Decalin Dehydrogenation for In-Situ Hydrogen Production To

Home , Decalin

DECALIN DEHYDROGENATION FOR IN-SITU HYDROGEN PRODUCTION TO

INCREASE CATALYTIC CRACKING RATE OF N-DODECANE

Thesis

Submitted to

The School of Engineering of the

UNIVERSITY OF DAYTON

In Partial Fulfillment of the Requirements for

The Degree of

Master of Science in Chemical Engineering

Christopher Bruening

Dayton, Ohio

May, 2018 DECALIN DEHYDROGENATION FOR IN-SITU HYDROGEN PRODUCTION TO

INCREASE CATALYTIC CRACKING RATE OF N-DODECANE

Name: Bruening, Christopher Robbins

APPROVED BY:

Matthew J. DeWitt, Ph.D. Donald K. Phelps, Ph.D. Advisory Committee Chairman Committee Member Distinguished Research Engineer Senior Research Chemist University of Dayton Research Institute Air Force Research Laboratory

Michael Elsass, Ph.D. Kevin Myers, D.Sc., P.E. Committee Member Committee Member Lecturer Professor Department of Chemical and Materials Department of Chemical and Materials Engineering Engineering

Robert J. Wilkens, Ph.D., P.E. Eddy M. Rojas, Ph.D., M.A., P.E. Associate Dean for Research and Innovation Dean, School of Engineering Professor School of Engineering School of Engineering

ABSTRACT

DECALIN DEHYDROGENATION FOR IN-SITU HYDROGEN PRODUCTION TO

INCREASE CATALYTIC CRACKING RATE OF N-DODECANE

Name: Bruening, Christopher Robbins University of Dayton

Advisor: Dr. Matthew J. DeWitt

Catalytic cracking of paraffinic hydrocarbons is a widely utilized industrial process, but catalyst deactivation over time requires regeneration or replacement of the catalyst bed. A gaseous hydrogen co-feed can be used to promote hydrocracking and decrease deactivation of the catalyst due to coke formation or active site poisoning. One potential alternative approach to extend the lifetime of a cracking catalyst is to generate molecular hydrogen in-situ via catalytic dehydrogenation of a cycloparaffin. In this effort, studies were performed using model compounds to investigate the impact of catalyst configuration and operating conditions on overall performance. For the purpose of this testing, decalin was selected as a model cycloparaffin, with n-dodecane used as a model n-paraffin compound. A blended cycloparaffin/n-paraffin feed was studied in a dual catalyst flow reactor system, containing both a dehydrogenation and cracking catalyst. Testing was performed with either the dehydrogenation catalyst upstream or

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with the two catalysts physical mixed. Products and reactant conversion rates from these

studies were compared to those from a baseline n-paraffin cracking study, with no cycloparaffin or dehydrogenation catalyst present.

Several commercially available Zeolite catalysts were initially screened for n-dodecane cracking activity to identify an appropriate cracking catalyst for further study.

A Zeolite Y catalyst provided adequate n-dodecane reactivity for observation of reactor configuration impact. Prior to the dual-bed and mixed-bed studies, a synthesized

Pt/Al2O3 dehydrogenation catalyst was studied independently with neat decalin feed as

well as a blended decalin/n-dodecane feed, for the purpose of determining appropriate

reactor conditions to be used in subsequent testing. Using the selected Zeolite Y catalyst,

extended duration testing was performed at 400°C and 500 psig to characterize the

activity and deactivation rate for n-dodecane cracking, to provide a baseline for subsequent comparison. Similar testing was performed with the Zeolite Y catalyst to

investigate the impact of blending decalin or naphthalene with the normal paraffin.

A dual-bed reactor configuration utilized the Pt/Al2O3 dehydrogenation catalyst

in-line and upstream of the Zeolite Y catalyst, with a 1:1 volumetric blend of

decalin/n-dodecane at 400°C and 500 psig. The dehydrogenation bed successfully

promoted decalin dehydrogenation for generating in-situ molecular hydrogen. However,

minimal initial n-dodecane conversion was observed, which was speculated to be due to

strong adsorption of naphthalene dehydrogenation product onto the Zeolite Y catalyst,

reducing the number of available active sites. After 240 minutes, n-dodecane conversion

in the dual-bed reactor system was higher than the baseline, while decalin conversion was

very high for the entire duration. In a mixed bed configuration, the two catalyst beds

iv were physically mixed with identical testing conditions. This configuration was intended to reduce the absolute naphthalene concentration on the Zeolite Y catalyst and eliminate the inhibition which occurred in the dual bed configuration. The mixed bed configuration promoted higher initial conversion and lower deactivation rate compared to all other feed/reactor configurations. Overall, it was determined that catalytic dehydrogenation of a cycloparaffin can be successfully employed to increase conversion of a normal paraffin in a catalytic cracking reactor, although future work could further optimize catalyst selection/loading, reactor configuration, and reaction conditions.

ACKNOWLEDGEMENTS

I would like to extend thanks to my research advisor, Dr. Matthew DeWitt. His

advice and support during the thesis project helped to make this document possible. I

would also like to extend gratitude to the other members of my thesis committee: Dr.

Donald Phelps, Dr. Michael Elsass, and Dr. Kevin Myers, for reviewing my thesis and

providing guidance.

I would also like to acknowledge the support of University of Dayton Research

Institute employees in assisting me with the project. I would like to thank Richard

Striebich and Linda Schafer, who assisted in product analysis; as well as Jhoanna Alger

and David Gasper, who assisted in experimental testing. Finally, I would like to thank the Air Force Research Laboratory for funding and support.

TABLE OF CONTENTS

ABSTRACT ...... iii ACKNOWLEDGEMENTS ...... vi LIST OF FIGURES ...... ix LIST OF TABLES ...... xii CHAPTER 1 INTRODUCTION AND BACKGROUND ...... 1 1.1 Introduction ...... 1 1.2 Decalin as a Hydrogen Carrier ...... 2 1.3 Normal and Isomerized Paraffin Cracking Mechanisms utilizing Zeolite Catalysts ...... 7 1.4 Hydrocracking, Ring Opening, and Alkylation of Cycloparaffinic and Aromatic Species on Zeolite Catalysts...... 13 1.5 Presentation of Research ...... 20 CHAPTER 2 MATERIALS AND METHODS ...... 22 2.1 Reactor Setup ...... 22 2.2 Analysis Instrumentation ...... 26 2.3 Catalyst Synthesis ...... 27 CHAPTER 3 ZEOLITE ACTIVITY SCREENING...... 30 CHAPTER 4 DEHYDROGENATION OF DECALIN ...... 34 4.1 Experimental Conditions of Study Segments ...... 34 4.2 Experimental Results...... 36 4.3 Primary Conclusions from Decalin Dehydrogenation Study ...... 42 CHAPTER 5 DEHYDROGENATION OF DECALIN BLENDED WITH N- DODECANE...... 44 5.1 Experimental Conditions of Study Segments ...... 44 5.2 Experimental Results...... 45

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5.3 Primary Conclusions from Decalin/n-Dodecane Dehydrogenation Study ...... 47 CHAPTER 6 DETERMINATION OF BASELINE DEACTIVATION RATE FOR N-DODECANE CRACKING OVER ZEOLITE Y ...... 49 6.1 Experimental Conditions of Study Segments ...... 49 6.2 Experimental Results...... 52 6.3 Overall Conclusions from n-dodecane Cracking Study ...... 57 CHAPTER 7 CRACKING OF N-DODECANE AND DECALIN BLENDED FEED OVER ZEOLITE Y ...... 59 7.1 Experimental Conditions of Study Segments ...... 59 7.2 Experimental Results...... 60 7.3 Primary Conclusions from n-Dodecane/Decalin Cracking Study...... 68 CHAPTER 8 CRACKING OF N-DODECANE AND NAPHTHALENE BLENDED FEED OVER ZEOLITE Y ...... 70 8.1 Experimental Conditions of Study Segments ...... 70 8.2 Results and Conclusions ...... 72 8.3 Primary Conclusions from n-Dodecane/Naphthalene Cracking Study ...... 77 CHAPTER 9 DEHYDROGENATION/CRACKING OF N-DODECANE AND DECALIN BLENDED FEED IN DUAL-BED REACTOR ...... 79 9.1 Experimental Conditions of Study Segments ...... 79 9.2 Experimental Results...... 80 9.3 Primary Conclusions from n-Dodecane/Decalin Dual-Bed Study ...... 89 CHAPTER 10 DEHYDROGENATION/CRACKING OF N-DODECANE AND DECALIN BLENDED FEED IN MIXED-BED REACTOR ...... 90 10.1 Experimental Conditions of Study Segments...... 91 10.2 Experimental Results...... 91 10.3 Primary Conclusions from n-Dodecane/Decalin Mixed-Bed Study ...... 99 CHAPTER 11 SUMMARY, CONCLUSIONS, AND FUTURE WORK ...... 101 REFERENCES ...... 110 APPENDIX A Sample Calculations ...... 115

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LIST OF FIGURES

Figure 1.2.1: Decalin (left) dehydrogenation to tetralin (center), followed by dehydrogenation to naphthalene (right)...... 4 Figure 1.3.1: Examples of β–scission reaction...... 9 Figure 1.3.2: Examples of Haag-Dessau mechanism...... 10 Figure 1.4.1: Ring contraction from decalin (left) to methyl-hydrindane (center) followed by ring opening to butyl-cyclohexane (right)...... 15

Figure 1.4.2: Non-branched and single-branched C10 products obtained from single bond cleavage in selective ring opening of butyl-cyclohexane...... 16 Figure 2.1.1: Diagram of reactor system configuration used for experimentation...... 23 Figure 2.1.2: Photo of packed bed reactor system furnace and feed gas mass flow controllers...... 24 Figure 2.1.3: Diagram of flow through hot or cold product trap...... 26

Figure 2.3.1: Photo of synthesized Pt/Al2O3 catalyst after final calcination...... 28 Figure 3.1.1: Conversion of n-dodecane with varying type of Zeolite and hydrogen/n-dodecane molar ratio at 350°C, 500 psig, and LHSV of 3 h-1...... 33 Figure 4.2.1: Conversion of decalin and selectivity to tetralin and naphthalene -1 over Pt/Al2O3 as function of temperature at LHSV of 1.5 and 0.75 h and 500 psig...... 37

Figure 4.2.2: Conversion of decalin and selectivity of products over Pt/Al2O3 as function of LHSV at temperature of 400°C and 500 psig...... 38

Figure 4.2.3: Hydrogen yield of decalin over Pt/Al2O3 as a function of LHSV at a temperature of 400°C and pressure of 500 psig...... 40

Figure 4.2.4: Hydrogen production via dehydrogenation of decalin over Pt/Al2O3 as a function of time and LHSV at a temperature of 400°C and pressure of 500 psig...... 41 Figure 5.2.1: Conversions of decalin and n-dodecane with selectivity of products over Pt/Al2O3 as function of LHSV at a temperature of 400°C and pressure of 500 psig...... 46

Figure 5.2.2: Hydrogen yield of decalin/n-dodecane blend over Pt/Al2O3 as a function of LHSV at a temperature of 400°C and pressure of 500 psig...... 47 Figure 6.2.1: Conversion of n-dodecane and Liquid-to-Gas (LTG) conversion over Zeolite Y as a function of time-on stream at an LHSV of 0.75 h-1, 400°C and 500 psig...... 53 Figure 6.2.2: Gaseous hydrocarbon product selectivity over 120 minute intervals for n-dodecane over Zeolite Y at an LHSV of 0.75 h-1, 400°C and 500 psig...... 55 Figure 6.2.3: Liquid hydrocarbon product selectivity over 120 minute intervals for n-dodecane over Zeolite Y at an LHSV of 0.75 h-1, 400°C and 500 psig...... 57 Figure 7.2.1: Conversion of n-dodecane and decalin in blended feed over Zeolite Y as function of time-on stream at LHSV of 0.75 h-1, 400°C, and 500 psig...... 60 Figure 7.2.2: Comparing n-dodecane and LTG conversion over time with a 9.14% decalin by mass blended feed to the baseline n-dodecane study at 400°C and 500 psig...... 62 Figure 7.2.3: Gaseous hydrocarbon product selectivity over time for n- dodecane/decalin blend over Zeolite Y at LHSV of 0.75 h-1, 400°C, and 500 psig...... 63 Figure 7.2.4: Liquid hydrocarbon product selectivity over time for n- dodecane/decalin blend over Zeolite Y at LHSV of 0.75 h-1, 400°C, and 500 psig...... 64 Figure 7.2.5: Hydrogen, total aromatic, and cycloparaffin product selectivity over time for n-dodecane/decalin blend over Zeolite Y at LHSV of 0.75 h-1, 400°C, and 500 psig...... 66 Figure 7.2.6: Aromatic species product selectivity over time for n- dodecane/decalin blend over Zeolite Y at LHSV of 0.75 h-1, 400°C, and 500 psig...... 67 Figure 8.2.1: Catalytic activity over time during dodecane cracking testing with a feed containing 9.14% by mass of decalin...... 73 Figure 8.2.2: Gaseous hydrocarbon product yield for n-dodecane/naphthalene blend over Zeolite Y at LHSV of 0.75 h-1, 400°C, and 500 psig...... 74 Figure 8.2.3: Liquid hydrocarbon product yield for n-dodecane/naphthalene blend over Zeolite Y at LHSV of 0.75 h-1, 400°C, and 500 psig...... 75 Figure 8.2.4: Aromatic product selectivity for n-dodecane/naphthalene blend over Zeolite Y at LHSV of 0.75 h-1, 400°C, and 500 psig...... 76

Figure 9.2.1: Conversion of n-dodecane and decalin in blended feed over Pt/Al2O3 upstream of Zeolite Y as a function of time-on stream at an LHSV of 0.75 h-1 (with respect to the Zeolite bed) and reaction conditions of 400°C and 500 psig...... 81 Figure 9.2.2: Comparing n-dodecane conversion over time for a dual-bed reactor with blended feed to the baseline n-dodecane study at 400°C and 500 psig...... 83 Figure 9.2.3: Hydrogen, total aromatic, and cycloparaffin product selectivity over time for n-dodecane/decalin blend over dual-bed reactor system at 400°C and 500 psig...... 84 Figure 9.2.4: Aromatic product selectivity over time for n-dodecane/decalin blend over dual-bed reactor system at 400°C and 500 psig...... 85 Figure 9.2.5: Gaseous hydrocarbon product selectivity over time for n-dodecane/decalin blend over dual-bed reactor system at a Zeolite LHSV of 0.75 h-1, 400°C, and 500 psig...... 86 Figure 9.2.6: Liquid hydrocarbon product selectivity over time for n-dodecane/decalin blend over dual-bed reactor system at a Zeolite LHSV of 0.75 h-1, 400°C, and 500 psig...... 88 Figure 10.2.1: Conversion of n-dodecane and decalin in blended feed over mixed Pt/Al2O3 and Zeolite Y as a function of time-on stream at an LHSV of 0.75 h-1 (with respect to the Zeolite Y volume) and reaction conditions of 400°C and 500 psig...... 92 Figure 10.2.2: Hydrogen, total aromatic, and cycloparaffin product selectivity over time for n-dodecane/decalin blend over mixed-bed reactor system at 400°C and 500 psig...... 94 Figure 10.2.3: Aromatic product selectivity over time for n-dodecane/decalin blend over mixed-bed reactor system at 400°C and 500 psig...... 95 Figure 10.2.4: Gaseous hydrocarbon product selectivity over time for n-dodecane/decalin blend over mixed-bed reactor system at 400°C and 500 psig...... 96 Figure 10.2.5: Liquid hydrocarbon product selectivity over time for n- dodecane/decalin blend over mixed-bed reactor system at 400°C and 500 psig...... 97 Figure 10.2.6: Comparison of n-dodecane conversion for single-bed, dual-bed, and mixed-bed reactor systems at the same LHSV (with respect to Zeolite Y volume) and reactor conditions of 400°C and 500 psig...... 99

LIST OF TABLES

Table 2.3.1: Synthesized Pt/Al2O3 catalyst physical properties...... 29 Table 3.1.1: Properties of Zeolite catalysts used in preliminary activity screening...... 31 Table 4.1.1: Run segments used for decalin dehydrogenation LHSV optimization at reaction conditions of 400°C and 500 psig...... 36 Table 6.1.1: Run segments used for n-dodecane cracking testing over Zeolite Y at an LHSV of 0.75 h-1 and reaction conditions of 400°C and 500 psig...... 52

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CHAPTER 1

INTRODUCTION AND BACKGROUND

1.1 Introduction

Hydrocarbon cracking is used in a variety of industries to create smaller and more

desirable species. This can be done through either thermal or catalytic cracking

reactions, with catalysts being used to increase cracking rate and selectivity towards

desired products. A major issue associated with the use of cracking catalysts is

deactivation, which is a loss of activity that can occur through pathways such as fouling,

poisoning, or thermal degradation.1 Zeolites are a class of common cracking catalysts,

comprised of silicon, aluminum, and oxygen present in different ratios and structures.2

Zeolites utilize acid-site chemistry to crack hydrocarbons, with the rate-limiting step

being initial protonation of the reactant.3 The addition of a gaseous hydrogen co-feed can

enhance cracking activity through activation of molecular hydrogen on the acid-sites, known as hydrocracking.4 In the absence of a hydrogen co-feed, it has been speculated that hydrogen-donating molecules could be utilized to facilitate hydrogen transfer in cracking reactions.5 Previously studied hydrogen donor molecules have included decalin,

cyclohexane, and methylcyclohexane.6,7 The purpose of this study was to investigate the

use of a hydrogen-donor feed blend to enhance catalytic cracking rate of a n-paraffin,

while also reducing the catalyst deactivation rate. This was done by utilizing a fixed-bed

reactor system in which catalytic dehydrogenation of the hydrogen-donor compound

would generate molecular hydrogen in-situ. The hydrogen would then be available for use in a downstream catalytic cracking bed. The in-situ hydrogen generation could also facilitate hydrogen transfer and hydrocracking to improve the capabilities of a zeolite catalyst.

1.2 Decalin as a Hydrogen Donor

Cycloparaffins have potential for use as a hydrogen source in applications where gaseous molecular hydrogen would be impractical. Decahydronaphthalene (decalin or

C10H18) in particular is valued as a hydrogen carrier due to its high gravimetric and

volumetric hydrogen content.8 Catalytic dehydrogenation of decalin has primarily been

investigated as an alternative to use of compressed hydrogen in fuel cell vehicles, but it

has also been studied for use in fuel upgrading9, Fischer-Tropsch synthesis6, and high-

temperature pyrolytic coke mitigation5.

