toluene Toluene DPE Toluene DPE Conc. ratio, φ Area ratio, φ ′ 1 245.716 665.174 26.975 12.853 681.312 3.451 0.495 0.497 4,661,512 5,309,275 0.995 0.878 2 245.716 665.174 26.975 12.895 688.898 2.257 0.321 0.494 3,113,331 5,239,430 0.649 0.594 3 245.716 665.174 26.975 12.895 688.898 2.257 0.321 0.494 3,155,526 5,356,137 0.649 0.589 4 245.716 665.174 26.975 12.809 680.497 4.511 0.646 0.495 6,076,890 5,448,990 1.306 1.115 5 245.716 665.174 26.975 12.862 683.110 2.860 0.409 0.496 3,945,113 5,370,835 0.824 0.735 6 245.716 665.174 26.975 12.843 683.896 4.073 0.581 0.494 5,421,251 5,308,569 1.176 1.021 7 200.469 603.124 24.947 13.917 686.879 0.633 0.090 0.495 936,594 5,114,786 0.182 0.183

340 8 200.469 603.124 24.947 13.875 686.725 0.961 0.137 0.493 1,275,945 4,972,450 0.278 0.257 9 200.469 603.124 24.947 13.933 685.710 2.033 0.290 0.495 2,729,404 5,156,575 0.585 0.529 10 200.469 603.124 24.947 13.890 685.295 2.979 0.424 0.493 3,914,514 5,115,121 0.860 0.765 11 200.469 603.124 24.947 13.904 681.976 4.149 0.593 0.495 5,361,116 5,134,233 1.196 1.044 12 200.469 603.124 24.947 13.874 682.949 5.175 0.737 0.493 6,525,471 5,102,867 1.495 1.279 13 200.469 603.124 24.947 13.898 681.807 6.175 0.880 0.494 7,713,206 5,105,073 1.781 1.511 14 200.469 603.124 24.947 13.931 685.469 5.734 0.813 0.493 7,123,896 5,171,007 1.650 1.378 15 200.469 603.124 24.947 13.726 679.948 7.214 1.029 0.489 8,888,683 5,105,855 2.107 1.741 16 200.469 603.124 24.947 13.943 680.745 8.188 1.165 0.495 9,968,157 5,197,672 2.354 1.918 17 3.309 685.015 4.086 0.590 0.478 5,597,086 5,284,038 1.235 1.059 18 3.374 687.350 4.086 0.588 0.486 5,741,374 5,532,217 1.211 1.038 19 3.362 682.241 3.762 0.546 0.488 5,296,159 5,563,796 1.119 0.952 20 Solid DPE 100 3.371 684.548 4.038 0.584 0.487 5,645,219 5,525,425 1.198 1.022 21 3.544 683.323 4.089 0.592 0.513 5,810,875 5,978,547 1.154 0.972 22 3.398 688.000 4.051 0.583 0.489 5,454,038 5,344,533 1.192 1.020 23 3.424 683.790 4.037 0.584 0.495 5,016,213 4,887,447 1.179 1.026 D.2.4 Diphenylethane

0.6

0.4

0.2 DPEconcentration (wt%)

0.0

0 5 10 15

DPE GCMS area (counts) x10

Figure D.4: Calibration plot for diphenylethane in Shimadzu GCMS- QP2010.

341 Table D.6: Calibration data for diphenylethane in Shimadzu GCMS-QP2010.

Standard Sample preparation GCMS areas Mixed (mg) Conc. (wt%) (counts) DPE Decalin DPE DPE 1 0.764 698.791 0.109 2,755,523 2 0.712 698.418 0.102 2,540,666 3 0.796 699.551 0.114 2,856,945 4 1.032 689.974 0.149 3,677,308 5 1.113 696.520 0.160 4,032,473 6 1.076 698.440 0.154 3,827,786 7 1.462 695.939 0.210 4,709,405 8 1.462 695.939 0.210 4,928,414 9 1.462 695.939 0.210 5,066,159 10 1.506 698.707 0.215 5,375,785 11 1.404 697.756 0.201 5,043,322 12 2.149 700.298 0.306 7,385,842 13 2.207 697.340 0.315 7,666,651 14 2.208 698.117 0.315 7,701,429 15 2.939 696.886 0.420 10,024,197 16 2.946 696.224 0.421 10,096,595 17 2.869 697.989 0.409 9,893,400 18 3.557 686.930 0.515 12,123,191 19 3.532 696.718 0.504 12,158,810 20 3.535 695.227 0.506 11,955,526 21 4.360 689.535 0.628 14,366,275 22 4.372 695.384 0.625 14,502,203 23 4.254 693.745 0.609 14,113,274

342 D.2.5 Diphenylpropane DPP was evaluated during model compound screening but was eliminated due to elevated conversion under reaction conditions. Not being used as an internal standard for the more numerous DPM experiments, as DPE was, the calibration data for DPP was not repeated with the internal standard or precision weighing techniques developed later in this study. The lower individual sample accuracy of the technique used in this calibration is accounted for by the preparation of multiple samples at each target concentration level.

1 DPPconcentration (wt%)

0 2 4

DPP GCMS area (counts) x10

Figure D.5: Calibration plot for diphenylpropane in Shimadzu GCMS- QP2010.

343 Table D.7: Calibration data for diphenylpropane in Shimadzu GCMS- QP2010.

Standard DPP conc. GCMS areas (wt%) (counts) 1 0.50 884,169 2 0.50 752,344 3 0.50 722,840 4 0.50 752,346 5 1.00 1,668,361 6 1.00 1,486,631 7 1.00 1,611,593 8 1.00 1,478,398 9 1.50 2,109,167 10 1.50 2,227,456 11 1.50 2,150,622 12 2.00 2,974,816 13 2.00 2,967,186 14 2.50 3,669,059 15 2.50 3,319,280 16 3.00 4,440,866 17 3.00 4,139,100 18 3.00 4,604,176

344 Appendix E

Detailed Experimental Results

This appendix contains unabridged data for the results provided in Chapter 4 and is arranged in the same order as presented therein. Many tables contain references to “sample codes” which were simply an experiment designation system used during this study.

E.1 Stirred Batch Reactor

E.1.1 Model Compound Screening Much of the data shown in Section 4.1.1 is tabulated and self-contained with the exception being the DPP conversion as a function of reaction temperature in Figure 4.1. This data is presented in Table E.1.

Table E.1: Conversion results for diphenylpropane hydroconversion in the ◦ stirred batch reactor at 420 - 445 C, 13.8 MPa H2, 1 h, 600 ppm Mo, 700 RPM at 3 wt% in decalin.

Reaction temperature DPP conversion (◦C) (wt%) 420 96.52 425 98.50 430 99.52 435 99.60 445 99.57

345 E.1.2 Diluted Diphenylmethane The first and second order kinetic fits for DPM hydroconversion in the decalin- diluted reactions are presented in Figure E.1 with the resulting kinetic coefficients in Table E.2.

100

2 order

st DPMconversion (wt%)

25 1 order

0 2 4 6 8

Reaction time (h)

Figure E.1: Conversion results obtained for diphenylmethane hydroconversion experiments performed in the stirred batch reactor at 0 - 600 ppm ◦ Mo, 445 C, 13.8 MPa H2, 0 - 8 h, 700 RPM at 3 wt% in decalin together with ﬁrst and second order kinetic models. ◦ - 0 ppm Mo. - 600 ppm Mo.

Table E.2: Coefﬁcients for the kinetic models of diphenylmethane hydroconversion for data obtained in the stirred batch reactor at 0 - 600 ppm Mo, ◦ 445 C, 13.8 MPa H2, 0 - 8 h, 700 RPM at 3 wt% in decalin, depicted in Figure E.1.

Catalyst loading 1st order 1 2nd order 2 (ppm Mo) (h−1) (wt%−1.h−1) 0 0.041 ± 0.001 0.049 ± 0.002 600 3 0.103 ± 0.009 0.155 ± 0.021 1 - rDPM 0ppm = k0ppm.CDPM, rDPM 600ppm = k600ppm.CMo.CDPM. 2 ′ 2 ′ 2 - rDPM 0ppm = k0ppm.CDPM, rDPM 600ppm = k600ppm.CMo.CDPM. 3 ′ - Reporting k600ppm.CMo and k600ppm.CMo for comparison in context of the experiments.

346 The data for Figures E.1 through 4.5 of Section 4.1.3 is presented in Tables E.3 through E.6 below.

E.1.3 Undiluted Diphenylmethane The first and second order kinetic fits for undiluted DPM hydroconversion are presented in Figure E.2 with the resulting kinetic coefficients in Table E.7.

1 order

2 order

20 DPMconversion (wt%)

10 st

1 order

0 2 4 6

Reaction time (h)

Figure E.2: Conversion results obtained for undiluted diphenylmethane hydroconversion experiments performed in the stirred batch reactor at 0 - ◦ 600 ppm Mo, 445 C, 13.8 MPa H2, 0 - 6 h, 700 RPM together with ﬁrst and second order kinetic models (coefﬁcients provided in Table E.7). Error bars indicate standard deviation. ◦ - 0 ppm Mo. - 600 ppm Mo.

The data for Figures 4.6 through 4.16 of Section 4.1.3 is presented in Tables E.8 through E.13 below. Particle size and stack height distribution data for Figures 4.21 and 4.22 is presented in Table E.14.

347 Table E.3: Results for diphenylmethane hydroconversion in the stirred batch reactor at 0 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 8 h, 700 RPM at 3 wt% in decalin.

Sample code Reaction time DPM conversion Liquid composition (h) (wt%) (wt%) Benzene Toluene DPM All other DPM-B-5 1 32.62 0.12 0.33 2.02 0.53 DPM-BT-18 4 47.44 0.19 0.30 1.58 0.94 DPM-BT-16 6 52.79 0.30 0.45 1.42 0.83 DPM-BT-19 8 58.21 0.61 0.91 1.25 0.23 348

Table E.4: Results for diphenylmethane hydroconversion in the stirred batch reactor at 0 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 8 h, 700 RPM at 3 wt% in decalin continued.

Sample code Mass yield Molar yield B:T molar ratio (g/gDPM reacted) (mol/molDPM reacted) (mol/mol) Benzene Toluene Benzene Toluene DPM-B-5 0.121 0.340 0.262 0.622 0.42 DPM-BT-18 0.134 0.209 0.289 0.381 0.76 DPM-BT-16 0.189 0.285 0.407 0.520 0.78 DPM-BT-19 0.348 0.520 0.749 0.950 0.79 Table E.5: Results for diphenylmethane hydroconversion in the stirred batch reactor at 600 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 8 h, 700 RPM at 3 wt% in decalin.

Sample code Reaction time DPM conversion Liquid composition (h) (wt%) (wt%) Benzene Toluene DPM All other DPM-BC600-9 0 30.64 0.00 0.00 2.08 0.92 DPM-BC600-13 0.5 30.77 0.01 0.01 2.08 0.90 DPM-BC600-10 1 35.69 0.02 0.03 1.93 1.02 DPM-BC600-11 1.5 43.68 0.04 0.06 1.69 1.21 DPM-BC600-12 2 42.87 0.07 0.10 1.71 1.11 DPM-BC600-15 4 72.78 0.33 0.51 0.82 1.34 DPM-BC600-17 6 83.42 0.53 0.78 0.50 1.19 DPM-BC600-20 8 80.19 0.70 0.98 0.59 0.73 349

Table E.6: Results for diphenylmethane hydroconversion in the stirred batch reactor at 600 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 8 h, 700 RPM at 3 wt% in decalin continued.

Sample code Mass yield Molar yield B:T molar ratio (g/gDPM reacted) (mol/molDPM reacted) (mol/mol) Benzene Toluene Benzene Toluene DPM-BC600-9 0.001 0.002 0.003 0.004 0.66 DPM-BC600-13 0.007 0.013 0.016 0.023 0.69 DPM-BC600-10 0.019 0.031 0.041 0.056 0.73 DPM-BC600-11 0.031 0.047 0.066 0.086 0.77 DPM-BC600-12 0.055 0.080 0.119 0.146 0.81 DPM-BC600-15 0.151 0.236 0.325 0.431 0.76 DPM-BC600-17 0.213 0.311 0.460 0.568 0.81 DPM-BC600-20 0.291 0.406 0.627 0.740 0.85 Table E.7: Coefﬁcients for the kinetic models of undiluted diphenylmethane hydroconversion for data obtained in the stirred batch reactor at 0 - 600 ◦ ppm Mo, 445 C, 13.8 MPa H2, 0 - 6 h, 700 RPM, depicted in Figure E.2.

Catalyst loading 1st order 1 2nd order 2 (ppm Mo) (h−1) (wt%−1.h−1) 0 0.088 ± 0.004 - 600 3 0.087 ± 0.008 0.098 ± 0.011 1 - rDPM 0ppm = k0ppm.CDPM, rDPM 600ppm = k600ppm.CMo.CDPM. 2 ′ 2 - Not modelled for 0 ppm Mo, rDPM 600ppm = k600ppm.CMo.CDPM. 3 ′ - Reporting k600ppm.CMo and k600ppm.CMo for comparison in context of the experiments.

350 Table E.8: Results for diphenylmethane hydroconversion in the stirred batch reactor at 0 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 6 h, 700 RPM.

Sample code Reaction time DPM conversion Pressure change Product composition (DPM-free) (h) (wt%) (Pa) ×105 (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-BT-0-48 0 1.67 0.00 14.17 30.30 1.92 29.27 24.34 DPM-BT-24 1 6.38 -0.55 31.76 36.25 0.78 1.33 29.87 DPM-BT-27 1 10.05 -3.72 38.34 37.84 1.71 0.39 21.71 DPM-BT-29 1 18.16 -2.28 40.56 28.09 0.37 0.41 30.57 DPM-BT-1-44 1 8.85 -8.48 41.90 47.27 1.24 0.00 9.58 DPM-BT-25 6 20.76 -9.03 44.29 44.27 0.00 1.18 10.26 DPM-BT-28 6 40.11 -13.58 38.30 35.30 0.30 1.24 24.86 DPM-BT-31 6 45.66 -13.79 31.97 28.06 0.28 1.37 38.33 DPM-BT-6-45 6 24.19 -16.13 42.19 48.79 0.59 1.16 7.26 351 Table E.9: Results for diphenylmethane hydroconversion in the stirred batch reactor at 0 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 6 h, 700 RPM continued.

Sample code Mass yield Molar yield B:T molar ratio (g/gDPM reacted) (mol/molDPM reacted) (mol/mol) Other cracking Other isom./cond. Benzene Toluene CHMB DPM-BT-0-48 0.305 0.553 0.019 0.293 0.243 0.55 DPM-BT-24 0.684 0.662 0.008 0.013 0.299 1.03 DPM-BT-27 0.826 0.691 0.017 0.004 0.217 1.20 DPM-BT-29 0.874 0.513 0.004 0.004 0.306 1.70 DPM-BT-1-44 0.903 0.863 0.012 0.000 0.096 1.05 DPM-BT-25 0.954 0.808 0.000 0.012 0.103 1.18 DPM-BT-28 0.825 0.645 0.003 0.012 0.249 1.28 DPM-BT-31 0.688 0.512 0.003 0.014 0.383 1.34 DPM-BT-6-45 0.909 0.891 0.006 0.012 0.073 1.02 Table E.10: Results for diphenylmethane hydroconversion in the stirred batch reactor at 600 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 6 h, 700 RPM.

Sample code Reaction time DPM conversion Pressure change Product composition (DPM-free) (h) (wt%) (Pa) ×105 (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-BC600-0-46 0 3.36 0.00 41.66 47.70 2.57 1.85 6.22 DPM-BC600-23 1 6.35 -2.07 29.10 36.19 8.78 3.35 22.58 DPM-BC600-26 1 21.69 -8.07 41.77 41.36 0.62 0.33 15.92 DPM-BC600-30 1 13.43 -4.96 36.36 36.46 0.67 0.00 26.51

352 DPM-BC600-1-34 1 3.51 -1.10 26.22 35.93 6.44 6.21 25.20 DPM-BC600-1-35 1 7.77 -2.76 31.61 40.76 4.92 1.44 21.27 DPM-BC600-1-39 1 13.08 -3.93 38.53 41.58 1.78 1.23 16.88 DPM-BC600-1-40 1 17.44 -6.14 37.00 40.69 1.82 1.10 19.40 DPM-BC600-1-41 1 12.70 -5.10 41.66 48.33 1.73 2.81 5.47 DPM-BC600-1-42 1 20.32 -6.21 45.05 45.14 0.60 0.00 9.21 DPM-BC600-1-43 1 17.70 -5.86 40.07 43.17 0.95 0.73 15.09 DPM-BC600-22 6 22.01 -12.48 41.25 51.25 3.05 0.59 3.86 DPM-BC600-32 6 47.21 -14.55 32.44 30.91 0.00 1.39 35.26 DPM-BC600-33 6 42.70 -15.72 35.69 36.56 0.35 1.75 25.65 DPM-BC600-6-36 6 25.90 -13.72 37.61 45.15 1.22 2.20 13.81 DPM-BC600-6-37 6 24.97 -11.65 41.93 47.05 0.88 1.26 8.88 DPM-BC600-6-47 6 38.19 -18.62 38.90 45.99 0.61 1.75 12.76 Table E.11: Results for diphenylmethane hydroconversion in the stirred batch reactor at 600 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 6 h, 700 RPM continued.

Sample code Mass yield Molar yield B:T molar ratio (g/gDPM reacted) (mol/molDPM reacted) (mol/mol) Other cracking Other isom./cond. Benzene Toluene CHMB DPM-BC600-0-46 0.018 0.062 0.897 0.871 0.026 1.03 DPM-BC600-23 0.034 0.226 0.627 0.661 0.088 0.95 DPM-BC600-26 0.003 0.159 0.900 0.755 0.006 1.19 DPM-BC600-30 0.000 0.265 0.783 0.666 0.007 1.18 DPM-BC600-1-34 0.062 0.252 0.565 0.656 0.064 0.86 353 DPM-BC600-1-35 0.014 0.213 0.681 0.744 0.049 0.91 DPM-BC600-1-39 0.012 0.169 0.830 0.759 0.018 1.09 DPM-BC600-1-40 0.011 0.194 0.797 0.743 0.018 1.07 DPM-BC600-1-41 0.028 0.055 0.897 0.882 0.017 1.02 DPM-BC600-1-42 0.000 0.092 0.970 0.824 0.006 1.18 DPM-BC600-1-43 0.007 0.151 0.863 0.788 0.009 1.09 DPM-BC600-22 0.006 0.039 0.888 0.936 0.030 0.95 DPM-BC600-32 0.014 0.353 0.699 0.564 0.000 1.24 DPM-BC600-33 0.017 0.256 0.769 0.668 0.004 1.15 DPM-BC600-6-36 0.022 0.138 0.810 0.824 0.012 0.98 DPM-BC600-6-37 0.013 0.089 0.903 0.859 0.009 1.05 DPM-BC600-6-47 0.017 0.128 0.838 0.840 0.006 1.00 Table E.12: Results for diphenylmethane hydroconversion in the stirred batch reactor at 1800 ppm Mo, 415 - 445◦C, 13.8 MPa H2, 1 h, 700 RPM.

