Kinetics of the Hydro-Deoxygenation of Stearic Acid Over

KINETICS OF THE HYDRO-DEOXYGENATION OF STEARIC ACID OVER

PALLADIUM ON CARBON CATALYST IN FIXED-BED REACTOR FOR THE

PRODUCTION OF RENEWABLE DIESEL

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

Albert Vam

Dayton, Ohio

August, 2013

KINETICS OF THE HYDRO-DEOXYGENATION OF STEARIC ACID OVER

PALLADIUM ON CARBON CATALYST IN FIXED-BED REACTOR FOR THE

PRODUCTION OF RENEWABLE DIESEL

Name: Vam, Albert

APPROVED BY:

Kevin J. Myers, D.Sc., P.E. Heinz J. Robota, Ph.D. Advisory Committee Chairman Research Advisor Professor; Graduate Chemical Engineering Ohio Research Scholar in Alternative Fuels Program Coordinator Alternative Fuels Group Leader Department of Chemical and Materials University of Dayton Research Institute Engineering

Amy R. Ciric, Ph.D. Committee Member Senior Lecturer Department of Chemical and Materials Engineering

John G. Weber, Ph.D. Tony E. Saliba, Ph.D. Associate Dean Dean, School of Engineering School of Engineering & Wilke Distinguished Professor

©Copyright by

Albert Vam

2013

iii

ABSTRACT

KINETICS OF THE HYDRO-DEOXYGENATION OF STEARIC ACID OVER

PALLADIUM ON CARBON CATALYST IN FIXED-BED REACTOR FOR THE

PRODUCTION OF RENEWABLE DIESEL

Name: Vam, Albert University of Dayton

Advisor: Dr. Heinz J. Robota Biological oils are potential sources of liquid transportation fuels. In the presence of a precious metal catalyst under reducing conditions, the transformation of biological oils to liquid fuels proceeds sequentially. First, any double bond is quickly saturated.

The double bond saturation is followed by the hydrogenolysis of the ester linkages which releases the saturated fatty acids from the propane backbone. The fatty acids are then deoxygenated producing n-alkanes. Typically, the n-alkanes are further treated to meet the required physical properties of a particular fuel fraction. Although important, the deoxygenation of the fatty acid had not yet been studied in production-like conditions.

For this reason, in this study, a comprehensive investigation of the deoxygenation of a representative fatty acid was carried out by studying stearic acid (S.A.) diluted in a highly isomerized C24 solvent. The deoxygenation of stearic acid was studied in the presence of hydrogen, in a trickle-bed reactor by using a 3 wt % carbon-supported palladium catalyst.

In order to simplify the study of the kinetics of the S.A. deoxygenation, a uniform S.A. concentration across the catalyst bed was desired. For this reason, the entire study was conducted under differential conditions, by limiting the S.A. conversion to 10 percent.

The limited conversion allowed me to assume uniform S.A. concentration across the catalyst bed.

The liquid products were identified early on as n-heptadecane, n-octadecane, and stearyl stearate. The rate of formation for each liquid product was examined over wide-ranging sets of temperature, initial S.A. concentration and hydrogen pressure. Kinetic data for the different products were graphically derived and rate expressions were developed and presented.

The S.A. deoxygenation reaction network in the presence of hydrogen was found to be rather complex. The observed rate of n-heptadecane production was a combined rate of two reactions, the stearic acid decarboxylation and the octadecanal decarbonylation.

Likewise, the observed rate of n-octadecane production was the sum of the rates of octadecanol reduction and stearyl stearate hydrogenolysis. In order to estimate the contribution of each of the rates of S.A. decarboxylation and octadecanal decarbonylation to the total rate of n-heptadecane production, the relative rate of CO/CO2 production was studied over a range of temperature and hydrogen pressure. And, in order to estimate the contribution of the rate of alcohol reduction and the rate of stearyl stearate hydrogenolysis to the total observed rate of n-octadecane production, a complete study of palmityl stearate hydrogenolysis was carried out to estimate the stearyl stearate hydrogenolysis partial contribution to the overall rate of n-octadecane production. v

Dedicated to my family

ACKNOWLEDGMENTS

This research was supported, in part, by the U. S. Air Force Cooperative Grant

Numbers F33615-03-2-2347 and FA8650-10-2-2934, Mr. Robert W. Morris Jr., Air

Force Grant Monitor. The research was also sponsored by the State of Ohio Subrecipient

Award No. COEUS # 005909 to the University of Dayton (Dr. Heinz Robota as the Grant

Monitor) under the “Center for Intelligent Propulsion and Advanced Life Management,” program with the University of Cincinnati (Prime Award NO. TECH 09-022). The authors gratefully acknowledge this grant support.

I would also like to express my genuine gratitude to my research advisor and professional mentor Dr. Heinz Robota. First for allowing me to be a member of his outstanding team, and second, for his time, expertise, patience, and for the invaluable chemistry and engineering knowledge that he has shared with me. This has been truly by far the best education I have ever received.

Special thanks go to Dr. Kevin. Myers, my advisory committee chairman and graduate studies advisor who was one of the reasons I chose to pursue my graduate studies at UD.

Thank you for all your support and for being so helpful. Thanks to Dr. Amy Ciric, my thesis committee member for taking the time to be on this committee and for sharing her knowledge in renewable fuels with me, inside and outside the classroom.

vii

Sincere thanks to Jhoanna Alger for all her help and contribution to this work in the lab, and for reviewing multiple drafts of this thesis.

Thanks to Linda Shafer for analyzing my samples. I also thank Dave Gasper to whom I always reached out with issues related to equipment design that were beyond my knowledge. Thanks to Sophia D’Angelo with whom I shared the office space, and all the people of UDRI who made it a pleasant work environment.

viii

TABLE OF CONTENTS

ABSTRACT ...... iv

DEDICATION…………………………………………………………………………. ..ix

ACKNOWLEDGMENTS…………………………………………………………….... vii

LIST OF ILLUSTRATIONS ...... xiv

LIST OF TABLES ...... xvi

LIST OF SYMBOLS/ABBREVIATIONS ...... xviii

CHAPTER 1 INTRODUCTION ...... 1

1.1. Biological Oils for Transportation Fuels ...... 1

1.2. Challenges to the Direct Use of Biological Oils as Fuels ...... 3

1.3. Methods of Upgrading ...... 3 1.3.1 Transesterification of Biological Oils ...... 5 1.3.2 Deoxygenation of Biological Oils ...... 6 1.3.3 Studying Fatty Acid Instead of a Triglyceride ...... 8

CHAPTER 2 LITERATURE REVIEW ...... 9

2.1. Catalyst Selection ...... 9

2.2. The Impact of Hydrogen ...... 10 2.2.1 Zero and Low Hydrogen Pressure ...... 10 2.2.2 Reaction under Pure Hydrogen ...... 13

2.3. Effect of the Initial Concentration of the Fatty Acid ...... 16

2.4. Effect of Chain Length ...... 16

2.5. Summary ...... 16

CHAPTER 3 EXPERIMENTAL APPROACH ...... 18

3.1. Studying the Kinetics of S.A. Deoxygenation in Place of Triglyceride ...... 18

3.2. Reactor Type, Differential Operation, and S.A. Concentration ...... 19

3.3. Importance of Hydrogen in Biological Oil Upgrading ...... 20

3.4. Materials ...... 22 3.4.1 Catalyst Preparation ...... 22 3.4.2 Catalyst Activation...... 22 3.4.3 Feed Materials ...... 23 3.4.4 Fixed-Bed Reactor ...... 24 3.4.5 Liquid Analysis ...... 27

3.5. Study Framework ...... 28

CHAPTER 4 APPARENT KINETICS OF THE S.A. DEOXYGENATION AND LIQUID SPECIES PRODUCTION ...... 29 4.1 Effect of Temperature ...... 29 4.1.1 Experiment ...... 29 4.1.2 Results ...... 31 4.1.2.1 Effect of Temperature on the Deoxygenation of Stearic Acid ...... 31 x

4.1.2.2 Effect of Temperature on the Rate of Liquid Species Production .. 37

4.2 Effect of Stearic Acid Initial Concentration ...... 45 4.2.1 Experiment ...... 45 4.2.2 Results ...... 47 4.2.2.1 Effect of S.A. Initial Concentration on the Rate of S.A. Deoxygenation ...... 47 4.2.2.2 Effect of S.A. Initial Concentration on the Rate of Liquid Species Production ...... 49

4.3 Effect of Hydrogen Pressure ...... 51 4.3.1 Experiment ...... 51 4.3.2 Results ...... 53

4.3.2.1 Effect of Hydrogen Pressure on the Rate of S.A. Deoxygenation .. 53

4.3.2.2 Effect of Hydrogen Pressure on the Rate of Liquid Species Production ...... 56

4.4 Effect of Hydrogen Pressure on the Rate of n-alkane Production at 300 °C ..... 59 4.4.1 Experiment ...... 59 4.4.2 Results ...... 59 4.4.2.1 Effect of Hydrogen Pressure on the Rate of S.A. Deoxygenation at 300˚C...... 59 4.4.2.2 Effect of Hydrogen Pressure on the Rate of Liquid Product Species Production at 300 ˚C ...... 62

4.5 Discussion ...... 63 4.5.1 Stearic Acid Deoxygenation Reaction Pathways ...... 63 4.5.2 Effect of Reaction Parameters on the Rate of Stearic Acid Deoxygenation and Liquid Species Production ...... 65 4.5.2.1 Effect of Temperature ...... 65 4.5.2.2 Effect of Stearic Acid Initial Concentration ...... 67 4.5.2.3 Effect of Hydrogen Pressure ...... 70 4.5.2.4 Highest Temperature and Lowest Hydrogen Pressure ...... 73

CHAPTER 5 INDIVIDUAL REACTION PATHWAY CONTRIBUTION TO THE RATE OF n-ALKANE PRODUCTION ...... 74

5.1 Decarboxylation Contribution to the Total Rate of n-C17 Production ...... 74 5.1.1 Introduction ...... 74 5.1.2 Experiment ...... 75 5.1.3 Results ...... 77 5.1.4 Discussion ...... 78

5.2 Stearyl Stearate Hydrogenolysis Contribution to the Total Rate of n-C18 Production ...... 80 5.2.1 Introduction ...... 80 5.2.2 Feed Materials ...... 82 5.2.3 Effect of Temperature on P.S. Hydrogenolysis ...... 83 5.2.3.1 Experiment ...... 83 5.2.3.2 Results ...... 84 5.2.4 Effect of Initial Palmityl Stearate Concentration ...... 86 5.2.4.1 Experiment ...... 86 5.2.4.2 Results ...... 87 5.2.5 Effect of Hydrogen Pressure ...... 89 5.2.5.1 Experiment ...... 89 5.2.5.2 Results ...... 89 5.2.6 Discussion ...... 92