A potential application for catalytic decalin dehydrogenation would be in

hydrogen fuel cell vehicles (HFCVs), because a fuel tank could be filled with liquid

decalin as opposed to compressed hydrogen. Transportation and storage of compressed

gaseous hydrogen is much more complex than use of a liquid hydrocarbon fuel, which

could utilize existing infrastructure. Decalin could be dehydrogenated within the vehicle, with product hydrogen separated and fed directly to the fuel cell. The remaining dehydrogenation product, naphthalene, could be collected and rehydrogenated, allowing repeated use. Liquid decalin has a hydrogen storage density of 64.8 kg of hydrogen per cubic meter. This can allow a vehicle to have a range of 250 miles on approximately 13 gallons of decalin fuel.8 Besides HFCVs, another use of decalin is as a hydrogen carrier

in fuel upgrading. Generally, molecular hydrogen is introduced as a gas-phase co-feed

with hydrocarbons over a catalytic bed for the purpose of hydrogenating species and

stabilizing products. From an industrial perspective, hydrogen gas has inherent safety

disadvantages of pressurized storage needs and high flammability potential when exposed

to air. Catalytic hydrogen transfer from donor species, such as decalin or tetralin, has the

ability to be utilized in low-pressure bio-oil upgrading as an alternative to gaseous

molecular hydrogen, which requires the use of high-pressure equipment.9 Decalin has

also been investigated as a hydrogen-donating molecule for coke mitigation in the pyrolysis of aviation fuels at high temperatures. Addition of small amounts of decalin to a fuel could allow termination of hydrocarbon coke-precursor radicals which would otherwise cause eventual blockage of fuel transfer lines and nozzles.5

Decalin is generally comprised of a mixture of trans and cis isomers, which have

boiling points of 187°C and 196°C, respectively. Naphthalene has a boiling point of

218°C and is a solid at room temperature. The conversion from decalin to naphthalene is

endothermic at complete conversion with an enthalpy change of approximately 297.3 kJ

per mole of decalin converted, depending on the cis-trans ratio.8 Decalin first undergoes

complete dehydrogenation of one cyclohexane-ring, forming tetralin and evolving three

dihydrogen molecules. This is followed by dehydrogenation of the adjacent

cyclohexane-ring, forming naphthalene and two additional dihydrogen molecules; the pathway is shown in Figure 1.2.1. The potential yield of five moles of molecular

hydrogen for each mole of decalin makes it a very attractive hydrogen donor compound

for use in reactions which would otherwise require free molecular hydrogen. Catalytic

dehydrogenation generally occurs on metal-site catalysts, and has been shown to take

place at temperatures as low as 200°C. Increased temperatures have the potential for

coke formation under certain conditions, which would deactivate the catalyst.8

Therefore, there is trade-off in temperature between maximizing dehydrogenation rate

and minimizing coke-forming side reactions. Platinum or palladium are the most

commonly studied catalysts for decalin dehydrogenation and these generally utilize a

carbon-based or alumina support material.6,8,9,10,11,12 Catalytic activity can be increased

by increasing metal loading, decreasing metal particle size, or altering the structure of the

support material. In particular, advanced high surface area carbon structures, such as

carbon nanotubes and carbon nanofibers, have allowed very high metal dispersion and

superior catalytic activity.10,12 Activity can also be increased by the addition of secondary metals to a platinum catalyst, such as iridium, tungsten, and rhenium.6,8

Figure 1.2.1: Decalin (left) dehydrogenation to tetralin (center), followed by dehydrogenation to naphthalene (right).

For maximum hydrogen production, product selectivity favoring complete dehydrogenation to naphthalene, rather than intermediate tetralin, is desired. Wang et al. concluded that the kinetic rate of tetralin dehydrogenation is much faster than the rate of tetralin formation from decalin at 240°C.10 They stated that aromatic rings are known to

be slow in desorbing from platinum catalyst active sites and favor quick adsorption onto

these same active sites. The strong adsorption tendency of tetralin allows for a high

selectivity for complete dehydrogenation of decalin to naphthalene. It also has been

observed that the rate of dehydrogenation for cis-decalin is much higher than that of its

isomer, trans-decalin. This is possibly due to the more stable nature of trans-decalin, leading to a higher energy barrier to reaction, compared to the strained-ring configuration of cis-decalin.10 Therefore, the decalin isomer ratio must be considered when utilizing

the cycloparaffin as a hydrogen source.

A variety of catalyst types and modifications can be employed to favorably increase kinetic rates of decalin dehydrogenation. Hodoshima et al. compared catalytic decalin dehydrogenation activity for a carbon-supported platinum (Pt/C) catalyst to

versions of the catalyst modified by the addition of iridium, tungsten, and rhenium.

These catalysts were tested in a batch reactor at temperatures between 210 and 280°C.

The Pt-Re/C catalyst was found to be the most active in decalin dehydrogenation, followed by Pt-W/C and Pt-Ir/C, with Pt/C exhibiting the lowest activity.8 Alternatively, dehydrogenation activity for platinum catalysts can be increased by the use of carbon nanotube (CNT) or carbon nanofiber (CNF) supports. Use of CNT as a platinum support greatly enhanced dehydrogenation activity compared to other carbon supports (granular carbon and carbon black). In the same study, alumina-supported platinum was shown to be slightly less active than the CNT catalyst, but still largely superior to the other carbon- supported catalysts.10 Similarly, CNF-supported platinum exhibited higher decalin dehydrogenation activity than activated carbon-supported platinum, likely due to superior metal dispersion characteristics.12

Rather than being used solely to generate molecular hydrogen, decalin can be utilized as an in-situ hydrogen donator for reactions that would otherwise require a

gaseous hydrogen feed. Shafaghat et al. recently demonstrated the use of both platinum

and palladium catalysts for hydrogen transfer of decalin coupled with bio-oil upgrading.9

This process uses the generated hydrogen to remove oxygen from other hydrocarbon

species. When these reactions occur at the same active site, it is known as transfer

hydrogenation. Transfer hydrogen reactions were studied with a batch reactor system at a

temperature of 275°C with various residence times. The carbon-supported noble metal

catalysts used in the study were able to both remove a hydrogen atom from the carrier

decalin as well as transfer it to the oxygen compounds, although this occurred at a higher

rate on the Pt/C catalyst than the Pd/C catalyst. It was observed that the dehydrogenation

rate for decalin was reduced with the addition of oxygen containing compounds, both due

to competition for active sites as well as coke formation from the latter. It was

hypothesized that there were two mechanisms through which the hydrogen transfer was

occurring. The first mechanism involved abstraction of a hydrogen atom from the donor

molecule being transferred to the acceptor at the active site. The other mechanism

involved the creation of molecular hydrogen from the donor. The free hydrogen is then

separately adsorbed onto an active site where it can be transferred to an acceptor

molecule.9 This study demonstrated the viability of gaseous hydrogen replacement via

in-situ hydrogen transfer, facilitated by catalytic decalin dehydrogenation.

Decalin hydrogen donation also has relevant application to aircraft fuel systems.

Modern turbine engines require integrated thermal management systems and as

performance is increased, temperature increases significantly. Creating a heat sink

through integration of heat exchangers directly between fuel transfer lines and an engine

provides an opportunity to mitigate this heat generation and increase engine

performance.13 However, high temperatures cause fuel to break down and form coke, which has potential to clog valves and fuel nozzles. Corporan et al. investigated the

6 impact of blending decalin into an aviation fuel on carbon deposition properties at pyrolytic temperatures.5 It was suspected that decalin or tetralin could act as a hydrogen donor in fuel pathways to mitigate fouling of aircraft fuel systems. Hydrogen radicals produced from decalin or tetralin could combine with coke-forming hydrocarbon radicals to create stable compounds that are less likely to further react to coke. Pyrolysis of fuel occurs at temperatures above 400°C, so a decalin-fuel blend was studied in a flowing reactor tube at an average temperature of 540°C. At 1.0% by mass of decalin in the blend, carbon deposition was reduced by a small amount while there was no benefit in a

0.05% blend. It was suggested that a homogeneous, or in-solution, catalyst could be added to the fuel blend to promote hydrogen donation by the decalin, but it is otherwise not an effective additive to reduce coke deposits under the studied reaction conditions.5

1.3 Normal and Isomerized Paraffin Cracking Mechanisms Utilizing Zeolite

Catalysts

Cracking refers to breaking down large hydrocarbons into lower carbon-number species. If a catalyst is used to promote cracking reactions, this is known as catalytic cracking, but in the absence of a catalyst, high temperatures can still cause thermal cracking at a relatively slower rate (i.e. pyrolysis). Catalytic cracking can be performed in a heated reactor without additional co-feed, but the addition of hydrogen

(hydrocracking) or steam (steam cracking) is regularly used for increasing reaction rate and minimizing catalyst deactivation. The types of catalyst active sites which are generally utilized for cracking reactions include both metal and acid sites. Metal sites are usually comprised of noble metal particles, such as platinum or palladium, while acid sites primarily act as proton donors and are most commonly found on Zeolites.4

Zeolite catalysts are synthetic acid-site catalysts with defined structures and are

commonly used to crack or isomerize many hydrocarbon compounds, including n-

paraffins.2 Cracking and isomerization reactions can be endothermic processes, and are

commonly used in petroleum processing to convert heavy hydrocarbons into desired fuel

components. Isomerization can be used to improve fuel quality, such as increasing the

octane rating of gasoline fuels.14 Zeolite active sites are contained within the highly porous framework, with pore sizes ranging from 0.3 to 1.2 nm. There are many different

Zeolite structures, but all contain silica (SiO2) and alumina (Al2O3). The ratio of these

two components can vary between Zeolite type and within different formulations of a

single Zeolite structure. Zeolite activity and product selectivity are affected by structure

type, surface area, and acidity, which all vary based on SiO2/Al2O3 ratio and synthesis

procedure. Additionally, various metals can be impregnated into the structure of the

Zeolite to alter catalytic activity or function.2

Cracking can occur via thermal or catalytic pathways. Non-catalytic thermal

cracking is initiated at high temperatures when radical species are produced via

homolytic bond dissociation, which further decomposes fuel via propagation reactions.

Alternatively in bifunctional catalysts, where a metal and acid site are both present, the

mechanism involves dehydrogenation of saturated compounds on the metal sites, with

isomerization and carbon-carbon bond scissions occurring on the acid sites. When

carbon-carbon bonds are dissociated directly on a metal site in the presence of hydrogen,

this is known as hydrogenolysis. In an acid site monofunctional catalyst, both

bimolecular and monomolecular cracking reactions involving radical species and protons

are involved.4

In bimolecular cracking reactions, a hydride molecule is abstracted from a carbon,

creating a carbenium ion (positively charged trivalent carbon) on a paraffin. The carbon-

carbon bond two sites away from the carbenium ion then undergoes bond scission, known

as β–scission, producing a double bond on the parent molecule adjacent to the original carbenium ion, along with formation of a new carbenium ion from the released hydrocarbon species, as shown in Figure 1.2.1. As a result of non-adjacent bond

splitting, propane is typically the smallest molecule produced via this type of

decomposition mechanism.15

Figure 1.3.1: Examples of β–scission reaction.16

Alternatively, monomolecular cracking on acid sites occurs via the more recently

proposed Haag-Dessau mechanism.15 In this mechanism, a paraffin is first protonated to

form a carbonium ion (positively charged pentavalent carbon) via a catalyst acid site.

The carbonium ion is an intermediate compound which then decomposes to produce a

normal paraffin (or dihydrogen) and another carbonium compound. The resulting

carbonium can crack further, or form an alkene via the release of a proton. An example

9 of this mechanism is shown in Figure 1.3.2. The Haag-Dessau mechanism allows the formation of ethane, methane, and dihydrogen, which would not be present from cracking that occurs via bimolecular pathways. If a high amount of alkenes are present in the hydrocarbon product mixture, the primary cracking pathway becomes bimolecular since alkenes are more readily protonated. This in turn inhibits the initiation required for the monomolecular reactions.

Figure 1.3.2: Examples of Haag-Dessau mechanism.15

Utilizing both metal and acid sites for catalytic cracking can be more advantageous than using only monofunctional acid catalysts. These advantages include lower required reaction temperatures and less coke formation. A proposed mechanism begins with paraffin dehydrogenation at a metal site, producing an alkene. The alkene then diffuses to an acid site where it is protonated to form a carbenium ion. These carbenium ions then isomerize and crack via β–scission pathways. β–scission decomposition does not occur with normal-isomers of carbenium compounds, so iso- alkylcarbenium ions are first produced. These iso-alkylcarbenium ions can alternatively collapse and desorb from the acid site as branched alkenes, which then diffuse back to a metal site for hydrogenation, resulting in iso-paraffin formation. At higher temperatures, more branching occurs at the acid sites, which is very favorable to increasing the kinetic rate of β–scission. Again due to the limitations of β–scission, products smaller than

10 propane are not generally formed from this mechanism. In what is referred to as ideal hydrocracking, the carbon number of the products is normally distributed around the reactant carbon number. However, due to slow desorption of cracked products, observed product distributions tend to be skewed to lower carbon numbers since the products are further cracked before desorbing to the bulk phase. Catalyst supports that allow quick desorption and diffusion from the pores are therefore more likely to resemble ideal hydrocracking. Rather than diffusion between metal and acid sites, bifunctional catalysts can alternatively exhibit spillover of adsorbed hydrogen from the metal sites onto the acid catalyst support, allowing more rapid hydrogen transfer to the cracking species.4

The largest issue associated with cracking in the production of fuels and other desirable hydrocarbon compounds is the subsequent coking that occurs. Coking and deactivation of the catalysts contributes significantly to high maintenance and operational costs of cracking reactors. The most common coke precursors include olefins and aromatics. In the absence of free molecular hydrogen, hydrogen-deficient compounds will isomerize and condense. This leads to the formation of heavy, carbon-rich species that precipitate from the solution, as they are no longer soluble in the bulk hydrocarbon stream.17,18 Hydrocarbon species that have low solubility are generally formed through the thermal decomposition of reactants as well as secondary reactions from the desired products.17 Olefins directly contribute to oligomerization of coke species, but polycyclic aromatic hydrocarbons (PAH) are a much larger contributor to coke formation due to their propensity for generating and interacting with radical species.19 Coke formation can reduce the desired product yield in industrial hydrocarbon processes and lead to

costly downtime, so reducing the rate of coke formation during cracking has major

interest from chemical refining businesses.

For Zeolites specifically, due to their highly porous nature, the deactivation from

coke primarily occurs due to blockage of pores. Operating at lower pressures, reducing

residence time, and avoiding hot spots are general guidelines for reducing coke

formation. Cracking reactions are generally modeled as first order, so gas phase

hydrocarbons and coke precursors are at a higher concentration with higher pressures.

Even though cracking can be first order, coke formation has shown to be exponentially

related to fuel conversion. Increased concentration of reactants at high pressure leads to

faster cracking activity, but increases the propensity for coke formation. However, at

supercritical conditions, coke has the potential to become soluble in the reaction products.

This would allow coke species to more easily migrate from the Zeolite pores and, keep

more available active sites for cracking reactions.19

Although coking is the primary contributor to Zeolite catalyst deactivation,

alternative pathways can also contribute to deactivation. Deactivation is typically used to

describe a decline in catalyst activity, but it can also refer to a change in selectivity with

operational time, which is commonly caused by catalytic poisoning. Poisoning can occur

in Zeolites as the result of hydrocarbons strongly adsorbing onto active sites. This will

initially cause a decrease in available active sites, reducing catalyst activity. But over

time, the strongly adsorbed species could alter the catalytic framework and functionality.

Therefore, species that have a very strong adsorption tendency can shift Zeolite catalyst selectivity over time, unrelated to deactivation from coke deposition.1

One Zeolite, ZSM-5, has been shown to have very high activity with respect to

paraffinic cracking, even without adding a metal to the structure. Li et al.20 studied

cracking of n-dodecane over ZSM-5 at supercritical conditions (550°C, 4 MPa/580 psi).

A coated tube flow reactor was used to quantify the conversion of n-dodecane as a

function of time, for determination of Zeolite deactivation rates with different

SiO2/Al2O3 ratios. Dodecane flow rate was relatively high, leading to a liquid hourly space velocity value of approximately 1000 h-1 or higher. Initial activity for HZSM-5

with a SiO2/Al2O3 ratio of 50 was approximately 40% and dropped to approximately

25% by the end of the run. The study showed rapid deactivation of the catalyst over the

course of 37.5 minutes, with the fastest drop occurring within the first 12.5 minutes. This

study demonstrated that a model n-paraffin feed (n-dodecane) can be utilized in a flowing packed-bed reactor to observe Zeolite catalyst deactivation at supercritical conditions.

1.4 Hydrocracking, Ring Opening, and Alkylation of Cycloparaffinic and Aromatic

Species on Zeolite Catalysts

One specific and extensively studied area of catalytic cracking deals with cyclic saturated (cycloparaffinic) and unsaturated 6-member (aromatic) hydrocarbons. Zeolites have the ability to facilitate splitting of intra-ring carbon-carbon bonds in cycloparaffins, thus “opening” the ring. These reactions are mostly performed with Zeolites containing an impregnated metal, but can occur with non-metal containing Zeolites. Chareonpanich et al.21 investigated the reaction pathways of single and multi-ring aromatic compounds

over a monofunctional Y-structure Zeolite in a batch reactor, under a hydrogen pressure

of 5.0 MPa. Hydrogenation and cracking reactions were observed for both naphthalene

and tetralin at 400°C. Light hydrocarbon yield increased with higher temperature, up to

600°C, and higher hydrogen pressure increased light hydrocarbon product yield for

testing with 1-methylnaphthalene. The proposed mechanism for the observed results

included initial adsorption of hydrogen onto the Zeolite active sites to create protonic

sites, which then initiated hydrogenation of aromatic species. The reaction pathways for

ring opening proposed by the researchers showed that naphthalene is first hydrogenated

to tetralin. The resulting saturated ring could be cracked, or the remaining aromatic ring

could be further hydrogenated, which would yield decalin as the product. However,

direct cracking of aromatic rings did not occur.

Proposed pathways for the cracking of paraffinic rings, such as those found in decalin or tetralin, is very similar to the previously discussed normal- and iso-paraffinic cracking mechanisms (Section 1.3). However, due to configurational stability of cycloalkylcarbenium ions (cycloparaffins with a positively charged carbon radical), endocyclic carbon-carbon bonds do not readily decompose via β–scission. Instead, molecular hydrogen must first be utilized for ring opening, which produces a paraffinic chain that can be cracked.4 It is suggested that one possible catalytic ring opening

pathway is a ring contraction reaction prior to ring opening. For example, decalin would

first convert into methyl-hydrindane, and then cleavage could occur between the

branched carbon and the cyclohexane ring, as shown in Figure 1.4.1. The cleavage can

occur on either a metal or acid site. If the remaining cyclohexane ring is opened with no

further reaction, it would be an example of what is referred to as selective ring opening.

If dissociation of a carbon-carbon bond occurred on the butyl chain or between the butyl

chain and the cyclohexane ring, these would be examples of nonselective dealkylation. If

both selective ring opening and nonselective dealkylation occur simultaneously, or in

14 sequential order, these would be examples of nonselective ring opening. The high cracking potential of acid site catalysts generally promotes nonselective ring opening.

Alternatively, another possible ring opening pathway is from non-catalytic thermal cracking, caused by bond fission at high temperatures, generally above 400°C.22

Figure 1.4.1: Ring contraction from decalin (left) to methyl-hydrindane (center) followed by ring opening to butyl-cyclohexane (right).

Metal site catalysts are preferential for selective ring opening. Metal site ring opening occurs through a hydrogenolysis (carbon bond cleavage utilizing hydrogen) reaction, which creates a product concentration high in ring opening species and low in dealkylation products. The possible products obtained from selective ring opening of butyl-cyclohexane are shown in Figure 1.4.2. This selectivity of metal sites towards ring opening products is in contrast with the non-selective ring opening tendencies of thermal and acid site cracking. The major benefit of increased selective ring opening is formation of cracked products with the same carbon number as the feed, which is desired in many circumstances.22

Figure 1.4.2: Non-branched and single-branched C10 products obtained from single bond cleavage in selective ring opening of butyl-cyclohexane.

In catalytic decalin reactions such as cracking and ring opening, there can be large differences in functionality between the cis and trans isomers based on catalyst functionality. Generally, cis-decalin is less stable, and therefore more reactive, compared to trans-decalin due to ring strain differences. The cis configuration is also more compact, with regard to cross-sectional diameter, and can therefore diffuse through pores more easily. This contributes to higher reactivity in highly porous catalysts, such as

Zeolites. Zeolites also have the ability to promote isomerization reactions between the

two conformations, which can change the equilibrium from the feed. For example, if a

catalyst is highly selective towards isomerization from cis to trans, then the overall

conversion of decalin will be reduced. When pure isomer feeds were separately run over

HY and Pt/HY catalysts with a hydrogen co-feed at 533 K, Santikunaporn et al.23

confirmed that the conversion of the cis-decalin was up to three times higher than that of the trans-decalin isomer. The selectivity towards ring opening products for cis-decalin

was also significantly higher, while trans-decalin was more selective towards cracking

products.