Sample code Reaction conditions DPM conversion Pressure change Product composition (DPM-free) Time Temperature (wt%) (Pa) ×105 (wt%) (h) (◦C) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-BC1800-1-50 1 445 29.77 -15.93 40.49 51.15 3.61 1.72 3.02 DPM-BC1800-1-51 1 430 18.12 -12.07 37.61 49.50 8.21 1.79 2.88 DPM-BC1800-1-52 1 415 10.82 -7.86 32.21 42.16 15.97 5.41 4.25 354

Table E.13: Results for diphenylmethane hydroconversion in the stirred batch reactor at 1800 ppm Mo, 415 - 445◦C, 13.8 MPa H2, 1 h, 700 RPM continued.

Sample code Mass yield Molar yield B:T molar ratio (g/gDPM reacted) (mol/molDPM reacted) (mol/mol) Other cracking Other isom./cond. Benzene Toluene CHMB DPM-BC1800-1-50 0.872 0.934 0.036 0.017 0.030 0.93 DPM-BC1800-1-51 0.810 0.904 0.082 0.018 0.029 0.90 DPM-BC1800-1-52 0.694 0.770 0.160 0.054 0.043 0.90 Table E.14: Sheet size and stack height distributions of MoS2 crystallites using transmission electron microscopy data for solid material from diphenylmethane hydroconversion in the stirred batch reactor at 600 ppm ◦ Mo, 445 C, 13.8 MPa H2, 1 h, 700 RPM.

Sheets per stack Number of stacks Sheet width range Number of sheets (-) (-) (nm) (-) 1 31 0 - 1 5 2 57 1 - 2 48 3 37 2 - 3 80 4 21 3 - 4 91 5 7 4 - 5 60 6 3 5 - 6 58 7 4 6 - 7 49 8 2 7 - 8 26 9 1 8 - 9 14 9 - 10 13 10 - 11 3 11 - 12 6 12 - 13 5 13 - 14 2

355 E.2 Stainless Steel Batch Micro-reactors

E.2.1 Inclined Stainless Steel Micro-Reactor The wall activation data shown in Figures 4.23 and 4.24 of Section 4.2.1 is presented in Tables E.15 and E.16. Data obtained with stabilised wall activity, Fig- ures 4.25 through 4.31, is presented in Tables E.17 through E.22.

E.2.2 Vertical Stainless Steel Micro-Reactor Data obtained in the vertical stainless steel micro-reactor, shown in Figures 4.32 through 4.37 in Section 4.2.2, is presented in Tables E.23 through E.26 below.

E.3 Glass Insert Batch Micro-reactor

E.3.1 Comparison with Stainless Steel Micro-Reactor Data shown in Figures 4.40 through 4.45 in Section 4.2.3, is presented in Ta- bles E.27 through E.30 below.

356 Table E.15: Results for diphenylmethane hydroconversion performed in the inclined micro-reactor at 1800 ppm Mo, ◦ 445 C, 13.8 MPa H2, 1 h, 0 RPM to study wall activation.

Sample code Reaction time DPM conversion Product composition (DPM-free) (h) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-M1800-1-68 1 6.06 26.18 42.80 11.89 8.08 11.05 DPM-M1800-74 1 6.53 30.14 42.01 13.38 5.29 9.18 DPM-M1800-1-72 1 6.78 29.83 41.22 12.52 5.46 10.97 DPM-M1800-1-70 1 6.90 30.89 42.10 12.11 5.51 9.39 DPM-M1800-1-64 1 6.98 15.64 24.02 2.71 10.81 46.83 DPM-M1800-1-75 1 7.33 31.04 42.03 14.63 4.03 8.26 DPM-M1800-1-69 1 7.36 31.73 41.75 16.11 4.34 6.07 357 DPM-M1800-1-76 1 7.50 31.71 42.15 17.07 4.09 4.99 DPM-M1800-1-71 1 7.56 30.83 43.63 16.40 3.89 5.26 DPM-M1800-1-73 1 8.03 29.91 43.80 17.25 3.94 5.11 DPM-M1800-1-66 1 8.60 31.54 39.77 19.14 4.45 5.11 DPM-M1800-1-67 1 10.31 33.12 42.62 15.45 3.68 5.13 DPM-M1800-1-62 1 2.07 0.81 23.23 2.27 14.18 59.53 DPM-M1800-2-82 2 10.99 36.67 45.12 10.62 2.75 4.84 DPM-M1800-2-83 2 14.24 36.00 45.38 13.03 1.81 3.78 DPM-M1800-2-79 2 11.58 34.81 45.58 11.63 2.67 5.31 DPM-M1800-3-85 3 15.81 37.05 45.76 11.94 2.27 2.99 DPM-M1800-3-78 3 14.56 34.94 45.47 13.06 2.16 4.36 DPM-M1800-3-77 3 19.57 33.76 44.60 14.42 2.28 4.94 DPM-M1800-4-84 4 19.65 37.50 46.64 10.72 2.09 3.05 DPM-M1800-4-80 4 19.30 37.57 46.32 9.91 2.35 3.85 Table E.16: Results for diphenylmethane hydroconversion performed in the inclined micro-reactor at 1800 ppm Mo, ◦ 445 C, 13.8 MPa H2, 1 h, 0 RPM to study wall activation, continued.

Sample code Molar yield Mass yield B:T molar ratio (mol/molDPM reacted) (g/gDPM reacted) (mol/mol) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-M1800-1-68 0.564 0.781 0.013 0.081 0.111 0.72 DPM-M1800-74 0.649 0.767 0.062 0.053 0.092 0.85 DPM-M1800-1-72 0.643 0.753 0.050 0.055 0.110 0.85 DPM-M1800-1-70 0.665 0.769 0.052 0.055 0.094 0.87 DPM-M1800-1-64 0.337 0.438 0.054 0.108 0.468 0.77 DPM-M1800-1-75 0.669 0.767 0.059 0.040 0.083 0.87 DPM-M1800-1-69 0.683 0.762 0.050 0.043 0.061 0.90 358 DPM-M1800-1-76 0.683 0.770 0.066 0.041 0.050 0.89 DPM-M1800-1-71 0.664 0.797 0.046 0.039 0.053 0.83 DPM-M1800-1-73 0.644 0.800 0.053 0.039 0.051 0.81 DPM-M1800-1-66 0.679 0.726 0.056 0.045 0.051 0.94 DPM-M1800-1-67 0.713 0.778 0.050 0.037 0.051 0.92 DPM-M1800-1-62 0.017 0.424 0.062 0.142 0.595 0.04 DPM-M1800-2-82 0.790 0.824 0.070 0.028 0.048 0.96 DPM-M1800-2-83 0.775 0.828 0.020 0.018 0.038 0.94 DPM-M1800-2-79 0.750 0.832 0.064 0.027 0.053 0.90 DPM-M1800-3-85 0.798 0.836 0.037 0.023 0.030 0.96 DPM-M1800-3-78 0.753 0.830 0.063 0.022 0.044 0.91 DPM-M1800-3-77 0.727 0.814 0.058 0.023 0.049 0.89 DPM-M1800-4-84 0.808 0.851 0.033 0.021 0.031 0.95 DPM-M1800-4-80 0.809 0.846 0.063 0.023 0.039 0.96 Table E.17: Results for diphenylmethane hydroconversion performed in the inclined micro-reactor at 0 ppm Mo, ◦ 445 C, 13.8 MPa H2,0-4h,0RPM.

Sample code Reaction time DPM conversion Product composition (DPM-free) (h) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MT-0-114 0 0.17 0.06 33.83 1.31 7.83 56.97 DPM-MT-0-105 0 0.56 0.00 56.03 1.49 7.79 34.70 DPM-MT-0-104 0 0.72 0.00 49.25 1.13 7.07 42.56 DPM-MT-1-89 1 11.71 30.97 43.50 5.59 3.60 16.35 DPM-MT-1-90 1 13.05 31.59 43.35 4.85 3.65 16.57 DPM-MT-1-91 1 13.69 31.37 41.99 4.64 7.68 14.32 DPM-MT-1-92 1 14.09 36.24 46.14 5.19 1.33 11.11 DPM-MT-1-98 1 15.45 39.47 46.65 5.60 0.79 7.50 DPM-MT-1-97 1 14.00 33.99 44.20 5.23 1.86 14.72 DPM-MT-1-95 1 12.57 32.38 43.54 5.51 2.64 15.94 DPM-MT-1-96 1 13.25 34.21 44.88 5.45 2.08 13.38 359 DPM-MT-1-87 1 9.18 35.59 49.44 7.01 1.70 6.26 DPM-MT-1-88 1 11.38 34.48 47.15 7.26 1.72 9.41 DPM-MT-1-99 1 15.38 38.62 44.95 5.20 0.71 10.52 DPM-MT-1-100 1 14.86 36.79 42.60 5.05 1.82 13.74 DPM-MT-1-101 1 15.11 35.40 42.52 4.81 1.64 15.62 DPM-MT-1-102 1 15.49 37.13 42.52 4.66 2.00 13.69 DPM-MT-1-86 1 7.24 33.29 49.91 7.01 1.69 8.10 DPM-MT-2-113 2 21.81 34.51 42.18 5.14 1.95 16.22 DPM-MT-2-94 2 20.36 35.53 43.95 5.64 2.03 12.86 DPM-MT-2-106 2 26.11 40.95 44.56 5.25 1.86 7.39 DPM-MT-2-112 2 20.25 33.99 41.12 5.81 1.69 17.40 DPM-MT-3-93 3 32.20 38.74 46.84 5.87 1.88 6.67 DPM-MT-3-109 3 26.10 35.74 44.20 6.09 1.73 12.24 DPM-MT-3-110 3 27.31 34.26 45.28 6.38 1.74 12.34 DPM-MT-3-103 3 30.16 40.44 45.26 4.97 1.63 7.71 DPM-MT-4-111 4 32.13 36.03 44.15 5.54 2.61 11.67 DPM-MT-4-108 4 33.60 36.68 43.71 6.19 2.18 11.23 DPM-MT-4-107 4 39.08 36.64 44.81 7.28 2.81 8.45 Table E.18: Results for diphenylmethane hydroconversion performed in the inclined micro-reactor at 0 ppm Mo, ◦ 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM continued.

Sample code Molar yield Mass yield B:T molar ratio (mol/molDPM reacted) (g/gDPM reacted) (mol/mol) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MT-0-114 0.001 0.618 0.013 0.078 0.570 0.00 DPM-MT-0-105 0.000 1.237 0.017 0.094 0.420 0.00 DPM-MT-0-104 0.000 0.988 0.012 0.078 0.467 0.00 DPM-MT-1-89 0.667 0.794 0.054 0.036 0.163 0.84 DPM-MT-1-90 0.680 0.792 0.047 0.036 0.166 0.86 DPM-MT-1-91 0.676 0.767 0.045 0.077 0.143 0.88 DPM-MT-1-92 0.781 0.842 0.050 0.013 0.111 0.93 DPM-MT-1-98 0.850 0.852 0.054 0.008 0.075 1.00 DPM-MT-1-97 0.732 0.807 0.051 0.019 0.147 0.91 DPM-MT-1-95 0.697 0.795 0.053 0.026 0.159 0.88 DPM-MT-1-96 0.737 0.819 0.053 0.021 0.134 0.90 360 DPM-MT-1-87 0.766 0.903 0.068 0.017 0.063 0.85 DPM-MT-1-88 0.743 0.861 0.070 0.017 0.094 0.86 DPM-MT-1-99 0.832 0.821 0.050 0.007 0.105 1.01 DPM-MT-1-100 0.792 0.778 0.049 0.018 0.137 1.02 DPM-MT-1-101 0.763 0.776 0.046 0.016 0.156 0.98 DPM-MT-1-102 0.800 0.776 0.045 0.020 0.137 1.03 DPM-MT-1-86 0.717 0.911 0.068 0.017 0.081 0.79 DPM-MT-2-113 0.743 0.770 0.050 0.020 0.162 0.97 DPM-MT-2-94 0.765 0.802 0.054 0.020 0.129 0.95 DPM-MT-2-106 0.882 0.814 0.051 0.019 0.074 1.08 DPM-MT-2-112 0.732 0.751 0.056 0.017 0.174 0.98 DPM-MT-3-93 0.834 0.855 0.057 0.019 0.067 0.98 DPM-MT-3-109 0.770 0.807 0.059 0.017 0.122 0.95 DPM-MT-3-110 0.738 0.827 0.062 0.017 0.123 0.89 DPM-MT-3-103 0.871 0.826 0.048 0.016 0.077 1.05 DPM-MT-4-111 0.776 0.806 0.054 0.026 0.117 0.96 DPM-MT-4-108 0.790 0.798 0.060 0.022 0.112 0.99 DPM-MT-4-107 0.789 0.818 0.070 0.028 0.084 0.97 Table E.19: Results for diphenylmethane hydroconversion performed in the inclined micro-reactor at 600 ppm Mo, ◦ 445 C, 13.8 MPa H2,0-4h,0RPM.

Sample code Reaction time DPM conversion Product composition (DPM-free) (h) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-M600-0-132 0 0.62 5.26 42.32 4.07 13.56 34.79 DPM-M600-0-131 0 0.69 8.52 38.89 3.81 19.22 29.57 DPM-M600-0-130 0 0.00 17.95 17.95 3.43 24.83 35.84 DPM-M600-0-129 0 0.00 21.04 21.04 2.10 17.76 38.06 DPM-M600-1-119 1 18.26 36.38 43.96 5.51 2.00 12.14

361 DPM-M600-1-115 1 22.22 35.58 42.77 5.13 1.93 14.60 DPM-M600-1-118 1 18.43 36.57 43.87 5.22 1.79 12.56 DPM-M600-1-116 1 21.33 34.75 42.03 5.42 1.40 16.41 DPM-M600-1-117 1 17.02 35.82 46.11 4.74 1.78 11.56 DPM-M600-2-123 2 23.02 37.80 45.04 6.01 1.53 9.62 DPM-M600-2-121 2 23.38 37.08 45.27 6.14 1.95 9.56 DPM-M600-2-127 2 19.79 36.51 45.11 6.83 1.48 10.07 DPM-M600-3-126 3 25.41 37.30 45.91 6.55 1.08 9.16 DPM-M600-3-122 3 29.23 37.73 46.32 6.87 1.63 7.45 DPM-M600-3-120 3 35.43 37.32 46.29 6.39 2.10 7.90 DPM-M600-4-124 4 35.18 38.30 46.80 6.52 1.66 6.72 DPM-M600-4-128 4 34.86 38.12 45.81 7.22 1.60 7.26 DPM-M600-4-125 4 31.73 37.72 47.07 6.68 2.04 6.49 Table E.20: Results for diphenylmethane hydroconversion performed in the inclined micro-reactor at 600 ppm Mo, ◦ 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM continued.

Sample code Molar yield Mass yield B:T molar ratio (mol/molDPM reacted) (g/gDPM reacted) (mol/mol) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-M600-0-132 0.113 0.773 0.039 0.136 0.348 0.15 DPM-M600-0-131 0.183 0.710 0.037 0.192 0.296 0.26 DPM-M600-0-130 0.387 0.328 0.033 0.248 0.358 1.18 DPM-M600-0-129 0.453 0.384 0.020 0.178 0.381 1.18 DPM-M600-1-119 0.784 0.803 0.053 0.020 0.121 0.98

362 DPM-M600-1-115 0.766 0.781 0.050 0.019 0.146 0.98 DPM-M600-1-118 0.788 0.801 0.050 0.018 0.126 0.98 DPM-M600-1-116 0.749 0.767 0.052 0.014 0.164 0.98 DPM-M600-1-117 0.771 0.842 0.046 0.018 0.116 0.92 DPM-M600-2-123 0.814 0.822 0.058 0.015 0.096 0.99 DPM-M600-2-121 0.799 0.827 0.059 0.019 0.096 0.97 DPM-M600-2-127 0.786 0.824 0.066 0.015 0.101 0.96 DPM-M600-3-126 0.803 0.838 0.063 0.011 0.092 0.96 DPM-M600-3-122 0.813 0.846 0.066 0.016 0.074 0.96 DPM-M600-3-120 0.804 0.845 0.062 0.021 0.079 0.95 DPM-M600-4-124 0.825 0.855 0.063 0.017 0.067 0.97 DPM-M600-4-128 0.821 0.836 0.070 0.016 0.073 0.98 DPM-M600-4-125 0.812 0.860 0.064 0.020 0.065 0.95 Table E.21: Results for diphenylmethane hydroconversion performed in the inclined micro-reactor at 1800 ppm Mo, ◦ 445 C, 13.8 MPa H2,0-4h,0RPM.

Sample code Reaction time DPM conversion Product composition (DPM-free) (h) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-M1800-0-151 0 0.18 31.69 68.31 0.00 0.00 0.00 DPM-M1800-0-144 0 0.22 15.04 27.38 5.26 28.96 23.37 DPM-M1800-1-136 1 18.16 34.86 40.53 16.65 1.83 6.13 DPM-M1800-1-137 1 15.57 35.70 40.25 16.88 2.20 4.97 DPM-M1800-1-138 1 16.91 35.46 40.72 16.19 1.72 5.91

363 DPM-M1800-1-139 1 16.07 34.18 39.94 17.17 2.33 6.38 DPM-M1800-1-134 1 22.43 36.92 41.90 13.85 1.46 5.88 DPM-M1800-1-135 1 19.77 36.27 40.96 14.46 1.91 6.40 DPM-M1800-2-149 2 21.16 36.02 39.59 18.05 2.34 4.01 DPM-M1800-2-146 2 17.93 35.40 40.59 15.91 2.28 5.82 DPM-M1800-2-148 2 21.82 34.95 39.66 18.22 1.78 5.40 DPM-M1800-2-142 2 19.95 36.82 40.91 15.46 2.05 4.76 DPM-M1800-3-140 3 35.24 36.09 41.13 15.23 3.37 4.18 DPM-M1800-3-147 3 29.12 35.57 40.25 16.79 3.21 4.18 DPM-M1800-3-150 3 29.04 35.69 40.01 16.74 3.12 4.45 DPM-M1800-4-143 4 33.81 34.66 43.05 15.74 3.56 2.99 DPM-M1800-4-141 4 35.89 35.59 42.30 15.01 3.55 3.56 DPM-M1800-4-145 4 41.65 35.61 41.31 15.48 3.73 3.87 Table E.22: Results for diphenylmethane hydroconversion performed in the inclined micro-reactor at 1800 ppm Mo, ◦ 445 C, 13.8 MPa H2, 0 - 4 h, 0 RPM, continued.