CHAPTER 6 CONCLUSION AND FUTURE STUDIES RECOMMENDATIONS 95

6.1 Conclusion …………………………………………………………………….95

6.2 Future Work Recommendation……………………………………………….. 98

BIBLIOGRAPHY……………………………………………………………………...101

xii

APPENDIX ...... 105

VARIAN 660 FTIR SPECTROMETER GAS PHASE ...... 105

xiii

LIST OF ILLUSTRATIONS

1. Figure 1.1.1 Triglyceride molecule...... 2

2. Figure 1.3.1 Deoxygenation of a triglyceride on Pd/C...... 7

3. Figure 2.2.1 Proposed S.A. reaction pathways on Pd/C catalyst in zero to low PH2

...... 12

4. Figure 2.2.2 Proposed S.A. reaction pathways on Pd/C catalyst under high PH2 15

5. Figure 3.4.1 Schematic of the reactor system ...... 26

6. Figure 4.1.1 The effect of temperature on the rate of S.A. deoxygenation in

Arrhenius form ...... 36

7. Figure 4.1.2 Reactive pathways in the deoxygenation of S.A. conversion over a

Pd/C catalyst in the presence of H2 ...... 38

8. Figure 4.1.3 Effect of temperature on the rate of n-C17 and n-C18 production in

Arrhenius form ...... 41

9. Figure 4.2.1 Effect of S.A. initial concentration on the rate of S.A. deoxygenation

determined by plotting ln(rS.A.disapp.) versus ln[S.A.]...... 48

10. Figure 4.2.2 Effect of S.A. concentration on the rate of n-C17, n-C18, and S.S.

production determined by plotting ln(rp) versus ln[S.A.] ...... 50

11. Figure 4.3.1 Effect of hydrogen pressure on the rate of S.A. deoxygenation

determined by plotting ln(r S.A.disapp.) versus ln(PH2) ...... 55

xiv

12. Figure 4.3.2 Effect of PH2 on the rate of n-C17, n-C18 and S.S. production

determined by plotting ln(rp) versus ln(PH2) ...... 58

13. Figure 4.4.1 Effect of PH2 on the rate of S.A. deoxygenation at 300 °C ...... 61

14. Figure 4.4.2 Effect of PH2 on the rates of liquid species production at 300 °C .... 63

15. Figure 5.1.1 The observed odd-numbered product n-C17 is the combined product

of the decarboxylation of S.A. and the decarbonylation of octadecanal...... 75

16. Figure 5.2.1 The observed even-numbered product n-C18 is the combined product

of the alcohol reduction and the S.S. hydrogenolysis ...... 81

17. Figure 5.2.2 Effect of temperature on the rate of P.S hydrogenolysis in Arrhenius

form ...... 85

18. Figure 5.2.3 Effect of P.S. initial concentration on the rate of P.S. hydrogenolysis

...... 88

19. Figure 5.2.4 Effect of hydrogen pressure on the rate of P.S. hydrogenolysis ..... 91

LIST OF TABLES

1. Table 1.3.1 Comparative assessment of the two methods of triglyceride upgrading

...... 3

2. Table 3.1.1 Fatty acid composition in a selected group of biological oils ...... 19

3. Table 4.1.1 Effect of temperature on the S.A. conversion and the rate of S.A.

deoxygenation ...... 31

4. Table 4.1.2 Effect of temperature on the rate of liquid species production ...... 37

5. Table 4.2.1 Effect of S.A. initial concentration on the rate of S.A. deoxygenation

...... 47

6. Table 4.2.2 Effect of S.A. initial concentration on the rate of liquid species

production and selectivity to products ...... 49

7. Table 4.3.1 Effect of hydrogen pressure on the rate of S.A. deoxygenation ...... 53

8. Table 4.3.2 Effect of hydrogen pressure on the rate of liquid species production

and product selectivity ...... 56

9. Table 4.4.1 Effect of PH2 on the rate of S.A. deoxygenation at 300 °C ...... 60

10. Table 4.4.2 Effect of PH2 on the rate of liquid species production and selectivity

to products at 300 °C ...... 62

11. Table 5.2.1 The effect of temperature on the rate of P.S. hydrogenolysis...... 84

xvi

12. Table 5.2.2 Effect of P.S. initial concentration on the rate of P.S. hydrogenolysis

...... 87

13. Table 5.2.3 Effect of hydrogen pressure on the rate of P.S. hydrogenolysis ...... 90

xvii

LIST OF SYMBOLS/ABBREVIATIONS

[ ] Concentration or Units of

0 Pre-exponential factor

α Reaction Order with Respect to S.A.

App. Apparent

β Reaction Order with Respect to Hydrogen disapp. Disappearance

Ea Activation Energy

FTIR Fourier Transform Infra-Red

GC Gas Chromatography k Rate Constant

KH Henry's constant

L.A. Lauric Acid n-C17 n-Heptadecane n-C18 n-Octadecane p Product

P.A. Palmityl Acid

P.S. Palmityl Stearate

PH2 Hydrogen Partial Pressure

xviii

r Rate

S.A. Stearic Acid

S.S. Stearyl Stearate t Time

WHSV Liquid Hourly Space Velocity wt% Weight Percent

xix

CHAPTER 1

INTRODUCTION

1.1. Biological Oils for Transportation Fuels

Petroleum resources are diminishing while the global use of liquid fuel is expected to increase 31 percent by the year 2035. [1] The scarcity of petroleum oil makes it an expensive commodity. The price of petroleum oil has an immediate impact on the price of consumer goods. Any increase in the price of petroleum oil puts a strain on consumer spending which ends up affecting the global economy as a whole. That has required the world to look into renewable alternatives. The need demanded exploring new possibilities and reawakened interest in biological resources. Oils derived from biological sources have potential to become a valued source of transportation fuels due to their high heating value [2], renewable nature [3], and relatively clean combustion compared to petroleum fuels. [4]

The use of biological oils as fuel in diesel engines is not a new discovery; in fact, it is just as old as the diesel engine itself. When the diesel engine was introduced in1900, it ran on unrefined vegetable oil, until the 1920s when the petroleum industry started promoting what became petro-diesel. Since then, the design of modern compression ignition

engines has been based on petro-diesel because of the favorable economy of the production of fuels from petroleum, fewer storage challenges, and higher energy density of petro-diesel compared with vegetable oils. [5]

Biological oils are natural triglycerides that can be extracted from vegetable seeds, animal fat, and what recently seems most appealing, algae. They have high content of saturated fatty acids that are the source of their high heating value. All natural triglycerides are chemically similar in structure. The chemical structure of a typical triglyceride molecule is shown in Figure 1.1.1. Triglycerides are composed of three acyl groups, which are typically 13-21 carbons in length and have varying degrees of unsaturation, from 0 to 6 double bonds. The three acyl groups are attached to the glycerol backbone. [6]

R1 –C–O–CH2

R2 –C–O–CH

O R –C–O–CH 3 2

Figure 1.1.1 Triglyceride molecule. R1, R2, and R3 are alkyl groups that have odd numbers of C atoms, typically 13-21 C’s. Any alkyl group may contain 0 to 6 double bonds

1.2. Challenges to the Direct Use of Biological Oils as Fuels

Biological oils are promising substitutes to petro-diesel. However, some of their physical properties prohibit their direct use in modern diesel engines. The most troubling properties of biological oils are their low volatility and high viscosity. These poor properties can quickly lead to engine damage. The issues they may cause include: oil ring sticking and plugging and gumming of filters, lines, and injectors. [4] A good fuel must adhere to a range of properties defined by specification such as volatility, viscosity, and cold flow properties. [7]

The change in physical properties while maintaining the heating value of the biological oil molecules can only be achieved by rearranging the molecular structure of the biological oil through a chemical transformation to molecules that have desired properties, while avoiding or minimizing cracking in order to maintain the heating value.

That’s what is commonly known as “upgrading”.

1.3. Methods of Upgrading

Transesterification and deoxygenation are two methods that are currently practiced in upgrading biological oils to transportation fuels. Table 1.3.1 shows a comparative assessment of the two methods. [5] Although, the terms biodiesel and renewable diesel seem to be applicable to the product of either method, in order to distinguish between them, it has been established that the product of transesterfication is known as “biodiesel” while the deoxygenation product is known as “renewable-diesel” or “green-diesel”.

Table 1.3.1 Comparative assessment of the two methods of triglyceride upgrading [2]

Transesterfication Deoxygenation Feed material Refined oil or fat Refined, unrefined or used oil

Other feed input Methanol Mixture of hydrogen and inert gas; or pure hydrogen

Main Product Fatty Acid Methyl Ester (FAME) n- alkanes“ Green- Diesel” “Biodiesel”

Catalyst Homogeneous Alkali alkoxide Supported precious metal, or hydroxide conventional hydrotreating sulfided or zeolite supported metal catalysts

Process temperature 25 - 50°C 250-450°C

Product treatment Requires large amounts of water Not required to separate the biodiesel from the catalyst

Fuel combustion Higher cetane than fossil diesel Higher cetane than fossil diesel

Fuel properties Poor Good

Storage Higher flash point therefore safer Similar to fossil diesel

Shelf life Lower compared to fossil-diesel Similar to fossil diesel due to rancidity reaction and corrosion tendency

1.3.1 Transesterification of Biological Oils

Transesterification is the most common and commercially practiced method for upgrading biological oils. [6] Transesterification is favorable due to its feasibility at essentially any scale with minimum equipment. Transesterification is widely accepted and promoted due to the high cetane rating and low sulfur content of the product. [5] The main objective of transesterification is to reduce the viscosity of the natural oil. In transesterification, typically methanol is used as a co-reactant in the presence of NaOH to separate the three acyl groups attached to the glycerol backbone. The reaction can be achieved at room temperature. However, mild heating can be applied in order to increase the reaction rate. The products of transesterification are fatty acid methyl esters, or

FAMEs, which are widely used as biodiesel. [6]

Although FAMEs have a lower viscosity than the parent triglyceride, their viscosity is still much higher than that of petro-diesel. [5] And more importantly, FAMEs still have cold-flow related issues (as discussed in the introduction).

In order to overcome these issues, several solutions have been proposed, including the following: blending with petro- diesel, additives, and using branched-chain esters.

Blending the methyl esters of vegetable oils is currently the most applied technique. The most common ratio is 80% petro-diesel and 20% methyl esters of vegetable oils

(commercially known as “B20”, 20 vol% biodiesel). Significant emission reductions

have been reported for these blends. [4] No engine problems were reported in larger scale tests. [4] The ester blends have reportedly been stable.

Although transesterification improves the viscosity of the natural oils, it is clear that it is not exactly the ideal substitute to petro- diesel. Ideally, direct conversion of triglycerides to hydrocarbon mixtures that are similar to those of petro-diesel is desired. Recent research has been focused on how to transform natural oils to hydrocarbons with similar properties to those of petro-diesel in a practical and economical way.

1.3.2 Deoxygenation of Biological Oils

Petro-diesel-like hydrocarbons can be obtained from biological oils by the elimination of the oxygen in the acid functional groups, in other words, by the “deoxygenation” of the biological oils to normal hydrocarbon, and subsequently applying some measure of structural rearrangements to ultimately obtain the desired fuel properties.

Triglyceride deoxygenation is currently achieved by three different methods. One way is by using conventional sulfided catalysts such as NiMo or CoMo at 300-450 °C in the presence of H2. [8] However, these catalysts deactivate quickly due to sulfur leaching, especially in the presence of any traces of water. [8] Another method is by using a zeolite supported metal catalyst at temperatures in the 250-300 °C range, also in the presence of H2. [8] The third method, which is the focus of this work, is by using a noble metal supported catalyst such as Pd/C or Pt/C at 300-360 °C in the presence or absence of

H2. [8]

The conversion of triglycerides to liquid fuels on Pd/C proceeds in a series of chemical transformations. [9] First, any double bonds are quickly saturated, followed by the hydrogenolysis of the ester linkages resulting in the release of propane and 3 saturated fatty acids, which is shown in Figure 1.3.1. This step is followed by the deoxygenation of the fatty acid yielding n-hydrocarbon. The hydrocarbon product may further be isomerized and/or cracked to achieve the desired fuel properties.

CH2- COO- R1 R4-COOH CH3 3

CH – COO-R2 + H2 R5-COOH + CH2 2

CH2- COO-R3 R6-COOH CH3 3

Figure 1.3.1 Deoxygenation of a triglyceride on Pd/C R1, R2, and R3 are alkyl groups that have odd numbers of C atoms, typically from 13-21 carbons. Any of the alkyl groups may contain 0 to 6 double bonds. R4, R5 and R6 are saturated alkyl groups.