Differences in Zeolite structure can alter product selectivity and reaction rate with

respect to cycloparaffin cracking. Corma et al.24 evaluated the effectiveness of different monofunctional Zeolite types on the activity of decalin and tetralin cracking. Flowing packed bed reactors were used with Zeolites of varying pore size at a reaction temperature of 450°C. The study determined the highest yield products from tetralin cracking over ZSM-5 to be C1-C4, naphthalene, benzene, and C10 aromatics. These

products are consistent with isomerization, dealkylation, hydrogen transfer, and protolytic

cracking reactions. In decalin cracking, there were high concentration of C10 aromatics

formed, but low amounts of tetralin or naphthalene formed. This would suggest that

direct dehydrogenation of decalin is not favored compared to ring opening. Ring opening

reactions of decalin are significantly faster than ring opening of the saturated ring of

tetralin. In general, the data suggested that larger pore Zeolites (B and Y) will open rings

more readily compared to smaller pore Zeolites, such as ZSM-5. On the other hand,

ZSM-5 is more active for cracking of alkyl branches on cycloparaffins or aromatics.

Santikunaporn et al.23 studied the ring opening of decalin and tetralin that occurs on a Y Zeolite with a hydrogen co-feed (60-65 mol H2/mol reactant) present. A decalin

feed containing both isomers was tested at 533 K over a range of HY Zeolites, as well as

one with platinum added via incipient wetness impregnation. The catalysts showed

product distributions primarily consisting of species produced via C10 ring opening and ring contraction, as well as C1-C5 cracked products. The addition of platinum lowered

the total acidity of the catalyst by 31.7%, which in turn decreased the observed decalin

conversion. However, the platinum increased the stability of the Zeolite, and a much

lower deactivation rate of the reactions resulted. In all cases, there were negligible

amounts of direct dehydrogenation (tetralin and naphthalene) products formed. The primary products observed from tetralin were decalin isomers, while other products were formed via secondary reaction pathways. Another interesting observation was that tetralin yielded up to nearly 5 wt% naphthalene on non-metal containing HY as compared to negligible dehydrogenation observed for decalin.23 This indicates that the barrier for

tetralin dehydrogenation is much lower than that for decalin dehydrogenation, which is

consistent with dehydrogenation studies utilizing Pt/C and other catalysts.10

Santikunaporn et al.23 used dual-bed and mixed-bed reactor configurations with 1 wt% Pt/SiO2 and non-metal HY Zeolite to study the difference in role of both metal and

acid-sites on tetralin conversion. While using only HY Zeolite produced many ring-

opening alkylbenzene products, the Pt/SiO2 catalyst (when physically mixed into a single

bed with HY) resulted in hydrogenation of the aromatic rings and produced

alkylcyclohexanes. This is to be expected since supported platinum is commonly used to

facilitate hydrogen transfer for a wide range of aromatic compounds. Mixing the Pt/SiO2

at a higher ratio increased tetralin conversion, ring-opening product yield, alkylcyclohexane selectivity, ring-contraction product yield, and decalin yield, but decreased the naphthalene yield. When a Pt/SiO2 bed was placed in-line after an HY bed, similar ring-opening and ring-contraction product yields were observed as compared to testing with only the Zeolite, but there was a significant conversion of the tetralin to decalin via hydrogenation reactions. When the Pt/SiO2 bed was placed before the HY

bed, higher amounts of ring-opening and ring-contraction products were promoted.

Ma et al.25 also studied the properties of metal and acid sites on the hydrogenation

and ring opening of tetralin. They found that conversion to decalin at 200°C was much

more prevalent in a platinum impregnated USY Zeolite catalyst compared to the pure

Zeolite, suggesting that metal sites are more effective in general for dehydrogenation.

However, they also discovered that ring-contraction reactions required both metal and acid sites, with an optimal Pt loading of 0.5 wt%. When the acid strength of the Pt/USY catalyst was decreased, more hydrocracking products were created, which also indicated that less ring-opening products were created, as hydrocracking mainly occurs from the opened rings.

As an alternative to ring opening, naphthalene can be alkylated in a Zeolite catalyst reactor system. Brzozowski and Tecza26 demonstrated alkylation of naphthalene

with propylene on an HY Zeolite catalyst at 150-300°C and 4.6 MPa. When an undiluted

6:1 molar ratio of naphthalene to propylene feed was used, the catalyst exhibited nearly complete propylene conversion in a fixed-bed flowing reactor at 300°C. The major products were isopropyl-naphthalene and diisopropyl-naphthalene, indicating direct

addition of the feed constituents. Sugi27 demonstrated similar high activity for alkylation

of naphthalene on Zeolite catalysts, including Zeolite Y, at temperatures between 150 and

300°C. In addition to propylation, butene and methyl-propene were used to demonstrate butylation capabilities with naphthalene. It was also speculated that bulky alkylnaphthalenes have the potential to deactivate Zeolite catalysts by blocking pores and access to active sites.

These studies show that naphthalene produced from decalin dehydrogenation would not be a stable, unreactive compound in the Zeolite cracking portion of a dual-bed

reactor system. Naphthalene has the potential for hydrogenation and ring opening on

Zeolite catalysts, forming alkylbenzene and cycloparaffinic products. Additionally,

naphthalene can be alkylated on Zeolites by the addition of olefins, even at temperatures

as low as 150°C. In addition to naphthalene reactions, unreacted decalin and partially

dehydrogenated product (tetralin) can exhibit ring opening and cracking reactions. These

reaction pathways have the potential to lower n-paraffinic cracking activity through

active site competition as well as through pore blockage from bulky isonaphthalene

production.

1.5 Presentation of Research

The ultimate goal of experimentation performed in this study was to test a dual-bed flowing reactor system in which a hydrogen-donor cycloparaffin would be blended into a n-paraffin liquid feed in a fixed-bed flowing reactor. The cycloparaffin could be catalytically dehydrogenated upstream of a cracking bed to generate in-situ molecular hydrogen for enhancing cracking activity. However, interactions between feeds and individual catalyst beds were first studied independently to better understand reactions that would occur in the dual-bed system. Model feed hydrocarbons were used to simplify reaction pathways and reduce confounding effects from possible side-reactions. Decalin was selected as the hydrogen-donor compound and n-dodecane was selected as the model cracking compound.

Initial activity screening was performed for various Zeolites for the purpose of selecting an adequate catalyst for use in subsequent studies, with results being described in Chapter 3. Each Zeolite was tested for n-dodecane cracking activity in hydrogen-

deficient and hydrogen-rich conditions. Upon completion of the Zeolite screening,

acceptable decalin dehydrogenation reactor conditions were determined in Chapters 4 and

5. These studies utilized a synthesized platinum catalyst to evaluate dehydrogenation

activity at various reaction temperatures and feed flow rates. Chapter 4 discusses a study

with a neat decalin feed while Chapter 5 discusses studies which introduced a n-dodecane

co-feed. After suitable reactor conditions were determined, baseline activity and

deactivation rate (at those conditions) were established for the selected catalyst. Chapter

6 presents baseline studies for a neat n-dodecane feed, while Chapter 7 discusses studies with a decalin co-feed. Chapter 8 discusses an investigation of the impact of decalin

dehydrogenation products on cracking activity by dissolving naphthalene into a

n-dodecane feed and observing impact on reaction behavior.

After reactor conditions and effects of individual catalyst beds were determined, a dual-bed reactor system was studied. In this system, a 1:1 volumetric blend of n- dodecane and decalin was reacted over a dehydrogenation catalyst bed prior to reaction in a cracking catalyst bed, with results discussed in Chapter 9. Similarly, Chapter 10 discusses a study which utilized the same feed, but the catalyst bed contained a physical mixture of dehydrogenation and cracking catalysts. These results were then compared to results from prior studies for cracking activity and catalyst deactivation, to determine feasibility of the reactor-configurations in enhancing performance.

CHAPTER 2

MATERIALS AND METHODS

2.1 Reactor Setup

Experimental studies were performed using a fixed-bed reactor configuration with a vertical cylindrical reactor situated within a heated 3-zone furnace. The experimental reactor configuration allowed gaseous and liquid feeds to be fed simultaneously into the top of a heated catalytic bed with products either analyzed on-line or collected for off-line analysis. Figure 2.1.1 shows a flow diagram of the reactor system configuration used during experimental studies. The gaseous feed consisted of a mixture of nitrogen, with or without hydrogen, while the liquid feed consisted of the desired liquid hydrocarbon reactant. The reaction zone was comprised of catalyst loaded in a ¼" or ½" outer diameter (0.18" and 0.37" inner diameter, respectively) 316 stainless steel reactor tube placed in a 3-zone clamshell furnace. The catalyst bed was positioned in the center zone; the inlet zone was used to preheat the feed to the target temperature. The inlet zone and remaining lower portion of the reactor tube were filled with inert Silicon Carbide (SiC) with contents of each zone separated by a glass wool plug. The SiC inlet zone improved heat transfer to ensure that the feed was at the target temperature prior to entering the catalyst bed. Gaseous reactants were fed using Brooks 500 sccm (standard cubic centimeter per minute) mass flow controllers while liquid reactants were fed using a

Teledyne 500D ISCO syringe pump. The feed streams were combined at the inlet to the

22 furnace, followed by mixing in the SiC inlet zone as the liquid vaporized. All three furnace zones were controlled to the same temperature to minimize any heat gradient in the reaction zone. After exiting the furnace, the outlet stream passed through hot and/or cold traps to collect condensable samples for off-line analysis. Gaseous species were either vented or directed to a MicroGC for on-line quantitation. Downstream of the product traps, but before the MicroGC, a Tescom backpressure regulator (1500 psi maximum) was used to control the overall pressure in the system. System pressure was set to 500 psig for all experimental studies, with operating temperature ranging from

350°C to 450°C.

H2 Liquid N 2 Feed Pump

3-Zone Catalyst Reactor Bed Furnace

Liquid and/or Cold and/or Hot Trap Solid Product Collection

Backpressure Gaseous Regulator Products to Micro GC

Figure 2.1.1: Diagram of reactor system configuration used for experimentation.

Figure 2.1.2 shows a photo of the furnace and gas mass flow controllers. A steel heat distributor was attached around the reactor tube to maintain a uniform temperature

throughout all furnace zones and reduce the chance of hot spots. Insulation was used at

the reactor inlet to reduce heat losses from the furnace and the exit line was

insulated/heated when a hot trap was in use to prevent solidification of products.

Feed Gas Controllers

Heat Distributor

Insulation

Figure 2.1.2: Photo of packed bed reactor system furnace and feed gas mass flow controllers.

The liquid product collection trap could be configured as a cold trap, hot trap, or a hot/cold trap combination in series. The traps were custom made from ½" outer-diameter stainless steel tubing of sufficient length to hold a liquid volume of approximately 10 mL.

The cold trap was used to collect condensable species for off-line analysis which

otherwise would condense downstream in gaseous-product transfer lines. The hot trap was only utilized in experiments where the product stream included species which would solidify at ambient temperature. In a cold trap configuration, the product trap was maintained at 20°C via a Lauda E100/RE106 chiller system. Coolant was circulated through a copper tube coil affixed around the trap. Nitrogen, hydrogen, or light

hydrocarbons (C1-C6) that remained in a gaseous state after the cold trap were directed

towards an on-line MicroGC for analysis. In a hot trap configuration, the product trap

was heated with electrical tape to a temperature ranging from 80-100°C, with subsequent gaseous compounds sent to the MicroGC. In a dual hot/cold trap configuration, the exit products from the reactor were directed into the previously described hot trap located in series directly upstream of the cold trap, with gaseous species directed to the MicroGC.

The small volume of the trap vessels helped to reduce lag-time between the true product gas composition and the composition observed in the MicroGC, since gaseous products must first enter and mix into the trap dead-volume. A 3-way ball valve allowed gas products to be diverted from the MicroGC and alternatively directed into a bubble flow meter for room-temperature volumetric flow rate determination. A diagram showing the path of product flow through a single trap is shown in Figure 2.1.3. Flow was directed into the lower portion of the trap body to ensure proper heating or cooling before gas could escape into the exit transfer line. During sampling from the hot/cold traps, the sampling valve remained open until release of gas occurred (indicative of complete liquid purge from the trap).

Product from Reactor

Gaseous Species

Condensable Products

Ball Valve

To Sample Collection Vial

Figure 2.1.3: Diagram of flow through hot or cold product trap.

2.2 Analysis Instrumentation

An Inficon 4-channel 3000 Micro Gas Chromatograph (MicroGC) with thermal conducting detectors (TCDs) was used for on-line analysis of hydrogen, nitrogen, and gaseous hydrocarbon species with a carbon number of six or less. The unit was calibrated with two gas standards, the first was used to quantify molar composition of normal and branched paraffins, normal and branched olefins, and hydrogen. The second calibration gas standard was used for quantifying nitrogen.

Liquid and solid products were analyzed using Gas Chromatography with Mass

Spectrometry detection (GC-MS). Prior to analysis, samples were diluted in hexane to

solubilize solid species and prevent detector saturation. The GC-MS unit was comprised of an Agilent 7890A gas chromatograph and Agilent Technologies 5975C VL mass spectrometer. Calibration was performed for reactants and primary products by subsequent dilution of the species in various volumes of hexane.

Reaction products were also analyzed using a two-dimensional gas chromatograph with simultaneous flame ionization and mass spectral detection (GCxGC-

FID/MS). This analysis was conducted using an Agilent 7890-5975 GC-MS system

equipped with Capillary Flow Technology (CFT) flow modulation. GCxGC has been

shown to accurately provide identification and quantitation of hydrocarbon compounds

separated by carbon number and group classification such as n-paraffin, iso-paraffin,

aromatic, and diaromatic, among others.29

2.3 Catalyst Synthesis

A Pt/Al2O3 catalyst was prepared via incipient wetness impregnation. Catapal C1

alumina, obtained from Sasol, with a reported surface area of 230 m2/g and bulk density of 670-750 g/L was used as the support. Prior to catalyst synthesis, the incipient wetness

of the support was determined by adding distilled water to a small amount of the alumina

until it was fully saturated. The incipient wetness value was found to be 0.8986 mL/g

[0.656 mL of water saturating 0.73 g of alumina], and is equal to the pore volume per

mass of the support. When a metal containing solution was added to a support at this

level, capillary action draws it into the pores. The mixture was then dried to remove the

27 liquid, leaving metal dispersed throughout the support. Tetraamineplatinum(II) Chloride

(5.258% Pt by mass) was used as the metal precursor; 8.49 g of which was diluted in distilled water to produce a total 79.8 mL of solution. This solution was thoroughly mixed with 88.81g of the alumina to impregnate the support with a target loading of 0.5% platinum by mass. The mixture was dried at 150°C overnight to evaporate volatile compounds, followed by calcination in air at 400°C for 4 hours. A photo of the post- calcination catalyst is shown in Figure 2.3.1. Black residue on the side of the ceramic bowl was a remnant of the initial drying step and was not included with the final catalyst sample. The bowl on the left side of the photo contains no residue because its contents were partitioned from the initial bowl after the drying step, for the purpose of increasing exposed surface area during calcination.

Figure 2.3.1: Photo of synthesized Pt/Al2O3 catalyst after final calcination.

Physisorption and chemisorption analyses were performed using a Micromeritics

ASAP 2020 (Accelerated Surface Area and Porosity) system. Physisorption analysis, utilizing nitrogen gas, was used to determine BET surface area, pore volume, average

pore diameter, and nanoparticle size. Chemisorption, utilizing hydrogen gas, was used to determine metal dispersion and metallic surface area. Properties obtained from the

ASAP 2020 are summarized in Table 2.3.1.

Table 2.3.1: Synthesized Pt/Al2O3 catalyst physical properties.

Catalyst Property Value BET Surface Area 244 m2/g Pore Volume 0.484 cm3/g Average Pore Diameter 7.9 nm Nanoparticle Size 24.6 nm Metal Dispersion 44.5 % Metallic Surface Area 0.55 m2/g of sample Metallic Surface Area 110 m2/g of metal

CHAPTER 3

ZEOLITE ACTIVITY SCREENING

The primary objective of this study was to identify a suitable Zeolite cracking catalyst for subsequent evaluation of deactivation rate in hydrogen deficient conditions.

An acceptable catalyst should exhibit moderate-to-high cracking activity with quantifiable deactivation occurring on a reasonable time scale. Subsequent tests will compare baseline deactivation rate of the catalyst during cracking of n-alkane feed to that observed with the in-situ production of molecular hydrogen from decalin dehydrogenation. A catalyst with excessively high activity may cause difficulty in observing benefits from hydrogen production.

Zeolite catalysts are well established as being effective in cracking straight-chain hydrocarbons due to their acid site structures, however, variation in acid-site density and structure between different Zeolite types results in significant variance in cracking functionality and activity.2 Zeolites can have varying physical structures and pore dimensions that exist within the structures. The SiO2/Al2O3 ratio can also vary within

Zeolite types and affects properties, such as acidity and thermal stability. Due to the high prevalence of Zeolite use in industrial applications, many types of Zeolites are commercially available. Various types of commercially available Zeolites were evaluated for performance and activity to identify a suitable candidate for use in a n-dodecane (n-C12H24) cracking study. n-Dodecane was chosen as a model cracking

feed as it allows a large range of cracking sites and contains sufficient molecular size to

form a variety of different normal, isomerized, and cyclic product types. Tested Zeolites

included ZSM-5, Zeolite B, and Zeolite Y with varying SiO2/Al2O3 ratios. Properties for

the specific catalysts studied are shown in Table 3.1.1. Pore dimensions listed as a range are due to multiple channel structures occurring within a Zeolite type framework (data obtained from the Database of Zeolite Structures28). The selected catalysts were chosen

for the wide variation in surface area, structure type classifications, and pore sizes. A

high catalyst surface area can increase activity due to an increased number of available

active sites, while structure type and SiO2/Al2O3 ratio can significantly alter catalytic

activity and selectivity. Pore size is an important property because it can impact

accessibility of molecules within the pore structure, but very small pores have the

potential to cause diffusion limitations and reduce the effective catalyst activity. All

Zeolite used in this study were obtained from Zeolyst International. Reported

SiO2/Al2O3 ratios and surface areas were provided by the supplier.

Table 3.1.1: Properties of Zeolite catalysts used in preliminary activity screening.

SiO2/Al2O3 Surface Area Pore Zeolite Type Mole Ratio (m2/g) Diameter (Å) ZSM-5 50 425 5.1-5.6 Zeolite Y 30 780 7.4 Zeolite Y 5.2 660 7.4 Zeolite B 38 710 5.6-7.7

Prior to testing, the Zeolites were calcined in air at 500°C for four hours. Catalyst

activity was inferred by measuring the extent of cracking of n-dodecane at 350°C.

Reactor pressure was held constant at 500 psig by the addition of nitrogen gas, and n-

dodecane flow rate was set to 0.08 mL/min. With a catalyst bed volume of 1.6 mL

(diluted in 0.8 mL of SiC), liquid hourly space velocity (LHSV), defined as the liquid

volumetric flow rate divided by the active catalyst bed volume, was equal to 3 h-1. Due

to subsequent studies investigating a hydrogen-rich environment as a driving factor in

reducing deactivation rate, studies were performed with various levels of molecular

hydrogen co-feed with a constant n-dodecane feed rate. Hydrogen co-feed molar ratios

(defined as molar flow rate of hydrogen divided by molar flow rate of n-dodecane) were studied at nominal levels of 0, 1, 5, and 10. A chilled trap was used to collect condensable products while gaseous species were analyzed on-line. A gaseous nitrogen co-feed was included to create more uniform heat and mass transfer in the system. The nitrogen co-feed also provided a higher exit gas product flow rate and facilitated real-

time data gathering. Nitrogen co-feed flow rate was varied to maintain the overall

gaseous feed flow rate at 160 mL/min when hydrogen was also present in the feed.

The ZSM-5 catalyst showed the highest cracking activity with almost complete conversion of n-dodecane for all conditions studied, as shown in Figure 3.1.1. This behavior was consistent with similar studies.20 Due to the exceptionally high cracking

activity of the ZSM-5 catalyst, it would be difficult to investigate the impact of in-situ

hydrogen production on reactivity. Comparing activity of the other catalysts in Figure

3.1.1, Zeolite Y (SiO2/Al2O3 ratio: 5.2) exhibited a higher conversion rate than the other

two candidates at low hydrogen co-feed ratios. Zeolite B demonstrated a greater cracking

capability than the Zeolite Y (SiO2/Al2O3 ratio: 5.2) catalyst when a hydrogen co-feed ratio of 10 was present, however, this ratio would not be achieved via complete

32 dehydrogenation of decalin in a 1:1 volumetric blend with n-dodecane (utilized in subsequent studies).