Sample code Molar yield Mass yield B:T molar ratio (mol/molDPM reacted) (g/gDPM reacted) (mol/mol) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-M1800-0-151 0.683 1.247 0.000 0.000 0.000 0.55 DPM-M1800-0-144 0.324 0.500 0.051 0.290 0.234 0.65 DPM-M1800-1-136 0.751 0.740 0.161 0.018 0.061 1.02 DPM-M1800-1-137 0.769 0.735 0.163 0.022 0.050 1.05 DPM-M1800-1-138 0.764 0.743 0.156 0.017 0.059 1.03

364 DPM-M1800-1-139 0.736 0.729 0.166 0.023 0.064 1.01 DPM-M1800-1-134 0.795 0.765 0.134 0.015 0.059 1.04 DPM-M1800-1-135 0.781 0.748 0.140 0.019 0.064 1.05 DPM-M1800-2-149 0.776 0.723 0.174 0.023 0.040 1.07 DPM-M1800-2-146 0.762 0.741 0.154 0.023 0.058 1.03 DPM-M1800-2-148 0.753 0.724 0.176 0.018 0.054 1.04 DPM-M1800-2-142 0.793 0.747 0.149 0.021 0.048 1.06 DPM-M1800-3-140 0.777 0.751 0.147 0.034 0.042 1.04 DPM-M1800-3-147 0.766 0.735 0.162 0.032 0.042 1.04 DPM-M1800-3-150 0.769 0.730 0.162 0.031 0.044 1.05 DPM-M1800-4-143 0.747 0.786 0.152 0.036 0.030 0.95 DPM-M1800-4-141 0.766 0.772 0.145 0.035 0.036 0.99 DPM-M1800-4-145 0.767 0.754 0.149 0.037 0.039 1.02 Table E.23: Results for diphenylmethane hydroconversion performed in the vertical stainless steel micro-reactor at 0 ◦ ppm Mo, 445 C, 13.8 MPa H2,1-4h,0RPM.

Sample code Reaction time DPM conversion Product composition (DPM-free) (h) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MVT-1-172 1 8.56 36.93 45.35 12.42 0.00 5.29 DPM-MVT-2-175 2 15.00 39.95 48.33 7.54 0.00 4.18 DPM-MVT-2-174 2 15.30 39.21 48.02 8.60 0.00 4.17 DPM-MVT-2-176 2 13.93 39.90 48.58 7.20 0.00 4.33 DPM-MVT-3-182 3 21.80 41.62 50.82 4.28 0.00 3.29 DPM-MVT-3-177 3 20.16 40.60 50.07 5.67 0.00 3.66 DPM-MVT-3-183 3 24.67 42.35 51.18 3.25 0.00 3.22 DPM-MVT-4-180 4 27.20 41.39 51.47 3.98 0.00 3.17 DPM-MVT-4-184 4 29.82 41.82 51.70 2.93 0.87 2.69 DPM-MVT-4-178 4 26.40 40.97 50.95 4.92 0.00 3.16 365

Table E.24: Results for diphenylmethane hydroconversion performed in the vertical stainless steel micro-reactor at 0 ◦ ppm Mo, 445 C, 13.8 MPa H2, 1 - 4 h, 0 RPM continued.

Sample code Molar yield Mass yield B:T molar ratio (mol/molDPM reacted) (g/gDPM reacted) (mol/mol) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MVT-1-172 0.795 0.828 0.120 0.000 0.053 0.96 DPM-MVT-2-175 0.860 0.882 0.073 0.000 0.042 0.98 DPM-MVT-2-174 0.844 0.877 0.083 0.000 0.042 0.96 DPM-MVT-2-176 0.859 0.887 0.069 0.000 0.043 0.97 DPM-MVT-3-182 0.896 0.928 0.041 0.000 0.033 0.97 DPM-MVT-3-177 0.874 0.914 0.055 0.000 0.037 0.96 DPM-MVT-3-183 0.912 0.934 0.031 0.000 0.032 0.98 DPM-MVT-4-180 0.891 0.940 0.038 0.000 0.032 0.95 DPM-MVT-4-184 0.901 0.944 0.028 0.009 0.027 0.95 DPM-MVT-4-178 0.882 0.930 0.047 0.000 0.032 0.95 Table E.25: Results for diphenylmethane hydroconversion performed in the vertical stainless steel micro-reactor at ◦ 1800 ppm Mo, 445 C, 13.8 MPa H2,1-4h,0RPM.

Sample code Reaction time DPM conversion Product composition (DPM-free) (h) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MV1800-0-168 0 0.00 15.02 15.02 0.00 0.00 69.95 DPM-MV1800-0-167 0 0.00 27.51 27.51 0.00 0.00 44.97 DPM-MV1800-0-169 0 0.00 22.74 22.74 0.00 0.00 54.52 DPM-MV1800-1-155 1 7.29 32.18 42.29 16.04 0.00 9.49 DPM-MV1800-1-154 1 5.25 31.12 40.73 18.42 0.00 9.74

366 DPM-MV1800-1-163 1 6.48 36.32 45.71 12.63 0.00 5.35 DPM-MV1800-1-186 1 11.71 42.61 49.47 4.87 0.00 3.06 DPM-MV1800-1-170 1 9.62 36.00 44.65 14.41 0.00 4.95 DPM-MV1800-1-157 1 8.02 33.92 43.81 17.48 0.00 4.78 DPM-MV1800-1-156 1 6.47 30.87 41.38 16.83 0.00 10.92 DPM-MV1800-2-158 2 15.45 35.48 44.25 16.46 0.00 3.81 DPM-MV1800-2-160 2 13.40 37.00 45.33 13.79 0.00 3.87 DPM-MV1800-3-161 3 20.51 35.82 44.83 13.69 0.00 5.67 DPM-MV1800-3-166 3 19.08 37.29 46.35 11.04 0.00 5.33 DPM-MV1800-3-165 3 21.66 38.22 48.18 10.42 0.00 3.18 DPM-MV1800-4-164 4 24.99 38.11 47.80 11.22 0.00 2.87 DPM-MV1800-4-159 4 25.35 37.80 47.40 12.14 0.00 2.67 DPM-MV1800-4-162 4 28.19 39.12 48.46 8.99 0.93 2.49 Table E.26: Results for diphenylmethane hydroconversion performed in the vertical stainless steel micro-reactor at ◦ 1800 ppm Mo, 445 C, 13.8 MPa H2, 1 - 4 h, 0 RPM continued.

Sample code Molar yield Mass yield B:T molar ratio (mol/molDPM reacted) (g/gDPM reacted) (mol/mol) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MV1800-0-168 0.324 0.274 0.000 0.000 0.700 1.18 DPM-MV1800-0-167 0.593 0.502 0.000 0.000 0.450 1.18 DPM-MV1800-0-169 0.490 0.415 0.000 0.000 0.545 1.18 DPM-MV1800-1-155 0.693 0.772 0.155 0.000 0.095 0.90 DPM-MV1800-1-154 0.670 0.744 0.178 0.000 0.097 0.90

367 DPM-MV1800-1-163 0.782 0.835 0.122 0.000 0.053 0.94 DPM-MV1800-1-186 0.918 0.903 0.047 0.000 0.031 1.02 DPM-MV1800-1-170 0.775 0.815 0.139 0.000 0.049 0.95 DPM-MV1800-1-157 0.731 0.800 0.169 0.000 0.048 0.91 DPM-MV1800-1-156 0.665 0.756 0.162 0.000 0.109 0.88 DPM-MV1800-2-158 0.764 0.808 0.159 0.000 0.038 0.95 DPM-MV1800-2-160 0.797 0.828 0.133 0.000 0.039 0.96 DPM-MV1800-3-161 0.772 0.818 0.132 0.000 0.057 0.94 DPM-MV1800-3-166 0.803 0.846 0.107 0.000 0.053 0.95 DPM-MV1800-3-165 0.823 0.880 0.101 0.000 0.032 0.94 DPM-MV1800-4-164 0.821 0.873 0.108 0.000 0.029 0.94 DPM-MV1800-4-159 0.814 0.865 0.117 0.000 0.027 0.94 DPM-MV1800-4-162 0.843 0.885 0.087 0.009 0.025 0.95 E.3.2 Visual Mixing Studies The data for the relationship between “vortex” (actually a circulating wave) height with mixing speed shown in Figure 4.47 in Section 4.2.3, is presented in Table E.31.

E.3.3 Comparison of Liquid Loading Volumes Full data relating to that shown in Table 4.15 in Section 4.2.3, is presented in Ta- bles E.32 and E.33.

E.3.4 Thermocouple Wall Activity The data relating to thermocouple wall activation, shown in Figures 4.51 through 4.55 in Section 4.2.3, is presented in Tables E.34 and E.35.

E.3.5 Effect of Mixing Speed Mixing speed data, shown in Figures 4.56 through 4.60 in Section 4.2.3, is presented in Tables ?? through E.39 below.

E.3.6 Optimum Mixing Speed Evaluation The first order kinetic fits for DPM hydroconversion are presented in Figure E.3 with the resulting kinetic coefficients in Table E.40. Reaction data at the optimal mixing speed of 2000 RPM, shown in Figures 4.69 through 4.74 in Section 4.2.3, is presented in Tables E.41 through E.44.

E.3.7 Spent Residue Hydroconversion Catalyst Evaluation Deactivated coke-catalyst agglomerate data, shown in Table 4.19 in Section 4.2.3, is presented in full in Tables E.45 and E.46.

368 30

0 RPM

2000 RPM

10 DPM conversion (wt%)

0 1 2 3 4

Reaction time (h)

Figure E.3: Conversion results obtained for diphenylmethane hydroconversion experiments performed in the glass insert micro-reactor at 0 - 1800 ◦ ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 - 2000 RPM, 150 µL liquid load together with first and second order kinetic fits. Error bars indicate standard deviation. Curves shown are first order kinetic fits with coefficients in Table E.40. • - 0 ppm Mo, 0 RPM. N - 1800 ppm Mo, 0 RPM. ◦ - 0 ppm Mo, 2000 RPM. △ - 1800 ppm Mo, 2000 RPM.

369 Table E.27: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 0 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 4 h, 0 RPM, 400µL liquid loading.

Sample code Reaction time DPM conversion Product composition (DPM-free) (h) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MGT-0-203 0 0.00 24.60 24.60 0.00 0.00 50.80 DPM-MGT-0-205 0 0.00 25.87 25.87 0.00 0.00 48.26 DPM-MGT-0-197 0 0.00 27.02 27.02 0.00 0.00 45.97 DPM-MGT-1-193 1 3.27 44.12 41.74 0.00 0.00 14.14 370 DPM-MGT-1-195 1 3.58 42.09 43.48 0.00 0.00 14.42 DPM-MGT-1-194 1 3.06 42.03 42.60 0.00 0.00 15.36 DPM-MGT-2-206 2 8.14 42.77 49.88 0.00 0.00 7.35 DPM-MGT-2-198 2 6.31 43.19 48.64 0.00 0.00 8.18 DPM-MGT-2-196 2 7.10 41.81 47.69 0.00 0.00 10.51 DPM-MGT-3-200 3 8.92 42.14 50.67 0.00 0.00 7.19 DPM-MGT-3-201 3 10.30 41.35 49.65 0.00 0.00 9.00 DPM-MGT-3-199 3 10.32 41.14 49.32 0.00 0.00 9.54 DPM-MGT-4-202 4 14.78 41.55 50.64 0.00 0.00 7.81 DPM-MGT-4-204 4 13.55 42.32 51.49 0.00 0.00 6.19 DPM-MGT-4-207 4 12.75 42.54 50.37 0.00 0.00 7.09 Table E.28: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 0 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 4 h, 0 RPM, 400µL liquid loading, continued.

Sample code Molar yield Mass yield B:T molar ratio (mol/molDPM reacted) (g/gDPM reacted) (mol/mol) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MGT-0-203 0.530 0.449 0.000 0.000 0.508 1.18 DPM-MGT-0-205 0.557 0.472 0.000 0.000 0.483 1.18 DPM-MGT-0-197 0.582 0.493 0.000 0.000 0.460 1.18 DPM-MGT-1-193 0.950 0.762 0.000 0.000 0.141 1.25 371 DPM-MGT-1-195 0.907 0.794 0.000 0.000 0.144 1.14 DPM-MGT-1-194 0.905 0.778 0.000 0.000 0.154 1.16 DPM-MGT-2-206 0.921 0.911 0.000 0.000 0.074 1.01 DPM-MGT-2-198 0.930 0.888 0.000 0.000 0.082 1.05 DPM-MGT-2-196 0.900 0.871 0.000 0.000 0.105 1.03 DPM-MGT-3-200 0.908 0.925 0.000 0.000 0.072 0.98 DPM-MGT-3-201 0.891 0.907 0.000 0.000 0.090 0.98 DPM-MGT-3-199 0.886 0.901 0.000 0.000 0.095 0.98 DPM-MGT-4-202 0.895 0.925 0.000 0.000 0.078 0.97 DPM-MGT-4-204 0.911 0.940 0.000 0.000 0.062 0.97 DPM-MGT-4-207 0.916 0.920 0.000 0.000 0.071 1.00 Table E.29: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 1800 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 4 h, 0 RPM, 400µL liquid loading.

Sample code Reaction time DPM conversion Product composition (DPM-free) (h) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MG1800-0-210 0 0.00 32.89 32.89 0.00 0.00 34.21 DPM-MG1800-0-219 0 0.00 19.64 19.64 0.00 0.00 60.72 DPM-MG1800-0-218 0 0.00 29.46 29.46 0.00 0.00 41.08 DPM-MG1800-1-208 1 5.88 38.21 48.04 4.35 0.00 9.41 372 DPM-MG1800-1-222 1 3.78 36.06 47.92 4.43 0.00 11.59 DPM-MG1800-1-209 1 5.59 39.79 49.42 4.85 0.00 5.93 DPM-MG1800-2-215 2 7.94 39.69 49.87 4.50 0.00 5.94 DPM-MG1800-2-212 2 9.94 40.00 49.51 5.95 0.00 4.53 DPM-MG1800-2-216 2 7.99 39.79 49.86 5.32 0.00 5.03 DPM-MG1800-3-217 3 11.22 40.35 50.55 4.33 0.00 4.77 DPM-MG1800-3-211 3 10.95 39.76 49.93 5.09 0.00 5.22 DPM-MG1800-3-213 3 12.43 40.18 50.13 5.01 0.00 4.68 DPM-MG1800-4-214 4 14.98 40.14 51.20 4.36 0.00 4.30 DPM-MG1800-4-221 4 13.77 39.54 50.34 4.35 0.00 5.77 DPM-MG1800-4-220 4 17.10 39.64 51.49 5.21 0.00 3.66 Table E.30: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 1800 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 4 h, 0 RPM, 400µL liquid loading, continued.

Sample code Molar yield Mass yield B:T molar ratio (mol/molDPM reacted) (g/gDPM reacted) (mol/mol) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MG1800-0-210 0.708 0.601 0.000 0.000 0.342 1.18 DPM-MG1800-0-219 0.423 0.359 0.000 0.000 0.607 1.18 DPM-MG1800-0-218 0.635 0.538 0.000 0.000 0.411 1.18 DPM-MG1800-1-208 0.823 0.877 0.042 0.000 0.094 0.94 373 DPM-MG1800-1-222 0.777 0.875 0.043 0.000 0.116 0.89 DPM-MG1800-1-209 0.857 0.902 0.047 0.000 0.059 0.95 DPM-MG1800-2-215 0.855 0.911 0.043 0.000 0.059 0.94 DPM-MG1800-2-212 0.862 0.904 0.057 0.000 0.045 0.95 DPM-MG1800-2-216 0.857 0.910 0.051 0.000 0.050 0.94 DPM-MG1800-3-217 0.869 0.923 0.042 0.000 0.048 0.94 DPM-MG1800-3-211 0.856 0.912 0.049 0.000 0.052 0.94 DPM-MG1800-3-213 0.865 0.915 0.048 0.000 0.047 0.95 DPM-MG1800-4-214 0.865 0.935 0.042 0.000 0.043 0.93 DPM-MG1800-4-221 0.852 0.919 0.042 0.000 0.058 0.93 DPM-MG1800-4-220 0.854 0.940 0.050 0.000 0.037 0.91 Table E.31: Effect of mixing speed on “vortex” height as studied in glass mock-up using reaction product obtained for diphenylmethane hydroconversion performed in the glass insert micro-reactor at 1800 ppm Mo, ◦ 445 C, 13.8 MPa H2, 1 h, 0 RPM.

Mixing speed “Vortex” height (RPM) (mm) 1500 10.1 2000 20.3 2250 23.8 2500 28.5

374 Table E.32: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 0 - 1800 ppm Mo, 445◦C, 13.8 MPa H2, 1 h, 0 RPM, 150 - 400 µL liquid loading.

Sample code Catalyst loading Liquid volume DPM conversion Product composition (DPM-free) (ppm Mo) (µL) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MGT-1-240 0 150 7.65 43.21 50.14 3.54 0.00 3.11 DPM-MGT-1-241 0 150 7.24 43.11 50.27 3.39 0.00 3.22 DPM-MGT-1-242 0 150 8.14 41.80 48.55 3.05 0.00 6.59 DPM-MGT-1-193 0 400 3.27 44.12 41.74 0.00 0.00 14.14 DPM-MGT-1-194 0 400 3.06 42.03 42.60 0.00 0.00 15.36 DPM-MGT-1-195 0 400 3.58 42.09 43.48 0.00 0.00 14.42 DPM-MG1800-1-226 1800 150 11.63 38.26 46.38 15.37 0.00 0.00 DPM-MG1800-1-227 1800 150 11.82 38.92 46.86 14.22 0.00 0.00 DPM-MG1800-1-228 1800 150 9.80 38.92 47.34 13.75 0.00 0.00 DPM-MG1800-1-208 1800 400 5.88 38.21 48.04 4.35 0.00 9.41 DPM-MG1800-1-209 1800 400 5.59 39.79 49.42 4.85 0.00 5.93 375 Table E.33: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 0 - 1800 ppm Mo, 445◦C, 13.8 MPa H2, 1 h, 0 RPM, 150 - 400 µL liquid loading, continued.