1.3.3 Studying Fatty Acid Instead of a Triglyceride

My involvement in algae oil upgrading prompted the need to examine the chemical transformation of biological oils to liquid fuels on the molecular level. Although algal triglycerides should be similar to any other triglyceride, during algal triglyceride upgrading, complications caused by foreign contaminants that interfered with the chemical transformation of the triglycerides were encountered. One of the issues was the accumulation of contaminants within the Pd/C bed that affected the catalyst performance.

[10] Since the triglyceride hydrogenolysis products are always saturated fatty acids, [11] in order to avoid complications that could arise due to the poor physical separation of the triglycerides which can potentially interrupt the process chemistry, this study was designed to focus on the fatty acid deoxygenation to n-hydrocarbon. This is the most complex and not yet fully understood step within the series of reactions. It is also the most critical step because the chemistry of the fatty acid deoxygenation is what dictates the alkane product composition.

CHAPTER 2

LITERATURE REVIEW

2.1. Catalyst Selection

The S.A. thermal decarboxylation at 300 ˚C and 6 bar of inert atmosphere in a stirred tank only achieved a very low conversion of 5 % after 6 hours of reaction time. [12]

Over 80 years ago, Bertram reported that a homogeneous selenium catalyst yielded a`50 % conversion of S.A. to heptadecane through decarboxylation, yet dehydrogenation of alkanes to olefins also took place. [13] Reports of the use of heterogeneous catalysts for the deoxygenation of fatty acids had not been widely available until 2005.

In 2006, in an effort to obtain the most efficient metal and support combination for S.A. deoxygenation, Murzin’s group screened several metal-support combinations. The metals studied were: Ni, Mo, Pd, Pt, Ir, Ru, Rh, and Os, and the supports were: Al2O3,

Cr2O3, MgO, SiO2, and activated carbons. [13] The performance of these catalysts was tested in a semi-batch reactor at 300 ˚C under 6 bar of inert atmosphere. The highest S.A. conversion was achieved on Pd/C followed by Pt/C.

In order to determine the optimum metal wt % on the support, three different metal loadings (1, 5, and 10 %) were compared. A 100% S.A. conversion was obtained on

5 wt% Pd/C, with 99 % selectivity to n-C17. The conversion on the 1 % and 10 % were

33.4 % and 48.1 % respectively. The highest conversion on Pt/C was 86 % with 95 % selectivity to n-C17. [13]

The Pd/C catalyst was further investigated in order to understand how the change in the different parameters during the catalyst preparation can influence its deoxygenation performance. It was reported that the metal dispersions decreased with increasing reduction temperature. [12] When the effect of metal dispersion on the deoxygenation of fatty acids was studied over a range 18-72 % [14], the catalysts with the lowest and highest dispersion showed the lowest initial deoxygenation rate, while the catalysts with medium dispersion (47 & 65 %) showed significantly higher initial reaction rates, which suggests that there is an optimum Pd crystallite size.

2.2. The Impact of Hydrogen

2.2.1 Zero and Low Hydrogen Pressure

The deoxygenation of fatty acids on Pd/ C catalyst has been studied mostly in the absence of H2 or at low H2 concentration of 5 or 10 mol %, primarily in a stirred tank reactor.

Murzin’s group reported that at 300 °C and 17 bar of total pressure, S.A. conversion reached 41% under pure He after 300 min of reaction time, whereas, under

5 vol % H2 in Ar and pure H2, the S.A. conversion reached 62% and 49% respectively after 360 min of reaction time. [15] The initial reaction rate of S.A. conversion was

higher in He than in 10% H2 in He. However, under inert atmosphere, the reaction rate dropped after 10 min [16] due to catalyst deactivation. At 40% conversion of S.A., the selectivity to olefins was 11.1% in He and decreased to 9.3% in H2–Ar. However, the selectivity to aromatics under H2–Ar (1.4%) was one third the selectivity to them under

He (4.5%). [15] It was proposed that the inhibition of S.A. deoxygenation at high hydrogen partial pressures (PH2) may be a consequence of competitive adsorption of S.A. and H2. [9] [16]

In a different study, the same group examined the deoxygenation of Lauric Acid (L.A.) on 5 wt % Pd/C, at 300 ˚C and 2 MPa of inert gas in a stirred tank reactor. The initial concentration of L.A. was 0.05mol/L. A total conversion of 43% with 91 % selectivity to hydrocarbons was achieved after 6 hours of reaction time, under inert atmosphere. The main liquid product was undecane. Undecene was also observed. The selectivity to undecane and undecene were 70 % and 19 % respectively. [17] Generally, in these studies, the gas phase data were more useful to understand the pathways at which the deoxygenation of the fatty acids took place on the catalyst surface. Murzin’s group reported that when they operated under inert atmosphere, CO2, the product of decarboxylation, was the predominant product. They also observed CO, the product of decarbonylation but in much lower concentrations. Immer and Lamb also reported a high selectivity of 95% to CO2 when they continuously injected S.A. into a stirred tank reactor that contained 5 wt % Pd/C. Their experiment was conducted at 300 °C and 5% H2 in 15 atm of total pressure. They also observed a sudden switch from decarboxylation to

decarbonylation when higher H2 pressure was applied. Initially, the decarboxylation inhibition was reversible by simply reducing the amount of H2. However, after 10 hours of operation under conditions that favored decarbonylation, decarboxylation activity could not be recovered. [16] This confirms that the presence of H2 drives deoxygenation in a different pathway.

Over the last few years, three reaction pathways have been considered to explain how the fatty acids deoxygenate on the surface of Pd/C. The three possible reactions are shown in

Figure 2.2.1.

C17H34 Direct Decarbonylation

C H C17H35-COOH Pd/C 17 36 Decarboxylation

C H 17 34 Decarboxylation

Figure 2.2.1 Proposed S.A. reaction pathways on Pd/C catalyst in zero to low PH2

Minor amounts of H2 were always observed during studies conducted under inert gas.

Murzin’s group reported that they found small amounts of CO when they operated under pure Ar. They suggested that decarbonylation of the carboxyl group occurred due to the

presence of a minor amount of hydrogen generated from the solvent or from the H2 saturating the Pd surface from catalyst reduction. In an attempt to investigate this further,

Immer and Lamb processed pure dodecane on Pd/C, under inert atmosphere at typical deoxygenation conditions. They reported that the amount of H2 generated corresponded to 2% dehydrogenation of the dodecane. The excess H2 that evolved was equivalent to

6+/- 2% of the S.A. moles converted. [16] The liberated H2 changed the pathway at which fatty acid deoxygenation proceeded. The presence of adsorbed hydrogen led to fast hydrogenation of the adsorbed carboxylate before it could undergo decarboxylation.

[9]

2.2.2 Reaction under Pure Hydrogen

Studying the deoxygenation of fatty acids under a pure H2 atmosphere had rarely been pursued. [15] Murzin’s group only recently investigated the liquid products obtained under pure H2 atmosphere in a stirred tank reactor. The group reported that at 300 ˚C and

2 MPa of pure H2, by using 0.1 g of 5 wt% Pd/C, the overall conversion of L.A. significantly increased to 65% (from 43% under inert) with the same selectivity to hydrocarbons of 91%. The products were only saturated hydrocarbons, predominantly undecane with 86% selectivity followed by dodecane with 4% selectivity. Lauryl alcohol was also observed in quantifiable amounts but only traces of lauryl aldehyde were detected, which suggests high aldehyde reactivity. [17] The group suggested that in the presence of H2, the deoxygenation of L.A. proceeds in two different pathways, direct decarboxylation and hydro-deoxygenation. In decarboxylation, a CO2 molecule cleaves from the L.A. molecule yielding undecane that is one carbon less than the parent fatty 13

acid. In hydrodeoxygenation, the group suggested that the H2 reduces the carboxyl group forming the aldehyde, which can either decarbonylate, producing a hydrocarbon and a

CO, or be reduced, producing alcohol. The alcohol can be reduced once more, yielding the hydrocarbon with the same number of carbons as the parent fatty acid. [17]

The group suggested that alcohol hydrogenation seemed to be unfavorable. Their suggestion was based on the experimental results they observed when the alcohol was fed as primary feed. The products were undecane with 86% yield versus 4.5% yield to dodecane, which was the highest yield of dodecane observed during their entire study.

Therefore, the group presumed that the deoxygenation of aldehydes is more favorable.

Their argument was supported by observations from reference [18], which confirms that aldehydes of boiling points higher than 200 °C decompose easily on Pd via decarbonylation at the range of temperatures used during the deoxygenation of the fatty acid. The product of the decarbonylation of aldehydes is a mixture of alkanes and olefins. The product distribution depends on the conditions and the nature of aldehyde

[18].

In another study conducted under 12 bar of pure H2 in a stirred tank, Peng investigated the deoxygenation of Palmitic Acid (P.A.) by using 1.00 g of it dissolved in 100 g of dodecane over 0.5 g of 5 wt % Pd/C. The experiment was conducted at 260 °C. After 6 hours of reaction time, 20% overall conversion was detected with 98% selectivity to n-

C15 and 1.9% n-C16. No other products were reported. [8]

With regard to the aldehyde reduction to alcohol, Peng reported an estimated equilibrium ratio (Calcohol/Caldehyde) of 57 at 260˚C and 40 bar of pure H2.

In conclusion, in the presence of H2, the reaction was recognized to be quite complex. In addition to decarboxylation and direct decarbonylation of the fatty acid, the presence of hydrogen opens the hydro-deoxygenation pathway which is just the start of a complex network of reactions that yield a different product distribution than the one observed under inert atmosphere. Figure 2.2.2 includes all suggested reaction pathways in the presence of H2.

C17H34 Direct Decarbonylation

Pd/C + H2

C17H35-COOH C17H36 Decarboxylation Pd/C

Pd/C + H2

C17H35-CHO Hydro-deoxygenation

Pd/C -/+ H2

C17H35-CH2OH

Pd/C + H2 H2O

C18H36

Figure 2.2.2 Proposed S.A. reaction pathways on Pd/C catalyst under high PH2

2.3. Effect of the Initial Concentration of the Fatty Acid

The initial reaction rate of S.A. deoxygenation was investigated by Murzin’s group in a fixed-bed reactor by using Pd/C at 300 °C and 17 bar of He. The rate of deoxygenation increased with the increase of the initial concentration of S.A. When the initial concentration was 0.8 mol/L, the reaction rate was 2.6-fold lower than when it was 1.6 mol/L. A complete conversion of S.A. was achieved after 180 min in the former case versus only 40% after 300 min in the latter case. [12] The initial concentration of S.A. was decisive for the catalyst deactivation. The higher the initial concentration, the faster the catalyst deactivation occurred. [19]

2.4. Effect of Chain Length

The influence of alkyl chain length on deoxygenation kinetics was once investigated by

Immer and Lamb for a series of C10–C18 fatty acids using a Pd/C catalyst. They observed that by decreasing the fatty acid carbon number, reaction time and H2 consumption increased. [20] In contrast, Murzin’s group reported that the rates of deoxygenation of series C17-C22 fatty acids on Pd/C were virtually identical. [14]

2.5. Summary

The highest level of fatty acid deoxygenation was achieved on Pd/C catalyst. Studies conducted to investigate the deoxygenation of fatty acids in the presence of Pd/C were often carried out at 300˚C and under inert atmosphere or low PH2. The main products were saturated and unsaturated hydrocarbons with one carbon less than the parent fatty acid.

Under inert atmosphere, the fatty acid deoxygenation proceeded at the fastest initial rate.

However, the rate of fatty acid deoxygenation declined within the first hour, due to rapid catalyst deactivation. Attempts to recover catalyst activity were never reported successful.