1.00

0.90

0.80

0.70

0.60 dodecane ZSM-5 0.50 Zeolite Y (low Si/Al) 0.40 Zeolite Y (high Si/Al) 0.30 Zeolite B Conversion of n - of Conversion 0.20

0.10

0.00 0.0 1.0 5.0 10.0 hydrogen/n-dodecane flow ratio (mol/mol)

Figure 3.1.1: Conversion of n-dodecane with varying type of Zeolite and hydrogen/n-dodecane molar ratio at 350°C, 500 psig, and LHSV of 3 h-1.

Based on the Zeolite catalyst screening study, Zeolite Y (SiO2/Al2O3 ratio: 5.2) was selected for future studies with dual-bed reactor configurations. Prior to dual-bed studies, the Zeolite Y catalyst was studied at longer reaction run times without a hydrogen co-feed for determination of adequate overall time-on-stream needed for observable deactivation. This deactivation rate was used as a baseline for comparison to dual-bed configuration testing. It was also necessary to determine cracking rates of n- dodecane over Zeolite Y in the presence of cycloalkane derived dehydrogenation products, as these would be present in a dual-bed reactor system.

CHAPTER 4

DEHYDROGENATION OF DECALIN

The primary objective of this study was to determine decalin dehydrogenation functionality of a synthesized Pt/Al2O3 catalyst at 500 psig and a temperature range of

350-450°C in a continuous flow reactor at various Liquid Hourly Space Velocities

(LHSV). The results were used to determine an appropriate reactor temperature and

LHSV for use in subsequent n-dodecane cracking reactor configurations.

4.1 Experimental Conditions of Study Segments

Dehydrogenation functionality of a synthesized Pt/Al2O3 catalyst (Chapter 3) was

examined as a function of temperature and LHSV with decalin used as a model

dehydrogenation compound. For identification of optimal reactor temperature, testing

was performed with a ¼" reactor tube containing 1.60 mL (1.08g) of the catalyst diluted

in 3.2 mL of silicon carbide. Initial dehydrogenation testing was performed at 350, 400,

and 450°C at a pressure of 500 psig and decalin liquid flow rates of 0.04 and 0.02

mL/min. These liquid flow rates corresponded to LHSV values of 1.5 and 0.75 h-1,

respectively. A 164 mL/min nitrogen co-feed was present for the entire duration of the

liquid flow to ensure uniform heat and mass transfer. A 6.86 mL/min hydrogen co-feed,

corresponding to a 1:1 molar ratio of hydrogen and decalin at an LHSV of 1.5 h-1, was

also present to promote maximum catalytic activity.

Liquid and gaseous product samples were collected and analyzed during the

decalin dehydrogenation study. Liquid product samples were collected from the heated

product trap at the mid-point and completion of the test duration, while on-line gaseous

analysis was performed at five minute intervals. The liquid samples represented an

average product distribution for liquid-phase hydrocarbon species throughout each

sampling duration. The sampling times corresponded to 3 liquid volume changes over

the catalyst bed for a total 6 volume changes at each LHSV. The first sample collected at

each run condition was considered transient due to reactor startup and equilibration.

Following testing at each condition, decalin feed and nitrogen flow were shut off, and the

catalyst bed was maintained at 450°C for four hours in flowing hydrogen prior to cooling

to room temperature. This procedure was used to reactivate the catalyst bed via hydrogen

reduction and minimize the potential impact of deactivation.

Upon completion of temperature screening, a new ½" reactor tube containing the

same catalyst and diluent loading was prepared. Reactor diameter was increased since a

larger cross-sectional area was necessary for subsequent dual-bed configuration tests.

The same test procedure was used for testing at LHSV values of 0.75, 1.5, and 3.75 h-1,

operating at the previously determined temperature, for the purpose of identifying an

appropriate liquid flow rate for use in subsequent testing at that temperature. An ideal

liquid flow rate would maximize hydrogen production while minimizing catalyst

deactivation. Each condition was replicated with and without a 6.86 mL/min hydrogen

co-feed to determine if activity could be sustained in hydrogen-deficient conditions. For

testing without a hydrogen co-feed, an initial transient sample utilized a hydrogen co-feed

for the first 40 minutes of run time, in an attempt to minimize deactivation that could

35 occur during startup conditions. Table 4.1.1 describes the run schedule and sample collection times used during each screening segment.

Table 4.1.1: Run segments used for decalin dehydrogenation LHSV optimization at reaction conditions of 400°C and 500 psig.

Total Liquid Sample LHSV Hydrogen Liquid Run Segment Collection Times (Since (h-1) Co-Feed Flow Time Beginning of Segment) (min) 1 0.75 Full Duration 480 4h, 8h First 40 2 0.75 480 4h, 8h minutes 3 1.5 Full duration 240 2h, 4h First 40 4 1.5 240 2h, 4h minutes 5 3.75 Full Duration 120 1h, 2h First 40 6 3.75 120 1h, 2h minutes

4.2 Experimental Results

The conversion of decalin and selectivity towards tetralin and naphthalene were quantified as a function of temperature to characterize the corresponding performance in a ¼" reactor tube with a hydrogen co-feed, with results shown in Figure 4.2.1. At 400 and 450°C, decalin conversion was 94% and 99%, respectively, regardless of LHSV.

Selectivity towards naphthalene was very high at 0.94 moles/mole of decalin converted, while tetralin selectivity was low. This result indicated that the dehydrogenation preferentially proceeds to completion. This is consistent with a higher reaction rate of the tetralin dehydrogenation step compared to decalin dehydrogenation, as previously reported.10 Minor amounts of side reaction products were observed at 450°C, consisting primarily of aromatic species. At 350°C, conversion was 72% at an LHSV of 1.5 h-1, but increased to 78% at an LHSV of 0.75 h-1. This difference was likely due to a longer

36 residence time with the lower space velocity. Overall, a reaction temperature of greater than or equal to 400°C is required for nearly complete decalin conversion to naphthalene, for the specific reactor configuration. A reaction temperature of 400°C was selected for subsequent tests since higher temperatures could possibly result in catalyst deactivation via side reactions.8 Replicate data points were collected at 400 and 450°C for an LHSV of 1.5 h-1, with results indicating repeatability for decalin conversion within ±5%.

1.00 0.90 0.80 0.70 0.60 Decalin ConvserionConversion 0.50 Tetralin Selectivity 0.40 Naphthalene Selectivity LHSV = 1.5 h-1h-1 0.30 LHSV = 0.75 h-1h-1 Conversion or or Selectivity Conversion 0.20 0.10 (moles produced/mole decalin decalin (molesconverted) produced/mole 0.00 350 370 390 410 430 450 Temperature (°C)

Figure 4.2.1: Conversion of decalin and selectivity to tetralin and naphthalene over -1 Pt/Al2O3 as function of temperature at LHSV of 1.5 and 0.75 h and 500 psig.

Following identification of an appropriate reaction temperature, the next objective was to further evaluate the impact of LHSV on decalin conversion and naphthalene selectivity as well as to investigate the stability of dehydrogenation without a gaseous hydrogen co-feed in a ½" reactor tube. All other reaction conditions were identical to the temperature screening study, and results are shown in Figure 4.2.2. For LHSV values of

0.75 h-1 and 1.5 h-1, conversion was nearly complete, independent of hydrogen co-feed,

but was reduced to approximately 80% at an LHSV of 3.75 h-1. This indicates that an

LHSV of 1.5 h-1 or less was necessary for optimal decalin dehydrogenation in the current

reactor configuration. Selectivity towards naphthalene decreased and selectivity towards

tetralin increased with an increase in LHSV. This was likely due to an inability of the

reaction to fully proceed to equilibrium with a shorter residence time. Replicate data points were collected at LHSV values of 0.75 and 3.75 h-1 without a hydrogen co-feed,

with results indicating repeatability for decalin conversion within ±3%. A replicate for a

space velocity of 3.75 h-1 with a hydrogen co-feed indicated repeatability for decalin conversion within ±10%.

0.9

0.8

0.7

0.6 Decalin ConvsersionConversion 0.5 Tetralin Selectivity Naphthalene Selectivity 0.4 With H2H2 CoCo-Feed-Feed

0.3 Without H2H2 CoCo-Feed-Feed Conversion or or Selectivity Conversion

0.2

(moles produced/mole decalin decalin (molesconverted) produced/mole 0.1

0 0.5 1 1.5 2 2.5 3 3.5 4 LHSV (h-1)

Figure 4.2.2: Conversion of decalin and selectivity of products over Pt/Al2O3 as function of LHSV at temperature of 400°C and 500 psig.

Quantitation of the hydrogen produced via dehydrogenation of decalin can

provide additional insight into the selectivity of direct dehydrogenation compared to

formation of secondary reaction products. Hydrogen production was determined using

the average of the on-line data collected during each test segment, and is shown in Figure

4.2.3. For segments containing a hydrogen co-feed, hydrogen production in the exit stream was corrected via subtraction of the co-feed flow rate. For space velocities of

0.75 and 1.5 h-1, hydrogen production was very close to the ideal ratio of 5. This

confirms that at these lower flow rates, the reaction was not limited by residence time

restrictions and side reactions were minimal. However, at a higher space velocity,

hydrogen yield decreased. This was a result of lower decalin conversion and a higher

selectivity towards tetralin. There was very little deviation over the test duration for all

conditions studied with no indication of a decrease in activity over time, suggesting

minimal deactivation. The data also demonstrated that hydrogen generation rates were

very similar regardless of whether a hydrogen co-feed was present, indicating that it was

not a significant contributor to activity.

5 4.5 4 3.5 3 2.5 With Hydrogen 2 Co-Feed 1.5 Without Hydrogen Co-Feed 1 Produced/Mole Decalin Decalin Fed) Produced/Mole

Hydrogen Yield (Moles Hydrogen Hydrogen Yield (Moles Hydrogen 0.5 0 0.75 1.5 3.75 Liquid Hourly Space Velocity (h-1)

Figure 4.2.3: Hydrogen yield of decalin over Pt/Al2O3 as a function of LHSV at a temperature of 400°C and pressure of 500 psig.

Although there was limited deviation in hydrogen production observed during the

LHSV screening study, further investigation of catalytic stability over a longer run time

was desired. Figure 4.2.4 shows the hydrogen production for extended duration testing of

twelve hours performed at the two lower space velocities (0.75 and 1.5 h-1). These

studies were conducted at 400°C and 500 psig with a 6.86 mL/min hydrogen co-feed

present for the first hour of liquid flow. After the initial six and a half hours of run time, liquid feed was stopped and the reactor was cooled to room temperature overnight in a nitrogen blanket to prevent possible catalyst degradation. The following morning, the

reactor temperature was returned to 400°C and decalin flow was reinitiated after the

temperature equilibrated. This testing allowed observation of space velocity impact on

catalyst deactivation rate by monitoring hydrogen output, since a decrease in hydrogen

production would indicate a lower rate of dehydrogenation. During the extended

duration study, there was a minor decline in hydrogen production at a space velocity of

1.5 h-1 over a twelve hour test duration, decreasing by approximately 7%. For a space

velocity of 0.75 h-1, hydrogen production remained steady for approximately 500 minutes, before a gradual decline, resulting in approximately 34% reduction after twelve hours. This indicates that an LHSV of 1.5 h-1 would be more appropriate for extended test durations in subsequent experimentation. Initial hydrogen productions in this study were consistent with those observed in the LHSV screening study to within ±9%.

40 35 30 25 20 15

Exit (mL/min) Exit 10 5 0 Hydrogen FlowRate At Reactor 0 100 200 300 400 500 600 700 Time on Stream (min)

Figure 4.2.4: Hydrogen production via dehydrogenation of decalin over Pt/Al2O3 as a function of time and LHSV at a temperature of 400°C and pressure of 500 psig.

Multiple factors could have contributed to the observed deactivation during the

extended duration study. A longer residence time, inherent with the lower LHSV, could

allow for a larger extent of thermal cracking side reactions and coke precursor formation.

Alternatively, deactivation could have occurred via strong adsorption of naphthalene onto

catalytic active sites. Naphthalene has been shown to exhibit strong adsorption

tendencies towards platinum10, which could either prevent decalin from reaching the

active sites or potentially cause catalytic poisoning through structural rearrangements.1

Based on the impact of temperature and LHSV on decalin dehydrogenation activity and

stability, it was determined that reaction conditions at 400°C and 1.5 h-1 space velocity

are adequate for subsequent studies. These parameters are appropriate for satisfying

subsequent testing requirements for decalin dehydrogenation and hydrogen production,

while minimizing catalyst deactivation over an extended time-on-stream.

4.3 Primary Conclusions from Decalin Dehydrogenation Study

Testing was performed to identify suitable reactor conditions for decalin

dehydrogenation over synthesized Pt/Al2O3 catalyst to be used in subsequent evaluation

of coupled dehydrogenation and cracking reactions. Based on results from an initial

temperature screening, it was determined that a temperature of at least 400°C was

required to achieve nearly complete dehydrogenation to naphthalene. A lower reaction

temperature resulted in a significant reduction in decalin conversion and desired product

selectivity. A space velocity screening study at 400°C was then performed and the

results indicated that a value of less than 1.5 h-1 was required for high decalin conversion.

Catalytic activity for decalin dehydrogenation would correlate to hydrogen

production in the product gas stream, which was used as a measure of deactivation. A

long-duration study was conducted for LHSV values of 0.75 and 1.5 h-1 to determine if hydrogen production would decline with increasing time-on-stream. At an LHSV of 1.5 h-1, only minor deactivation was observed for the entire test duration. However,

deactivation did occur after 500 minutes with an LHSV of 0.75 h-1. Based on review of

the conditions studied during the dehydrogenation testing performed, it was determined

42 that a reaction temperature of 400°C and LHSV of 1.5 h-1 (with respect to the dehydrogenation catalyst bed), would be appropriate for use in subsequent dual- and mixed-bed studies investigating the coupling of catalytic dehydrogenation and cracking chemistries.

CHAPTER 5

DEHYDROGENATION OF DECALIN BLENDED WITH N-DODECANE

The dehydrogenation of cycloparaffins over noble metal catalysts can proceed

with high selectivity to produce molecular hydrogen and aromatics. However, other

chemical classes, such as n-paraffins, can also react over these catalysts. If multiple

chemical classes are combined in the feed stream, various factors, such as competition for

active sites, could impact the desired dehydrogenation of the cycloparaffin to produce

hydrogen. Decalin and n-dodecane were selected as model cycloparaffin and n-paraffin

compounds for investigating the impact of co-feeding on selective dehydrogenation.

5.1 Experimental Conditions of Study Segments

The impact of varying feed composition on the dehydrogenation functionality of

the synthesized Pt/Al2O3 catalyst was studied as a function of LHSV with a 1:1

volumetric ratio of decalin to n-dodecane (molar ratio of approximately 1.5:1). This

study was also used to provide insight into the effect of an n-paraffin addition on the

catalyst dehydrogenation functionality, which could potentially be reduced via active site

competition. During this testing, reactor conditions, including gaseous co-feeds, catalyst

loading, LHSV values, and total time-on-stream, were identical to those used for the decalin dehydrogenation LHSV screening study, presented in Chapter 4.

5.2 Experimental Results

The conversions of decalin and n-dodecane, as well as selectivity towards tetralin

and naphthalene, were quantified as a function of LHSV to characterize the

corresponding catalytic performance, and results are shown in Figure 5.2.1. Overall, the

fractional conversion of decalin was similar to studies without the n-paraffin co-feed

(Chapter 4), with observed values over 90% at lower LHSV values and a decline in conversion at 3.75 h-1. A high selectivity towards naphthalene indicates that an LHSV of

1.5 h-1 or less is sufficient for nearly complete decalin conversion when the

dicycloparaffin is blended with n-dodecane. This result suggests that there is minimal

competition for catalyst active sites at these conditions. At an LHSV of 1.5 h-1, the

corresponding n-dodecane conversion was at 20%, but increased at a lower LHSV.

However, the primary n-dodecane reaction pathway was isomerization rather than

cracking. This isomerization could potentially be beneficial to zeolite cracking activity in

subsequent dual-bed reactor studies, as the rate of β–scission is significantly increased

with isomerized paraffins.4

0.9

0.8

0.7

0.6 Decalin Conversion Tetralin Selectivity 0.5 Naphthalene Selectivity 0.4 Dodecane Conversion

0.3 With H2 Co-Feed Conversion or or Selectivity Conversion Without H2 Co-Feed 0.2

(moles produced/moles decalin decalin (molesconverted) produced/moles 0.1

0 0.5 1 1.5 2 2.5 3 3.5 4 LHSV (h-1)

Figure 5.2.1: Conversions of decalin and n-dodecane with selectivity of products over Pt/Al2O3 as function of LHSV at a temperature of 400°C and pressure of 500 psig.

Hydrogen yield, per mole of decalin fed, was analyzed as a function of LHSV,

shown in Figure 5.2.2. Similar to dehydrogenation of a neat decalin feed (Chapter 4),

hydrogen yield was essentially at the theoretical maximum of five moles produced per

mole of decalin fed. There was very little deviation over the test duration for all

conditions studied with no indication of a decrease in production over time, suggesting

minimal deactivation. It was also observed that hydrogen yield was nearly independent

of whether hydrogen was present as a co-feed. A long-duration experiment was not performed, but there was no observable decline in hydrogen production over the studied time periods. This steady trend, along with hydrogen yield at equivalent ratios to a neat decalin feed, indicated that decalin dehydrogenation functionality was not negatively impacted by n-dodecane addition.

5 4.5 4 3.5 3 2.5 With Hydrogen 2 Co-Feed Decalin Fed) Decalin

Hydrogen Yield 1.5 Without Hydrogen 1 Co-Feed

(Moles Hydrogen Produced/Mole Produced/Mole (Moles Hydrogen 0.5 0 0.75 1.5 3.75 Liquid Hourly Space Velocity (h-1)

Figure 5.2.2: Hydrogen yield of decalin/n-dodecane blend over Pt/Al2O3 as a function of LHSV at a temperature of 400°C and pressure of 500 psig.

5.3 Primary Conclusions from Decalin/n-Dodecane Dehydrogenation Study

The primary objective of this study was to evaluate the impact of n-dodecane co- feed addition on decalin dehydrogenation activity over a Pt/Al2O3 catalyst at 400°C and

500 psig. High conversion of decalin to naphthalene, with mild n-dodecane isomerization, were observed. At space velocities of 1.5 and 0.75 h-1, hydrogen yield was near the theoretical maximum for complete decalin dehydrogenation. This result was identical to hydrogen yield using a neat decalin feed (Chapter 4), indicating that catalytic dehydrogenation activity was minimally impacted by the n-paraffin addition. A similar hydrogen yield could also suggest that blend ratio is scalable at levels between 0:1 and

1:1 n-dodecane:decalin. Even though mild isomerization of n-dodecane occurred, the high hydrogen yield indicated that active site competition between the two reactants was minimal. Based on the reactor conditions studied, subsequent testing with an LHSV of

1.5 h-1 would be adequate for evaluation of coupled in-situ dehydrogenation and cracking in either dual- or mixed-bed catalyst configurations.

CHAPTER 6

DETERMINATION OF BASELINE DEACTIVATION RATE FOR N-DODECANE

CRACKING OVER ZEOLITE Y

The primary objective of this study was to determine deactivation rate and

product selectivity for n-dodecane cracking on Zeolite Y at 400°C and 500 psig in a

continuous flow reactor. This information was used to determine an appropriate time on

stream for comparison to subsequent catalyst deactivation experiments, which would

attempt to minimize the deactivation rate via reactor and feed manipulations.

6.1 Experimental Conditions of Study Segments

Experimental tests were performed using the fixed-bed flowing reactor system.