Sample code Molar yield Mass yield B:T molar ratio (mol/molDPM reacted) (g/gDPM reacted) (mol/mol) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MGT-1-240 0.931 0.915 0.034 0.000 0.031 1.02 DPM-MGT-1-241 0.928 0.918 0.033 0.000 0.032 1.01 DPM-MGT-1-242 0.900 0.886 0.029 0.000 0.066 1.02 DPM-MGT-1-193 0.950 0.762 0.000 0.000 0.141 1.25 DPM-MGT-1-194 0.905 0.778 0.000 0.000 0.154 1.16 DPM-MGT-1-195 0.907 0.794 0.000 0.000 0.144 1.14 DPM-MG1800-1-226 0.824 0.847 0.148 0.000 0.000 0.97 DPM-MG1800-1-227 0.838 0.856 0.137 0.000 0.000 0.98 DPM-MG1800-1-228 0.838 0.864 0.133 0.000 0.000 0.97 DPM-MG1800-1-208 0.823 0.877 0.042 0.000 0.094 0.94 DPM-MG1800-1-209 0.857 0.902 0.047 0.000 0.059 0.95 Table E.34: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 0 ppm Mo, 445◦C, 13.8 MPa H2, 1 h, 0 RPM, 400 µL liquid loading to study thermocouple wall activation.

Sample code DPM conversion Product composition (DPM-free) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MG1800-1-234 1.60 31.60 60.98 7.42 0.00 0.00 DPM-MGT-1-235 5.37 40.73 51.98 3.53 0.00 3.76 DPM-MGT-1-236 6.65 43.09 49.99 3.47 0.00 3.44 DPM-MGT-1-237 7.40 42.06 51.01 3.32 0.00 3.61 DPM-MGT-1-238 7.68 43.17 50.08 3.27 0.00 3.48 DPM-MGT-1-239 7.99 41.56 48.93 3.37 0.00 6.14 DPM-MGT-1-240 7.65 43.21 50.14 3.54 0.00 3.11 DPM-MGT-1-241 7.24 43.11 50.27 3.39 0.00 3.22 DPM-MGT-1-242 8.14 41.80 48.55 3.05 0.00 6.59 376

Table E.35: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 0 ppm Mo, 445◦C, 13.8 MPa H2, 1 h, 0 RPM, 400 µL liquid loading to study thermocouple wall activation, continued.

Sample code Molar yield Mass yield B:T molar ratio (mol/molDPM reacted) (g/gDPM reacted) (mol/mol) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MG1800-1-234 0.681 1.113 0.072 0.000 0.000 0.61 DPM-MGT-1-235 0.877 0.949 0.034 0.000 0.038 0.92 DPM-MGT-1-236 0.928 0.913 0.034 0.000 0.034 1.02 DPM-MGT-1-237 0.906 0.931 0.032 0.000 0.036 0.97 DPM-MGT-1-238 0.930 0.914 0.032 0.000 0.035 1.02 DPM-MGT-1-239 0.895 0.893 0.033 0.000 0.061 1.00 DPM-MGT-1-240 0.931 0.915 0.034 0.000 0.031 1.02 DPM-MGT-1-241 0.928 0.918 0.033 0.000 0.032 1.01 DPM-MGT-1-242 0.900 0.886 0.029 0.000 0.066 1.02 Table E.36: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 0 ppm Mo, 445◦C, 13.8 MPa H2, 1 h, 0 - 2250 RPM, 150 µL liquid loading.

Sample code Mixing speed DPM conversion Product composition (DPM-free) (RPM) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MGT-1-240 0 7.65 43.21 50.14 3.54 0.00 3.11 DPM-MGT-1-241 0 7.24 43.11 50.27 3.39 0.00 3.22

377 DPM-MGT-1-242 0 8.14 41.80 48.55 3.05 0.00 6.59 DPM-MGT-1-243 1500 6.70 41.93 54.72 3.35 0.00 0.00 DPM-MGT-1-244 1500 7.34 42.08 55.38 2.54 0.00 0.00 DPM-MGT-1-245 1500 5.74 43.03 56.97 0.00 0.00 0.00 DPM-MGT-1-246 2000 3.01 35.54 59.62 4.84 0.00 0.00 DPM-MGT-1-247 2000 2.98 35.94 64.06 0.00 0.00 0.00 DPM-MGT-1-248 2000 3.78 37.68 62.32 0.00 0.00 0.00 DPM-MGT-1-249 2250 1.52 34.67 65.33 0.00 0.00 0.00 DPM-MGT-1-250 2250 1.10 29.75 70.25 0.00 0.00 0.00 DPM-MGT-1-252 2250 2.62 35.04 64.96 0.00 0.00 0.00 Table E.37: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 0 ppm Mo, 445◦C, 13.8 MPa H2, 1 h, 0 - 2250 RPM, 150 µL liquid loading, continued.

378 DPM-MGT-1-242 0.900 0.886 0.029 0.000 0.066 1.02 DPM-MGT-1-243 0.903 0.999 0.032 0.000 0.000 0.90 DPM-MGT-1-244 0.906 1.011 0.025 0.000 0.000 0.90 DPM-MGT-1-245 0.927 1.040 0.000 0.000 0.000 0.89 DPM-MGT-1-246 0.765 1.089 0.047 0.000 0.000 0.70 DPM-MGT-1-247 0.774 1.170 0.000 0.000 0.000 0.66 DPM-MGT-1-248 0.812 1.138 0.000 0.000 0.000 0.71 DPM-MGT-1-249 0.747 1.193 0.000 0.000 0.000 0.63 DPM-MGT-1-250 0.641 1.283 0.000 0.000 0.000 0.50 DPM-MGT-1-252 0.755 1.186 0.000 0.000 0.000 0.64 Table E.38: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 1800 ppm Mo, 445◦C, 13.8 MPa H2, 1 h, 0 - 2250 RPM, 150 µL liquid loading.

Sample code Mixing speed DPM conversion Product composition (DPM-free) (RPM) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MG1800-1-226 0 11.63 38.26 46.38 15.37 0.00 0.00 DPM-MG1800-1-227 0 11.82 38.92 46.86 14.22 0.00 0.00 DPM-MG1800-1-228 0 9.80 38.92 47.34 13.75 0.00 0.00 DPM-MG1800-1-256 1500 11.88 40.24 50.10 7.94 1.72 0.00 DPM-MG1800-1-257 1500 10.74 39.20 49.16 9.65 1.99 0.00 DPM-MG1800-1-258 2000 9.20 38.39 48.98 10.66 1.97 0.00 DPM-MG1800-1-259 2000 11.61 39.81 50.73 9.47 0.00 0.00 DPM-MG1800-1-260 2250 7.56 36.60 50.16 13.23 0.00 0.00 DPM-MG1800-1-261 2250 8.07 38.21 50.32 11.47 0.00 0.00 DPM-MG1800-1-262 2250 8.72 34.22 46.10 19.68 0.00 0.00 379

Table E.39: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 1800 ppm Mo, 445◦C, 13.8 MPa H2, 1 h, 0 - 2250 RPM, 150 µL liquid loading, continued.

Sample code Molar yield Mass yield B:T molar ratio (mol/molDPM reacted) (g/gDPM reacted) (mol/mol) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MG1800-1-226 0.824 0.847 0.148 0.000 0.000 0.97 DPM-MG1800-1-227 0.838 0.856 0.137 0.000 0.000 0.98 DPM-MG1800-1-228 0.838 0.864 0.133 0.000 0.000 0.97 DPM-MG1800-1-256 0.867 0.915 – 0.017 0.000 0.95 DPM-MG1800-1-257 0.844 0.898 0.079 0.020 0.000 0.94 DPM-MG1800-1-258 0.827 0.894 0.096 0.020 0.000 0.93 DPM-MG1800-1-259 0.857 0.926 0.107 0.000 0.000 0.93 DPM-MG1800-1-260 0.788 0.916 0.095 0.000 0.000 0.86 DPM-MG1800-1-261 0.823 0.919 0.132 0.000 0.000 0.90 DPM-MG1800-1-262 0.737 0.842 0.115 0.000 0.000 0.88 Table E.40: Coefﬁcients for the kinetic models of diphenylmethane hydroconversion for data obtained in the glass insert micro-reactor at 0 - 1800 ◦ ppm Mo, 445 C, 13.8 MPa H2, 0 - 4 h, 0 - 2000 RPM, 150 µL liquid load, depicted in Figure E.3.

Catalyst loading Mixing speed 1st order 1 (ppm Mo) (RPM) (h−1) 0 0.073 ± 0.004 0 2000 0.056 ± 0.004 0 0.078 ± 0.007 1800 2 2000 0.061 ± 0.008 1 - rDPM 0ppm = k0ppm.CDPM, rDPM 1800ppm = k1800ppm.CMo.CDPM. 2 - Reporting k1800ppm.CMo for comparison in context of the experiments.

380 Table E.41: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 0 ppm Mo, 445◦C, 13.8 MPa H2, 1 - 4 h, 0 - 2000 RPM, 150 µL liquid loading.

Sample code Mixing speed Reaction time DPM conversion Product composition (DPM-free) (RPM) (h) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MGT-1-240 0 1 7.65 43.21 50.14 3.54 0.00 3.11 DPM-MGT-1-241 0 1 7.24 43.11 50.27 3.39 0.00 3.22 DPM-MGT-1-242 0 1 8.14 41.80 48.55 3.05 0.00 6.59

381 DPM-MGT-2-267 0 2 17.06 43.69 51.58 3.21 0.00 1.52 DPM-MGT-2-268 0 2 16.48 43.20 50.95 2.76 0.00 3.09 DPM-MGT-4-269 0 4 24.03 41.86 51.18 2.77 0.00 4.19 DPM-MGT-4-270 0 4 23.11 42.47 52.67 2.95 0.00 1.92 DPM-MGT-1-246 2000 1 3.01 35.54 59.62 4.84 0.00 0.00 DPM-MGT-1-247 2000 1 2.98 35.94 64.06 0.00 0.00 0.00 DPM-MGT-1-248 2000 1 3.78 37.68 62.32 0.00 0.00 0.00 DPM-MGT-2-271 2000 2 13.96 41.46 51.54 7.00 0.00 0.00 DPM-MGT-2-272 2000 2 12.46 42.31 52.93 4.76 0.00 0.00 DPM-MGT-4-273 2000 4 19.11 40.09 51.55 5.00 0.00 3.36 DPM-MGT-4-274 2000 4 20.12 39.95 53.33 3.79 0.00 2.93 Table E.42: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 0 ppm Mo, 445◦C, 13.8 MPa H2, 1 - 4 h, 0 - 2000 RPM, 150 µL liquid loading, continued.

382 DPM-MGT-2-267 0.941 0.942 0.031 0.000 0.015 1.00 DPM-MGT-2-268 0.930 0.930 0.027 0.000 0.031 1.00 DPM-MGT-4-269 0.902 0.935 0.027 0.000 0.042 0.97 DPM-MGT-4-270 0.915 0.962 0.028 0.000 0.019 0.95 DPM-MGT-1-246 0.765 1.089 0.047 0.000 0.000 0.70 DPM-MGT-1-247 0.774 1.170 0.000 0.000 0.000 0.66 DPM-MGT-1-248 0.812 1.138 0.000 0.000 0.000 0.71 DPM-MGT-2-271 0.893 0.941 0.068 0.000 0.000 0.95 DPM-MGT-2-272 0.911 0.966 0.046 0.000 0.000 0.94 DPM-MGT-4-273 0.864 0.941 0.048 0.000 0.034 0.92 DPM-MGT-4-274 0.860 0.974 0.037 0.000 0.029 0.88 Table E.43: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 1800 ppm Mo, 445◦C, 13.8 MPa H2, 1 - 4 h, 0 - 2000 RPM, 150 µL liquid loading.

Sample code Mixing speed Reaction time DPM conversion Product composition (DPM-free) (RPM) (h) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MG1800-1-226 0 1 11.63 38.26 46.38 15.37 0.00 0.00 DPM-MG1800-1-227 0 1 11.82 38.92 46.86 14.22 0.00 0.00

383 DPM-MG1800-1-228 0 1 9.80 38.92 47.34 13.75 0.00 0.00 DPM-MG1800-2-277 0 2 18.76 39.53 49.45 7.65 0.98 2.38 DPM-MG1800-2-278 0 2 17.47 40.20 50.10 6.98 0.00 2.72 DPM-MG1800-4-279 0 4 25.50 40.30 50.68 6.50 0.00 2.53 DPM-MG1800-4-280 0 4 22.17 40.32 52.59 6.05 0.00 1.04 DPM-MG1800-1-258 2000 1 9.20 38.39 48.98 10.66 1.97 0.00 DPM-MG1800-1-259 2000 1 11.61 39.81 50.73 9.47 0.00 0.00 DPM-MG1800-2-281 2000 2 13.11 39.39 51.34 7.83 1.44 0.00 DPM-MG1800-2-282 2000 2 12.32 38.98 51.51 8.14 1.37 0.00 DPM-MG1800-4-283 2000 4 18.62 37.89 52.43 8.61 1.07 0.00 DPM-MG1800-4-284 2000 4 – 36.66 52.76 9.48 – – Table E.44: Results for diphenylmethane hydroconversion in the glass insert micro-reactor at 1800 ppm Mo, 445◦C, 13.8 MPa H2, 1 - 4 h, 0 - 2000 RPM, 150 µL liquid loading, continued.

Sample code Molar yield Mass yield B:T molar ratio (mol/molDPM reacted) (g/gDPM reacted) (mol/mol) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MG1800-1-226 0.824 0.847 0.148 0.000 0.000 0.97 DPM-MG1800-1-227 0.838 0.856 0.137 0.000 0.000 0.98 DPM-MG1800-1-228 0.838 0.864 0.133 0.000 0.000 0.97 384 DPM-MG1800-2-277 0.851 0.903 0.074 0.010 0.024 0.94 DPM-MG1800-2-278 0.866 0.915 0.067 0.000 0.027 0.95 DPM-MG1800-4-279 0.868 0.925 0.063 0.000 0.025 0.94 DPM-MG1800-4-280 0.869 0.960 0.058 0.000 0.010 0.90 DPM-MG1800-1-258 0.827 0.894 0.103 0.020 0.000 0.93 DPM-MG1800-1-259 0.857 0.926 0.091 0.000 0.000 0.93 DPM-MG1800-2-281 0.848 0.937 0.076 0.014 0.000 0.91 DPM-MG1800-2-282 0.840 0.940 0.079 0.014 0.000 0.89 DPM-MG1800-4-283 0.816 0.957 0.083 0.011 0.000 0.85 DPM-MG1800-4-284 0.790 0.963 0.091 0.000 0.011 0.82 Table E.45: Results for diphenylmethane hydroconversion performed in the glass insert micro-reactor at 1800 ppm ◦ Mo, 445 C, 13.8 MPa H2, 1 h, 2000 RPM, 150 µL liquid load using various coke-catalyst agglomerates from residue hydroconversion experiments.

Sample code Coke type DPM conversion Product composition (DPM-free) (wt%) (wt%) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MGC74-1-286 Heat treated 11.06 34.50 44.20 21.31 0.00 0.00 DPM-MGC54-1-287 Fresh 0.54 29.98 70.02 0.00 0.00 0.00 DPM-MGC44-1-289 Recycled 1.85 8.63 74.27 10.57 0.00 6.53 385

Table E.46: Results for diphenylmethane hydroconversion performed in the glass insert micro-reactor at 1800 ppm ◦ Mo, 445 C, 13.8 MPa H2, 1 h, 2000 RPM, 150 µL liquid load using various coke-catalyst agglomerates from residue hydroconversion experiments, continued.

Sample code Molar yield Mass yield B:T molar ratio (mol/molDPM reacted) (g/gDPM reacted) (mol/mol) Benzene Toluene CHMB Other cracking Other isom./cond. DPM-MGC74-1-286 0.743 0.807 0.206 0.000 0.000 0.92 DPM-MGC54-1-287 0.646 1.278 0.000 0.000 0.000 0.51 DPM-MGC44-1-289 0.186 1.356 0.102 0.000 0.065 0.14 Appendix F

Data Processing and Analysis

F.1 Data Acquisition and Analysis Two programs were used for the extraction of data from GCMS chromatograms and its subsequent processing (removing noise, scaling peaks, etc.). The former was an AutoHotkey (v1.1.14.03) script shown in Listing F.1. The MATLAB (R2012a) code, presented in Listing F.2, used a graphical user interface, Figure F.1, to import and process the peak data as desired.

Listing F.1: AutoHotkey (v1.1.14.03) script used for extraction of data from GCMS chromatograms.

; Ross Kukard, May 2013 ; Script used to collect data from GCMS files and trasnfer to Excel #NoEnv ; Recommended for performance and compatibility with f uture AutoHotkey releases . #Warn ; Recommended f o r catching common e r r o r s . SendMode Input ; Recommended f o r new s c r i p t s due to i t s s u p e r ior speed and reliability . SetWorkingDir %A ScriptDir% ; Ensures a consistent starting directory. ; Starting the actual script ; Will commandeer the F keys from 12 through 5−i sh to use $F5:: ; copy filename and open file send {F2} send ˆc send { enter } send { enter } send ! l h r e t u r n $F6:: ; paste filename and move down one row send ˆv

386 send {down} r e t u r n $F7:: ; copy data and open new file send ˆc send {esc} send ˆ o r e t u r n $F8:: ; paste data and move over one column send ˆv send { right } send { right } send { right } send {up} r e t u r n $F9:: ; copy, tab forward send ˆc send ˆ{ pgdn} r e t u r n $F10:: ; paste transpose, down send { lalt } send {e} send {s} send {e} send { enter } send {down} send ˆ{ pgup} r e t u r n

Listing F.2: MATLAB (R2012a) code used in conjunction with graphical user interface for processing of GCMS chromatogram results.