In the presence of H2, even in small amounts, the initial rate of fatty acid deoxygenation was slower than under inert atmosphere; however, the catalyst lifetime was substantially extended. The main products were also saturated and unsaturated n-hydrocarbons with one carbon less than the fatty acid, but the yield to saturated hydrocarbons was higher than when operated under inert atmosphere.

The proposed reaction pathways for the deoxygenation of fatty acids in inert or low PH2 atmosphere are: decarboxylation, which yields the saturated n-hydrocarbon with one carbon less than the parent fatty acid and the co-product CO2; direct decarbonylation, which yields an olefin with one carbon less than the parent fatty acid and the co-product

CO; and hydrogenation, which yields the aldehyde.

In pure H2, the products were mainly saturated hydrocarbons, with 86% selectivity to the hydrocarbon with one carbon less than the fatty acid; the hydrocarbon with the same carbon number also appeared but at a rate of about 1/20th of the rate of the one carbon less than the fatty acid hydrocarbon. These yields were reported after 300 min of reaction time. A relatively high amount of alcohol was also observed. In the presence of H2,

Murzin’s group suggested that the deoxygenation of the fatty acid proceeds via decarboxylation and hydro-deoxygenation.

CHAPTER 3

EXPERIMENTAL APPROACH

3.1. Studying the Kinetics of S.A. Deoxygenation in Place of

Triglyceride

The initial deoxygenation in the transformation of triglycerides involves saturated fatty acids dissolved in saturated triglycerides. Table 3.1 1 shows that 18 carbon-long fatty acids are the most common fatty acids in biological oils, which makes S.A. a typical reaction intermediate in the deoxygenation of biological oils. For this reason, S.A. was elected to be the molecule of study. The S.A. was dissolved in an isomerized C24 paraffinic solvent. The solvent role is to perform as a substitute to the saturated triglycerides.

Table 3.1.1 Fatty acid composition in a selected group of biological oils [15] [16] [17]

Biological Oil Fatty Acid composition [ wt%] 16:0 18:0 18:1 18:2 18:3 Algae oil 10.11 2.66 66.92 14.56 1.09 Beef Tallow 23.6 19.4 42.4 2.9 0.9 Corn 11.67 1.85 25.16 60.6 0.48 Cottonseed 28.33 0.89 13.27 57.51 0.00 Peanut 11.38 2.39 48.28 31.95 0.93 Rapeseed 3.49 0.85 64.4 22.3 8.23 Soybean 11.75 3.15 23.26 55.53 6.31 Sunflower 6.08 3.26 16.93 73.73 0.00

3.2. Reactor Type, Differential Operation, and S.A. Concentration

This study was schemed to investigate the kinetics of S.A deoxygenation by using 3wt %

Pd/C catalyst in a biological oil refinery-like setting. For this reason, the experiments were conducted in continuous mode by using a fixed-bed reactor. In order to simplify the study of the kinetics of S.A. deoxygenation, a uniform S.A. concentration across the reactor bed was required. For this reason, the entire study was conducted under differential conditions, by limiting the S.A. conversion to 10 percent. The limited conversion allowed me to assume a uniform S.A. concentration across the catalyst bed.

Given the low overall S.A. conversion of 10 percent, in order to be able to accurately quantify the liquid products, a relatively high initial S.A. concentration was required. For this reason, a 20 wt. % S.A concentration was used. At 20 wt % S.A. concentration and

10% S.A. conversion, the concentration of all products combined would make only 2 wt

% of the total liquid collected.

3.3. Importance of Hydrogen in Biological Oil Upgrading

Fatty acids are not typically used as feedstock in refineries but biological oils are. In biological oil refining, hydrogen is an essential element. Its importance is due to three reasons. First, as Table 3.1.1 shows, the majority of fatty acids found in naturally- occurring triglycerides have at least one double bond. The double bond saturation always precedes the hydrogenolysis of the unsaturated acyl groups in the triglycerides, which means that in the absence of H2, double bond saturation will not occur causing a substantial decrease in the rate of the fatty acid deoxygenation. This was confirmed by an investigation conducted by Immer and Lamb. Immer and Lamb studied the kinetics of stearic and oleic acid deoxygenation. The rates of deoxygenation of the two fatty acids with different degrees of saturation were found to be indistinguishable in 10% H2 in He.

However, under He atmosphere, the deoxygenation of oleic acid was considerably slower. Only 11% oleic acid conversion was achieved after 3 hours while the S.A. conversion reached 92% at the same operating conditions. [16] The authors proposed that the lower deoxygenation rate in the absence of hydrogen is due to the cis-C–C double bond adsorption which could compete with carboxylic group adsorption. [9] The second reason that requires the use of H2 is the catalyst lifetime. In the absence of H2, the products of deoxygenation contain significant amounts of olefins and aromatics (at higher temperatures) [19, 13] that quickly adhere to the catalyst surface blocking the catalytic sites from further adsorbing and eventually disabling the catalyst. Under inert atmosphere, catalyst deactivation was inevitable. [11] Finally, the third reason, it was also evident that hydrogen drives the transformation of the fatty acids into reaction

pathways that yielded other products that might be desired for certain fuel fractions.

These products include the alcohol and the hydrocarbon with the same number of carbons as the parent fatty acid. These products were not observed under inert atmosphere. [17]

Therefore, H2 is strongly believed to be an essential component to this process and its effect needed to be examined in this study. The presence of hydrogen helps the reaction to proceed at a faster rate, extends the catalyst lifetime, and yields products that may be desirable for particular fuel fractions.

3.4. Materials

3.4.1 Catalyst Preparation

A 3 wt % Palladium on activated carbon support catalyst was used throughout this study.

The catalyst was synthesized by the incipient wetness impregnation method. First, a

-2 PdCl4 solution was prepared by dissolving PdCl2 crystals purchased from Sigma-

Aldrich in 3 M HCl. Norit RX3 EXTRA activated carbon extrudates were crushed and sieved to 40-60 mesh size. The pore volume of the crushed carbon was estimated to be

1.60 grams H2O per gram of carbon. The density of the solution was assumed to be equal

-2 to the density of H2O. The solution was made so that 1.6 grams of PdCl4 solution contained 0.03 grams of Pd. The solution was then added slowly to the freshly crushed carbon in the amount of 1.60 grams of solution per1 gram of support while mixing. The

Pd impregnated carbon was then placed in the furnace at 150 ˚C and left to dry overnight.

3.4.2 Catalyst Activation

The catalyst activation temperature program was set as: ramp to 120 ˚C in 30 min, dwell for 4 hr, ramp to 350 ˚C in 45 min, dwell for 4 hr, step to the desired operating temperature and dwell. The catalyst was reduced in 100 mL/min H2 flow during the activation. The system pressure was set to 100 psig. The Pd dispersion was determined to be 22%. The Pd dispersion was measured by a third party, by using CO as the adsorbing gas in a Micromeritics volumetric adsorption instrument.

3.4.3 Feed Materials

S.A. (97% pure) was purchased from Fisher Scientific. The S.A. was solid at room temperature. The S.A. was dissolved in Synfluid PAO 2.5 cst (99% pure), a highly isomerized C24 hydrocarbon purchased from Chevron Phillips Chemical Company. The solvent needed for this study ought to meet two requirements. First, it had to have a high boiling point in order to stay liquid at the high temperature used throughout the study, and to have a low melting point, in order to stay melted at room temperature. Since isomerized C24 meets the two requirements, it was chosen to be the solvent.

The solvent was added to the S.A. to achieve the target S.A. concentration in the solution.

Although the solvent was liquid at room temperature, when it was added and mixed with the S.A., the mixture became solid at room temperature. For this reason, the S.A. in the

C24 solution was kept stirred and heated at 70 ˚C just until before it was used.

Estimating the C24 solvent vapor-liquid equilibrium behavior at operating temperature and pressure was important for two reasons. First, the increase in the solvent mole fraction in the gas phase would decrease the mole fraction of all the other gaseous species, including the hydrogen mole fraction (yH2), which consequently decreases the gaseous species partial pressure. And since the influence of PH2 was a major part of this study, it was important to estimate the solvent mole fraction in order to estimate the actual hydrogen mole fraction and subsequently the PH2.

The second reason was that the normal boiling points of C24 solvent and S.A. were 277 ˚C and 383 ˚C respectively. Therefore, at any set of conditions, the liquid C24 solvent would

have a higher evaporation rate than S.A. As the liquid C24 would turn to vapor, the S.A. concentration would increase in the liquid. Since the influence of S.A. concentration is another critical parameter in this study, it was necessary to estimate the amount of solvent that could potentially turn into vapor at the operating temperature and pressure.

Since vapor-liquid equilibrium data on the solvent are not available but the normal boiling of C24 solvent (277˚C) was close to the normal boiling point of n-C16 (281˚C), it was reasonable to assume that the two hydrocarbons would behave similarly under the same temperature and pressure. Therefore, the hexadecane vapor liquid equilibrium data which were available were used to estimate the mole fraction of the solvent in the gas phase. By using reference [21] and further extrapolation, the solvent mole fraction of C24 in the vapor phase was estimated at 0.07 at 280 ˚C and 100 psig of total system pressure.

3.4.4 Fixed-Bed Reactor

A schematic of the reactor system used in this study is shown in Figure 3.4.1. A ½ inch,

316 smooth-bore stainless steel tubular fixed-bed reactor was used in this investigation.

The reactor tube is 24 inches long, oriented vertically inside a 3-zone Applied Testing

Systems Inc. furnace with an 18 inch heated zone. Each 6 inch section of the heated zone can be controlled independently. The temperature within the 18 inch heated zone can be monitored by inserting a thermocouple in the 1/8 inch stainless steel thermo-well that is located in the center of the reactor tube. Liquid feed is poured into the liquid feed reservoir which feeds a heated 500mL Teledyne ISCO syringe pump that supplies the liquid to the reactor at the specified flow rate. Hydrogen and argon flow rates are controlled by Brooks mass flow controllers. The liquid and gas are mixed just before 24

entering the reactor. The liquid leaving the reactor is collected in a 500 mL stainless steel reservoir from which the samples are collected. The gas exiting the reactor passes through a back pressure regulator, and then can be directed by a switching valve to the

FTIR or the GC depending on the type of analysis required.

In order to keep the S.A. in C24 solvent solution melted, the feed reservoir, syringe pump, and the effluent liquid reservoir are kept at 72 ˚C. The heated section of the reactor system is surrounded by the dashed line in Figure 3.4.1.

The reactor tube was loaded as follows: first, the bottom fitting that included the thermo- well was attached to the reactor tube. Second, the reactor tube was filled up with silicon carbide (SiC), which is an inert material, to approximately 2 inches below the center.

Third, the experiment-specified amount of the undiluted 3 wt % Pd/C catalyst was added into the reactor tube. Fourth, the reactor tube was topped off with more SiC. Fifth, a small piece of glass wool was added to secure the tube contents. Finally, the top fitting was attached, and the tube was placed in the reactor system in preparation for catalyst activation.

A PI

B Tc Tc

D H C E

FTIR

PI K L Vent Tc I PI

GC (Offsite)

Figure 3.4.1 Schematic of the reactor system

A: Liquid feed reservoir; B: Hydrogen mass flow controller; C: Argon mass flow controller; D:500 mL ISCO high pressure syringe pump; E: 1/2" reactor tube; F: Three zone furnace; G: Heated liquid product collector (65-80 ˚C); H: Dashed lined zone is heated zone; I: Cold Trap; K: Back-pressure regulator; L: Two- position valve to flow to FTIR or GC. Tc: Thermocouple, PI: Pressure indicator

3.4.5 Liquid Analysis

The liquid feeds and samples were analyzed at the Air Force Research Laboratory by using gas chromatography-ﬂame ionization detection (GC-FID) using an Agilent model

7890 gas chromatograph fitted with a 30 m DB-5MS column having 0.25 mm inner diameter (ID) and a 0.25-μm film. The GC temperature program began with an initial temperature of 40 °C (3-min hold) followed by ramping (10 °C/min) to 325 °C (20- minute hold). The samples were diluted 1 to 10 in carbon disulfide. Samples (1 μL) were injected onto the column using a 100:1 split with a constant column H2 carrier gas flow rate of 1 mL/min. The GC injector temperature was 300 °C, and the detector was held at

350 °C.