For extended duration tests, a reactor tube configuration containing 3.2 mL (0.96 g) of catalyst diluted in 1.6 mL of silicon carbide was used. Catalyst bed volume was twice

that used in the zeolite screening study (discussed in Chapter 3) to increase the initial

conversion rate and provide better sensitivity to the deactivation rate. Prior to loading the

reactor bed, the Zeolite Y catalyst was calcined in a furnace in static air at 500°C for four

hours. Testing was performed at 400°C and 500 psig with a n-dodecane flow rate of 0.04

mL/min. The reactor temperature and liquid flow rates were based on conditions determined during decalin dehydrogenation testing, discussed in Chapters 4 and 5.

Subsequent dual-bed reactor studies would require identical temperatures for

dehydrogenation and cracking beds. Therefore, a temperature of 400°C was used for this

study as that was found to be optimal for decalin conversion (in Chapter 4). The larger

catalyst bed and lower liquid flow rate, compared to conditions tested in the zeolite

screening study (Chapter 3), corresponded to a lower tested LHSV value of 0.75 h-1. This

LHSV value, with respect to the volume of zeolite catalyst, was the same used in

subsequent dual-bed and mixed-bed reactor studies. A 164 mL/min nitrogen co-feed was

present for the entire duration of testing, while a hydrogen co-feed of 32 mL/min was

used during initial run segments. Nitrogen was set to a constant flow rate for all run

segments, regardless of whether hydrogen was present as a co-feed, to provide an internal standard when calculating the total exit gas flow rate. A hydrogen flow rate of 32 mL/min corresponded to the amount of hydrogen produced during decalin dehydrogenation over the Pt/Al2O3 catalyst at a liquid feed flow rate of 0.04 mL/min

(determined in Chapter 4). Therefore, the hydrogen/n-dodecane ratio used in this study

was equivalent to the ratio produced during testing with a 1:1 volume ratio of n-dodecane

and decalin under ideal reaction conditions.

Three run segments were used to study cracking activity and deactivation, with

and without a hydrogen co-feed. In the first segment, hydrogen was co-fed to simulate

conditions that would be present with an upstream dehydrogenation stage, but without

decalin related byproducts present. An additional goal of initiating testing with a

hydrogen co-feed was to minimize catalyst deactivation when obtaining an estimate of

the maximum cracking rate when the catalyst was not limited by a hydrogen-deficient

environment. During subsequent testing without hydrogen present in the feed stream, it

was expected that catalyst deactivation would result in a loss of activity, which would

impact the potential for use of the zeolite as a cracking catalyst. Liquid product samples were collected from the chilled trap at both two and four hours after starting liquid flow, with on-line gaseous analysis every five minutes. The liquid samples represented an average product distribution for liquid-phase hydrocarbon species throughout each two- hour span. The initial two-hour sample was considered a transient sample due to stabilizing conditions during reactor startup, while the four-hour sample was indicative of steady-state performance. Following the four-hour sample collection, n-dodecane feed and nitrogen flow were shut off, and the catalyst bed was maintained at 450°C for four hours in flowing hydrogen prior to cooling to room temperature. This procedure was used to simulate the reactivation of a dehydrogenation catalyst, which could be implemented in a dual-bed configuration. The hydrogen reactivation was intended for a

dehydrogenation catalyst, but also had the potential to impact the zeolite catalyst, and

therefore was used for consistency. This first segment of testing, consisting of the first

four hours of time-on-stream, was used as the initial baseline data point where catalyst

activity was considered to be at a maximum, with no deactivation occurring.

The second segment of testing was used to determine the rate of catalyst

deactivation without a hydrogen co-feed. Initial operating conditions were similar to the prior segment, but hydrogen co-feed was terminated 40 minutes after n-dodecane feed was initiated. The initial hydrogen co-feed was used to minimize the potential for coke formation while initial reaction conditions stabilized. Liquid product samples were collected every two hours over an eight hour run time. Upon completion of this test duration, liquid feed was stopped and the reactor temperature was allowed to cool overnight to ambient temperature under a nitrogen blanket. The second segment was

representative of catalyst deactivation rate in hydrogen-deficient reaction conditions. The

third run segment used the same reaction conditions and methodology as the second

segment, but without an initial hydrogen co-feed. Liquid flow was re-initiated after the temperature had stabilized at the system set-point. Liquid samples were collected two and four hours after the liquid flow was initiated with the intent of adding additional data on deactivation rate in hydrogen-deficient conditions. This resulted in a total catalytic bed run time of twelve hours in the combined second and third segments of the experiment. Table 6.1.1 describes the run schedule and sample collection times used during each experiment segment.

Table 6.1.1: Run segments used for n-dodecane cracking testing over Zeolite Y at an LHSV of 0.75 h-1 and reaction conditions of 400°C and 500 psig.

Total Liquid Sample Collection Hydrogen Run Segment Day Liquid Times (Since Beginning of Co-Feed Flow Time Segment) Initial 1 Full Duration 4h 2h, 4h Extended First 40 2 8h 2h, 4h, 6h, 8h Duration Part 1 minutes Extended 3 N/A 4h 2h, 4h Duration Part 2

6.2 Experimental Results

The conversion of n-dodecane was quantified as a function of the overall time-on-

stream to characterize the corresponding activity and catalyst stability. Time-on-stream was defined as reaction time since the end of the first segment, corresponding to time in hydrogen-deficient reaction conditions. After an unstable startup period during the first sample of the first segment, stabilization in gas product composition was achieved, and therefore the second sample was considered to be representative of the steady maximum

52 catalyst activity, used as time zero. Figure 6.2.1 shows the n-dodecane conversion as a function of time-on-stream; with test segments labeled. Liquid-to-gas (LTG) conversion, defined as the mass percent of liquid reactant which was converted into gaseous products during the reactions, was also determined. An example LTG calculation is provided in the sample calculations appendix. This value was needed in calculating overall n-dodecane conversion because certain light-weight hydrocarbon products (C6 and below) formed from cracking reactions are not present in the liquid product sample and must be accounted for in the product gas stream to close the overall mass balance. The n-dodecane conversion was defined as the fraction of feed n-dodecane that was reacted to product species.

40 Segment 2 Segment 3 35

20 LTG 15 n-C12 Conversion (%) 10

0 0 100 200 300 400 500 600 700 Time on Stream (min)

Figure 6.2.1: Conversion of n-dodecane and Liquid-to-Gas (LTG) conversion over Zeolite Y as a function of time-on stream at an LHSV of 0.75 h-1, 400°C and 500 psig.

With a hydrogen co-feed, the n-dodecane conversion was approximately 34%.

This is higher than conversion observed in Chapter 3 due to the lower LHSV used in this

53 portion of the study. A gradual decrease in activity was observed following hydrogen cutoff, eventually reaching an asymptotic value of approximately 10% conversion after

12 hours. There appeared to be two rates of deactivation; the first being a rapid decrease in conversion (up to 240 minutes) followed by a subsequent slower deactivation rate.

Since coke is formed when hydrogen deficient compounds polymerize and condense, the absence of a hydrogen co-feed may promote a large increase in coke formation reactions.18 Coke deposition can cause a decline in catalytic activity, but another possible deactivation pathway would be a decrease in the rate of acid site proton regeneration, which requires sufficient hydrogen ion availability at the active sites. LTG conversion

(Figure 6.2.1) showed that roughly half of products were gaseous species, having a carbon number of C6 or lower. This ratio held as the overall conversion dropped. Light hydrocarbons formation from the parent n-dodecane requires cracking for formation, so these products indicate the prevalence of primary and possible secondary cracking reactions.

Product gas analysis showing selectivity over time is presented in Figure 6.2.2.

Gaseous products in the highest yield (with a carbon number under six) were isobutane, propylene, propane, and isopentane. The selectivity of these major gas species is normalized to n-dodecane conversion to determine if reaction pathways differed at later run times. Saturated hydrocarbon species are indicative of cracking and isomerization reactions while propylene, a hydrogen deficient compound, would indicate the potential

18 for coke precursor formation. Olefin species with a carbon number of C3 or higher could be produced via β–scission, but would protonate to form paraffins in hydrogen-rich environments.4

0.40

0.35

0.30

0.25

Dodecane Converted) Dodecane Propylene - 0.20 trans-2-Butene Isobutene 0.15 IsoPentane Product Selectivity Product Hexane 0.10

0.05 (Moles Produced/Moles n (Moles Produced/Moles 0.00 0 200 400 600 800 Time on Stream (min)

Figure 6.2.2: Gaseous hydrocarbon product selectivity over 120 minute intervals for n-dodecane over Zeolite Y at an LHSV of 0.75 h-1, 400°C and 500 psig.

The largest selectivity shift occurred with propylene, steadily rising from 0.12 to

0.33 moles produced per mole of n-dodecane converted over the 12 hour time-on-stream.

This trend suggests hydrogen-deficient cracking mechanisms were increasing as overall activity decreased, and is consistent with a reduction in activity due to lack of acid site proton regeneration. Similarly supporting this observation, isobutene and trans-2-butene selectivity values increased over time, while most iso-paraffin and n-paraffin selectivity values, such as those for isopentane and hexane, remained relatively steady. The high

LTG conversion values and increase in olefin selectivity without a hydrogen co-feed observed in this testing will be compared to dual and mixed bed data sets to determine if deactivation pathways contrast. A decrease in cracking activity with higher selectivity to

55 hydrogen-deficient cracked species could suggest a similar deactivation method involving acid-site protonation, while a decrease in cracking activity without a change in selectivity may indicate deactivation through coke deposition.

Liquid product analysis showed that isomerized paraffin species were produced in the highest yield, with lower selectivity towards aromatics, n-paraffins, cyclo-paraffins, and olefins, as shown in Figure 6.2.3. Liquid species include all hydrocarbons with a carbon number of 5 or higher and account for C5 and C6 species that remained solubilized in the liquid phase rather than exiting the system with the gaseous products.

The iso-paraffin class only included species with a carbon number below twelve, indicating that both cracking and isomerization occurred. Since isomerization occurs via skeletal rearrangement reactions on acid sites and cracking occurs via β–scission on acid sites4, an overall decline in n-dodecane conversion would be consistent with a decrease in available acid sites, possibly due to coke deposition or, similarly, a decrease in the number of protonated active sites. Iso-dodecane selectivity was relatively low for the entire test duration, which is consistent with low isomerization capabilities of monofunctional acid catalysts described in the literature.4 Aromatic, n-paraffin, and cyclo-paraffin selectivity values did not exhibit large changes during the overall time-on- stream, suggesting that the catalyst deactivation pathway did not significantly hinder their mechanism of formation. Liquid olefin selectivity increased without a hydrogen co-feed, but the rise was minimal compared to gaseous species. However, the trend of increasing olefin selectivity with decreasing iso-paraffin yield could be a result of decreasing catalytic protonation capabilities of β–scission products.

0.6

0.5

0.4 n-paraffins (C5-C11) Dodecane Converted) Dodecane

- Isoparaffins (C5-C11) 0.3 Aromatics (C6-C15) Cycloparaffins (C6-C9)

Product Selectivity Selectivity Product 0.2 Olefins (C5-C12) Isododecane (C12) 0.1 (Moles Produced/Mole n (Moles Produced/Mole 0 0 200 400 600 800 Time on Stream (min)

Figure 6.2.3: Liquid hydrocarbon product selectivity over 120 minute intervals for n-dodecane over Zeolite Y at an LHSV of 0.75 h-1, 400°C and 500 psig.

6.3 Overall Conclusions from n-Dodecane Cracking Study

Deactivation rate was examined for cracking of n-dodecane over a Zeolite Y catalyst. Based on the decline in n-dodecane conversion during experimentation without a hydrogen co-feed, the testing showed that there is a quantifiable decline in Zeolite Y catalytic activity for an extended duration study of 12 hours. This duration was determined to be sufficient for subsequent deactivation studies since a decrease in catalyst deactivation rate should be noticeably distinct from the drop observed in the present study.

In studying deactivation of the Zeolite Y catalyst for n-dodecane cracking in hydrogen-deficient conditions, it was found that a decrease in n-dodecane conversion

coincided with a decrease in cracked iso-paraffin species and an increase in olefin

product selectivity. This suggests a primary cracking reaction pathway via β–scission,

which was increasingly unable to protonate olefin products over the duration of the

studied time-on-stream. However, the increase in unsaturated hydrocarbons could have

also resulted in coke formation, which has been hypothesized to be the main pathway of

zeolite deactivation.1 This could mean that competing deactivation pathways of coke

deposition and loss of protonating capacity are occurring within the catalyst. The

increase in olefin selectivity can be used to compare differences in deactivation pathways

in subsequent reactor or feed configurations, with extent of olefin formation inversely

coinciding with protonation capability or possibly an alternate cracking mechanism.

CHAPTER 7

CRACKING OF N-DODECANE AND DECALIN BLENDED FEED

OVER ZEOLITE Y

The primary objective of this study was to determine the deactivation rate and product selectivity for cracking of a blended n-dodecane/decalin feed on Zeolite Y at

400°C and 500 psig in a continuous flow reactor. Impact of decalin addition on catalyst deactivation rate and product selectivity were compared to the neat n-dodecane cracking study (Chapter 6). In a dual-bed reactor configuration where decalin is dehydrogenated to naphthalene prior to entering a cracking stage, incomplete dehydrogenation would result in decalin feed to the cracking bed. The decalin could compete for active sites with the n-dodecane and possibly impact the selective cracking activity of the n-paraffin.

7.1 Experimental Conditions of Study Segments

The stability and deactivation rate of Zeolite Y cracking functionality with a blended n-dodecane/decalin feed were examined at 400°C. A blend containing a low percentage of decalin in n-dodecane was used as the feed. A 10% by volume (9.14% by mass) blend of decalin was used because high conversion rates of decalin to naphthalene were previously demonstrated in Chapters 4 and 5. Therefore, large amounts of decalin would not be expected to enter the cracking bed. The reactor conditions, including

gaseous co-feeds, catalyst loading, LHSV (0.75 h-1), and total time-on-stream, were consistent with those used for the neat n-dodecane cracking study discussed in Chapter 6.

7.2 Experimental Results

The conversions of n-dodecane and decalin were quantified as a function of the overall time-on-stream to characterize the impact of decalin addition to the reactant feed.

Consistent with the neat n-dodecane cracking study (Chapter 6), the first data point

contained a hydrogen co-feed (promoting maximum activity). Figure 7.2.1 shows the

n-dodecane, decalin, and Liquid-to-Gas (LTG) conversion as a function of time-on-

stream; the specific test segments are also identified. LTG conversion was defined as the

total percent mass of liquid reactant, including n-dodecane and decalin, which was

converted into gaseous products at sampling conditions.

100 Segment 2 Segment 3

60 LTG 40 n-C12

Conversion (%) Decalin 20

0 0 200 400 600 800 Time on Stream (min)

Figure 7.2.1: Conversion of n-dodecane and decalin in blended feed over Zeolite Y as function of time-on stream at LHSV of 0.75 h-1, 400°C, and 500 psig.

At the initial data point, where a hydrogen co-feed was present, the catalyst

converted approximately 19% of the feed into low molecular weight gaseous products,

with a n-dodecane conversion of approximately 37% and a decalin conversion of 86%.

Conversion of n-dodecane gradually declined after the hydrogen co-feed was terminated, eventually reaching an asymptotic value of approximately 20% conversion after 720 minutes. Decalin conversion remained stable for approximately 360 minutes, before declining to approximately 75% after 12 hours. The high conversion of decalin and minor decrease in activity was consistent with a previous study which reported monofunctional acid sites were shown to be very effective for cracking cycloparaffins.24

However, decalin conversion could have also occurred via dehydrogenation reactions,

which will be analyzed further. Decline in n-dodecane conversion could possibly be attributed to either coke formation or loss of acid site functionality in the absence of hydrogen.

Conversion of n-dodecane with the blended feed was compared to the baseline

conversion (Chapter 6) and is shown in Figure 7.2.2. The addition of decalin did not

cause a dramatic change in initial n-dodecane conversion, but decreased the observed deactivation rate. An explanation for a lower deactivation rate could be that dehydrogenation of decalin produced free molecular hydrogen, which could facilitate protonation of active sites and increase n-dodecane cracking activity, similar to the effect from utilizing a hydrogen co-feed. Additionally, cracked intermediates could interact with n-dodecane and help to promote faster thermal cracking reactions through propagation of carbonium or carbenium ions.15

n-Dodecane/Decalin Blend 20 Neat n-Dodecane Feed 15 Overall Conversion Conversion (%) LTG Conversion 10

0 0 200 400 600 Time on Stream (min)

Figure 7.2.2: Comparing n-dodecane and LTG conversion over time with a 9.14% decalin by mass blended feed to the baseline n-dodecane study at 400°C and 500 psig.

Based on LTG conversion values from Figure 7.2.2, light hydrocarbons were

created at a higher rate as compared to the Chapter 6 neat n-dodecane data, contributing to the higher n-dodecane conversion. LTG conversion for the blended feed was initially

19% and declined to 11% while neat n-dodecane cracking resulted in an initial LTG conversion of 16% which declined to 5%. This indicates that decalin addition has a net positive effect on light hydrocarbon production compared to cracking of only n-dodecane. Gas analysis indicated that the gaseous products with the highest selectivity in the blended feed were isobutane, followed by propane, isopentane, n-butane, and propylene, as shown in Figure 7.2.3. However, these species could be formed via decalin cracking as well. Hexane was produced at a higher rate when a hydrogen co-feed was

62 present, but dropped immediately after co-feed was stopped. This was possibly from high activity of decalin ring opening and cracking rather than production via n-dodecane cyclicization.22 Besides hexane, there were no significant shifts in gaseous product selectivity over the test duration. Propylene selectivity did not increase as significantly as observed during the neat n-dodecane study, suggesting lower coke precursor formation.

High iso-butane and propane selectivity could have been formed during decalin cracking, but propane formation could also be derived from propylene hydrogenation via in-situ generated hydrogen.

0.5

0.45

0.4

0.35

0.3 Dodecane Converted) Dodecane - Propylene 0.25 Propane 0.2 IsoButane

Product Selectivity Product 0.15 Hexane

0.1

0.05 (Moles Produced/Moles n (Moles Produced/Moles 0 0 200 400 600 800 Time on Stream (min)

Figure 7.2.3: Gaseous hydrocarbon product selectivity over time for n-dodecane/decalin blend over Zeolite Y at LHSV of 0.75 h-1, 400°C, and 500 psig.

Liquid product selectivity analysis, with respect to n-dodecane conversion, showed the highest selectivity was towards isomerized paraffins and aromatics, as shown

63 in Figure 7.2.4. The aromatic content was hypothesized to likely be derived from dehydrogenation and cracking of decalin, rather than from n-dodecane, since dehydrogenation reactions can occur on acid sites.24 Analysis of specific aromatic species with respect to decalin conversion will be subsequently discussed. Isoparaffin selectivity increased after the hydrogen co-feed was terminated and olefin selectivity remained low, indicating cracking functionality without major coke-precursor formation.

0.7 )

0.6

0.5

Dodecane Converted Dodecane 0.4 n-paraffins (C5-C11) - Isoparaffins (C5-C11) 0.3 Aromatics (C6-C15) Olefins (C5-C12) Product Selectivity Product 0.2 Isododecane (C12)

0.1 Moles Produced/Moles n Moles Produced/Moles ( 0 0 200 400 600 800 Time on Stream (min)

Figure 7.2.4: Liquid hydrocarbon product selectivity over time for n-dodecane/decalin blend over Zeolite Y at LHSV of 0.75 h-1, 400°C, and 500 psig.

For identification of decalin decomposition pathways, selectivity of hydrogen, aromatics, and cycloparaffins, with respect to decalin conversion, were examined, as shown in Figure 7.2.5. Selectivity towards aromatic products was higher than for cycloparaffinic products, indicating that dehydrogenation was a predominant conversion

pathway for decalin, rather of cracking of ring-opened products. The addition of these two compound classes indicates that they contributed to nearly all of the decalin conversion, and secondary ring opening (producing non-cyclic species) could not have been a significant reaction pathway. Hydrogen selectivity increased to approximately 1.7 moles produced per mole of decalin fed at 120 minutes on-stream, indicating partial dehydrogenation activity in the catalyst. Complete selective dehydrogenation to naphthalene would result in a hydrogen selectivity of 5 moles per mole of decalin converted. Since zeolites are known to hydrogenate naphthalene prior to ring opening21,

the reverse hydrogenation reaction can also occur until equilibrium is reached. Hydrogen

selectivity rapidly declined after 120 minutes, possibly suggesting a decrease in Zeolite Y

dehydrogenation functionality over time. However, stable selectivity towards aromatics

and cycloparaffins indicate that major reaction pathways were not changing as the

hydrogen selectivity dropped. Therefore, it is hypothesized that hydrogen was consumed

through secondary reactions, such as active site protonation.