% Ross Kukard − Department of Chemical and Biological Engineering − UBC % 2013 % Simple GUID for extracting and analysing data from GCMS data stored i n % spreadsheets

% Most recent updates % − Now with an increment button :D

function varargout = Data processor 1(varargin) %DATA PROCESSOR 1 M−file for Data processor 1 . f i g % DATA PROCESSOR 1, by itself , creates a new DATA PROCESSOR 1 or raises the e x i s t i n g % s i n g l e t o n * . % % H = DATA PROCESSOR 1 returns the handle to a new DATA PROCESSOR 1 or the handle to % the existing singleton * . % % DATA PROCESSOR 1(’Property ’,’Value ’ ,...) creates a new DATA PROCESSOR 1 using the % given property value pairs. Unrecognized properties are passed vi a % varargin to Data processor 1 OpeningFcn. This calling syntax produces a % warning when there is an existing singleton * .

387 % % DATA PROCESSOR 1( ’CALLBACK ’ ) and DATA PROCESSOR 1( ’CALLBACK’ ,hObject ,...) c a l l the % l o c a l f u n c t i o n named CALLBACK i n DATA PROCESSOR 1.M with the given input % arguments. % % *See GUI Options on GUIDE’s Tools menu. Choose ”GUI allows only one % instance to run (singleton)”. % % See also : GUIDE, GUIDATA, GUIHANDLES

% Edit the above text to modify the response to help Data processor 1

% Last Modified by GUIDE v2.5 30−Aug−2013 11:50:18

% Begin initialization code − DO NOT EDIT gui Singleton = 1; gui State = struct(’gui Name’, mfilename, ... ’gui Singleton’, gui Singleton , ... ’gui OpeningFcn ’ , @Data processor 1 OpeningFcn, ... ’gui OutputFcn’, @Data processor 1 OutputFcn, ... ’gui LayoutFcn’, [], ... ’gui Callback’, []); if nargin && ischar(varargin {1}) gui State.gui Callback = str2func(varargin {1}) ; end

if nargout [ varargout {1: nargout }] = gui mainfcn(gui State , varargin {:}) ; else gui mainfcn(gui State , varargin {:}) ; end % End initialization code − DO NOT EDIT

% −−− Executes just before Data processor 1 is made visible. function Data processor 1 OpeningFcn(hObject , eventdata , handles, varargin) % This function has no output args, see OutputFcn. % hObject handle to figure % eventdata reserved − to be defined in a future version of MATLAB % handles structure with handles and user data (see GUIDATA) % varargin unrecognized PropertyName/PropertyValue pair s from the % command l i n e ( see VARARGIN)

% First define the file and sheet handles.filename = ’Sample repreps − microreactor2.xlsx ’; handles.sheetname = ’BT Data ’ ;

% Define the range of the data handles. col start = ’mk’; handles. row start = 3; handles. col end = ’mm’ ; handles.row end = 500;

handles. key areas = zeros (1:5) ; handles.data = []; handles . names = [ ] ;

388 % Assign these default values to the GUI edit boxes set (handles. edit filename , ’String ’ ,handles.filename) set (handles.edit sheetname , ’String ’ ,handles.sheetname) set (handles. edit start col , ’String ’,handles.col start) set (handles. edit start row , ’String ’,handles.row start) set (handles. edit e n d col , ’String ’ ,handles.col end) set (handles. edit end row , ’String ’ ,handles.row end)

% Choose default command line output for Data processor 1 handles.output = hObject;

% Update handles structure guidata(hObject , handles);

% UIWAIT makes Data processor 1 wait for user response (see UIRESUME) % uiwait(handles.figure1);

% −−− Outputs from this function are returned to the command line. function varargout = Data processor 1 OutputFcn(hObject , eventdata , handles) varargout {1} = handles.output;

function edit start col Callback(hObject , eventdata , handles)

% −−− Executes during object creation , after setting all propert ies . function edit start col CreateFcn(hObject , eventdata , handles) i f ispc && isequal( get (hObject , ’BackgroundColor ’) , get (0, ’ defaultUicontrolBackgroundColor ’)) set (hObject , ’BackgroundColor ’ , ’white ’); end

function edit start r o w Callback(hObject , eventdata , handles)

function edit start r o w CreateFcn(hObject , eventdata , handles) i f ispc && isequal( get (hObject , ’BackgroundColor ’) , get (0, ’ defaultUicontrolBackgroundColor ’)) set (hObject , ’BackgroundColor ’ , ’white ’); end

function edit e n d col Callback(hObject , eventdata , handles)

% −−− Executes during object creation , after setting all propert ies . function edit end col CreateFcn(hObject , eventdata , handles) i f ispc && isequal( get (hObject , ’BackgroundColor ’) , get (0, ’ defaultUicontrolBackgroundColor ’)) set (hObject , ’BackgroundColor ’ , ’white ’); end

function edit end row Callback(hObject , eventdata , handles)

% −−− Executes during object creation , after setting all propert ies . function edit end row CreateFcn(hObject , eventdata , handles) i f ispc && isequal( get (hObject , ’BackgroundColor ’) , get (0, ’ defaultUicontrolBackgroundColor ’)) set (hObject , ’BackgroundColor ’ , ’white ’); end

389 % −−− Executes on button press in button extract . function b u t t o n extract Callback(hObject , eventdata , handles) % Get latest values from edit boxes handles.filename = get (handles. edit filename , ’String ’); handles.sheetname = get (handles.edit sheetname , ’String ’); handles. col start = get (handles. edit start col , ’String ’); handles. row start = get (handles. edit start row , ’String ’); handles. col end = get (handles. edit e n d col , ’String ’); handles.row end = get (handles. edit end row , ’String ’);

% Set peak locations from previous runs to zero set (handles. edit dpe row , ’String ’,’0’) set (handles.edit benzene row, ’String ’,’0’) set (handles. edit t o l u e n e row , ’String ’,’0’) set (handles.edit chmb row, ’String ’,’0’) set (handles.edit dpm row, ’String ’,’0’)

% Clear old handles handles. key areas = zeros (1,5); handles.data = []; handles . names = [ ] ;

% Collate into coordinates for extraction coords = [handles.col start num2str (handles. row start) ’: ’ handles.col end num2str (handles.row end) ];

% Extract from Excel file [handles.data, handles.names] = xlsread(handles.filename , handles .sheetname , coords ) ;

% Test data set %handles.data = [1 0.1; 3 4; 5 6; 7 8]; %handles.names = [’abc’; ’def’; ’ghi’; ’jkl ’]; f o r tab = [num2cell(handles.data) ,cellstr(handles.names) ];

% Display data in the table set (handles.table1 , ’data ’ ,for t a b , ’ColumnName ’ ,{ ’ Retention Time (min)’, ’Area (−) ’ , ’Name ’ }) ;

guidata(hObject , handles);

function edit f i l e n a m e Callback(hObject , eventdata , handles)

% −−− Executes during object creation , after setting all propert ies . function edit filename CreateFcn(hObject , eventdata , handles) i f ispc && isequal( get (hObject , ’BackgroundColor ’) , get (0, ’ defaultUicontrolBackgroundColor ’)) set (hObject , ’BackgroundColor ’ , ’white ’); end

function edit sheetname Callback(hObject , eventdata , handles)

% −−− Executes during object creation , after setting all propert ies . function edit sheetname CreateFcn(hObject , eventdata , handles)

390 i f ispc && isequal( get (hObject , ’BackgroundColor ’) , get (0, ’ defaultUicontrolBackgroundColor ’)) set (hObject , ’BackgroundColor ’ , ’white ’); end

function edit d e s i r e d m i n Callback(hObject , eventdata , handles)

% −−− Executes during object creation , after setting all propert ies . function edit desired min CreateFcn(hObject , eventdata , handles) i f ispc && isequal( get (hObject , ’BackgroundColor ’) , get (0, ’ defaultUicontrolBackgroundColor ’)) set (hObject , ’BackgroundColor ’ , ’white ’); end

% −−− Executes on button press in button analyze . function button analyze Callback(hObject , eventdata , handles) % Call in the handles so we can work with them data = handles.data; names = handles . names ;

% Collect values from the relevant edit boxes min area = str2num ( get (handles. edit d e s i r e d min , ’String ’)); benz row = str2num ( get (handles.edit benzene row, ’String ’)); tol row = str2num ( get (handles. edit t o l u e n e row , ’String ’)); chmb row = str2num ( get (handles.edit chmb row, ’String ’)); dpm row = str2num ( get (handles.edit dpm row, ’String ’)); dpe row = str2num ( get (handles. edit dpe row , ’String ’)); solv start = str2num ( get (handles. edit solv start , ’String ’)); solv end = str2num ( get (handles. edit solv end , ’String ’));

flagged = 0;

index = true(1, size (data,1));

for i = 1: length (data) % Run through all of the rows i f data (i ,1) > solv start & data (i,1) < solv end % If the retention time is in the solvent range index(i) = false; set (handles. edit dpe row , ’String ’,’0’) set (handles.edit benzene row, ’String ’,’0’) set (handles. edit t o l u e n e row , ’String ’,’0’) set (handles.edit chmb row, ’String ’,’0’) set (handles.edit dpm row, ’String ’,’0’) flagged = 1; end end

% Knock out the DPE internal standard i f dpe row ˜= 0 % If it is zero, then it hasn’t been added index ( dpe row) = false; set (handles. edit dpe row , ’String ’,’0’) % Set the edit box back to zero to prevent unintentionally removing other rows % Set all other boxes to zero too to prevent confusion set (handles.edit benzene row, ’String ’,’0’) set (handles. edit t o l u e n e row , ’String ’,’0’) set (handles.edit chmb row, ’String ’,’0’)

391 set (handles.edit dpm row, ’String ’,’0’) flagged = 1; end

key areas = handles.key areas ;

% Extract benzene, toluene , CHMB and DPM areas i f benz row ˜= 0 % If it is zero, then it hasn’t been added key areas (1,1) = data (benz row,2) ; % Collect the appropriate area end i f tol r o w ˜= 0 key areas (1,2) = data (tol row ,2); end i f chmb row ˜= 0 key areas (1,4) = data (chmb row,2) ; end i f dpm row ˜= 0 %DPM gets located and removed only once, so don’t overwrite i t once we have it key areas (1,3) = data (dpm row,2) ; end % Calculate the cracking area (everything less than chmb, not B or T) i f chmb row ˜= 0 subindex = index(1:(chmb row−1)) ; % Narrow the index to only the range of interest subdata = data(subindex,2) ; % Get rid of the extra data leaving only cracking products key areas (1,5) = sum(subdata)− key areas (1,1) − key areas (1,2); % Add the cracking products and remove B and T elseif dpm row ˜= 0 % Maybe there is no CHMB, then use the DPM peak instead subindex = index(1:(dpm row−1)) ; subdata = data(subindex,2) ; key areas (1,5) = sum(subdata)− key areas (1,1) − key areas (1,2); else % E.g. if those peaks haven’t been entered yet key areas (1,5) = 0; end % Isom and condensation area calculated after offenders removed

% Need to remove the offeding rows, then calculate area and %, then remove % those rows with too small an area i f flagged == 1 data = data(index ,:) ; % Remove offenders from data and names list before calculating the area percents names = names ( index , : ) ; end

% Calculate total area and percentages total area = sum(data); total area = total area(2) − key areas (1,3); %Do not include DPM ... but do not delete it from the table, doing so gives you no cracking calc reference if there is no CHMB data(:,3) = data(:,2)/total a r e a *100; % Need to re−index data matrix as it changed size when we kicked out the % offenders above index = true(1, size (data,1)); % Knock out areas that are too small for i = 1: length (data) % Run through all of the rows

392 i f data(i ,3) < min area % If one of the area percentages is too small ... index(i) = false; % Flag it for removal flagged = 1; end end

% Knock out the ones that had too small an area i f flagged == 1 data = data(index ,:) ; % Remove offenders from data and names list before calculating the area percents names = names ( index , : ) ; flagged = 0; msgbox( ’Minimum area , solvent , DPM and/ or DPE peaks removed . Please re−enter peak locations. ’ , ’Warning’ , ’Warn’) end % And finally calculate the total DPM, DPE, solvent free area and % total area = sum(data); total area = total area(2) − key areas (1,3); % Again less the DPM, although it is still in the table data(:,3) = data(:,2)/total a r e a *100;

% Calculate isom/condensation area ... i.e. total less everything else key areas (1,6) = total a r e a − (key areas(1,1) + key areas(1,2) + key areas(1,4) + key areas(1,5)); %DPM has already been removed, hence no (1,3)

% Display data in the tables f o r tab = [num2cell(data),cellstr(names)]; set (handles.table1 , ’data ’ ,for t a b , ’ColumnName ’ ,{ ’ Retention Time (min)’, ’Area (−) ’ , ’ Area Percent ’ , ’Names’ }) ; set (handles.table2 , ’data ’ ,key areas , ’ColumnName ’ ,{ ’Benzene’, ’Toluene’, ’DPM’, ’ CHMB’ , ’Cracking ’ , ’Isom,cond’ }) ;

% And display in the text set (handles. text total a r e a value , ’String ’,total area) % Set the total area display set (handles. text min area value , ’String ’, num2str ( min (data(: ,3)))) set (handles. text max area value , ’String ’, num2str (max(data(: ,3))))

handles.data = data; handles . names = names ; handles. key areas = key areas ; % Save handles guidata(hObject , handles);

function edit benzene row Callback(hObject , eventdata , handles)

% −−− Executes during object creation , after setting all propert ies . function edit benzene row CreateFcn(hObject , eventdata , handles) i f ispc && isequal( get (hObject , ’BackgroundColor ’) , get (0, ’ defaultUicontrolBackgroundColor ’)) set (hObject , ’BackgroundColor ’ , ’white ’); end

function edit t o l u e n e r o w Callback(hObject , eventdata , handles)

% −−− Executes during object creation , after setting all propert ies .

393 function edit toluene row CreateFcn(hObject , eventdata , handles) i f ispc && isequal( get (hObject , ’BackgroundColor ’) , get (0, ’ defaultUicontrolBackgroundColor ’)) set (hObject , ’BackgroundColor ’ , ’white ’); end

function edit chmb row Callback(hObject , eventdata , handles)

% −−− Executes during object creation , after setting all propert ies . function edit chmb row CreateFcn(hObject , eventdata , handles) i f ispc && isequal( get (hObject , ’BackgroundColor ’) , get (0, ’ defaultUicontrolBackgroundColor ’)) set (hObject , ’BackgroundColor ’ , ’white ’); end

function edit dpm row Callback(hObject , eventdata , handles)

% −−− Executes during object creation , after setting all propert ies . function edit dpm row CreateFcn(hObject , eventdata , handles) i f ispc && isequal( get (hObject , ’BackgroundColor ’) , get (0, ’ defaultUicontrolBackgroundColor ’)) set (hObject , ’BackgroundColor ’ , ’white ’); end

function edit dpe row Callback(hObject , eventdata , handles)

% −−− Executes during object creation , after setting all propert ies . function edit dpe row CreateFcn(hObject , eventdata , handles) i f ispc && isequal( get (hObject , ’BackgroundColor ’) , get (0, ’ defaultUicontrolBackgroundColor ’)) set (hObject , ’BackgroundColor ’ , ’white ’); end

function edit solv start Callback(hObject , eventdata , handles)

% −−− Executes during object creation , after setting all propert ies . function edit solv start CreateFcn(hObject , eventdata , handles) i f ispc && isequal( get (hObject , ’BackgroundColor ’) , get (0, ’ defaultUicontrolBackgroundColor ’)) set (hObject , ’BackgroundColor ’ , ’white ’); end

function edit solv e n d Callback(hObject , eventdata , handles)

% −−− Executes during object creation , after setting all propert ies . function edit solv end CreateFcn(hObject , eventdata , handles) i f ispc && isequal( get (hObject , ’BackgroundColor ’) , get (0, ’ defaultUicontrolBackgroundColor ’)) set (hObject , ’BackgroundColor ’ , ’white ’); end

% −−− Executes on button press in button3. function button3 Callback(hObject , eventdata , handles)

key areas = handles.key areas clipboard( ’copy’ , num2str (handles.key areas));

394 % −−− Executes on button press in button increment . function button increment Callback(hObject , eventdata , handles) % char converts numbers to letters , double is the reverse % Keep everything lowercase for simplicity %a is 97, z is 122 % Each extraction is 3 columns wide, for example a−>c % This would start at 97 and end at 99 % The increment would add 3 to each to get d−>f as 100−>102

% Extract current values from the GUI handles. col start = get (handles. edit start col , ’String ’); handles. col end = get (handles. edit e n d col , ’String ’);

% First we pull the current columns from the system current start = handles.col start ; current end = handles.col end ;

% Check if we are operating in single letter columns or double if length (current s t a r t ) == 1 % i.e. still in the single digits i f double( current s t a r t ) +3 > 122 % i.e. it increments beyond z current start = char([97 double(current s t a r t ) +3−123+97]) ; % Add ’ a ’ to the beginning, see how far over the second value goes and increment from a else current start = char (current start+3); end

%Now onto the end column, need to check if it has hit double dig i t s if length (current end ) == 2 % Simplified test as if start is not double, end will not be incrementing it ’s first letter current end(2) = char (current end(2)+3); elseif double(current end ) +3 > 122 % If it ’s single but will increment beyond z current end = char([97 double(current end ) +3−123+97]) ; else % It ’s single and will not increment beyond z current end = char (current end+3); end

elseif length (current s t a r t ) == 2 % i.e. double digits ... implies the the end is also double digits i f double( current start(2))+3 > 122 % i.e. if the second letter increments beyond z % Increment the first and second letter i f double( current start(1))+3 > 122 %i.e. if the first letter would go beyond z msgbox( ’ First column identifier increments outside range . Please increment manually. ’ , ’Warning ’ , ’Warn’) else % Increment first one by ONE !!!! NOT BY 3 !!!! and the second one by three current start = char([double(current start(1))+1 double(current start ( 2 ) ) +3−123+97]) ; end else % If the second letter does not increment beyond z, only increment the second letter current start(2) = char (current start (2)+3);

395 end

% Again we repeat for the end column, if start is in doubles, the end % will be too so we don’t need to test it i f double(current end ( 2 ) ) +3 > 122 % i.e. if the second letter increments beyond z % Increment the first and second letter i f double(current end ( 1 ) ) +3 > 122 %i.e. if the first letter would go beyond z msgbox( ’ First column identifier increments outside range . Please increment manually. ’ , ’Warning ’ , ’Warn’) else % Increment first one by ONE !!!! NOT BY 3 !!!! and the second one by three current end = char([double(current end(1))+1 double(current end(2)) +3−123+97]) ; end else % If the second letter does not increment beyond z, only increment the second letter current end(2) = char (current end(2)+3); end else % If it can’t figure out the number of letters in the column name msgbox( ’Could not determine column parameters . Please increment manually. ’ , ’ Warning ’ , ’Warn’) end

% Finally we set the new handles for the start ad end columns handles. col start = current start ; handles. col end = current end ; set (handles. edit start col , ’String ’,handles.col start) set (handles. edit e n d col , ’String ’ ,handles.col end)

% Save the handles guidata(hObject , handles);

396 Figure F.1: MATLAB (R2012a) graphical user interface used for processing of GCMS chromatogram results.

397 F.2 Gas and Liquid Mass Balances

F.2.1 Liquid Product Liquid product mass balances were performed for both the stirred batch and micro- reactors. For the stirred batch system, the mass of liquid loaded into the reactor, mLTotal.in, was determined by weighing the empty reactor and re-weighing once the reaction mixture had been added. Similarly, the mass of liquid after the reaction, mLTotal.out, was determined by the difference between the initial empty reactor mass and the mass of the reactor with liquid after the reaction. The difference in these masses, ∆mLTotal, is then used to determine the overall mass balance, mBal, and subsequently the mean,m ¯ Bal, and standard deviation, sm.Bal, therein. These loaded and recovered masses together with their mass balances, the mean and standard deviation, are shown in Table F.1 for ten experiments with a sample calculation of mBal shown in Equation F.1 using the ﬁrst set of data. Note that some of the recovered masses are actually higher than the loaded masses. This may be due to various experimental factors such as inaccurate positioning of the reactor (and the heating element cables) on the balance or remnants of the graphite gasket sticking to the sealing face or falling into the liquid during opening. The procedure was identical for the micro-reactor systems, with the mean mass balance for all experiments (spanning inclined stainless steel, vertical stainless steel and glass insert reactors) being 96.8 wt% ± 2.2 wt%. The greater uncertainty in these results is thought to be due to the liquid left, on the thermocouple after the reaction, contributing toward the perceived mass lost. Whilst a similar effect would also occur in the stirred batch reactor, the amount in that system would be negligible compared to the total mass loaded.

mLTotal.in −|mLTotal.out| ∆mLTotal = 1 − mLTotal.in 80.36 − 81.38 = 1 − 80.36 (F.1) = 0.987 = 98.7%

398 Table F.1: Mass balance results for DPM hydroconversion in the stirred batch ◦ reactor at 445 C, 13.8 MPa H2, 0 - 6 h, 0 - 600 ppm Mo, 700 RPM.