The feed solution composition was always confirmed before the feed was used. The S.A. was 97% pure and the C24 solvent was 99% pure. The feed analysis often showed traces of some of the hydrocarbons of interest but always <0.02%. For this reason, when the liquid sample analysis was received, any hydrocarbon concentration of interest present in the feed was always deducted from the hydrocarbon concentration in the product liquid sample.

3.5. Study Framework

This study consists of two parts. The first part that is presented in Chapter 4 discusses the effect of operating conditions on the rate of S.A. deoxygenation and the rate of liquid products formation. The liquid products were identified early on and a reaction network was established for the process based on the observed gas and liquid product species.

The goal in the first part was to study the impact of temperature, S.A. concentration, and hydrogen pressure on the rate of S.A. deoxygenation and the rates of product formation in a systematic way. For this reason, the apparent activation energies and the reaction orders with respect to S.A. and hydrogen were estimated based on the overall rates observed. These parameters were used to derive rate expressions that can be used to predict the rate of S.A. deoxygenation of S.A and the rates of production of the different products at any set of conditions. In this part of the study, the measurements were obtained on a lined out catalyst that had been on stream for several days. The catalyst stability was tested from time-to-time by reproducing results that were established at reference points.

The goal of the second part of the study which is discussed in Chapter 5 discusses the contribution of each individual reaction to overall rates of the hydrocarbons production.

This part was based on the proposed reaction pathways scheme suggested to explain the reaction network based on the liquid products observed.

CHAPTER 4

APPARENT KINETICS OF THE S.A. DEOXYGENATION AND LIQUID SPECIES PRODUCTION

4.1 Effect of Temperature

4.1.1 Experiment

This part of the study was conducted under pure H2. The H2 entered the reactor at 20 mL/min. The system pressure was set to 117 +/- 3 psig and 5.00 g of the 3 wt % Pd/C were used. The 19.40 wt % S.A. in C24 solvent was fed at 0.104 mL/min, delivering

6.01∙10-5 mol/min of S.A. At this liquid flow rate, the liquid Weight Hourly Space

Velocity (WHSV) = 1/hr.

A temperature range of 230-260˚C was found suitable for this part of the investigation

(4.1). The S.A. conversion fell between 1.92 and 9.73% over this temperature range, which satisfied the differential operation target conversion of less than 10%.

This part of the study (4.1) was conducted on one charge of catalyst. At the beginning of the first run, the catalyst had been on stream for 183 hours. The last sample was taken at the 416th hour. During the study, the catalyst stability was tested from time-to-time by reproducing results generated at established reference points.

In order to examine the temperature impact on the rate of S.A. deoxygenation and the rate of product formation, the operating temperature was varied by 5 °C increments within the selected range of 230-260 °C, while the system pressure, liquid flow, and H2 flow were unchanged.

Typically, before the end of the day, an operating temperature would be established and the system was allowed to reach steady state overnight. The following morning, the transient liquid processed overnight was collected and discarded. The liquid was processed for another 8 hours at steady conditions, after which, the sample was collected and sent for analysis. The cycle started again by setting the next desired operating temperature.

In addition to the 6 data points collected between 230-260 °C, the process was studied at

300 °C. The decision to study the process at 300 °C was due to two reasons. First, as mentioned earlier, the goal is to understand the kinetics of the transformation at similar conditions to those used in commercial settings. Typically, biological oil upgrading is conducted at a temperature of 300 to 330 °C. [8] Second, most of the references studied the process at 300 °C. Therefore, operating at the same temperature would facilitate evaluating and comparing results. In order to maintain the level of S.A. conversion under

10 % at 300 °C, only 0.50 g of the 3 wt % Pd/C was used. Also, a much higher WHSV =

30/hr. was used, in order to decrease the reaction time.

4.1.2 Results

4.1.2.1 Effect of Temperature on the Deoxygenation of Stearic Acid

Table 4.1.1 shows the effect of temperature on the S.A. conversion and the rate of S.A. deoxygenation (disappearance) over the range of 230-260 °C, and at 300 °C.

Table 4.1.1 Effect of temperature on the S.A. conversion and the rate of S.A. deoxygenation T S.A. Rate of S.A. disappearance -1 -1 [˚C] % conversion [mol∙ gcat ∙min ] 230 1.92 2.30E-07 235 2.77 3.32E-07 240 4.5 5.41E-07 245 5.66 6.79E-07 255 8.52 1.00E-06 260 9.73 1.21E-06 300† 5.50 1.99E-05

Measurements were taken over the range of 230-260 °C, at 117 +/- 3 psig, under pure H2 and by using 5.00 g of 3 wt % Pd/C. The 19.40 wt % S.A. in C24 solution entered the system at 0.104 mL/min, which is equivalent to 6.01∙10-5 mol/min of S.A. The WHSV = 1/hr. †The data collected at 300 °C were taken by using 0.50 g of 3 wt % Pd/C, under 117+/- 3 psig, WHSV=30/hr.

In order to confirm the absence of intraparticle mass transfer resistance in the catalyst, the observed rate of S.A. deoxygenation at 230˚C presented in Table 4.1.1 was compared to the rate of palm stearin hydrogenation that Manoj and Mahajani reported from a study that was conducted in a similar system. In reference [22], Manoj and Mahajani confirmed that the rate of palm stearin hydrogenation was not limited by intraparticle mass transfer resistance. This was confirmed by using two different methods. First, they estimated the effective diffusivities of hydrogen and fatty acids by using the Wilke-

Chang Equation. [23] They found that the effect of mass transfer inside the catalyst pores was negligible. Second, they examined the effect of the particle size on the rate of hydrogenation. They observed that the catalyst particle size had no effect on the rate of reaction, and hence, pore diffusion was deemed to be absent.

Since the rate of palm stearin hydrogenation reported in their study was 4 orders of magnitude faster than the observed rate of S.A. deoxygenation, I am confident that these observed rates were not restricted by intraparticle mass transfer either, which means that the observed rates are based on the true kinetics.

In order to derive kinetic expressions for the process, the following approach was used:

Since, the overall reaction can be written as:

S.A. + H2 P1 +P2 +P3+ …..Pn

where, P1, P2, P3 …Pn are product 1, product 2, product 3 to product n.

According to the law of mass action, the rate of S.A. disappearance can be written as:

. [24] Equation 4.1-1

And, the rate of product (P) formation can be written as:

. [24] Equation 4.1-2

Where,

is the rate of S.A. disappearance.

[S.A.] is the S.A. concentration. k S.A. is the rate constant for the S.A. deoxygenation.

αS.A. is the reaction order with respect to S.A. for the S.A. deoxygenation.

[H2] is the H2 concentration.

βS.A. is the reaction order with respect to H2 for the deoxygenation of S.A. is the rate of product formation. [P] is the product concentration. kP. is the apparent rate constant for the S.A. deoxygenation.

αp is the reaction order with respect to [S.A.] for product P.

βp is the reaction order with respect to H2 for product P.

Since, KH∙[H2] = yH2 P =PH2 , [25] Equation 4.1-3

where, KH is the H2 Henry’s constant in liquid solution, yH2 is the H2 mole fraction in the gas phase, and P is the system pressure. 33

∴ [H2] = . Equation 4.1-4

Substituting from equation 4.1-4 in equation 4.1-1 yields:

Equation 4.1-5

Similarily, substituting from equation 4.1-4 in equation 4.1-2, the rate expression for product P could be written as:

Equation 4.1-6

And since, - where, A0 is the pre-exponential factor, Ea is the activation energy and

R is the gas constant = 8.314∙10-3 kJ∙mol-1∙K-1, substituting from equation 4.1-7 in equation 4.1-5 yields:

Equation 4.1-8

Taking the natural logarithm of both sides of equation 4.1-8 yields:

Equation 4.1-9

Although, the S.A. concentration was slightly changing as it was consumed during the process, operating under differential conditions by limiting the S.A. conversion to 10 percent allowed me to assume that the S.A. concentration was unchanged. At this point

since, A0 S.A., [S.A.], PH2, KH, and the reaction orders with respect to [S.A.] and H2 were invariant with respect to temperature; equation 4.1-9 can be rewritten as:

Equation 4.1-10

The activation energy is a parameter that indicates the sensitivity of a single reaction rate to temperature, but the observed rate of S.A. deoxygenation is a combined rate of two reactions (as shown in Figure 4.1.2). For this reason, the “apparent” activation energy was estimated instead of the actual activation energy by using the observed overall rate of

S.A. disappearance as a function of temperature.

In order to estimate the apparent activation energy (App.Ea S.A.disapp.), the natural logarithm of the rate of S.A. disappearance shown in table 4.1.1 was plotted versus

1/T [K] in Figure.4.1.1. The slope of the line is = . Multiplying the slope by (–R) yields App.Ea. S.A.disapp. =148 kJ/ mol.

-10.50

]) ])

1 -

-11.50

∙min

- cat -12.50

[mol∙g y = -17805x + 20.091

-13.50

S.A.disapp. ln(r -14.50

-15.50 0.0017 0.00175 0.0018 0.00185 0.0019 0.00195 0.002 0.00205

1/T[K]

Figure 4.1.1 The effect of temperature on the rate of S.A. deoxygenation in Arrhenius form Measurements were taken over the range of 230-260 °C, at 117 +/- 3 psig of pure H2, and by using 5.00 g of 3 wt % Pd/C. The 19.40 wt % S.A. in C24 solution entered the system at 0.104 mL/min, delivering 6.01∙10-5 mol/min of S.A. The WHSV = 1/hr. † Data collected at 300 °C were taken by using 0.50 g of 3 wt % Pd/C, at 117+/- 3 psig, WHSV=30/hr

4.1.2.2 Effect of Temperature on the Rate of Liquid Species Production

In this section, the effect of temperature on the rate of liquid species production was investigated over the temperature range of 230-260 °C. N-heptadecane, n-octadecane, and stearyl stearate (S.S.) were the three products observed in the liquid sample. Table

4.1.2 shows the effect of temperature on the rate of liquid species production and the selectivity to products. Based on the liquid products observed, I propose that the deoxygenation of S.A. proceeded as shown in Figure 4.1.2.

Table 4.1.2 Effect of temperature on the rate of liquid species production

T Rate of Liquid Species Production Selectivity

-1 -1 [˚C] [mol∙gcat .min ] %

n-C17 n-C18 S.S. n-C17 n-C18 S.S. 230 1.94E-07 1.32E-08 2.30E-08 84 6 10 235 2.76E-07 2.09E-08 3.58E-08 83 6 11 240 4.34E-07 2.62E-08 8.09E-08 80 5 15 245 5.68E-07 3.06E-08 8.08E-08 84 5 12 255 9.20E-07 4.77E-08 6.03E-08 90 5 6 260 9.40E-07 5.60E-08 1.70E-07 81 5 14 300† 1.94E-05 2.13E-07 3.24E-07 97 1 2

Measurements were taken over the range of 230-260 °C, at 117 +/- 3 psig of pure H2, by using 5.00 g of 3 wt % Pd/C. The 19.40 wt % S.A. in C24 solution entered the system at 0.104 mL/min, delivering 6.01∙10-5 mol/min of S.A. The WHSV = 1/hr. † Data collected at 300 °C were taken by using 0.50 g of 3 wt % Pd/C, at 117+/- 3 psig, WHSV=30/hr.