1.8 ) 1.6

1.4

1.2

1 Hydrogen 0.8 Total Aromatics 0.6 Cycloparaffins (C6-C9) Product Selectivity Product

0.4

0.2 Moles Produced/Moles Decalin Converted Decalin Moles Produced/Moles ( 0 0 200 400 600 800 Time on Stream (min)

Figure 7.2.5: Hydrogen, total aromatic, and cycloparaffin product selectivity over time for n-dodecane/decalin blend over Zeolite Y at LHSV of 0.75 h-1, 400°C, and 500 psig.

To provide improved insight into the primary decalin dehydrogenation pathways, selectivity values for primary products likely formed via decalin dehydrogenation were compared individually with respect to moles of decalin converted, and are shown in

Figure 7.2.6. The largest selectivity was towards alkylbenzenes, followed by diaromatics and tetralin. Tetralin is formed through dehydrogenation of decalin, while alylbenzenes are formed from ring opening and cracking of tetralin.22 Naphthalene is formed from

further dehydrogenation of tetralin10, which could precede paraffin addition to form

alkylated diaromatics, such as methyl-naphthalene or dimethyl-naphthalene.

Alkylbenzene and diaromatic selectivities appeared to correlate with hydrogen selectivity

trends (Figure 7.2.5) since both of these classes require dehydrogenation during their

formation. Tetralin selectivity increased over time, opposing the trend of alkylbenzenes

and diaromatics. These product trends suggest that reactions involving ring opening of

tetralin to produce alkylbenzenes, or tetralin dehydrogenation to diaromatics, were

decreasing. However, hydrogen selectivity decreased more rapidly over time compared

to the decrease in those dehydrogenation products, further suggesting that it was

consumed in secondary reactions. An alternative deactivation pathway would be through

active-site poisoning, possibly due to high adsorption tendencies of products, which has been reported for aromatic species such as naphthalene.1

0.3 )

0.25

0.2

Alkylbenzenes (C6-C9) 0.15 C4-benzene (C10) Naphthalene Diaromatics (>C10)

Product Selectivity Product 0.1 Tetralin

0.05 Moles Produced/Moles Decalin Converted Decalin Moles Produced/Moles (

0 0 200 400 600 800 Time on Stream (min)

Figure 7.2.6: Aromatic species product selectivity over time for n-dodecane/decalin blend over Zeolite Y at LHSV of 0.75 h-1, 400°C, and 500 psig.

7.3 Primary Conclusions from n-Dodecane/Decalin Cracking Study

Deactivation rate was examined for cracking of a n-dodecane/decalin blended feed over a Zeolite Y catalyst. Compared to neat n-dodecane cracking data (Chapter 6), a n-dodecane/decalin blended feed demonstrated higher n-dodecane conversion with a lower rate of catalyst deactivation. It was speculated that dehydrogenation of decalin generated in-situ hydrogen, which contributed to protonation of the active sites. Normal-

and iso-olefin product selectivity did not increase dramatically with time-on-stream as it

had with a neat n-dodecane feed, indicating that β–scission unsaturated products were

becoming saturated, possibly at the newly protonated active sites. The low olefin

selectivity also indicated lower coke precursor formation, which could have partially

contributed to a slower deactivation rate. The results from this study suggest that in a

subsequent dual-bed reactor configuration, incomplete dehydrogenation of decalin

upstream of the cracking bed should not cause a significant negative impact on

n-dodecane cracking.

The addition of decalin in the feed stream resulted in the formation of hydrogen,

most likely produced via dehydrogenation and subsequent formation of aromatic species.

An analysis of possible decalin dehydrogenation-derived products demonstrated that

hydrogen formation correlated with alkylbenzene and diaromatic selectivities, which all

declined over time. An increase in selectivity towards tetralin during the same time

period then likely indicated a decrease in secondary dehydrogenation and tetralin ring-

opening reactions. Very low hydrogen observed in the gaseous product stream after 480

minutes, despite the presence of other dehydrogenation products, revealed that in-situ

hydrogen was likely being consumed in secondary reactions, such as active site

68 protonation. However, protonation was not sufficient to prevent overall catalyst deactivation during the duration of the study, as n-dodecane conversion still decreased.

Alternatively, certain aromatic species are known to have the potential to poison active sites due to their strong adsorption tendencies; this could have partially contributed to selectivity differences as time-on-stream progressed.1

CHAPTER 8

CRACKING OF N-DODECANE AND NAPHTHALENE BLENDED FEED

OVER ZEOLITE Y

The primary objective of this study was to determine the product selectivity for

cracking of a blended n-dodecane/naphthalene fed over Zeolite Y at 400°C and 500 psig

in a continuous flow reactor. Impact of naphthalene addition on the deactivation rate and

product selectivity of the cracking catalyst were compared to the neat n-dodecane

(Chapter 6) and the n-dodecane/decalin blended feed study (Chapter 7). In a dual-bed reactor configuration where decalin is dehydrogenated to naphthalene upstream of a cracking bed, the dehydrogenated compound would enter the cracking bed with n-dodecane as a co-feed. Naphthalene has previously been observed to exhibit high adsorption potential on catalyst active sites10, which could hinder n-dodecane adsorption

rate onto active sites and inhibit cracking activity.

8.1 Experimental Conditions of Study Segments

The stability and deactivation rate of Zeolite Y cracking functionality with a blended n-dodecane/naphthalene feed were examined at 400°C. In a dual-bed reactor

system utilizing a 1:1 volumetric blend of these two feed components, the mass

composition of n-dodecane and naphthalene would be roughly equal upon entering the cracking bed under complete dehydrogenation of decalin. However, a naphthalene blend

of 11.9% by mass was used for this study because this concentration was near the

maximum solubility limit of naphthalene in n-dodecane at room temperature. Reactor

conditions, such as LHSV, sampling schedule, and product analysis methods were

consistent with the Segment 1 cracking portion of the neat n-dodecane study (Chapter 6).

This segment included a hydrogen co-feed for the entire duration, with a total time-on- stream of four hours with liquid sampling at two hour intervals. Segments 2 and 3, studying hydrogen-deficient catalyst deactivation, were not replicated as naphthalene would not be present in a system without hydrogen, as they are co-products of decalin dehydrogenation.

Two levels of hydrogen co-feed were used to investigate the impact on catalytic activity using the n-dodecane/naphthalene blended feed. The hydrogen co-feed levels

(labeled as “High Hydrogen Co-Feed” and “Low Hydrogen Co-Feed” in all plots contained in the present chapter) were achieved using volumetric flow rates of 32 and 3.6 mL/min, respectively. This allowed comparison of the impact of naphthalene addition at various hydrogen levels, since in-situ produced hydrogen could be consumed via secondary reactions, such as protonation of active sites or hydrocarbon hydrogenation pathways. The lower hydrogen co-feed level (3.6 mL/min) was used as it corresponded to a 5:1 molar flow ratio with the 11.9% mass of naphthalene flow in the liquid feed. A

5:1 hydrogen:naphthalene molar ratio simulated complete dehydrogenation of decalin to naphthalene in a theoretical upstream dehydrogenation catalyst bed. The higher hydrogen co-feed level (32 mL/min) was consistent with that used in prior deactivation studies (Chapters 6 and 7), for comparisons to be made under similar reaction conditions.

8.2 Results and Conclusions

The conversions of n-dodecane and naphthalene were quantified as functions of hydrogen feed level to characterize the corresponding performance. This data was compared to Segment 1 data from prior n-dodecane cracking studies (Chapters 6 and 7) in Figure 8.2.1. The primary observation was that naphthalene appeared to decrease n-dodecane cracking functionality. At equivalent hydrogen co-feed levels, Zeolite Y converted approximately 34% of the neat n-dodecane feed (Chapter 6), but converted only 13% of the n-dodecane in a naphthalene-blended feed. Although n-dodecane conversion was low, naphthalene conversion was over 50% (naphthalene-derived products are subsequently discussed). The high naphthalene conversion indicates that diaromatic species could be occupying the catalyst active sites and inhibiting n-dodecane cracking. In a dual-bed system where the feed is comprised of around approximately

50% decalin, a higher naphthalene product concentration could potentially limit n-dodecane reaction activity further. Conversion of n-dodecane dropped slightly to approximately 11% when the hydrogen co-feed was lowered in the naphthalene-blended feed, while naphthalene conversion fell more significantly to 26%. Since naphthalene reactions occurring on zeolites have been shown to first include a rehydrogenation step21,

it would be expected that a reduction in hydrogen partial pressure would reduce the

naphthalene conversion. However, naphthalene adsorption would still potentially result

in n-dodecane adsorption inhibition, and literature indicating a strong adsorption potential

for naphthalene would be consistent with these assertions.10 This strong adsorption

potential could possibly contribute to catalyst poisoning in a dual-bed or mixed-bed reactor configurations.1

40 Neat n-Dodecane Feed (High Hydrogen Co-Feed) n-Dodecane/Decalin Blended Feed 30 (High Hydrogen Co-Feed) n-Dodecane/Naphthalene Blended Feed (High H2 Co-Feed) Conversion (%) 20 n-Dodecane/Naphthalene Blended Feed (Low H2 Co-Feed)

0 n-Dodecane Naphthalene

Figure 8.2.1: Catalytic activity over time during dodecane cracking testing with a feed containing 9.14% by mass of decalin.

Light molecular weight gaseous product yield demonstrated a decline in normal- and isoparaffin production when the hydrogen co-feed level was lowered in the n-dodecane/naphthalene blend, as shown in Figure 8.2.2. If the difference were simply a result of hydrogenation of hydrogen-deficient olefin products, it would be expected that olefin production would increase at the lower hydrogen co-feed level. However, olefin yield was consistent among the two data sets, indicating that the difference in paraffin yield was due to higher overall cracking functionality rather than hydrogenation of product olefins. This would suggest that excess hydrogen in the system is increasing active site cracking activity through direct interaction, possibly through protonation.

0.0018

0.0016

0.0014

0.0012

0.001 High Hydrogen Co-Feed 0.0008 Low Hydrogen Co-Feed

0.0006 Product Product Yield (Moles)

0.0004

0.0002

0 n-Paraffin Iso-Paraffin Olefin

Figure 8.2.2: Gaseous hydrocarbon product yield for n-dodecane/naphthalene blend over Zeolite Y at LHSV of 0.75 h-1, 400°C, and 500 psig.

Liquid product analysis revealed high yield of aromatics (neglecting naphthalene)

and isoparaffins, as shown in Figure 8.2.3. Isoparaffins were likely derived from

n-dodecane, as observed during neat n-dodecane cracking (Chapter 6). The high level of

aromatic production was most likely derived from naphthalene, since it is itself an

aromatic compound. Specific naphthalene decomposition pathways are examined in

more detail in a subsequent discussion. When the hydrogen co-feed level was reduced,

isoparaffin yield decreased, which was consistent with the observed impact on gaseous

products. Similarly, aromatic yield decreased when the hydrogen co-feed level was reduced, which is consistent with the naphthalene conversion data. Very low yield of cycloparaffins suggests that complete hydrogenation of naphthalene to decalin was not favorable, even with the higher hydrogen co-feed present.

0.002

0.0018

0.0016

0.0014

0.0012

0.001 High Hydrogen Co-Feed 0.0008 Low Hydrogen Co-Feed

Product Product Yield (Moles) 0.0006

0.0004

0.0002

0 n-paraffins Isoparaffins Aromatics, Cycloparaffins Olefins (C5- Isododecane (C5-C11) (C5-C11) Neglecting (C6-C11) C12) (C12) Naphthalene (C6-C15)

Figure 8.2.3: Liquid hydrocarbon product yield for n-dodecane/naphthalene blend over Zeolite Y at LHSV of 0.75 h-1, 400°C, and 500 psig.

To better quantify naphthalene decomposition pathways, selectivity values for aromatic products with respect to naphthalene conversion were compared individually, as shown in Figure 8.2.4. The largest compound class was alkylnaphthalenes, with alkylbenzenes being the only other significant compound class. Alkylnaphthalenes differ from naphthalene only in that hydrocarbon chains are substituted for C-H bonds, while alkylbenzenes could be formed via dehydrogenation to tetralin following by ring opening and cracking. A large increase in alkylnaphthalene selectivity with lowered hydrogen co- feed was observed, indicating that reactions involving hydrocarbon addition to naphthalene were favored at lower hydrogen concentration. It has previously been demonstrated that Zeolite Y facilitates alkylation of naphthalene in hydrogen-deficient conditions.26,27 The decrease in paraffin selectivity observed in Figures 8.2.2 and 8.2.3

75 with a concurrent increase in alkylnaphthalene selectivity could suggest that the feed hydrogen is involved with an alkylation terminating mechanism. C4-benzene selectivity, formed via ring opening of tetralin, was much lower than selectivity towards lower molecular weight alkylbenzenes. This indicates a high level of cracking activity with respect to paraffinic chains attached to benzene, consistent with literature.22 Low cycloaromatic selectivity indicates that further hydrogenation of tetralin or alkylbenzenes was not a significant reaction.

0.7

0.6

0.5

0.4

0.3

Product Selectivity Product 0.2 High Hydrogen Co-Feed Low Hydrogen Co-Feed 0.1 (Moles/Mole Naphthalene Naphthalene (Moles/Mole Converted) 0

Figure 8.2.4: Aromatic product selectivity for n-dodecane/naphthalene blend over Zeolite Y at LHSV of 0.75 h-1, 400°C, and 500 psig.

8.3 Primary Conclusions from n-Dodecane/Naphthalene Cracking Study

The primary conclusion from this study was that naphthalene can have a significant inhibition impact on cracking of n-dodecane with the tested Zeolite Y catalyst.

Even with a high hydrogen co-feed, the addition of naphthalene caused n-dodecane conversion to drop from 34% to 13%. With a lower hydrogen co-feed level, n-dodecane conversion was suppressed further to 11%. This inhibition was determined to likely be from high adsorption potential of naphthalene, which decreased the number of available active sites for n-dodecane to react on.

Naphthalene conversion was approximately 52% when using the higher hydrogen co-feed, but decreased to 26% with the low hydrogen co-feed. This result indicates that free molecular hydrogen concentration has a significant effect on naphthalene conversion. Literature suggests that naphthalene reactions must first undergo a hydrogenation reaction with one of the aromatic rings21, the rate of which would be hindered by low amounts of hydrogen in the system.

Analysis of light molecular weight gaseous hydrocarbon products indicated that paraffin product yield was decreased with a lower hydrogen co-feed, but olefin yield remained constant. This would suggest that the excess hydrogen was not involved with directly hydrogenating olefin products, but rather helped to increase selective cracking rate at the active sites, possibly through protonation. Liquid product analysis indicated that overall naphthalene-related aromatic product yield was reduced with a lower hydrogen co-feed, but alkylnaphthalene selectivity was increased. An increase in

77 alkylnaphthalene selectivity at lower hydrogen concentration could indicate the importance of hydrogen being used to terminate these reactions.

The significant reduction in n-dodecane cracking observed when naphthalene was added to the feed should be taken into consideration when analyzing results from a dual- bed system. The feed in this study only contained a small amount of naphthalene, but if that is increased to 50%, n-dodecane cracking functionality could be severely hindered.

Even though this study did not investigate naphthalene effects at long times-on-stream, a high naphthalene concentration could also cause a large shift in catalyst selectivity through eventual active-site poisoning, as was speculated to occur in the n-dodecane/decalin cracking study (Chapter 7). These impacts may negate any intended positive effect from hydrogen produced in the first catalyst bed of a dual-bed reactor system.

CHAPTER 9

DEHYDROGENATION/CRACKING OF N-DODECANE AND DECALIN BLENDED

FEED IN DUAL-BED REACTOR

The primary objective of this study was to evaluate cracking activity and

durability of Zeolite Y, utilizing a blended n-dodecane/decalin feed with a catalytic dehydrogenation bed located upstream of the cracking bed, both at 400°C and 500 psig.

Impact of decalin addition on catalyst deactivation rate and product selectivity were compared to results observed for the baseline neat n-dodecane cracking study (Chapter

6). The availability of molecular hydrogen from decalin dehydrogenation was hypothesized to increase cracking functionality and minimize activity loss during the test duration.

9.1 Experimental Conditions of Study Segments

The cracking activity and deactivation rate of Zeolite Y with an upstream

Pt/Al2O3 dehydrogenation zone, was examined at 400°C and 500 psig, when utilizing a

blended n-dodecane/decalin feed. A blend containing a 1:1 volumetric ratio of

n-dodecane and decalin was selected with the goal of generating in-situ hydrogen from

the dicycloparaffin to be utilized with n-dodecane in the cracking zone. Results from a blended feed with only a dehydrogenation bed (Chapter 5) indicated that nearly complete dehydrogenation of decalin to naphthalene should occur in the dehydrogenation portion

of the dual-bed system, with mild isomerization of n-dodecane also occurring. Chapter 6 results established a baseline deactivation rate for a pure n-dodecane feed over Zeolite Y,

however, it was observed in Chapter 8 results that naphthalene significantly inhibited

n-dodecane cracking on the catalyst. This study will determine if dehydrogenated decalin

is beneficial to cracking as a hydrogen donor, or if the effects are ultimately negated by

naphthalene co-production.

The loading and dilution of the Pt/Al2O3 bed was identical to that used in the

decalin dehydrogenation studies, described in Chapters 4 and 5, while loading of the

Zeolite Y cracking bed was identical to that used in the baseline n-dodecane cracking study (Chapter 6). A small amount of SiC separated the two catalyst beds to prevent mixing and to facilitate uniform mass transfer between reaction zones. The remaining reactor conditions, including gaseous co-feeds, LHSV (0.75 h-1 with respect to the Zeolite

catalyst volume), and total time-on-stream, were also consistent with those used for the neat n-dodecane cracking study.

9.2 Experimental Results

The conversions of n-dodecane and decalin were quantified as a function of the overall time-on-stream to characterize the impact of a dual-bed configuration on n-dodecane cracking. Consistent with the neat n-dodecane cracking study (Chapter 6), the first data point contained a hydrogen co-feed (promoting maximum activity). Figure

9.2.1 shows the n-dodecane, decalin, and LTG conversion as a function of time-on- stream; the specific test segments are also identified and are consistent with the segments described in Chapter 6. LTG conversion was defined as the total mass percent of liquid

80 reactant, including n-dodecane and decalin, which was converted into gaseous products at sampling conditions.

100

Segment 2 Segment 3 60 LTG 40 n-C12

Conversion (%) Decalin 20

0 0 200 400 600 800 Time on Stream (min)

At the initial data point, where a hydrogen co-feed was present, the catalyst converted approximately 7% of the feed into low molecular weight gaseous products, with a minimal n-dodecane conversion of less than 1% and a decalin conversion of nearly

100%. Conversion of n-dodecane increased after the hydrogen co-feed was terminated, remaining between 16% and 20% for all data points at or following 240 minutes. Decalin conversion remained stable for the first 240 minutes, before declining to approximately

76% at 480 minutes. The decalin conversion then steadily increased for the remainder of the study, resulting in approximately 84% conversion after 720 minutes (12 hours). The initial low n-dodecane conversion is consistent with suppression due to the presence of naphthalene, which was observed in the n-dodecane/naphthalene cracking study, discussed in Chapter 8. The increase in n-dodecane conversion could indicate that

naphthalene inhibition was being reduced, which could be due to a shift in catalyst

product selectivity and functionality. The initial 7% LTG conversion, with negligible

n-dodecane conversion, suggests that some of the decalin conversion was related to

cracking reactions rather than being completely selective towards complete

dehydrogenation to naphthalene.