Sample code Liquid masses (g) Mass balance Loaded Recovered (wt%) DPM-BT-27 80.36 81.38 98.7 DPM-BT-1-44 80.35 78.79 98.1 DPM-BT-6-45 80.33 78.30 97.5 DPM-BT-0-48 80.33 80.18 99.8 DPM-BC600-1-34 80.34 79.58 99.1 DPM-BC600-6-37 80.34 81.94 98.0 DPM-BC600-1-39 80.33 80.36 100.0 DPM-BC600-1-41 80.33 79.47 98.9 DPM-BC600-1-42 80.33 81.30 98.8 DPM-BC600-0-46 80.33 80.38 99.9 Mean 98.9 Standard deviation 0.9

F.2.2 Gas Product Indicated when dealing with gaseous products in Chapter 4, the mass of species reporting to the gas phase was negligible. This was determined by considering the composition of a representative gas sample and the volume of that sample in the reactor. As an example, consider the gas product from the vertical stainless steel micro-reactor for 1800 ppm Mo after 4h (Table 4.13), one of the highest gas yields observed in the micro-reactor for DPM hydroconversion with the hydrocarbon species comprising a total of approximately 0.564 wt%. Per the data below and Equation F.2, the mass contribution of the gaseous products is only approximately 2.0 % that of the total liquid mass (assuming no gaseous products leave the reactor and enter the gas supply lines during reaction). 3 Loaded liquid volume (cm ) = VLTotal = 0.400

Loaded liquid mass (g) = mLTotal = 0.402

Hydrocarbon gas concentration (wt%) = CGTotal = 0.564 Reactor length (cm) = L = 25.0 Reactor inner diameter (cm) = ID = 0.3048

399 π.ID2 V = × L GLTotal 4 π.0.3048 = × 25.0 4 = 1.82cm3

VGTotal = VGLTotal −VLTotal = 1.82 − 0.40 = 1.42cm3 C (F.2) m = V × GTotal GTotal GTotal 100 0.564 = 1.42 × 100 = 0.008g mGTotal φGL = mLTotal 0.008 = 0.402 = 0.020 = 2.0%

For the sake of argument, assume that in the case of the glass insert micro- reactor with 150 µL of liquid, the gaseous hydrocarbon species mix freely with the total volume of both the insert and the rest of the shell. The highest hydrocarbon gas concentrations from DPM hydroconversion in this system are with 1800 ppm Mo, 4 h at 0 RPM (Table 4.18), totaling 0.0809 wt%. Per the data below, as shown in Equation F.3, this scenario results in the gaseous products contributing only 2.9 % of the mass that the liquid does.

400 3 Loaded liquid volume (cm ) = VLTotal = 0.150

Loaded liquid mass (g) = mLTotal = 0.151

Hydrocarbon gas concentration (wt%) = CGTotal = 0.0809

Insert length (cm) = LInsert = 25.0

Insert inner diameter (cm) = IDInsert = 0.40

Insert outer diameter (cm) = ODInsert = 0.60

Shell length (cm) = LShell = 50.0

Shell inner diameter (cm) = IDShell = 0.6223

Glass bead outer diameter (cm) = ODBead = 0.6223

Close-packing factor (−) = FBead = 0.74048

Bead packing height (cm) = hBead = 25.0

401 VGLTotal = VG.Insert +(VG.Shell) π.ID2 = Insert × L + 4 Insert π.ID2 ( Shell × L − 4 Shell π.OD2 Insert × L − 4 Insert π.ID2 Shell × h × F ) 4 Bead Bead π.0.402 = × 25.0+ 4 π.0.622 ( × 50.0− 4 π.0.602 × 25.0− 4 π.0.622 × 25.0 × 0.74) 4 = 3.14 +(15.21 − 7.07 − 5.63) (F.3) = 5.65cm3

VGTotal = VGLTotal −VLTotal = 5.65 − 0.15 = 5.5cm3 C m = V × GTotal GTotal GTotal 100 0.0809 = 5.5 × 100 = 0.004g mGTotal φGL = mLTotal 0.004 = 0.151 = 0.029 = 2.9%

402 F.3 Kinetic Analyses Widely known and used, simple ﬁrst and second order kinetics were utilised during this study in an attempt to quantify the rate of reaction.

F.3.1 First Order Defined per Equation F.4, the first order rate equation may be rearranged and integrated as shown. Thus, a plot of ln[A] versus t yields a slope of −k and an intercept of ln[A]0. As an example, the DPM conversion data for 600 ppm Mo from Figure 4.6 is presented in Table F.2. Aside from the reaction time and DPM conversion, three other sets of numbers may be seen. The zeroed conversion, X′ is calculated per Equation F.5 such that calculations begin with X0 = 0. The next two columns are for plotting to determine kinetic coefficients with ln(1 − X′/100) of interest in first order kinetics. As the kinetic equations are defined in terms of compositions, it is necessary to convert the DPM composition data for it to be applicable. This is achieved by assuming initially pure DPM and subtracting the conversion from that to determine a normalised concentration. Figure F.2 shows the data from Table F.2 plotted for first order kinetic analysis. The slope of a linear ′ −1 fit gives k = 0.087 h whilst the intercept gives Xt0 = 0 wt%, as defined. Note that this coefficient is defined as k′ and not k, the former indicating this to be a ′ catalytic system and hence k = CMo.k.

d[A] − = k[A] dt d[A] = −kdt [A] [A] 1 t (F.4) d[A] = − kdt [A] Z[A]0 Zt0 ln[A] − ln[A]0 = −kt

ln[A] = −kt + ln[A]0

′ Xt = Xt − Xt0 (F.5)

403 0.0

Intercept = 0

-0.2

Slope = -0.087 ± 0.008 ln(1-X'/100)

-0.4

-0.6

0 2 4 6

Reaction time (h)

Figure F.2: Plot of conversion results for diphenylmethane hydroconversion ◦ in the stirred batch reactor at 600 ppm Mo, 445 C, 13.8 MPa H2,0-6 h, 700 RPM for first order kinetic analysis. F.3.2 Second Order The second order kinetic model is defined, rearranged and integrated as shown in Equation F.6. A plot of 1/[A] versus t would thus yield a slope of k and an intercept of 1/[A]0. As discussed above, the mass composition data is converted to normalised concentration and plotted, Figure F.3. From the linear fit is may be seen that k′ = 0.098 wt%−1.h−1 (and as per the first order fit, this being a catalytic ′ system, k = CMo.k) and [A]0 = 0.934. This latter value corresponds to a model- ′ predicted X0 = 6.6 wt%.

d[A] − = k[A]2 dt d[A] = −kdt [A]2 [A] 1 t d[A] = − kdt (F.6) [A]2 Z[A]0 Zt0 1 1 − = kt [A] [A]0 1 1 = kt + [A] [A]0

404 1.8

1.6

1.4

Slope = 0.098 ± 0.011 1/(1-X'/100)

1.2

Intercept = 1.071 ± 0.043

1.0

0 2 4 6

Reaction time (h)

Figure F.3: Plot of conversion results for diphenylmethane hydroconversion ◦ in the stirred batch reactor at 600 ppm Mo, 445 C, 13.8 MPa H2,0-6 h, 700 RPM for ﬁrst order kinetic analysis.

F.3.3 The Arrhenius Law Application of the Arrhenius Law to batch reactor data begins by assuming constant reactor volume, V, ﬁxed reaction time, t, and a ﬁrst order reaction as per Equation F.7.

d[X] N = −r V A0 dt A

= kCA

= kCA0(1 − X) 1 (F.7) N d[X] = kC dt A0 1 − X A0 −ln(1 − X) = kt = k′′

The Arrhenius Law, Equation F.8, may thus be substituted into Equation F.7 as per Equation F.9. A plot of ln(−ln(1 − X)) versus 1/T would thus yield an ′′ intercept of ln(k0 ) and a slope of −Ea/R.

405 −Ea k = k0e RT −Ea (F.8) ln(k) = ln(k ) − 0 RT

−Ea ln(k′′) = ln(k′′) − 0 RT −Ea (F.9) ln(−ln(1 − X)) = ln(k′′) − 0 RT

F.4 Thermodynamic Simulations To lend some support to the radical reaction mechanisms proposed in Section 5.1.2, a number of thermodynamic simulations were conducted to calculate the Gibbs free energies of reaction for the abstraction of H* from various positions on DPM and

CHMB molecules. For comparison, these were all based on an attack by *CH3 to affect either abstraction of H* to form CH4 or radical addition. It was found that abstraction of H* from one of the DPM rings was, for many positions, a spontaneous reaction. By far the most favourable position for abstraction was found to be from the bridge carbon as shown in Reaction A of Table F.3, with this radical likely to stabilise onto either of the aromatic rings (due to their resonance structures). The internal ortho location was seen to be the most favoured radical position on the ring, as shown in Reaction B. This molecule may then close the inter-ring bond, stabilising the radical on a tertiary carbon as seen in Reaction C. With one H* removed, attack by another *CH3 would result in either radical addition at the current radical or abstraction of a second H*. The former reaction, as per the mechanisms of Reactions D and E in Table F.3, is more favourable for a radical in the internal ortho position than on the bridge and supports the presence of the MBP species in the liquid product. The latter reaction, assuming Reaction C occurs first, is spontaneous and re-aromaticises the molecule to form fluorene as shown in Reaction F. In the case of CHMB, the abstraction of the first H* is again most favoured from the bridge carbon (Reaction G in Table F.4) but only slightly more so than positions on the saturated ring (such as the internal ortho position in Reaction H).

406 Abstraction of the ﬁrst H* from the aromatic ring is unlikely (exempliﬁed in Reac- tion I). Analogous to the reactions seen for DPM in Table F.3, with one H* removed from the saturated ring of CHMB it may close the inter-ring bond per Reaction J. Following this step, abstraction of a second H* from the molecule stabilises the hexF molecule as shown in Reaction K. Simulation of the above reported thermodynamic properties was performed in Accelrys Materials Studio (v4.4) using DMol3 [110, 111] geometry optimisation and frequency calculations. The values of interest were the Gibbs free energy of ∆ 718K reaction at reaction conditions, Gr , for various radical reactions as well as the molar volume of H2, VBP, used for diffusivity simulations (see Section F.7). The settings used in Materials Studio are presented in Figure F.4 with the calculations performed attempting to minimise the energy of each structure. Following convergence, results were presented as the total energy (Ha) and thermodynamic values from 25 to 1000 K in 25 K increments (simple linear interpolation used to obtain values at reaction temperature) including entropy (S in cal/mol.K), enthalpy (H in kcal/mol) and the Gibbs free energy (G, in kcal/mol). As the settings used for such iterative simulations have a great effect on the outcome, two reactions were simulated to compare the results with experimental data and with simulations by other researchers. These reactions were the formation of CH4 from C and H2 (compared with experimental data) and the opening of a cyclohexyl radical ring (compared with simulations reported in the literature). Methane formation data at 298.15 K is presented below with a sample calculation in Equation F.10. Note that as the carbon is simulated as 4C graphite crystal, the energy of a single carbon atom is 0.25 of the total. The value obtained by the simulation is approximately -81.7 kJ.mol−1 compared to an experimental -74.9 kJ.mol−1 [158]. A reasonable comparison accurate to within 9%. Similar simulations and calculations for open- ∆ 298K −1 ing of a cyclohexyl radical ring yielded Hr = 21.4 kcal.mol compared to those reported in the literature of 21.2 kJ.mol−1 [159]. It may thus be concluded that the results from the Materials Studio simulations are suitable for comparison. Tables F.5 through ?? provide a full listing of the various reactions discussed in Section 5.1.2.

407 Total energy of C (graphite) (Ha) = EC.Total = −152.469

Total energy of H2 (Ha) = EH2.Total = −1.16406

Total energy of CH4 (Ha) = ECH4.Total = −40.4926

Enthalpy correct for C (kcal/mol) = HcC = 12.828

Enthalpy correct for H2 (kcal/mol) = HcH2 = 8.361

Enthalpy correct for CH4 (kcal/mol) = HcCH4 = 29.943

Ha to kcal conversion ((kcal/mol)/Ha) = ConvHa.kcal = 627.51

kcal to kJ conversion (kJ/kcal) = Convkcal.kJ = 4.1868

HC = EC.Total ×ConvHa.kcal + HcC = −152.469 × 627.51 + 12.828 = −95663.2

HH2 = −722.098

HCH4 = −25379.5 (F.10) ∆ 298K Hr = HCH4 − (0.25 × HC + 2 × HH2) = −25379.5 − (−23915.8 − 1444.2) = −19.502

−1 = −19.502Convkcal.kJ = −81.65 kJ.mol

F.5 Phase Density and Composition Simulations Phase composition and density simulations were conducted in AspenTech Aspen Plus (v7.3) using the flowsheet shown in Figure F.5 and the Peng-Robinson equation of state. Initial stream flow rates were set to be equivalent to the appropriate batch amount on an hourly basis. For instance, to simulate the stirred batch reactor with no conversion, the flow rate of the “DPM” stream (comprising pure DPM) was set −1 to 80 g.h at room conditions and the “HYDROGEN” stream (pure H2) set to 170 cm3.h−1 at 445◦C and 13.8 MPa. The “FLASHER” block was set to operating conditions of 445◦C and 13.8 MPa. DPM conversion was simulated by altering

408 the ﬂow rates of the “DPM” and “BT” streams (the latter containing an equimolar mixture of benzene and toluene) using a sensitivity analysis, the results (shown in Figure 5.4) being presented in Tables F.8 through F.12. Table of results

F.6 Physical Property Simulations The changes to the physical properties of the reaction liquid with changes in temperature, pressure and composition were determined by simulation in AspenTech Aspen Plus (v7.3) using the ﬂowsheet shown in Figure F.5 and the Peng-Robinson equation of state. The properties of interest were the viscosity, density and surface tension of the mixture, reported in Aspen Plus as MUMX, RHOMX and SIG- MAMX respectively. The sensitivity analysis described in Section F.5 was used with the addition of FLASHER temperature and pressure variations with the aforementioned properties reported in the results.

F.7 Hydrogen Solubility and Diffusivity

F.7.1 Hydrogen Solubility Simulations

Simulation of H2 solubility in the reaction liquid with changes in temperature, pressure and composition were performed in AspenTech Aspen Plus (v7.3) using the

ﬂowsheet shown in Figure F.5 and the Peng-Robinson equation of state. The H2 solubility was determined as the composition (by mass and moles) in the LIQUID stream exiting the FLASHER block. Sensitivity analyses described in Section F.5 and Section F.6 were used to determine the change in solubility with conditions, the results being shown in Figures F.6 and F.7 (data in Tables F.8 through F.12). For solubility in decalin, a mixture of 50-50 (wt%) cis- and trans-decalin was substituted for DPM in the appropriate stream. It should be noted that the Peng-Robinson equation-of-state property method used in these simulation is not accurate for the estimation of H2 solubilities in hydrocarbons. The results obtained from these limited-range simulations do, however, correspond reasonably well with experimental literature values [113, 114, 121, 122], particularly for DPM at reaction conditions, and are thus considered suitable for illustrative purposes in this study.

409 Examples are given in Tables F.13, F.14 and F.15 for DPM, toluene and benzene respectively wherein linear interpolation and extrapolation of reported values was used to calculate the comparable numbers shown.

410 Table F.2: Conversion results for diphenylmethane hydroconversion in the stirred batch reactor at 600 ppm Mo, 445◦C, 13.8 MPa H2, 0 - 6 h, 700 RPM with calculated values for kinetic analyses. 1 Sample code Reaction time, t DPM conversion, X Zeroed conversion, X’ ln(1 − X′/100) 1 − X′/100 (h) (wt%) (wt%)

411 DPM-BC600-0-46 0 3.36 0.00 0.000 1.000 DPM-BC600-26 1 21.69 18.33 -0.202 1.224 DPM-BC600-1-40 1 17.44 14.08 -0.152 1.164 DPM-BC600-1-42 1 20.32 16.96 -0.186 1.204 DPM-BC600-1-43 1 17.70 14.33 -0.155 1.167 DPM-BC600-32 6 47.21 43.85 -0.577 1.781 DPM-BC600-33 6 42.70 39.34 -0.500 1.648 DPM-BC600-6-47 6 38.19 34.83 -0.428 1.534 Table F.3: Gibbs free energies of reaction for H* abstraction from and stabilisation of diphenylmethane simulated in Accelrys Materials Studio (v4.4) using DMol3 [110, 111] geometry optimisation and frequency calculations.