Hydrogenolysis C18H38 Decarboxylation

C17H35-COOH C17H36 Decarbonylation Pd/C

Pd/C + H2

C17H35-CHO Hydrodeoxygenation

Pd/C -/+ H2

C17H35-COOH Condensation C17H35-COO-C18H37 C17H35-CH2OH Reduction

Heat H2O Pd/C + H2

C18H38 Reduction

Figure 4.1.2 Reactive pathways in the deoxygenation of S.A. conversion over a Pd/C catalyst in the presence of H2

Figure 4.1.2 shows that the observed n-C17 was a combined product of the decarboxylation of the S.A. and the decarbonylation of the highly reactive aldehyde and that the observed n-C18 was a combined product of the complete hydrogenation of S.A. and the S.S. hydrogenolysis. At this point, given the information available, it was not yet possible to determine the individual reaction pathway contribution to the total observed rate of n-C17 or n-C18 production.

In order to understand the sensitivity of the rate of product formation to temperature, the apparent activation energy was to be estimated. As shown in Figure 4.1.2, the observed rate of each of the hydrocarbon products is a combined rate of two reactions. For this 38

reason and as mentioned in section 4.1.2, in order to determine the sensitivity of the combined rate of each hydrocarbon production to temperature, “apparent” activation energy was derived based on the observed rate of each of the hydrocarbon species as a function of temperature.

Figure 4.1 3 shows each of the n-C17 and n-C18 production rate dependence on temperature in Arrhenius format. The natural logarithms of the rate of n-C17 and n-C18 production were plotted versus 1/T [K]. The data shown in Figure 4.1.4 were combined from 3 different sets: A, B and C. Set A included 6 data points that were taken over the temperature range of 230-260 °C, all 6 data points were collected on one charge of 5.00 g of 3 wt% Pd/C, and WHSV = 1/hr. Data included in set B were one data point taken at

300 °C by using 0.50 g of catalyst and WHSV = 30/hr added to the 6 data points included in set A. Set C consisted of two data points (in the dashed line box) that were later added to the plot. The two data points were collected at 265 and 270 °C, by using 1.00 g of catalyst and WHSV =10/hr. The two points were originally collected during experiments related to gas phase measurements, which are discussed in Chapter 5. Those experiments were conducted on a different batch of catalyst and a much lower S.A. concentration of

1 2.5 wt %. Despite the differences, upon adjusting the rate of n-C17 and n-C18 production to 19.4% wt % S.A, by using the reaction orders for n-C17 and n-C18 production with

1 In order to add the two data points on the plot shown in figure 4.1.4, the n-C17 and n-C18 reaction order with respect to S.A. estimated in the section (4.2)was used to make the adjustment for the S.A concentration from 2.5 to 19.40 wt %

respect to S.A. that is derived in the next section (4.2), the rate of n-C17 and n-C18 production fell within reasonable deviation from the trendline. In addition to the possible variation in the catalyst characteristics, the deviation could also be due to inaccuracy in quantifying the small amounts of n-alkanes, particularly the n-C18, which was found at concentrations of less than 0.01 wt % in the two samples.

-10.5 C17-A C18-A -11.5 C18 B C17-B -12.5 C17-C C18-C

y = -19145x + 22.565

]) -13.5 R² = 0.9726

∙min 1 - -14.5

-15.5

[mol∙gcat

ln(r -16.5 y = -11248x + 4.4285 R² = 0.9654 -17.5

-18.5 0.00174 0.00179 0.00184 0.00189 0.00194 0.00199 1/ T[K]

Figure 4.1.3 Effect of temperature on the rate of n-C17 and n-C18 production in Arrhenius form

Three sets of data are included on this figure (Sets A, B and C) A: Data included 6 data points that were taken over the temperature range of 230-260 °C, all 6 data points were collected on one charge of 5.00 g of 3 wt% Pd/C, and WHSV = 1/hr. B: Data were one data point taken at 300 °C by using 0.50 g of catalyst and WHSV = 30/hr added to the 6 data points included in set A. C: Two data points (in the dashed lined box) were taken at 265 and 270 ˚C by using 1.00 g of catalyst and WHSV =10/hr. Data sets A and B were taken on two different catalyst charges but from the same batch of 3 wt % Pd/C catalyst, by using 19.40 wt % S.A. dissolved in isomerized C24 solvent at 117 +/-3 psig. Data set C was collected on a different batch of 3 wt % Pd/C catalyst by using 2.5 wt % S.A. All rates were adjusted to 19.40 wt % S.A concentration. All rates presented were normalized to 1.00 g of 3wt% Pd/C

In order to estimate the activation energy of the rate of liquid species production, equation 4.1-7 was substituted in equation 4.1-6 to yield:

Equation 4.1-11

And,

Equation 4.1-12

Taking the natural logarithm of both sides of Equation 4.1-11 yields:

. Equation 4.1-13

Since A0_n-alkane, [S.A.], , KH, and PH2 were invariant with respect to temperature, equation 4.1-13 could be rewritten as:

Equation 4.1-14

From Figure 4.1.4, the slope of the line of each product was where, R (Gas constant) = 8.314∙10-3 kJ/mol∙K. Multiplying the slope by (-R) yields:

The approach used to calculate the apparent activation energy for n-C17 and n-C18 was not applicable in the case of S.S. When the values of ln(r S.S.) were added to the plot shown in Figure 4.1.4 as a function of T, the data did not follow any particular trend. Therefore, no activation energy could be derived for the rate of S.S. production.

The relatively large amounts of S.S. observed required me to understand the nature of the

S.S. formation. So, as a first step, I decided to find out whether the alcohol condensation occurred catalytically or thermally. In order to do so, I designed an experiment in which a solution composed of 2 wt % octadecanol, 18 wt % S.A. and 80% C24 solvent was fed into a heated reactor tube that only contained SiC (no catalyst). The SiC was the same inert material used during the deoxygenation of S.A. experiments to fill the spaces under and above the catalyst bed in the reactor tube as explained in section 3.4.3. The homogeneous condensation was monitored at 260 °C and 100 psig of pure H2, which were the typical operating conditions used in the S.A. deoxygenation study.

The liquid sample analysis showed that octadecanol-S.A. condensation proceeded at a high rate. Ninety percent of the octadecanol fed condensed forming S.S. without catalyst.

This was the evidence that the observed S.S. was the product of the thermal condensation of the octadecanol with the unreacted S.A. in the reactor system. The high rate of S.S. production observed in the experiment was due to the abundance S.A. (18 wt% of the solution fed), with which the alcohol quickly condensed, forming the S.S.

In order to compare the amount of unreacted S.A. to the amount of alcohol produced, the ratio of the molar flux of S.A. leaving the reactor to the rate of S.S. production was

considered (which is equal to the rate of alcohol production). The ratio of S.A. molar flux leaving the reactor to the rate of S.S. production ranged between 102 to 105, which explains that at this high concentration of S.A, it is more likely for the alcohol to collide and condense with S.A.

With regard to where the condensation occurred in the reactor system, it was unknown.

However, the catalyst bed is positioned in the center of the reactor tube and the rest of the reactor tube is packed with SiC. When the alcohol was formed on the catalyst bed in the presence of S.A., the condensation might have occurred anywhere on the catalyst bed or on the SiC below the catalyst before exiting the tube, but based on the high rate of the thermal homogeneous condensation, I speculate that most of the condensation occurred within the catalyst bed. That is due to the higher liquid holdup in the catalyst compared to SiC. The high liquid holdup in the catalyst is due by the high internal volume of the carbon support (porosity).

4.2 Effect of Stearic Acid Initial Concentration

4.2.1 Experiment

In order to determine the effect of S.A. initial concentration on the rate of S.A. deoxygenation and the rate of liquid species production, three solutions with different

S.A. concentrations were prepared by diluting the S.A. with the isomerized C24 solvent.

The three solutions had 4.94, 10.30, and 20.64 wt % SA, which are equivalent to 0.14,

0.29, and 0.58 mol/L of S.A. in the solutions. This part of the study (section 4.2) was conducted at 260 °C, in pure H2, by using 5.00 g of 3 wt % Pd/C. The liquid flow was set so that it corresponded to a WHSV = 1/hr.

The system pressure in this section (4.2) was set to 930 +/- 1% psig. The reason for using such a high pressure was due to the fact that as the S.A. concentration decreased to 4.94 and 10.30 wt% in the feed solution, the rates of the products appearance as well as the

S.A. conversion were expected to decrease, which would make the product quantification difficult. Therefore, in order to make up for the decrease in S.A. concentration in the feed solution while maintaining the same amount of catalyst and temperature, the PH2 had to be raised. This will be demonstrated explicitly later.

The data presented in this section (4.2) were all collected on one charge of catalyst. When the first sample was taken, the catalyst had been on stream for 360 hours. The last sample was collected at the 433rd hour.

In order to examine the effect of the initial concentration of S.A. on the rate of S.A. deoxygenation and the rate of product formation, the sequence of concentration was fed 45

in a descending order. On the first day, the 20.64 wt % solution was fed and the system was allowed to reach steady state overnight. The following morning, the transient liquid processed overnight was collected and discarded. The liquid flowed for 8 hours after which the sample was collected and sent for analysis.

In preparation for the following concentration of 10.30 wt %, the syringe pump was emptied into the feed reservoir from which the liquid was pipetted out. About 50mL of the 10.30 wt % solution was used to rinse the syringe pump and feed reservoir, by repeatedly filling and emptying the pump into the feed reservoir. The liquid was pipetted out and discarded. The process was repeated 3 times. The same process was repeated before the 4.94 wt % solution was fed.

4.2.2 Results

4.2.2.1 Effect of S.A. Initial Concentration on the Rate of S.A. Deoxygenation

Table 4.2.1 shows the effect of S.A. initial concentration on the rate of S.A. deoxygenation.

Table 4.2.1 Effect of S.A. initial concentration on the rate of S.A. deoxygenation

S.A. S.A. Fed Rate of S.A. Disappearance -1 -1 -1 wt % [mol∙min ] [mol∙gcat ∙min ] 4.94 1.45E-05 8.00E-07 10.30 3.02E-05 9.59E-07 20.64 6.04E-05 1.65E-06

Measurements were taken at 260 ˚C, 930 +/-1% psig of total pressure, under pure H2 atmosphere, by using 5.00 g of 3 wt % Pd/C, and 0.104 mL/min of feed solution which corresponded to WHSV = 1/hr

In order to derive a reaction order with respect to [S.A.] for the deoxygenation of S.A., equation 4.1-5 developed in section 4.1 was used,

Equation 4.1-5 taking the natural logarithm of both sides of equation 4.1-5 yields:

Equation 4.2-1

Since k S.A., KH, and PH2 were invariant with respect to [S.A.], equation 4.2-1 could be rewritten as:

Equation 4.2-2

When was plotted versus ln [S.A.] as shown in Figure 4.2.1, was the slope of the line. Therefore, the reaction order of S.A. deoxygenation with respect to

[S.A.] was 0.50.

-13.20

-13.30

-13.40

])

1 - -13.50

y = 0.5039x - 8.4858

min ∙

- -13.60

cat g

∙ -13.70

-13.80 [mol

-13.90

disapp.

-14.00 S.A.

-14.10 ln(r -14.20 -11.40 -11.20 -11.00 -10.80 -10.60 -10.40 -10.20 -10.00 -9.80 -9.60 ln[S.A. [mol.mL-1]]

Figure 4.2.1 Effect of S.A. initial concentration on the rate of S.A. deoxygenation determined by plotting ln(rS.A.disapp.) versus ln[S.A.].