Conversion of n-dodecane with the dual-bed reactor configuration was compared to the baseline conversion (Chapter 6) and is shown in Figure 9.2.2. The increase in

conversion observed in the dual-bed reactor system is clearly differentiated from the decreasing trend of the baseline study. Although n-dodecane conversion was initially suppressed, conversions were nearly identical at 240 minutes, with the dual-bed system having superior n-dodecane conversion at longer run times. This discrepancy demonstrates that Zeolite Y cracking mechanisms were significantly altered by the upstream dehydrogenation catalyst bed. If a desired outcome is to increase n-dodecane conversion, the test results suggest that this could be optimized at longer run times by using a dual-bed configuration.

20 Single Bed (cracking) with 15 C12 Dual Bed with 10 C12+Decalin Dodecane Conversion (%) Conversion Dodecane - n 5

0 0 200 400 600 800 Time on Stream (min)

Figure 9.2.2: Comparing n-dodecane conversion over time for a dual-bed reactor with blended feed to the baseline n-dodecane study at 400°C and 500 psig.

For investigation of decalin conversion pathways, selectivity of hydrogen (right axis) along with aromatics and cycloparaffins (left axis), with respect to decalin conversion, were examined, as shown in Figure 9.2.3. Hydrogen selectivity was initially approximately 2 moles per mole of decalin converted, but increased to near the theoretical maximum of 5 moles produced at 120-240 minutes. Similarly, aromatic selectivity (including naphthalene) was initially at 0.84 and increased to 0.95 and 0.97 at

120 and 240 minutes, respectively. The lower initial aromatic product selectivity is consistent with the presence of observed gaseous products. This indicates that selectivity of decalin reaction pathways initially included cracking. The hydrogen co-feed present during the initial data point could have shifted decalin equilibrium from dehydrogenation products and allowed cracking to occur. After 240 minutes, aromatic and hydrogen selectivity began to decrease. However, hydrogen selectivity decreased much more

rapidly compared to aromatic production, suggesting hydrogen utilization in secondary

reactions, such as n-dodecane hydrocracking. Also, as aromatic selectivity decreased,

cycloparaffin selectivity increased. This indicates a shift in product selectivity towards

cracked species, which is consistent with the observed increase in n-dodecane

conversion.

1.2 6 )

1 5 Hydrogen Selectivity 0.8 4

0.6 3 Total Aromatics Cycloparaffins (C6-C9) 0.4 2 Hydrogen Product Selectivity Product

0.2 1 Moles Produced/Moles of Converted Decalin Moles Produced/Moles ( 0 0 0 200 400 600 800 Time on Stream (min)

Figure 9.2.3: Hydrogen, total aromatic, and cycloparaffin product selectivity over time for n-dodecane/decalin blend over dual-bed reactor system at 400°C and 500 psig.

Determining the specific decalin dehydrogenated product content can help to

understand impact on n-dodecane cracking functionality. Aromatic product selectivity,

with respect to decalin conversion, was examined, as shown in Figure 9.2.4. Naphthalene

was the only significant aromatic product during the first 240 minutes. After 240

minutes, tetralin, alkylbenzene, and C4-benzene selectivities began to increase. Low diaromatic selectivity suggests that alkylation was not a significant reaction, in contrast to

results from the blended n-dodecane/naphthalene study, described in Chapter 8. The

previous study had utilized only a small amount naphthalene in the feed, which would be

approximately one fifth the concentration in the hydrocarbon mixture entering the

cracking bed in the present dual-bed study. A higher naphthalene concentration could have contributed to vastly different reaction pathways on the Zeolite Y catalyst. Also, the blended n-dodecane/naphthalene study was performed with a shorter total duration, so long-term effects from naphthalene addition were not determined. Current results indicate that the proposed catalyst selectivity shift towards cracking was possibly related to the observed hydrogen consumption, which could have hydrogenated aromatic species

and initiated hydrocracking reactions.

) 1 0.9 0.8 0.7

0.6 Alkylbenzenes (C6-C9) 0.5 C4-benzene (C10) 0.4 Naphthalene Diaromatics (>C10)

Product Selectivity Product 0.3 Tetralin 0.2 0.1 Moles Produced/Moles Decalin Converted Decalin Moles Produced/Moles ( 0 0 200 400 600 800 Time on Stream (min)

Figure 9.2.4: Aromatic product selectivity over time for n-dodecane/decalin blend over dual-bed reactor system at 400°C and 500 psig.

LTG conversion was initially low, but began to increase after 240 minutes, which

was likely related to increasing conversion of n-dodecane. This can be observed by

quantifying selectivity towards gaseous products species with respect to n-dodecane conversion, as shown in Figure 9.2.5. Data prior to 240 minutes is not included, as minimal n-dodecane conversion during this time results in high uncertainty in selectivity values. All major gaseous species were normal- or iso-paraffins. Low olefin selectivity is consistent with the hydrogen consumption observed in Figure 9.2.3. This indicates that hydrogen was being utilized by secondary reaction pathways and successfully facilitated higher Zeolite Y cracking activity at the later data points.

2.5

1.5

Dodecane Converted) Dodecane Propylene - Propane 1 IsoButane n-Butane Product Selectivity Product IsoPentane 0.5

(Moles Produced/Moles n (Moles Produced/Moles 0 0 200 400 600 800 Time on Stream (min)

Figure 9.2.5: Gaseous hydrocarbon product selectivity over time for n-dodecane/decalin blend over dual-bed reactor system at a Zeolite LHSV of 0.75 h-1, 400°C, and 500 psig.

Liquid product analysis, similar to gaseous product data, revealed negligible

olefin selectivity, as shown in Figure 9.2.6. This result is consistent with the observed

hydrogen consumption, as hydrogen-deficient products were not selectively produced.

Significant isododecane selectivity was observed at 240 minutes, suggesting that minimal

cracking occurred at this data point. This is consistent with the low gaseous hydrocarbon

product selectivity. However, isododecane selectivity decreased as time-on-stream progressed, while lower molecular weight isoparaffin selectivity increased, which indicates that cracking functionality on the catalyst was increasing during this time. If the Zeolite Y catalyst was deactivating, cracked isoparaffin selectivity would not be expected to increase. This liquid product data along with gaseous product analysis supports a conclusion that, despite an initial suppression of dodecane cracking, the dual-

bed configuration does reduce the rate of Zeolite Y catalyst deactivation at 720 minutes

of run time.

0.9

0.8

0.7

0.6 Dodecane Converted) Dodecane - n-paraffins (C5-C11) 0.5 Isoparaffins (C5-C11) 0.4 Olefins (C5-C12)

Product Selectivity Product 0.3 Isododecane (C12)

0.2

0.1

(Moles Produced/Mole n (Moles Produced/Mole 0 0 200 400 600 800 Time on Stream (min)

Figure 9.2.6: Liquid hydrocarbon product selectivity over time for n-dodecane/decalin blend over dual-bed reactor system at a Zeolite LHSV of 0.75 h-1, 400°C, and 500 psig.

One possible cause of the apparent shift in catalyst activity and selectivity would

be through active site poisoning, which has the potential to cause a change in

functionality via alteration of catalyst framework structures.1 For poisoning to occur, a species must have strong adsorbing potential. It has already been observed in the n-dodecane/naphthalene study (Chapter 8) that naphthalene likely adsorbs at a much higher rate than n-dodecane. It is also known through literature that aromatics have a strong catalytic adsorption tendency.10 Therefore, it would be reasonable to assume that

naphthalene strongly adsorbed to the Zeolite Y catalyst, which then caused a structural

change and altered reaction pathways. Since these altered pathways were favorable with

respect to n-dodecane conversion, future experimentation could work to identify a pre-

treatment of the catalyst that causes a similar alteration prior to use in a reactor.

9.3 Primary Conclusions from n-Dodecane/Decalin Dual-Bed Study

Compared to the baseline n-dodecane cracking study (Chapter 6), utilizing a n-dodecane/decalin blended feed with the tested dual-bed reactor resulted in a lower initial n-dodecane conversion, but higher conversion at long run times. The dual-bed reactor system also resulted in a lower level of olefin production, indicating that hydrogen was successfully being utilized to enhance cracking rate and/or saturate the cracked products. Decalin production was very high through the entire run but there was an apparent shift in product selectivity from complete dehydrogenation to cracked species as run time increased. A possibility of the apparent shift in Zeolite Y reaction products during testing could have been due to active site poisoning, which is known to alter catalyst functionality even without deactivation in overall activity.1 Overall, the data

suggests that the dual-bed reactor system was able to enhance Zeolite Y cracking

functionality after 720 minutes compared to the baseline, even though initial activity was

suppressed. If any structural changes could be identified, it may be possible to specify a

pre-treatment to induce these changes before utilization in a cracking reactor system.

CHAPTER 10

DEHYDROGENATION/CRACKING OF N-DODECANE AND DECALIN BLENDED

FEED IN MIXED-BED REACTOR

The primary objective of this study was to evaluate cracking activity and

durability of Zeolite Y at 400°C and 500 psig, utilizing a blended n-dodecane/decalin

feed, with a catalytic bed comprised of physically mixed dehydrogenation and cracking

catalysts. Impact of bed mixing on catalyst activity and product selectivity were

compared to results observed with the neat n-dodecane cracking study (Chapter 6), the blended n-dodecane/decalin cracking study (Chapter 7), and the dual-bed reactor study

(Chapter 9). An initial cracking inhibition observed in the dual-bed reactor study was

speculated to be due to naphthalene adsorption on the Zeolite Y. By mixing the

dehydrogenation and cracking beds rather than placing them in series, it was

hypothesized that inhibition could be minimized by allowing naphthalene production to

occur over the full volume of the cracking catalyst, rather than introducing a high

concentration at the onset of the bed. Decalin addition was previously shown to increase

n-dodecane cracking activity in the blended feed cracking study (Chapter 7). Therefore,

the decalin could potentially provide cracking enhancement benefits in addition to

benefits from hydrogen production.

10.1 Experimental Conditions of Study Segments

The cracking activity and deactivation rate of Zeolite Y physically mixed with

Pt/Al2O3 was studied at 400°C and 500 psig, utilizing a 1:1 volumetric blend of

n-dodecane and decalin. It was hypothesized that dehydrogenating decalin in close

proximity with n-dodecane cracking would reduce the absolute local naphthalene

concentration, while generating in-situ hydrogen for use in cracking. It is possible that

the presence of cracked products at the Pt/Al2O3 catalyst could promote deactivation as

previous studies did not investigate the corresponding impact on dehydrogenation

activity.

The loading and dilution of the Pt/Al2O3 bed was identical to that used during

decalin dehydrogenation studies (Chapters 4 and 5), while the Zeolite Y cracking bed

loading was identical to that used in the baseline n-dodecane cracking study (Chapter 6).

The remaining reactor conditions, including use of gaseous hydrogen and nitrogen co- feeds, LHSV (0.75 h-1 with respect to the Zeolite catalyst volume), and total time-on- stream, were also consistent with those used for the neat n-dodecane cracking study.

These reactor conditions and catalyst loadings were therefore also consistent with these from the dual-bed reactor study (Chapter 9).

10.2 Experimental Results

The conversions of n-dodecane and decalin were quantified as a function of the overall time-on-stream to characterize the impact of a mixed-bed configuration on n-dodecane cracking activity. Consistent with the neat n-dodecane cracking study

(Chapter 6), the first data point contained a hydrogen co-feed (promoting maximum

activity). Figure 10.2.1 shows the n-dodecane, decalin, and LTG conversions as a

function of time-on-stream. LTG conversion was defined as the total mass percent of

liquid reactants which was converted to gaseous products at sampling conditions.

100

60 LTG 40 n-C12

Conversion (%) Decalin

0 0 200 400 600 800 Time on Stream (min)

Figure 10.2.1: Conversion of n-dodecane and decalin in blended feed over mixed -1 Pt/Al2O3 and Zeolite Y as a function of time-on stream at an LHSV of 0.75 h (with respect to the Zeolite Y volume) and reaction conditions of 400°C and 500 psig.

At the initial data point, where a hydrogen co-feed was present, the catalyst

converted approximately 20% of the feed to low molecular weight gaseous products, with

a n-dodecane conversion of 53% and a decalin conversion of 91%. Conversion of

n-dodecane decreased after the hydrogen co-feed was terminated, reducing to approximately 24% conversion after 720 minutes. Decalin conversion remained relatively stable for the entire run, remaining at approximately 89% after 720 minutes.

The initially high n-dodecane conversion suggests that inhibition observed in the dual- bed study was not present when using a mixed-bed configuration. In fact, this was the

highest recorded n-dodecane conversion for any reactor configuration. However,

cracking decreased after the hydrogen co-feed was terminated, which could mean that in-

situ hydrogen generation was not sufficient to maintain maximum activity.

For investigation of decalin conversion pathways, selectivity of aromatics and

cycloparaffins (left axis), as well as hydrogen (right axis), with respect to decalin

conversion, were examined, as shown in Figure 10.2.2. Hydrogen selectivity was

initially at 2.4 moles per mole of decalin converted and decreased over the run duration,

resulting in approximately 0.3 moles per mole of decalin converted after 720 minutes.

Aromatic selectivity was initially at 0.7 and decreased after 120 minutes to less than 0.5

at 720 minutes. Hydrogen selectivity for the mixed-bed reactor never reached the

theoretical maximum of 5 as it did in the dual-bed reactor, which was expected due to

cracking of the feed decalin on Zeolite Y. However, the decrease in hydrogen selectivity to nearly zero, while aromatic selectivity did not completely disappear, suggests that hydrogen was being utilized in secondary reactions, as also speculated in the dual-bed

study. Similar to the dual-bed study results, as aromatic selectivity decreased,

cycloparaffin selectivity increased. This could be due to deactivation of the

dehydrogenation catalyst, or a result from naphthalene-induced structural change of the

Zeolite Y catalyst, as had been hypothesized to occur in the dual-bed reactor. Addition of

cycloparaffin and aromatic selectivities only sum to approximately 0.6 after 240 minutes,

indicating that remaining decalin products should result in increased normal and

isomerized paraffin/olefin production in gaseous or liquid products.

) 1 5 0.9 4.5 0.8 4 Hydrogen Selectivity 0.7 3.5 0.6 3

0.5 2.5 Total Aromatics 0.4 2 Cycloparaffins (C6-C9) Hydrogen

Product Selectivity Product 0.3 1.5 0.2 1 0.1 0.5 Moles Produced/Moles Decalin Converted Decalin Moles Produced/Moles ( 0 0 0 200 400 600 800 Time on Stream (min)

Figure 10.2.2: Hydrogen, total aromatic, and cycloparaffin product selectivity over time for n-dodecane/decalin blend over mixed-bed reactor system at 400°C and 500 psig.

Detailed characterization of the specific decalin dehydrogenated product content can help to understand impact on n-dodecane cracking functionality. Aromatic product selectivity, with respect to decalin conversion, is shown in Figure 10.2.3. Decalin conversion was high and stable for the entire duration, but naphthalene and tetralin selectivity values were relatively low and do not close the mass balance for decalin products. This indicates competition between cracking/ring opening on Zeolite Y and dehydrogenation on Pt/Al2O3. Low selectivity for all aromatics indicates that decalin conversion products should result in an increase in gaseous or liquid paraffin/olefin products. Also, the results demonstrated a higher selectivity towards tetralin than for naphthalene. Since high selectivity towards naphthalene is known to occur over

Pt/Al2O3, this suggests that there was either a high rate of naphthalene reaction occurring

on the Zeolite Y, or very fast primary reaction of decalin on the Zeolite compared to the

dehydrogenation catalyst. This interaction could have caused a poisoning effect similar

to what was speculated to occur in the dual-bed reactor study. Constant decalin conversion over the entire run duration could possibly be explained by higher selectivity towards direct decalin cracking at longer run times, but this increase in decalin cracking came at the expense of n-dodecane cracking.

0.25 )

0.2

0.15 Tetralin Naphthalene 0.1 Alkylbenzenes (C6-C9) Diaromatics (>C10) Product Selectivity Product C4-benzene (C10) 0.05 Moles Produced/Moles Decalin Converted Decalin Moles Produced/Moles ( 0 0 200 400 600 800 Time on Stream (min)

Figure 10.2.3: Aromatic product selectivity over time for n-dodecane/decalin blend over mixed-bed reactor system at 400°C and 500 psig.

LTG conversion was stable over the entire run duration and was higher than that observed in the dual-bed reactor configuration, indicating a higher cracking rate to light hydrocarbons. Cracking pathways of n-dodecane can be studied by quantifying selectivity towards gaseous products species, with respect to n-dodecane conversion, as shown in Figure 10.2.4. Selectivity towards saturated paraffins was much higher than

selectivity towards olefins, which is consistent with in-situ hydrogen generation and

consumption speculated to be occurring. After hydrogen co-feed was terminated, all

gaseous hydrocarbon product selectivities increased for the duration of the study. This is

likely due to an increase in decalin cracking products, since decalin conversion remained

high with decreasing direct dehydrogenation activity. Double ring opening of decalin

and subsequent cracking would result in isomerized and normal paraffin products. The

shift towards decalin cracking could indicate that preferential cracking of decalin occurs

more readily in hydrogen-deficient conditions compared to n-dodecane.

2 1.8 1.6

1.4 Propylene 1.2 Propane 1 IsoButane Dodecane Converted) Dodecane - 0.8 n-Butane 0.6 IsoPentane Product Selectivity Product 0.4 1-Pentene 0.2 Hexane (Moles/Mole n (Moles/Mole 0 0 200 400 600 800 Time on Stream (min)

Figure 10.2.4: Gaseous hydrocarbon product selectivity over time for n-dodecane/decalin blend over mixed-bed reactor system at 400°C and 500 psig.

Liquid product analysis (shown in Figure 10.2.5) revealed low olefin selectivity.

This result is consistent with the hydrogen consumption hypothesis, as hydrogen-

deficient products were not selectively produced. Isododecane and n-paraffin selectivities were relatively consistent over the run duration, but isoparaffin selectivity

decreased over the first 360 minutes before increasing over the remaining time-on- stream. The initial decrease could be due to decreasing n-dodecane conversion, with the subsequent increase attributed to the hypothesized selectivity shift towards decalin cracking/ring opening. This liquid product data, along with gaseous product analysis, support a conclusion that the use of a mixed-bed reactor system resulted in an increase in decalin cracking selectivity over time, with a decrease in dehydrogenation activity.

However, unlike the dual-bed configuration, n-dodecane cracking decreased over time.

0.9

0.8

0.7

0.6 Dodecane Converted) Dodecane - n-paraffins (C5-C11) 0.5 Isoparaffins (C5-C11) 0.4 Olefins (C5-C12)

Product Selectivity Product 0.3 Isododecane (C12)

0.2

0.1

(Moles Produced/Mole n (Moles Produced/Mole 0 0 200 400 600 800 Time on Stream (min)

Figure 10.2.5: Liquid hydrocarbon product selectivity over time for n-dodecane/decalin blend over mixed-bed reactor system at 400°C and 500 psig.

When the mixed-bed n-dodecane conversion data is compared to the previously used reactor configurations, the mixed-bed system was superior, as demonstrated in

Figure 10.2.6. The single-bed blended feed data (Chapter 7) was the closest in activity to

the mixed-bed data, even though there was no Pt/Al2O3 catalyst. In that study, it was

observed that the Zeolite Y catalyst was able to promote significant decalin conversion to

aromatics, producing in-situ hydrogen. Product analysis in the single-bed blended feed study showed higher selectivity towards tetralin than for naphthalene, similar to that observed for the mixed-bed configuration. Hydrogen selectivity was also similar for the two studies, which suggests that the majority of decalin conversion in the mixed-bed reactor system was primarily a result of Zeolite Y catalytic activity, rather than an effect from the dehydrogenation catalyst. Therefore, higher n-dodecane conversion for the mixed-bed system compared to both single-bed systems could have simply been a function of increased decalin feed amount (50% volume compared to 10% and 0%) rather than a unique advantage of the bed mixing. This would indicate that the cracking catalyst contributed to reaction activity to a much greater extent in comparison to the dehydrogenation catalyst, either through much faster direct decalin conversion or through subsequent reaction of the Pt/Al2O3 naphthalene product. In comparing the tested

configurations, the dual-bed reactor system was unique in demonstrating an increase in

n-dodecane conversion with time, which therefore must have been related to the

configuration rather than just from decalin addition. However, initial inhibition in the

dual-bed system resulted in inferior n-dodecane conversion, although optimization of

various factors could be investigated in an attempt to mitigate the inhibition and increase

conversion in that system.