∆ 718K Reaction code Reaction mechanism Gr −1 (kJ.molreactant)

A + CH3 + CH4 -133.4

B + CH3 + CH4 -3.4

C -59.2

D + CH3 -159.3

E + CH3 -294.2

F + CH3 + CH4 -376.7

Abstraction by *CH3 chosen arbitrarily and is used for comparative purposes only.

412 Table F.4: Gibbs free energies of reaction for H* abstraction from and stabilisation of cyclohexylmethylbenzene simulated in Accelrys Materials Studio (v4.4) using DMol3 [110, 111] geometry optimisation and frequency calculations.

∆ 718K Reaction code Reaction mechanism Gr −1 (kJ.molreactant)

G + CH3 + CH4 -76.4

H + CH3 + CH4 -39.8

I + CH3 + CH4 11.1

J 20.8

K + CH3 + CH4 -369.6

Abstraction by *CH3 chosen arbitrarily and is used for comparative purposes only.

413 414

Figure F.4: Settings used in Accelrys Materials Studio (v4.4) simulations. Table F.5: Gibbs free energies of reaction for hydrogen abstraction or radical addition at various locations on diphenylmethane and cyclohexylmethylbenzene molecules as simulated in Accelrys Materials Studio (v4.4).

∆ 718K Number Reaction mechanism Gr −1 (kJ.molreactant)

1 + CH3 + CH4 0.8

2 + CH3 + CH4 -0.9

3 + CH3 + CH4 -2.4

4 + CH3 + CH4 -2.9

5 + CH3 + CH4 -3.4

6 + CH3 + CH4 -133.4

7 + CH3 -159.3

8 + CH3 -294.2

9 + CH3 + CH4 -116.8

10 + CH3 -137.7

415 Table F.6: Gibbs free energies of reaction for hydrogen abstraction or radical addition at various locations on diphenylmethane and cyclohexylmethylbenzene molecules as simulated in Accelrys Materials Studio (v4.4), continued.

∆ 718K Number Reaction mechanism Gr −1 (kJ.molreactant)

11 + CH3 + CH4 20.9

12 + CH3 -309.3

13 -59.2

14 + CH3 + CH4 -376.7

15 + CH3 + CH4 -45.0

16 + CH3 + CH4 -45.3

17 + CH3 + CH4 -19.9

18 + CH3 + CH4 -48.7

19 + CH3 + CH4 -39.8

416 Table F.7: Gibbs free energies of reaction for hydrogen abstraction or radical addition at various locations on diphenylmethane and cyclohexylmethylbenzene molecules as simulated in Accelrys Materials Studio (v4.4), continued.

∆ 718K Number Reaction Gr −1 (kJ.molreactant)

20 + CH3 + CH4 -68.1

21 + CH3 + CH4 -76.4

22 + CH3 + CH4 15.5

23 + CH3 + CH4 10.9

24 + CH3 + CH4 19.2

25 + CH3 + CH4 28.4

26 + CH3 + CH4 11.1

27 + CH3 -196.0

28 20.8

29 + CH3 + CH4 -369.6

417 Figure F.5: AspenTech Aspen Plus (v7.3) ﬂowsheet for phase composition and density simulation of diphenylmethane hydroconversion at 445◦C, 13.8 MPa H2.

418 30

13.8 MPa

13.1 MPa

12.4 MPa

Supercritical

boundary Hydrogencontent (mol%)

13.8 MPa

13.1 MPa

0.4 12.4 MPa

0.2 Hydrogencontent (wt%)

0.0

0 10 20 30 40

DPM conversion (wt%)

Figure F.6: Results for hydrogen solubility in the liquid phase during diphenylmethane hydroconversion (assuming an equimolar benzene/- toluene product) at 445◦C and 12.4 - 13.8 MPa as simulated in Aspen- Tech Aspen Plus (v7.3).

419 30

445°C

200°C

20°C 10 Hydrogencontent (mol%)

Supercritical

boundary

0.4

445°C

200°C

0.2

20°C Hydrogencontent (wt%)

0.0

0 10 20 30 40

DPM conversion (wt%)

Figure F.7: Results for hydrogen solubility in the liquid phase during diphenylmethane hydroconversion (assuming an equimolar benzene/- toluene product) at 20 - 445◦C and 13.8 MPa as simulated in AspenTech Aspen Plus (v7.3).

420 Table F.8: Results for hydrogen solubility, phase density and species separation between phases during diphenylmethane hydroconversion (assuming an equimolar benzene/toluene product) at 445◦C and 12.4 MPa as simulated in AspenTech Aspen Plus (v7.3).

Pressure DPM conversion Hydrogen solubility in liquid Phase density (g/cm3) Proportion of total species in vapour (%) (MPa) (wt%) (wt%) (mol%) Vapour Liquid Benzene Toluene DPM 0.0 0.27 18.17 0.076 0.414 - - 27.5 2.5 0.27 18.28 0.082 0.394 57.3 52.6 29.0 4.9 0.28 18.40 0.087 0.373 58.1 53.6 30.5 7.4 0.28 18.53 0.093 0.356 59.0 54.6 32.2 9.9 0.29 18.67 0.099 0.342 59.8 55.7 33.9

421 12.4 0.30 18.83 0.105 0.329 60.8 56.8 35.7 14.8 0.30 18.99 0.111 0.319 61.8 58.0 37.6 17.3 0.31 19.18 0.117 0.309 62.9 59.3 39.7 19.8 0.32 19.38 0.123 0.301 64.0 60.6 41.8 12.4 22.2 0.33 19.60 0.129 0.294 65.3 62.1 44.2 24.7 0.34 19.84 0.135 0.288 66.6 63.6 46.7 27.2 0.35 20.10 0.142 0.283 68.1 65.4 49.5 29.6 0.35 20.38 0.148 0.279 69.8 67.3 52.6 32.1 0.37 20.70 0.155 0.275 71.8 69.5 56.0 34.6 0.38 21.06 0.163 0.272 74.0 71.9 59.9 37.1 0.39 21.47 0.170 0.270 76.6 74.8 64.3 39.5 0.41 21.93 0.179 0.268 79.8 78.3 69.6 42.0 0.42 22.49 0.188 0.267 83.8 82.7 76.0 Table F.9: Results for hydrogen solubility, phase density and species separation between phases during diphenylmethane hydroconversion (assuming an equimolar benzene/toluene product) at 445◦C and 13.1 MPa as simulated in AspenTech Aspen Plus (v7.3).

Pressure DPM conversion Hydrogen solubility in liquid Phase density (g/cm3) Proportion of total species in vapour (%) (MPa) (wt%) (wt%) (mol%) Vapour Liquid Benzene Toluene DPM 0.0 0.28 19.02 0.076 0.404 - - 25.4 2.5 0.29 19.16 0.082 0.382 54.6 49.9 26.7 4.9 0.29 19.30 0.088 0.363 55.3 50.8 28.2 7.4 0.30 19.45 0.093 0.346 56.1 51.7 29.6 9.9 0.31 19.62 0.099 0.332 56.9 52.6 31.2

422 12.4 0.32 19.81 0.105 0.320 57.7 53.6 32.8 14.8 0.32 20.00 0.111 0.310 58.6 54.7 34.6 17.3 0.33 20.22 0.117 0.302 59.5 55.8 36.4 19.8 0.34 20.45 0.124 0.294 60.5 57.0 38.4 13.1 22.2 0.35 20.71 0.130 0.288 61.6 58.3 40.5 24.7 0.36 20.98 0.137 0.282 62.8 59.7 42.7 27.2 0.37 21.28 0.144 0.278 64.1 61.3 45.3 29.6 0.38 21.61 0.151 0.274 65.6 63.0 48.1 32.1 0.40 21.99 0.158 0.271 67.3 64.9 51.1 34.6 0.41 22.41 0.166 0.268 69.3 67.1 54.6 37.1 0.43 22.89 0.174 0.267 71.6 69.8 58.7 39.5 0.44 23.44 0.183 0.266 74.6 73.0 63.6 42.0 0.47 24.10 0.192 0.265 78.4 77.2 69.6 Table F.10: Results for hydrogen solubility, phase density and species separation between phases during diphenylmethane hydroconversion (assuming an equimolar benzene/toluene product) at 20◦C and 13.8 MPa as simulated in AspenTech Aspen Plus (v7.3).

Temperature DPM conversion Hydrogen solubility in liquid Phase density (g/cm3) Proportion of total species in vapour (%) (◦C) (wt%) (wt%) (mol%) Vapour Liquid Benzene Toluene DPM 0.0 0.04 3.42 0.011 0.983 - - 0.00 2.5 0.04 3.44 0.011 0.976 0.14 0.06 0.00 4.9 0.04 3.46 0.011 0.969 0.13 0.05 0.00 7.4 0.05 3.48 0.011 0.963 0.13 0.05 0.00 9.9 0.05 3.49 0.011 0.958 0.12 0.05 0.00

423 12.4 0.05 3.51 0.011 0.953 0.12 0.05 0.00 14.8 0.05 3.53 0.011 0.948 0.11 0.05 0.00 17.3 0.05 3.54 0.011 0.944 0.11 0.04 0.00 19.8 0.05 3.56 0.011 0.941 0.10 0.04 0.00 20 22.2 0.05 3.57 0.011 0.937 0.10 0.04 0.00 24.7 0.05 3.58 0.011 0.934 0.10 0.04 0.00 27.2 0.05 3.60 0.011 0.931 0.09 0.04 0.00 29.6 0.05 3.61 0.011 0.928 0.09 0.04 0.00 32.1 0.06 3.62 0.011 0.925 0.09 0.04 0.00 34.6 0.06 3.63 0.011 0.923 0.08 0.03 0.00 37.1 0.06 3.64 0.011 0.920 0.08 0.03 0.00 39.5 0.06 3.66 0.011 0.918 0.08 0.03 0.00 42.0 0.06 3.67 0.011 0.916 0.08 0.03 0.00 Table F.11: Results for hydrogen solubility, phase density and species separation between phases during diphenylmethane hydroconversion (assuming an equimolar benzene/toluene product) at 200◦C and 13.8 MPa as simulated in AspenTech Aspen Plus (v7.3).

Temperature DPM conversion Hydrogen solubility in liquid Phase density (g/cm3) Proportion of total species in vapour (%) (◦C) (wt%) (wt%) (mol%) Vapour Liquid Benzene Toluene DPM 0.0 0.09 6.74 0.008 0.818 - - 0.19 2.5 0.09 6.77 0.009 0.811 9.3 5.5 0.19 4.9 0.09 6.80 0.009 0.804 9.0 5.4 0.18 7.4 0.09 6.83 0.010 0.799 8.7 5.2 0.18 9.9 0.10 6.86 0.011 0.793 8.5 5.1 0.18

424 12.4 0.10 6.88 0.012 0.788 8.2 4.9 0.17 14.8 0.10 6.91 0.012 0.783 8.0 4.8 0.17 17.3 0.10 6.94 0.013 0.779 7.8 4.7 0.17 19.8 0.10 6.97 0.013 0.775 7.6 4.5 0.16 200 22.2 0.11 6.99 0.014 0.771 7.4 4.4 0.16 24.7 0.11 7.02 0.015 0.767 7.2 4.3 0.16 27.2 0.11 7.04 0.015 0.764 7.0 4.2 0.15 29.6 0.11 7.07 0.016 0.760 6.8 4.1 0.15 32.1 0.11 7.09 0.016 0.757 6.7 4.0 0.15 34.6 0.11 7.12 0.016 0.754 6.5 3.9 0.15 37.1 0.12 7.14 0.017 0.751 6.3 3.8 0.14 39.5 0.12 7.17 0.017 0.748 6.2 3.7 0.14 42.0 0.12 7.19 0.018 0.746 6.1 3.6 0.14 Table F.12: Results for hydrogen solubility, phase density and species separation between phases during diphenylmethane hydroconversion (assuming an equimolar benzene/toluene product) at 445◦C and 13.8 MPa as simulated in AspenTech Aspen Plus (v7.3).

Temperature DPM conversion Hydrogen solubility in liquid Phase density (g/cm3) Proportion of total species in vapour (%) (◦C) (wt%) (wt%) (mol%) Vapour Liquid Benzene Toluene DPM 0.0 0.30 19.85 0.076 0.394 - - 23.6 2.5 0.30 20.01 0.082 0.372 52.1 47.3 24.8 4.9 0.31 20.17 0.088 0.353 52.8 48.1 26.1 7.4 0.32 20.35 0.094 0.337 53.4 48.9 27.4 9.9 0.33 20.55 0.100 0.324 54.0 49.8 28.8

425 12.4 0.34 20.75 0.106 0.313 54.8 50.7 30.3 14.8 0.35 20.98 0.112 0.303 55.5 51.6 31.8 17.3 0.36 21.22 0.118 0.295 56.3 52.6 33.5 19.8 0.37 21.49 0.125 0.288 57.1 53.6 35.2 445 22.2 0.38 21.78 0.131 0.282 58.1 54.7 37.1 24.7 0.39 22.10 0.138 0.277 59.1 55.9 39.2 27.2 0.40 22.43 0.145 0.273 60.2 57.3 41.5 29.6 0.41 22.81 0.153 0.270 61.5 58.8 43.9 32.1 0.43 23.24 0.160 0.268 63.0 60.5 46.7 34.6 0.44 23.72 0.168 0.266 64.6 62.4 49.8 37.1 0.46 24.27 0.177 0.265 66.7 64.7 53.4 39.5 0.48 24.91 0.187 0.264 69.4 67.7 57.9 42.0 0.51 25.68 0.197 0.264 72.9 71.5 63.5 Table F.13: Literature values for hydrogen solubility in diphenylmethane [113] for comparison with Aspen Plus simulations.

Source Temperature Pressure xH2 (◦C) (atm) (MPa) (psig) 189.6 100 10.1 1470 0.0591 150 15.2 2204 0.083 268.7 100 10.1 1470 0.0785 150 15.2 2204 0.1105 Reported 348.6 100 10.1 1470 0.0985 150 15.2 2204 0.145 428.5 100 10.1 1470 0.138 150 15.2 2204 0.203 189.6 13.8 2000 0.076 268.7 13.8 2000 0.102 P interp. 348.6 13.8 2000 0.132 428.5 13.8 2000 0.185 T interp. 200 13.8 2000 0.080 T extrap. 445 13.8 2000 0.196

Table F.14: Literature values for hydrogen solubility in toluene [114] for comparison with Aspen Plus simulations.

Source Temperature Pressure xH2 (◦C) (atm) (MPa) (psig) 188.7 99.8 10.1 1467 0.0704 149.9 15.2 2203 0.1023 Reported 229 99.3 10.1 1459 0.0812 150.2 15.2 2207 0.1227 188.7 13.8 2000 0.0935 P interp. 229 13.8 2000 0.1112 T interp. 200 13.8 2000 0.0985

Table F.15: Literature values for hydrogen solubility in benzene [115] for comparison with Aspen Plus simulations.

Source Temperature Pressure xH2 (◦C) (atm) (MPa) (psig) 200 75.8 7.7 1114 0.0704 Reported 134 13.6 1969 0.1023 P extrap. 200 13.8 2000 0.1353

426 F.7.2 Hydrogen Dissolution Rate

To determine to rate at which H2 is able to dissolve into the reaction liquid, experiments were conducted whereby the pressure of the micro-reactor, loaded with DPM and at 20◦C, was rapidly increased and monitored with time (the pressure being sampled to within 7 kPa [1 psi] every 1 s by the pressure transducer). The results are shown in Figure 5.9 and tabulated in Table F.16. This data shows that, following pressurisation to 12.76 MPa, the gas pressure drops by 82.74 kPa within 40 s. Equation F.11 shows that, assuming ideal gas behaviour, this corresponds to ′ a H2 concentration, CH2, of approximately 14.0 mol% in the liquid.

Table F.16: Results for pressure monitoring of a diphenylmethane-loaded glass insert micro-reactor with time at 20◦C with 150 µL liquid load.

Time Pressure (s) (MPa) 1 12.755 16 12.693 31 12.679 46 12.673 2 12.755 17 12.693 32 12.679 47 12.673 3 12.728 18 12.693 33 12.679 48 12.673 4 12.728 19 12.686 34 12.679 49 12.673 5 12.721 20 12.686 35 12.679 50 12.673 6 12.721 21 12.686 36 12.679 51 12.673 7 12.700 22 12.686 37 12.673 52 12.673 8 12.700 23 12.686 38 12.673 53 12.673 9 12.700 24 12.686 39 12.673 54 12.673 10 12.700 25 12.686 40 12.673 55 12.673 11 12.700 26 12.686 41 12.673 56 12.673 12 12.700 27 12.679 42 12.673 57 12.673 13 12.693 28 12.679 43 12.673 58 12.673 14 12.693 29 12.679 44 12.673 15 12.693 30 12.679 45 12.673

Pressure change (Pa) = ∆PGTotal = 0.400 3 −6 Total gas volume (m ) = VGTotal = 0.402 × 10 Ideal gas constant (J.mol−1.K−1) = R = 8.314 Temperature (K) = T = 293.15

DPM load (mmol) = nDPM = 0.879

427 PGTotal.VGTotal = nGTotal.R.T ∆P .V ∆n = GTotal GTotal GTotal R.T 82740 × 0.402 × 10−6 = 8.314 × 293.15 = 0.143 mmol ∆ ′ nGTotal (F.11) CH2 = nLTotal ∆n = GTotal ∆nGTotal + nDPM 0.143 = 0.143 + 0.879 = 0.140 = 14.0%

F.7.3 Hydrogen Diffusion

With H2 diffusion through the liquid suspected to be a factor limiting conversion in the unmixed micro-reactors, information from both Aspen Plus (per above) and Accelrys Materials Studio (as in Section F.4) simulations were used to calculate the hydrogen concentration through the liquid with time. This was achieved by ′ solving Fick’s second law of diffusion shown in Equation F.12 where CH2 is the −3 concentration of H2 in mol.m , t is the time in s, DH2.DPM is the diffusion coeffi- 2 −1 cient of H2 in DPM at reaction conditions in m .s and x is the distance from the gas-liquid interface in m. A time period 0 - 3600 s (1 h) was used together with the solubility of H2 in DPM determined in Aspen Plus as the surface boundary condition concentration. Assuming the rate of H2 dissolution to be significantly faster than the diffusion of H2 through the liquid (an assumption validated by the simulation), the surface concentration of H2 was kept constant. ◦ With the diffusion coefficient of H2 in DPM at 445 C and 13.8 MPa not available in the literature, this value was estimated using the modified Wilke-Chang correlation [123, 160], Equation F.13. Herein Di in j is the diffusivity of species i in species j in m2.s−1, η is the Wilke-Chang parameter (1.173 × 10−16, [123]), φ ∗ is the association parameter of the solvent (1.0 for benzene, ether, heptane and “other

428 −1 unassociated solvents”, [160]), Mr j is the molar mass of the solvent in kg.kmol , −1 −1 T is the temperature in K, µ j is the solvent dynamic viscosity in kg.m .s (from

Aspen Plus simulations) and VBP.i is the molar volume of the solute at its boiling point in m3.kmol−1 (from Materials Studio simulations).