Measurements were taken at 260 ˚C, 930 +/- 1% psig of total pressure, under pure H2 atmosphere, over 5.00 g of 3 wt % Pd/C, and 0.104 mL/min of feed solution which corresponded to WHSV = 1/hr.

4.2.2.2 Effect of S.A. Initial Concentration on the Rate of Liquid Species

Production

Table 4.2.2 shows the effect of S.A. initial concentration on the rate of liquid species production and selectivity to products.

Table 4.2.2 Effect of S.A. initial concentration on the rate of liquid species production and selectivity to products

S.A Initial S.A. Rate of Liquid Species Selectivity Concentration Production wt [mol/mL] [mol∙g -1∙min-1] % % cat

n-C17 n-C18 S.S. n-C17 n-C18 S.S. 4.94 1.39E-04 3.96E-07 8.77E-08 1.29E-07 71 13 16 10.30 2.90E-04 5.69E-07 1.02E-07 3.80E-07 51 10 40 20.64 5.81E-04 7.68E-07 1.29E-07 7.53E-07 47 8 46

Measurements were taken at 260 ˚C, 930 +/-1% psig, under pure H2, over 5.00 g of 3 wt% Pd/C catalyst, and WHSV = 1/hr.

In order to derive a reaction order for each liquid product with respect to [S.A.], equation

4.1-6 developed in section 4.1 was used here as the starting point.

. Equation 4.1-6

Taking the natural logarithm of both sides of equation 4.1-6 yields:

Equation 4.2-4

Since kp, αp, , KH, and PH2 were invariant with respect to [S.A.], equation 4.2-4 could be rewritten as:

Equation 4.2-5

Equation 4.2-5 was used to graphically estimate the value of α for each product by plotting the values of ln(rP) of each product versus ln[S.A.] as shown in Figure 4.2.2. The apparent reaction order of n-C17, n-C18, and S.S. production with respect to [S.A.] are the slopes of the corresponding lines, 0.46, 0.27, and 1.24 respectively.

-14

y = 0.4629x - 13.821 R² = 0.9986

-14.5

)

]

1 -

y = 1.2375x - 13.368

∙min -15 R² = 0.9874 1

- n-C17 cat n-C18

-15.5 Stearyl Stearate

[mol∙g

ln (r ln y = 0.2689x - 15.735 -16 R² = 0.9796

-16.5 -2.2 -1.7 -1.2 -0.7 ln[S.A.[mol∙L-1]]

Figure 4.2.2 Effect of S.A. concentration on the rate of n-C17, n-C18, and S.S. production determined by plotting ln(rp) versus ln[S.A.]

Measurements were taken at 260 ˚C, 930 +/-1% psig, under pure H2, over 5.00 g of 3 wt% Pd/C catalyst and WHSV = 1/hr.

4.3 Effect of Hydrogen Pressure

4.3.1 Experiment

The effect of PH2 on the rate of S.A. deoxygenation and the rate of liquid species production was studied over PH2 range of 127- 931 (+/- 1%) psig. In this section (4.3.1),

5.00 g of 3 wt % Pd/C were used. The temperature was set to 260 °C and the system total pressure was at 931 +/- 1% psig. The 20.60 wt % S.A. solution was fed to the reactor at 0.104 mL/min, which corresponded to a WHSV = 1/hr. At this flow rate, S.A. entered the reactor at 6.40∙10-5 mol/min.

In order to obtain the desired PH2 without changing the total pressure, the H2 was diluted with argon. The desired PH2 was achieved by controlling the H2/Ar molar ratio (or volume ratio), while maintaining the total gas flow rate at 80 mL/min. The actual H2 mole fraction in the gas phase was determined by the online GC. The operating PH2 was determined by multiplying the H2 mole fraction obtained from the GC by the total pressure.

The data presented in this section (4.3) were all collected on one charge of catalyst. At the beginning of the first run, the catalyst had been on stream for 149 hours. The last sample was taken at the 360th hour. During the experiment, the catalyst stability was tested from time to time by reproducing results established at reference points. The rates of liquid species production were monitored at 6 different PH2.

Initially, the plan was to study the effect of PH2 on the rate of product formation over PH2 range of 0-930 psig. However, after purging the system and verifying the H2 removal by 51

GC, upon the start of liquid flow, small amounts of H2 were still observed in the system, even when the H2 supply valve was off. The same observations were mentioned in other studies. [14] [16] The H2 observed in the system was believed to be the H2 liberated from the mild dehydrogenation of the paraffinic C24 solvent on the Pd/C surface, a side reaction that took place simultaneously on the catalyst surface at the operating temperature.

In order to estimate the amount of H2 that could potentially be liberated from the solvent dehydrogenation during the S.A. deoxygenation study, the dehydrogenation of the neat solvent was investigated under pure Ar at 2 different operating temperatures, 260 and

280 °C. The neat C24 solvent was introduced to the 3 wt % Pd/C catalyst at

0.832 mL/min under pure Ar atmosphere. The Ar flowed into the reactor at 42 mL/min.

The system pressure was set to 100 psig and 5.00 g of 3 wt % Pd/C were used. The amount of H2 observed at 260 and 280 °C was 1.7 and 2.8 mole percent respectively of the total moles of the gas phase, which is equivalent to 3.9∙10-5 and 5.9∙10-5 mol/min of

-5 H2. This amount of H2 was substantial considering that S.A. was fed at 6.40∙10 mol/min, of which a maximum of 10% conversion was targeted. This means that the molar ratio of S.A. converted to liberated H2 was 1:6.09 at 260 °C and 1:9.21 at 280 °C.

Therefore, in reality, operating in a H2 free atmosphere was not achievable.

Although I attempted to examine the process at the lowest PH2 possible (below 20psig), the lowest PH2 I managed to operate under while maintaining 80 mL/min of total gas flow rate was approximately 127 psig. This is because the lowest H2 flow that the H2 mass flow controller could regulate was 2 % of its 500 mL/min capacity, which is 10 mL/min 52

of H2 making the lowest yH2 = 0.125. Given that the total pressure was set to 930 psig, the lowest PH2 would be just under 120 psig. The high total pressure was necessary to minimize the vaporization of the components at the high temperature used throughout the study.

4.3.2 Results

4.3.2.1 Effect of Hydrogen Pressure on the Rate of S.A. Deoxygenation

Table 4.3.1 shows the effect of hydrogen pressure on the rate of S.A. deoxygenation.

Table 4.3.1 Effect of hydrogen pressure on the rate of S.A. deoxygenation

PH2 Rate of S.A disappearance. -1 -1 [psig] [mol∙.gcat ∙min ] 127 1.18E-06 224 1.23E-06 331 1.32E-06 404 1.43E-06 556 1.39E-06 931 1.60E-06

Measurements were taken at 260 ˚C, 930 +/- 1% psig, by using 5.00 g of 3 wt % Pd/C. The 20.60 wt % S.A. in C24 solution was fed at 0.104 mL/min. The PH2 was achieved by controlling the H2/Ar molar ratio while maintaining the total gas flow rate at 80 mL/min.

In order to estimate the reaction order for the deoxygenation of S.A. with respect to H2,

Equation 4.2-1 developed in section 4.2 was used here as the starting point,

Equation 4.2-1

Since, kS.A., , [S.A.], and KH were invariant with respect to H2, equation 4.2-2 could be written as:

Equation 4.3-1

When was plotted versus as shown in Figure 4.3.1, βS.A., the reaction order with respect to H2, was found to be 0.15.

-13.3

-13.35

]) 1

- -13.4 min

∙ y = 0.15x - 14.412 1

- -13.45

cat

∙

1 -

-13.5 [mol -13.55

S.A. disapp. S.A. -13.6 ln(r -13.65

-13.7 4.50 5.00 5.50 6.00 6.50 7.00

ln(PH2 [psig])

Figure 4.3.1 Effect of hydrogen pressure on the rate of S.A. deoxygenation determined by plotting ln(r S.A.disapp.) versus ln(PH2)

4.3.2.2 Effect of Hydrogen Pressure on the Rate of Liquid Species Production

The effect of PH2 on the rates of liquid species production and product selectivity is shown in Table 4.3.2.

Table 4.3.2 Effect of hydrogen pressure on the rate of liquid species production and product selectivity

P yH2 Rate of Liquid Species Production Selectivity

-1 -1 [psig] [mol∙gcat ∙min ] %

n-C17 n-C18 S.S. n-C17 n-C18 S.S.

127 0.14 9.52E-07 5.68E-08 1.69E-07 81 5 14

224 0.24 9.06E-07 6.98E-08 2.54E-07 74 6 21

331 0.36 8.84E-07 8.84E-08 3.43E-07 67 7 26

404 0.43 9.68E-07 1.04E-07 3.60E-07 68 7 25

556 0.60 8.84E-07 1.01E-07 4.05E-07 64 7 29

931 1.00 7.68E-07 1.22E-07 7.53E-07 47 7 46

In order to understand the rate of liquid species production dependence on PH2, the same approach used in section 4.3.2.1 was followed. In order to derive a reaction order for each liquid product with respect to H2, Equation 4.2-4 developed in section 4.2 was used.

Equation 4.2-4

Since, kp, , [S.A.],and KH were invariant with respect to PH2, equation 4.2-4 could be written as, 56

Equation 4.3-3

Equation 4.3-3 was used to graphically estimate the apparent reaction order with respect to H2 for each rate of liquid species production, by plotting ln (rp) versus ln PH2 as shown in Figure 4.3.2. The slope of each line is β for that product. Thus, n-C17, n-C18 and S.S. reaction orders with respect to H2 were -0.09, 0.39, and 0.68 respectively.

-13.50

-14.00 y = -0.0876x - 13.416

]) -14.50

- ∙min

1 y = 0.6825x - 18.903

- -15.00

R² = 0.976 n-C17

cat g ∙ n- C18

-15.50 [mol

Stearyl Stearate p

-16.00 ln(r y = 0.3938x - 18.563 R² = 0.9433 -16.50

-17.00 4.80 5.30 5.80 6.30 6.80 ln(P [psig]) H2

Figure 4.3.2 Effect of PH2 on the rate of n-C17, n-C18 and S.S. production determined by plotting ln(rp) versus ln(PH2) Measurements were taken at 260 ˚C, 930 +/- 1% psig, by using 5.00 g of 3 wt % Pd/C. The 20.60 wt % S.A. in C24 solution was fed at 0.104 mL/min. The PH2 was achieved by controlling the H2/Ar molar ratio while maintaining the total gas flow rate at 80 mL/min.

4.4 Effect of Hydrogen Pressure on the Rate of n-alkane Production at

300 °C

4.4.1 Experiment

In order to derive kinetic data at similar conditions to those used in practice, I decided to investigate the effect of PH2 on the rate of S.A. deoxygenation and rate of product formation at 300 °C. [8]

During this part of the study, the system pressure was set to 315 psig +/- 1%. The desired

PH2 was achieved by controlling the H2/Ar molar ratio while maintaining constant gas flow rate of 150 mL/min. The S.A. concentration in the C24 solvent was 20.40 wt %. In order to keep the conversion under 10 %, the amount of catalyst was reduced to 0.50 g and the WHSV = 30/hr. The yH2 was monitored by the online GC.

Due to the high liquid flow rate, shorter processing time of 4 hours was allowed and was still sufficient due to the fast liquid turnover. The liquid processed over the first two hours was drained and discarded. The liquid was processed for another two hours at steady conditions, after which the sample was collected and sent for analysis.

4.4.2 Results

4.4.2.1 Effect of Hydrogen Pressure on the Rate of S.A. Deoxygenation at 300˚C

Table 4.4.1 shows the effect of PH2 on the rate of S.A. deoxygenation at 300 °C. The effect of PH2 on the rate of S.A. deoxygenation at 300 ˚C is presented in figure 4.4.1 by plotting ln rS.A versus ln(PH2).