Single Bed (C12 Feed) 50 Single Bed (Blended Feed) Dual Bed 40 Mixed Bed

Conversion (%) 20

0 0 100 200 300 400 500 600 700 800 Time on Stream (min)

Figure 10.2.6: Comparison of n-dodecane conversion for single-bed, dual-bed, and mixed-bed reactor systems at the same LHSV (with respect to Zeolite Y volume) and reactor conditions of 400°C and 500 psig.

10.3 Primary Conclusions from n-Dodecane/Decalin Mixed-Bed Study

In the mixed-bed reactor configuration, n-dodecane conversion was initially high and decreased over time, while decalin exhibited nearly complete conversion for the entire duration. Naphthalene and tetralin selectivites were low and declined further over time, suggesting that complete dehydrogenation on the Pt/Al2O3 catalyst, without

subsequent secondary reactions, was not a prominent reaction pathway. Decalin product

selectivity was more similar to that from the single-bed blended feed study, rather than the dual-bed study, suggesting that Zeolite Y functionality was the primary driver of catalytic activity in the system. The mixed-bed system outperformed the dual-bed and all other systems, with respect to n-dodecane conversion, over a 720 minute study duration.

However, it was speculated that the increased n-dodecane conversion was primarily a result of decalin addition, as this correlated to an increase in n-dodecane conversion among the mixed-bed and single-bed configurations. The dual-bed reactor was unique in exhibiting increased n-dodecane conversion with time, but an initial inhibition resulted in inferior activity over the run duration. Due to substantial differences in product distribution over time for the different reactor systems, it is difficult to determine whether one system has a distinct advantage over others in every case. Future experimentation could alter blend amount, catalyst types, or reactor temperature to optimize each configuration for desired results.

100

CHAPTER 11

SUMMARY, CONCLUSIONS, AND FUTURE WORK

As an alternative to a gaseous molecular hydrogen co-feed, hydrogen-donor

hydrocarbons can be catalytically dehydrogenated to produce hydrogen for the purpose of

promoting hydrocracking reactions. Utilizing a cycloparaffin as an in-situ hydrogen

donor to produce molecular hydrogen for enhancing n-paraffin cracking activity and

decrease deactivation rate on a Zeolite catalyst was investigated. Decalin and n-dodecane

were selected as model dehydrogenation and cracking compounds, respectively. A

synthesized Pt/Al2O3 catalyst was used for dehydrogenation and a commercial Zeolite

catalyst for cracking. Initial testing identified adequate reaction conditions for decalin

dehydrogenation over the Pt/Al2O3 catalyst, as well as a screening study to identify a

suitable Zeolite cracking catalyst for testing. These conditions were then used to study

deactivation rate of the Zeolite in various reactor configurations and with various

n-dodecane liquid feeds, either neat or containing a blend of either decalin or

naphthalene. Cracking of neat n-dodecane over a single Zeolite cracking bed was used as

an initial baseline study. Decalin/n-dodecane and naphthalene/n-dodecane blended feeds

were tested with a single cracking bed to investigate the impact of these compounds on

Zeolite Y activity. Finally, a dual-bed catalyst configuration (with the Pt/Al2O3 catalyst upstream of the Zeolite) and a mixed-bed catalyst configuration (with the two catalysts

101 physically mixed) were studied to determine if in-situ hydrogen could be utilized for increasing n-dodecane cracking rate and decreasing catalyst deactivation rate.

Initial screening of several Zeolite type catalysts at 350°C and 500 psig

(Chapter 3) determined that Zeolite Y, with a SiO2/Al2O3 ratio of 5.2, would be sufficient to quantify the impact of in-situ hydrogen production on n-dodecane cracking activity in subsequent studies. Although it exhibited lower cracking activity than a ZSM-

5 catalyst, moderate conversion with Zeolite Y would allow better resolution of variance in activity compared to the extremely high n-dodecane conversion obtained with ZSM-5.

All other screened Zeolite catalysts exhibited lower cracking functionality compared to the selected Zeolite Y, when testing at reasonable hydrogen co-feed levels.

Investigation into activity of a synthesized Pt/Al2O3 catalyst with a neat decalin feed was studied for determination of appropriate reactor conditions to be used in subsequent testing (Chapter 4). Results indicated that a reactor temperature of 400°C and

LHSV of 1.5 h-1 was sufficient for nearly complete decalin dehydrogenation to naphthalene, with minimal deactivation. At these conditions, dehydrogenation activity, without a hydrogen co-feed, remained highly stable over 720 minutes. The results also showed a hydrogen yield very close to the theoretical maximum value of five moles produced per mole of decalin converted. The Pt/Al2O3 catalyst was then screened at various LHSV values while utilizing a 1:1 volumetric blend of decalin with n-dodecane

(Chapter 5). High decalin conversion was observed with the blended feed and hydrogen yield was again near the theoretical maximum. Conversion of n-dodecane was approximately 20%, which was primarily due to isomerization of the parent molecule.

As with the neat decalin dehydrogenation testing, an LHSV of 1.5 h-1 was confirmed to

102

be sufficient for the blended feed, with minimal impact of n-dodecane addition on decalin dehydrogenation activity. Reactor conditions of 400°C, 500 psig, and a liquid flow rate

(0.04 mL/min) corresponding to an LHSV of 1.5 h-1, with respect to the volume of

Pt/Al2O3, were used in subsequent testing.

Deactivation rate and product selectivity for n-dodecane cracking on Zeolite Y at

400°C and 500 psig in a continuous flow reactor were determined to establish a baseline

rate during hydrogen-deficient conditions (Chapter 6). A larger cracking catalyst bed

volume, compared to the volume of dehydrogenation catalyst previously studied, resulted

in an LHSV of 0.75 h-1, when utilizing the appropriate liquid flow rate determined in dehydrogenation studies. The testing showed a quantifiable decline in Zeolite Y catalytic activity for an extended duration study of 720 minutes, which was determined to be sufficient for subsequent deactivation studies. It was observed that a decrease in n-dodecance conversion coincided with a decrease in cracked isoparaffinic product selectivity and an increase in olefin product selectivity, which suggests that saturation of hydrogen-deficient β–scission cracking products was increasingly unable to occur.

Olefin production would have the potential to form coke species and potentially deactivate the catalyst. All subsequent testing utilized identical reactor conditions for temperature, pressure, and LHSV (with respect to the zeolite catalyst bed).

Deactivation rate and product selectivity was then studied for a n-dodecane feed containing 10% by volume of decalin (Chapter 7), utilizing identical conditions from the baseline study. Decalin had the potential to compete with n-dodecane for active sites and possibly lower cracking functionality with respect to the normal paraffin. Instead, the blended feed demonstrated higher n-dodecane conversion compared to a neat n-dodecane

103

feed. It was speculated that dehydrogenation of decalin occurred, which allowed in-situ

hydrogen production to promote hydrocracking and saturation of olefin products. This

caused both an initially higher cracking rate as well as a decreased catalyst deactivation

rate. Product analysis confirmed the presence of aromatic species even as hydrogen

concentration in the product gas went to near zero. This indicated that in-situ hydrogen,

formed during decalin conversion to aromatics, was most likely being utilized in

secondary reactions. However, deactivation still occurred, so in-situ hydrogen generation

was not sufficient to fully mitigate reduced reaction rates when only utilizing the Zeolite

Y catalyst. Alternatively, aromatic species could have contributed to catalyst poisoning

and created structural changes, which could alter product selectivity.

Impact of naphthalene addition on n-dodecane cracking with Zeolite Y was studied at 400°C and 500 psig (Chapter 8). The purpose of this study was to determine possible negative effects that could occur in a subsequent dual-bed configuration where dehydrogenated decalin would result in naphthalene entering the cracking catalyst bed.

A concentration of 11.9% naphthalene by mass was added to the n-dodecane feed, as this

was near the maximum solubility limit at room temperature. An extended duration study

in hydrogen-deficient reaction conditions was not performed, as naphthalene would not

be present in a system without hydrogen formation as a co-product. Two hydrogen co-

feed levels were used to investigate impacts from varying hydrogen concentrations. The

primary conclusion from the study was that naphthalene can have a significant inhibition

effect on cracking of n-dodecane when utilizing a Zeolite Y catalyst. Conversion of

n-dodecane was significantly lower compared to testing with a neat n-dodecane feed,

with a lower hydrogen co-feed level in the blended feed also contributing further to

104

reduced conversion. The inhibition was speculated to likely be from strong adsorption of

naphthalene onto active sites, which decreased the number of available sites for reaction

of n-dodecane. A lower hydrogen co-feed level also resulted in increased selectivity

towards alkylnaphthalene products, suggesting that molecular hydrogen is involved with

termination of alkylation reactions. The observed naphthalene inhibition indicated that

n-dodecane conversion could be severely hindered in a dual-bed reactor configuration,

where a much higher concentration of naphthalene would enter the Zeolite Y catalyst

bed.

A dual-bed reactor was then studied utilizing a blended n-dodecane/decalin feed

at a 1:1 volumetric ratio, and with a dehydrogenation catalyst bed upstream of a cracking

bed, both at 400°C and 500 psig (Chapter 9). Pt/Al2O3 was used as the upstream

dehydrogenation catalyst and Zeolite Y was used as the cracking catalyst. Catalytic

activity and deactivation rate was compared to the baseline neat n-dodecane cracking study. Availability of molecular hydrogen from decalin dehydrogenation in the Pt/Al2O3

catalyst bed was expected to increase cracking functionality and minimize activity loss in

the Zeolite Y bed during testing. Utilization of the dual-bed reactor and blended

n-dodecane/decalin feed resulted in negligible initial n-dodecane conversion, but conversion increased over the duration of the study. This resulted in a higher n-dodecane conversion after 720 minutes compared to the baseline n-dodecane cracking study. Initial suppression of n-dodecane conversion was speculated to be a result of naphthalene inhibition, as was observed in the n-dodecane/naphthalene cracking study. Product analysis indicated that the dual-bed reactor system also resulted in a lower level of olefin production compared to the baseline study, indicating that hydrogen was successfully

105 being utilized to enhance cracking rate and/or saturate the cracked products, following the initial cracking suppression. Decalin product selectivity shifted from initial complete dehydrogenation to cracked species as run time increased. This could have been a result of active site poisoning, which may have caused a structural change in the Zeolite Y catalyst and altered functionality without decreasing activity, which is a catalytic deactivation pathway supported by literature.1 In general, it was determined that use of the dual-bed system enhanced Zeolite Y cracking functionality after 720 minutes compared to the baseline study. Future studies could be performed to identify possible structural changes that occurred via poisoning, for the purpose of developing a pre-treatment that induces these changes before utilization in a cracking reactor system.

Testing was performed using a configuration with the Pt/Al2O3 and Zeolite Y catalysts physically mixed into a single reaction bed. The mixed-bed reactor utilized reaction conditions at 400°C and 500 psig and a 1:1 volumetric blended feed of n-dodecane/decalin (Chapter 10). Impact of bed mixing on catalyst deactivation rate and product selectivity were compared to results observed for the baseline neat n-dodecane cracking study, the blended n-dodecane/decalin cracking study, and the dual-bed reactor study. It was speculated that the initial naphthalene inhibition observed in the dual-bed reactor could be mitigated by allowing decalin dehydrogenation to occur along the length of the catalyst bed, as would occur in the mixed-bed system. Unlike in the dual-bed reactor results, the mixed-bed configuration resulted in deactivation behavior similar to the single-bed cracking studies. Conversion of n-dodecane was initially very high, but decreased after hydrogen co-feed was terminated. Nearly complete decalin conversion was observed at all data points, but a selectivity shift from dehydrogenated to cracking

106

products was observed, similar to that observed in the other configurations. However,

aromatic species were still produced at longer run times, even though hydrogen output

was negligible, which indicated that in-situ hydrogen was consumed in secondary reactions. Due to low selectivity towards naphthalene and tetralin, it was also speculated that increased n-dodecane conversion was primarily a result of decalin addition rather than from inclusion of the Pt/Al2O3 catalyst. When comparing the mixed-bed, dual-bed,

and single-bed reactor configurations for n-dodecane cracking performance, it was determined that the mixed-bed system was superior at the studied reactor conditions and run duration. However, increased n-dodecane cracking activity also correlated with an increase in decalin blend amount, so it is possible that this was the primary factor in causing the increase. Future studies could identify impact of decalin blend amount on n-dodecane conversion, independent of reactor configuration. Overall, the primary result

of the present testing determined that hydrogen donor cycloparaffins have the ability to

catalytically generate molecular hydrogen and enhance normal paraffin cracking

functionality on Zeolite Y. This can prolong the practical life of a catalyst bed and

decrease required downtime from regeneration or replacement of the bed.

Subsequent experimentation utilizing these reactor configurations could

investigate the impact from varying conditions such as temperature, space velocity, and

catalyst loading. Selection of alternative catalysts or model feed compounds could also

be used to optimize reaction pathways. For example, ZSM-5 has pore sizes on the order

of the width of a naphthalene molecule, which could significantly limit diffusion to

active-sites and reduce the amount of poisoning. Additionally, a bifunctional metal/acid- site catalyst could be studied to eliminate the need for separate dehydrogenation and

107

cracking catalysts, performing both functions without the need for products to diffuse

between the different catalysts. Hydrogen donors, such as methylcyclohexane or

cyclohexane, could be used as an alternative to decalin, which would eliminate

complications from naphthalene product solidification at room temperature. Replacing

n-dodecane with an isomerized paraffin feed could also potentially result in increased β–

scission cracking rate.

The dual-bed reactor was unique in exhibiting increasing cracking functionality

over time; optimization of that configuration in future work could lead to improved

performance. This could be done by identifying any structural changes that occurred to

the Zeolite Y catalyst (via Temperature Programmed Desorption, X-Ray Diffraction,

Chemisorption, Physisorption, or Transmission Electron Microscopy analysis) during the

dual-bed study. These techniques could be used to characterize coke deposition as well

as changes in crystalline structure, porosity, and metal loading/dispersion that occurred

during testing. A pre-treatment methodology could then be developed to alter the catalyst

in a similar manner prior to use in a cracking system, which would produce a Zeolite

catalyst that exhibits high n-paraffin cracking activity with resistance to naphthalene

inhibition. Further alterations in the configuration of dehydrogenation and cracking beds

could also be studied. One solution to reduce naphthalene inhibition could be to separate

the catalyst beds as two unique reactor systems, feeding decalin and n-dodecane

separately to each respective bed. Neat decalin could be dehydrogenated over a Pt/Al2O3 bed, with product hydrogen separated by a membrane and diverted as a co-feed with n-dodecane to the cracking bed. This would also have the advantage of being able to

108 optimize reaction conditions, such as feed flow rate and temperature, for each catalyst bed individually.

109

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114

APPENDIX A

Sample Calculations

Liquid-to-gas (LTG) conversion was calculated from an average of the gas composition data points obtained from the MicroGC between liquid samples. The molar flow rate of each gaseous specie was determined via multiplication of the molar composition and the total exit gas flow. Total exit gas flow was calculated by dividing the known constant Nitrogen flow rate (164.14 mL/min) by the Nitrogen molar fraction in the exit gas. Molar flow rates were converted to mass flow rates using the molecular weight of each specie, and were summed for determining total gaseous mass flow rate from the reactor. The LTG conversion was determined using the ratio of gaseous product flow rate to the liquid feed flow rate. LTG conversion was calculated from gaseous analysis (rather than from the mass of liquid collected) because possible liquid sampling errors, such as spraying, could cause some samples to be inaccurate with regard to liquid product mass recovery. Therefore, use of gaseous analysis was considered to minimize sources of error.

, = × 164.14 𝑚𝑚𝑚𝑚 𝑣𝑣̇𝑡𝑡𝑡𝑡𝑡𝑡𝑡𝑡𝑡𝑡 𝑒𝑒𝑒𝑒𝑒𝑒𝑒𝑒 𝑥𝑥𝑁𝑁𝑁𝑁𝑁𝑁𝑁𝑁𝑁𝑁𝑁𝑁𝑁𝑁𝑁𝑁 𝑚𝑚𝑚𝑚𝑚𝑚 = × × , × 𝑃𝑃 𝑛𝑛̇ 𝑖𝑖 �𝑣𝑣̇𝑡𝑡𝑡𝑡𝑡𝑡𝑡𝑡𝑡𝑡 𝑒𝑒𝑒𝑒𝑒𝑒𝑒𝑒 𝑥𝑥𝑖𝑖� � � 𝑅𝑅 𝑇𝑇 , = ×

𝑚𝑚̇ 𝑔𝑔𝑔𝑔𝑔𝑔 𝑒𝑒𝑒𝑒𝑒𝑒𝑒𝑒 � 𝑛𝑛̇ 𝑖𝑖 𝑀𝑀𝑀𝑀𝑖𝑖

115

, = , 𝑚𝑚̇ 𝑔𝑔𝑔𝑔𝑔𝑔 𝑜𝑜𝑜𝑜𝑜𝑜 𝐿𝐿𝐿𝐿𝐿𝐿 𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶 𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑖𝑖𝑖𝑖 𝑚𝑚̇

Conversion of a reactant was determined by dividing the amount of reactant mass which was used to form products to the amount of reactant mass present in the feed. This ratio was manipulated via the following method for conversion of reactant specie “i”, with “in” and “out” subscripts corresponding to reactant and product values, respectively.

, , , , , = = 1 = 1 , , 𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑜𝑜𝑜𝑜𝑜𝑜 𝑖𝑖 𝑖𝑖𝑖𝑖 𝑖𝑖 𝑜𝑜𝑜𝑜𝑜𝑜 𝑖𝑖 𝑜𝑜𝑜𝑜𝑜𝑜 ̇ , � 𝑖𝑖,𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑜𝑜𝑜𝑜𝑜𝑜� 𝑖𝑖 𝑚𝑚̇ − 𝑚𝑚̇ 𝑚𝑚̇ 𝑚𝑚 𝑦𝑦 𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶 𝑖𝑖 𝑖𝑖𝑖𝑖 − 𝑖𝑖 𝑖𝑖𝑖𝑖 − 𝑚𝑚̇ 𝑚𝑚̇ ̇ 𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑖𝑖𝑖𝑖 � 𝑖𝑖 𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑖𝑖𝑖𝑖� , , 𝑚𝑚, 𝑦𝑦 , = = = 1 , , , 𝑚𝑚̇ 𝑔𝑔𝑔𝑔𝑔𝑔 𝑜𝑜𝑜𝑜𝑜𝑜 𝑚𝑚̇ 𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑖𝑖𝑖𝑖 − 𝑚𝑚̇ 𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑜𝑜𝑜𝑜𝑜𝑜 𝑚𝑚̇ 𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑜𝑜𝑜𝑜𝑜𝑜 𝐿𝐿𝐿𝐿𝐿𝐿 𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝑟𝑟𝑠𝑠𝑠𝑠𝑠𝑠𝑠𝑠 𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑖𝑖𝑖𝑖 𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑖𝑖𝑖𝑖 − 𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑖𝑖𝑖𝑖 𝑚𝑚, ̇ = , (1 𝑚𝑚̇ ) 𝑚𝑚̇

𝑚𝑚̇ 𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑜𝑜𝑜𝑜𝑜𝑜 𝑚𝑚̇ 𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑖𝑖𝑖𝑖 − 𝐿𝐿𝐿𝐿𝐿𝐿 𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶 , = 1 (1 ) 𝑦𝑦𝑖𝑖,𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑜𝑜𝑜𝑜𝑜𝑜 𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝑖𝑖 − � − 𝐿𝐿𝐿𝐿𝐿𝐿 𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶𝐶 � �� 𝑖𝑖 𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙𝑙 𝑜𝑜𝑜𝑜𝑜𝑜 𝑦𝑦

116