δC′ δ 2C′ H2 = D H2 (F.12) δt H2.DPM δx2

T D = η × (φ ∗.Mr )0.5 × (F.13) i in j j µ 0.6 j.VBP.i Using this information, modiﬁed MATLAB code for diffusion-reaction systems by Iber [161] was implemented as shown in Listings F.3 through F.6.

Listing F.3: Main MATLAB (R2012a) code used for simulations of hydrogen diffusion in diphenylmethane at 445◦C and 13.8 MPa.

% Ross Kukard − Department of Chemical and Biological Engineering , UBC % November/December 2013

% 1D Hydrogen diffusivity in diphenylmethane. % Examining the effect of liquid volume (and hence depth).

% This script simulates the diffusion of hydrogen through diphenymethane % down through the liquid of an unstirred microreactor. % The diffusion coefficient is calculated using the Wilke−Change % correlation and van der Waals molecular volumes from Accelrys Materials % Studio calculations.

% Initially there is no hydrogen in the DPM. % The top boundary condition is defined as the solubility of hydrogen i n %DPM at reaction conditions from Aspen simulations. % The volumes of fluid of interest are 150 uL and 400 uL. % The microreactor glass insert has an internal diameter of x mm.

% Using Wilke−Chang equation from Geankoplis, C. J. − Transport Processes and Unit Operations, pg 401

% We start by clearing the workspace and command window clc clear close all

% Enter some of the constants DPM volume = 150; % uL DPM volume = DPM volume / 10ˆ9; % mˆ3 insert ID = 3.84; % mm insert ID = insert ID / 10ˆ3; % m depth = (4 * DPM volume) / ( pi * insert IDˆ2); %m volume = pi.rˆ2.h = pi.(Dˆ2/4).h =>

429 h = 4.volume/(pi.Dˆ2)

% Values for Wilke−Chang eta = 1.173*10ˆ( −16); % Modified Wilke−Chang parameter phi = 1.0; % Association parameter of the solvent. 1.0 for benzene, ether , heptane and other unassociated solvents Mwj = 168.234; % Solvent (DPM) molar mass ... kg/kmol = g/mol T = 718.15; % Temperature in K muj = 0.128; % Solvent viscosity at T ... this is in cP but want kg/m.s muj = muj/100 * 10; %1 P = 100 cP ... 1 P = 1/10 Pa.s ... 1 Pa.s = 1 kg/m.s ... so this is now in kg/m.s VBPi = 14.3; % Molar volume of solute at it ’s normal boiling point ... cmˆ3 /mol ( Value from Wilke and Chang, 1955, pg 268) VBPi = VBPi * 1000 / 10ˆ6; % cmˆ3/mol to m3/kmol

% Calculate diffusion coefficient D i j = eta * (phi *Mwj) ˆ(0.5) * T / (muj * VBPiˆ0.6) ; % Diffusion coefficient of hydrogen in DPM ... mˆ2/s

% Now we can start setting up our PDEs % Fick’s second law simply states that dc/dt = D.dˆ2c/dxˆ2 % Code adapted from Dagmar Iber − Numerical Solution of reaction −d i f f u s i o n problems P(1) = Dij; % Diffusion coefficient P(2) = 1.37632; % c0 ... concentration of Hydrogen at surface ... assumed constant and equal to solubility of 1376.32 mol/mˆ3 = 0.00137632 mol/cm3 L = depth; % Domain length (m) maxt = 3600; %Time length (s) ... 60 = 1 min. 3600 = 1 h m= 0; % Symmetry parameter (see pdepe help for more) corresponding to slab % Create a mesh of the distance and time domains t = linspace (0,maxt,100) ; x = linspace (0,L,100);

% Now we can call pdepe which needs the following: % m % PDEfun − function containing PDEs % ICfun − function containing initial conditions for t = 0 for all x % BCfun − function containing boundary conditions for x = 0 and x = L % xmesh.tspan.(number of variables) sol = pdepe (m,@PDEfun,@ICfun,@BCfun,x,t ,[] ,P) ; u = sol;

% Plotting the results % 3D surface plot figure ( 1 ) surf (x,t ,u, ’edgecolor ’, ’none’); xlabel ( ’ Distance (m) ’) ylabel ( ’ Time (s) ’) zlabel ( ’ H 2 Concentration i n DPM (kmol/mˆ3) ’) axis ([0 L 0 maxt 0 P(2)]) set ( gcf () , ’Renderer’ , ’painters ’) view (45,30) % Set view azimuth and elevation set ( gca , ’FontSize ’ ,12, ’fontWeight ’ , ’bold ’) set (findall( gcf , ’type’, ’text ’), ’FontSize’,12,’fontWeight ’, ’bold’)

% 2D line plot

430 figure ( 2 ) hold all for n = linspace (1 , length (t) ,10) plot (x,sol(n,:) ,’linewidth ’ ,2) end xlabel ( ’ Distance (m) ’) ylabel ( ’ H 2 Concentration i n DPM (kmol/mˆ3) ’) axis ([0 L 0 0.1]) % axis([0 L 0 P(2)]) set ( gca , ’FontSize ’ ,12, ’fontWeight ’ , ’bold ’) set (findall( gcf , ’type’, ’text ’), ’FontSize’,12,’fontWeight ’, ’bold’) grid

% For the output, x is a row vector of distances (m) from surface of l i q u i d % sol is a matrix with each row representing one time step (in 100 % increments of maxt) and each column % represnting the distance (m) from the surface

Listing F.4: MATLAB (R2012a) code containing Fick’s second law partial differential equation used in conjunction with Run me.m for simulations of hydrogen diffusion in diphenylmethane.

% Ross Kukard − Department of Chemical and Biological Engineering , UBC % November/December 2013 % 1D Hydrogen diffusivity in diphenylmethane. % Examining the effect of liquid volume (and hence depth).

% Function containing PDEs

function [c, f, s] = PDEfun(x,t,u,dudx,P)

D = P(1);

c = 1; % c is the left side coefficient f = D. * dudx; % f is the flux term, i.e. diffusion. Here simply Fick’s second law s = 0; % for example −2*u; % s is the source term, so where the reaction gets added .

Listing F.5: MATLAB (R2012a) code containing boundary conditions used in conjunction with Run me.m for simulations of hydrogen diffusion in diphenylmethane.

% Function containing boundary conditions (BC)

function [pl, ql, pr, qr ] = BCfun(xl, ul, xr, ur, t, P)

431 c0 = P(2);

% Our boundary conditions are that there is no flux at the bottom (right) % boundary and that the concentration at the left (top) boundary is % constant pl = ul−c0 ; ql = 0; pr = 0; qr = 1;

Listing F.6: MATLAB (R2012a) code containing initial conditions used in conjunction with Run me.m for simulations of hydrogen diffusion in diphenylmethane.

% Function containing initial conditions (IC) function u0 = ICfun(x,P) u0 = 0; %So at time t = 0, the concentration is 0 for all x

The results of this simulation are presented in Figure 5.10 with the corresponding data in Tables F.17 through F.19 (note that for accuracy the simulation was run in 36 s increments but only the 15 min [900 s] data increments plotted are tabulated below).

F.8 Area:Volume Ratios The area:volume ratio of each reactor system was an important parameter during this study as it varied with reactor dimensions. Below diagrams of the stirred batch reactor, inclined and vertical stainless steel micro-reactors and glass insert micro- reactor accompany calculations of their A:V ratios. Beginning with the stirred batch reactor, the dimensions shown in Figure F.8 (the stirrer blade is 1/4” = 0.6 cm wide and 1/16” = 0.2 cm thick) may be used to calculate the total area exposed to the reaction liquid. These calculations are simple geometric areas with the results for each component shown in Table F.20. Knowing the total loaded volume to be 80 cm3, the A:V ratio may thus be calculated as 1.29 cm2/cm3.

432 Table F.17: Results for MATLAB simulations of hydrogen diffusion in diphenylmethane at 445◦C, 13.8 MPa.

Time (s) 0 900 1800 2700 3564 Distance from interface Hydrogen concentration (cm) (kmol/m3) 0.000 1.37632 1.376 1.376 1.376 1.376 0.013 0 1.274 1.304 1.317 1.325 0.026 0 1.173 1.232 1.259 1.274 0.039 0 1.074 1.161 1.200 1.223 0.052 0 0.977 1.091 1.143 1.173 0.065 0 0.884 1.022 1.085 1.122 0.078 0 0.795 0.955 1.029 1.073 0.092 0 0.710 0.889 0.974 1.024 0.105 0 0.630 0.825 0.919 0.976 0.118 0 0.555 0.763 0.866 0.928 0.131 0 0.486 0.704 0.815 0.882 0.144 0 0.423 0.647 0.764 0.836 0.157 0 0.365 0.593 0.716 0.792 0.170 0 0.313 0.541 0.669 0.749 0.183 0 0.267 0.493 0.623 0.707 0.196 0 0.225 0.447 0.580 0.666 0.209 0 0.189 0.404 0.538 0.627 0.222 0 0.158 0.364 0.498 0.588 0.235 0 0.131 0.327 0.460 0.552 0.249 0 0.107 0.292 0.424 0.516 0.262 0 0.087 0.260 0.390 0.483 0.275 0 0.071 0.231 0.358 0.450 0.288 0 0.057 0.205 0.328 0.419 0.301 0 0.045 0.180 0.299 0.390 0.314 0 0.036 0.158 0.273 0.362 0.327 0 0.028 0.139 0.248 0.335

Figure F.9 presents the dimensions of the vertical stainless steel micro-reactor. If constant liquid volume and perpendicular liquid-wall interfaces are assumed, the angle of inclination has no impact on the wall or thermocouple area exposed to the liquid. Table F.21 shows the contribution of the reactor and thermocouple walls to the total area and, with a liquid loading of 0.40 cm3, the resulting A:V ratio is 20.2

433 Table F.18: Results for MATLAB simulations of hydrogen diffusion in diphenylmethane at 445◦C, 13.8 MPa, continued.

Time (s) 0 900 1800 2700 3564 Distance from interface Hydrogen concentration (cm) (kmol/m3) 0.340 0 0.022 0.121 0.225 0.310 0.353 0 0.017 0.105 0.203 0.286 0.366 0 0.013 0.091 0.184 0.263 0.379 0 0.010 0.078 0.165 0.242 0.392 0 0.008 0.067 0.148 0.222 0.406 0 0.006 0.058 0.133 0.204 0.419 0 0.004 0.049 0.119 0.186 0.432 0 0.003 0.042 0.106 0.170 0.445 0 0.002 0.035 0.094 0.155 0.458 0 0.002 0.030 0.083 0.141 0.471 0 0.001 0.025 0.074 0.128 0.484 0 8.52E-04 0.021 0.065 0.116 0.497 0 6.06E-04 0.017 0.057 0.105 0.510 0 4.27E-04 0.014 0.050 0.095 0.523 0 2.99E-04 0.012 0.044 0.085 0.536 0 2.08E-04 0.010 0.039 0.077 0.549 0 1.43E-04 0.008 0.034 0.069 0.563 0 9.78E-05 0.007 0.029 0.062 0.576 0 6.64E-05 0.005 0.025 0.055 0.589 0 4.47E-05 0.004 0.022 0.049 0.602 0 2.99E-05 0.004 0.019 0.044 0.615 0 1.98E-05 0.003 0.016 0.039 0.628 0 1.31E-05 0.002 0.014 0.035 0.641 0 8.55E-06 0.002 0.012 0.031 0.654 0 5.55E-06 0.001 0.010 0.027 0.667 0 3.58E-06 0.001 0.009 0.024 cm2/cm3. Reducing the liquid volume to 150 µL changes the area contributions as shown in Table F.22, with a resulting A:V ratio of 20.7 cm2/cm3. In the glass-insert micro-reactor, the inner diameter is increased to 0.40 cm, reducing the liquid height as compared to the stainless steel micro-reactors. With the thermocouple now the only active surface exposed, the A:V ratios for 400 and

434 Table F.19: Results for MATLAB simulations of hydrogen diffusion in diphenylmethane at 445◦C, 13.8 MPa, continued.

Time (s) 0 900 1800 2700 3564 Distance from interface Hydrogen concentration (cm) (kmol/m3) 0.680 0 2.29E-06 9.00E-04 0.007 0.021 0.693 0 1.46E-06 7.07E-04 0.006 0.018 0.706 0 9.19E-07 5.53E-04 0.005 0.016 0.720 0 5.76E-07 4.31E-04 0.004 0.014 0.733 0 3.59E-07 3.35E-04 0.004 0.012 0.746 0 2.22E-07 2.59E-04 0.003 0.011 0.759 0 1.37E-07 1.99E-04 0.003 0.009 0.772 0 8.35E-08 1.53E-04 0.002 0.008 0.785 0 5.07E-08 1.17E-04 0.002 0.007 0.798 0 3.06E-08 8.89E-05 0.002 0.006 0.811 0 1.84E-08 6.74E-05 0.001 0.005 0.824 0 1.10E-08 5.09E-05 0.001 0.005 0.837 0 6.50E-09 3.83E-05 8.39E-04 0.004 0.850 0 3.84E-09 2.87E-05 6.89E-04 0.003 0.863 0 2.25E-09 2.14E-05 5.63E-04 0.003 0.877 0 1.31E-09 1.59E-05 4.60E-04 0.002 0.890 0 7.64E-10 1.18E-05 3.74E-04 0.002 0.903 0 4.41E-10 8.72E-06 3.04E-04 0.002 0.916 0 2.54E-10 6.41E-06 2.46E-04 0.002 0.929 0 1.45E-10 4.70E-06 1.98E-04 0.001 0.942 0 8.28E-11 3.43E-06 1.60E-04 0.001 0.955 0 4.70E-11 2.49E-06 1.28E-04 9.16E-04 0.968 0 2.65E-11 1.81E-06 1.03E-04 7.72E-04 0.981 0 1.49E-11 1.30E-06 8.21E-05 6.49E-04 0.994 0 8.36E-12 9.38E-07 6.54E-05 5.44E-04 1.007 0 4.67E-12 6.73E-07 5.20E-05 4.56E-04

150 µL liquid loadings (the exposed areas presented in Tables F.23 and F.24 respectively) are determined to be 4.0 and 4.1 cm2/cm3 respectively.

435 Figure F.8: Dimensions of stirred batch reactor in the context of 80 cm3 liquid loading for determination of exposed surface area. Not to scale. Table F.20: Surface area of stirred batch reactor components exposed to liquid reaction mixture.

Component Surface area (cm2) Reactor bottom 31.7 Reactor walls 50.4 Gas sparger 2.0 Cooling loop 6.0 Stirrer shaft 0.6 Stirrer blade 12.3 Total 103.0

F.8.1 Gas-Liquid Interfacial Area Whilst the angle of inclination of the stainless steel micro-reactor may not have an effect on the A:V ratio, it does affect the gas-liquid interfacial area. As shown in Figure F.10, inclination of the micro-reactor to 30◦ results in an elliptical interface. The area of this elliptical interface is given by A =(0.5 × L) × (0.5 × ID)= 0.146 cm2 whilst that of the circular interface (as for the vertical orientation is simply A = π × ID2 × 0.25 = 0.073 cm2. By comparison, the interface for the glass insert

436 Figure F.9: Dimensions of stainless steel micro-reactor in the context of 400 µL liquid loading for determination of exposed surface area. Not to scale. Table F.21: Surface area of stainless steel micro-reactor wall and thermocouple exposed to 400 µL liquid reaction mixture.

Component Surface area (cm2) Reactor bottom 0.07 Reactor wall 5.25 Thermocouple bottom 0.02 Thermocouple wall 2.75 Total 8.10 micro-reactor with ID = 0.4 cm is 0.126 cm2.

F.9 Coke Solubility It was desired to quantify the extent to which the coke, introduced with the coke- catalyst agglomerates, was able to dissolve in the DPM. This would be an indication of how much was available to react. As coke solubility is a function of tem-

437 Table F.22: Surface area of stainless steel micro-reactor wall and thermocouple exposed to 150 µL liquid reaction mixture.

Component Surface area (cm2) Reactor bottom 0.07 Reactor wall 1.97 Thermocouple bottom 0.02 Thermocouple wall 1.05 Total 3.11

Table F.23: Surface area of thermocouple exposed to 400 µL liquid reaction mixture in glass insert micro-reactor.

Component Surface area (cm2) Thermocouple bottom 0.02 Thermocouple wall 1.61 Total 1.63 perature, the solubility of each of the sample was tested at 20 and 100◦C to get an general indication of this temperature dependence. Ideally the solubility would be measured at the operating temperature of 445◦C, but elevated temperatures would certainly allow thermal cracking to proceed thereby obscuring solubility data. As only a relative comparison between the samples was needed, the lower temperature tests were sufﬁcient. For each test 2.5 wt% of coke-catalyst agglomerate (approximately 38 mg) was added to 1.5 g of DPM at 20◦C. The samples were agitated by shaking vigorously for 30 seconds. In the case of the 20◦C tests, the samples were left on the laboratory counter for 48 h whilst for 100◦C, the samples were placed within a furnace for the same period of time. After the time had elapsed, the liquid was pipetted out

Table F.24: Surface area of thermocouple exposed to 150 µL liquid reaction mixture in glass insert micro-reactor.

Component Surface area (cm2) Thermocouple bottom 0.02 Thermocouple wall 0.62 Total 0.63

438 Figure F.10: Effect of inclination on gas-liquid surface area in stainless steel micro-reactor. Not to scale. of the sample and the sediment dried for 48 h at 105◦C. Comparison of the ﬁnal dried mass with the initial mass added to the DPM was performed to determine the amount of coke which was removed as dissolved species in the liquid.

439