Table 4.4.1 Effect of PH2 on the rate of S.A. deoxygenation at 300 °C

PH2 Rate of S.A. disappearance -1 -1 [psig] [mol∙gcat ∙min ] 20 1.42E-05 62 1.82E-05 114 2.00E-05 196 2.39E-05 310 3.24E-05

Measurements were taken at 315 +/- 1% psig total system pressure, using 0.50 g of 3 wt % Pd/C. The 20.40 wt % S.A. in C24 solution was fed at 0.3 mL/min. The PH2 was achieved by manipulating the H2/Ar mole ratio while maintaining constant gas flow rate at 150 mL/min.

-10.2

-10.4 y = 0.2756x - 12.035

R² = 0.9254

]) 1

- -10.6

min

∙

1 -

cat -10.8

∙ [mol

-11

S.A. ln(r ln(r -11.2

-11.4 2.5 3 3.5 4 4.5 5 5.5 6

ln (PH2 [psig])

Figure 4.4.1 Effect of PH2 on the rate of S.A. deoxygenation at 300 °C

4.4.2.2 Effect of Hydrogen Pressure on the Rate of Liquid Product Species

Production at 300 ˚C

Table 4.4.2 shows the effect of PH2 on the rate of liquid species production and the selectivity to liquid products at 300 °C. Figure 4.4.2 shows the effect of PH2 on the rate of liquid species production at 300 °C in a ln(rp) versus ln(PH2) plot.

Table 4.4.2 Effect of PH2 on the rate of liquid species production and selectivity to products at 300 °C

P yH2 Rate of Liquid Species Production Selectivity [psig] [mole∙g -1 min-1] % cat n-C n-C S.S. n-C n-C S.S. 17 18 17 18 20 5 1.41E-05 1.16E-07 0.0 0E+00 99.2 0.8 0.0 62 12.3 1.80E-05 1.30E-07 1.05E-07 98.7 0.7 0.5 114 25 1.94E-05 2.12E-07 3.24E-07 97.3 1.1 1.6 196 50 2.26E-05 3.30E-07 9.51E-07 94.7 1.4 3.9 310 100 2.74E-05 1.29E-06 3.68E-06 84.4 4.1 11.6

Measurements were taken at 310 psig total system pressure, by using 0.50 g of 3 wt % Pd/C. The 20.40 wt% S.A. in C24 solution was fed at a rate of 0.3 mL/min. The PH2 was achieved by by manipulating the H2/Ar mole ratio while maintaining the total gas flow rate at 150 mL/min.

Combining Equation 4.3-3,

, Equation 4.3-3 and Figure 4.4.2 show that the reaction order of n-C17, n-C18, and S.S. with respect to H2 were 0.22, 0.78, and 1.94 respectively.

-8

-10 y = 0.2281x - 11.872

R² = 0.9683

]) 1

- -12

.min

-14 cat

y = 0.7793x - 18.696

R² = 0.7377 n-C17

[mol.g

p -16 n-C18

ln(r S.S. -18 y = 1.9432x - 23.963 R² = 0.9887 -20 2.5 3 3.5 4 4.5 5 5.5 6 ln (P [psig]) H2

Figure 4.4.2 Effect of PH2 on the rates liquid species production at 300 °C

Measurements were taken at 315 +/- 1% psig total system pressure, using 0.50 g of 3 wt % Pd/C. The 20.40 wt% S.A. in C24 solution was fed at 0.3 mL/min. The PH2 was achieved by by manipulating the H2/Ar mole ratio while maintaining constant gas flow rate at 150 mL/min.

4.5 Discussion

4.5.1 Stearic Acid Deoxygenation Reaction Pathways

The deoxygenation of S.A. in the presence of hydrogen exhibits a complex reaction network. In addition to the three liquid products, n-C17, n-C18, and S.S., the gas phase analysis showed that the process generated CO and CO2, which are the co-products of the

octadecanal decarbonylation and the S.A. decarboxylation respectively. The amount of

CO exceeded the amount of CO2 under all sets of conditions that were studied.

By looking at the liquid and gas products, I propose that the S.A. deoxygenation proceeds in two pathways, decarboxylation and hydro-deoxygenation, as shown in Figure 4.1.2.

This proposal is in agreement with the reaction pathways that Murzin’s group suggested when they studied the deoxygenation of L.A. in reference [17]. In decarboxylation, as explained earlier, a molecule of CO2 is directly cleaved from the carboxyl group yielding the terminal product n-C17. The second pathway is the hydro-deoxygenation which in the first step yields the reactive intermediate aldehyde (octadecanal).

The aldehyde reacts further via two possible routes. The first, and more favorable, is decarbonylation in which a CO splits from the aldehyde yielding n-heptadecane or 1- heptadecene. Any heptadecene produced is immediately reduced due to the H2 abundance in the system, which makes n-C17 always the terminal product of the octadecanal decarbonylation. The second route is when the octadecanal proceeds in a step-wise reduction yielding the alcohol (octadecanol), which was never detected as a free molecule in the liquid samples. The octadecanol either condensed with S.A. forming

S.S., or further reduced yielding n-C18. The S.S. hydrogenolysis is another potential source of n-C18.

4.5.2 Effect of Reaction Parameters on the Rate of Stearic Acid Deoxygenation and

Liquid Species Production

4.5.2.1 Effect of Temperature

Table 4.1.1 shows that S.A. conversion varied from 1.92 to 9.73% over the temperature range of 230 - 260 °C at 117 +/- 3 psig of pure H2, in the presence of 5.00 g of 3 wt %

Pd/C and WHSV = 1/hr. The rate of S.A. deoxygenation increased by a factor of 5.3 as the temperature increased from 230 to 260 °C. When the temperature increased from 230 to 300 °C, the rate of S.A. deoxygenation increased 87 fold. The sensitivity of the rate of

S.A. deoxygenation to temperature is shown in Figure 4.1.1. The apparent activation energy of the S.A. deoxygenation was estimated at148 kJ/mol.

Table 4.1.2 shows the temperature influence on the rate of the liquid species production. n-C17 was the predominant product over the entire temperature range of 230 - 260 °C.

The high activation energy of 159 kJ/mol estimated from Figure 4.1.4 reflects the high sensitivity of the rate of n-C17 production to temperature.

The selectivity to n-C17 (the percent of the moles of S.A. converted to n-C17) ranged from

80 to 90% over the temperature range studied, while the selectivity to S.S., the product of the alcohol-S.A. condensation ranged from 6 -14%. n-C18 had the lowest but consistent selectivity of 5 - 6%. The low apparent activation energy of 93 kJ/moL shows the low dependence of the rate of n-C18 on temperature.

As the temperature increased from 230 to 300 °C, the rate of n-C17 production increased by more than 100 fold, while the rate of n-C18 and S.S. production only increased by a

factor of 16 and 14 respectively. This shows that at 300 °C, the decarbonylation of octadecanal is much faster than the octadecanal reduction route, which is the route of alcohol and n-C18 formation. Furthermore, as the temperature increased to 300 °C, the rate of S.A. decarboxylation is also expected to increase yielding more contribution to the overall rate of n-C17 production.

The faster rate of octadecanal decarbonylation and S.A. decarboxylation at 300 °C led the selectivity to n-C17 to jump to 97%, reducing the selectivity to S.S. and n-C18 to 2 and 1% respectively.

With regard to the values of the activation energy estimated for S.A deoxygenation and the activation energy for the liquid species production, the estimated activation energy for the S.A. deoxygenation was slightly lower than the apparent activation energy estimated for n-C17 but much higher than the apparent activation energy estimated for n-C18 production. This is due to the fact that the majority of S.A. converted to n-C17 under all conditions.

4.5.2.2 Effect of Stearic Acid Initial Concentration

Table 4.2.1 shows that the rate of S.A. deoxygenation increased by a factor of 1.46 as

S.A. concentration increased from 4.94 to 20.64%. Figure 4.2.1 shows that the apparent reaction order of S.A. deoxygenation with respect to S.A. is 0.50 order.

With regard to the rate of the liquid species production dependence on S.A. initial concentration, Table 4.2.2 shows that the rate of all liquid species production increased as the initial concentration of S.A. increased and that the rate of n-C17 production was again the fastest over the entire range of S.A. concentration. The ratio of the rate of-C17 production to the rate of n-C18 production was 5.59, 5.12, and 5.95 at 4.94, 10.30, and

20.64 wt % initial S.A. concentration respectively, while the ratio of the rate of n-C17 production to the rate of S.S. production was 4.42, 1.28, and 1.02.

For the first time, at the operating temperature of 260 C, PH2 = 930psig, and 20.64 wt % of S.A. initial concentration, the rate of S.S. production came very close to the rate of n-

C17 production, the ratio of the rate of n-C17 production to the rate of S.S. production at these operating conditions was 1.02.

With respect to the impact of the S.A. initial concentration on the rate of n-C18 production, Table 4.2.2 shows that when the initial concentration of S.A. increased from

4.94 to 20.64 wt %, the rate of n-C18 only increased by 30%.

It was evident that the rate of S.S. production was the most influenced by the initial concentration of S.A. Figure 4.2.3 shows that S.S. reaction order with respect to S.A. was 1.24, while the apparent reaction order for n-C17 and n-C18 with respect to S.A. were

0.46 and 0.27 respectively. It was clear that the estimated reaction order for the n-C17 production with respect to S.A. is very close to the value of the estimated reaction order of 0.50 for S.A. deoxygenation. That is due to the fact that most of the S.A. converted to n-C17 over the entire range of S.A. concentration used in this section.

The rate of S.S. production dependence on the initial S.A. concentration really reflects the rate of alcohol production dependence on the initial S.A. concentration. The higher rate of alcohol production was followed by a faster rate of condensation or esterification.

Although the S.S. hydrogenolysis was a potential source of n-C18 production, by looking at Table 4.2.2 one can see that the increase in the rate of n-C18 production was very small compared to the strong increase in the rate of S.S. production. This could be explained by high S.A. concentration leading to more frequent collisions between the alcohol molecules and S.A molecules yielding a high rate of condensation. The high rate of condensation was followed by hydrogenolysis that took place at a slower rate than the direct alcohol reduction to n-C18. In Figure 4.5.3, the alcohol condensation proceeded at a higher rate than the alcohol reduction.

C17H36

Decarbonylation Pd/C + H2

C17H35-CHO

Hydrogenation/ Pd/C -/+ H Condensation 2 dehydrogenation

C17H35-COOH C17H35-COO-C18H38 C17H35-CH2OH

Heat Hydrogenolysis Pd/C + H H2O 2 Pd/C + H2 Reduction

H2O

C17H35-COOH + C18H38 C18H38

Figure 4.5.1 Octadecanol (C17H35-CH2OH) is involved in three reactions: Decarbonylation that yields the aldehyde (C17H35-CHO), condensation that yields S.S. (C18H37-COO-C17H37), and reduction that yields n-octadecane (C18H38)

In terms of selectivity, Table 4.2.2 shows that n-C17 still had the highest selectivity over the entire range of S.A. initial concentrations. The highest selectivity to n-C17 was at

71% and was obtained at the lowest concentration of S.A. in the feed. The selectivity to n-C18 was also at its highest of 13%, at the lowest initial S.A. concentration. As the initial concentration increased, the selectivity to hydrocarbons declined. The selectivity to n-

C17 and n-C18 decreased by 20% and 23% respectively when the initial S.A. concentration increased from 4.94 to 10.30 wt %. The selectivity to n-C17 and n-C18 decreased by another 7 and 28%, respectively, as the concentration of S.A. increased from 10.30 to

20.64 wt %.

Figure 4.2.2 shows that the selectivity to S.S., however, followed an opposite trend. As the S.A. concentration increased from 4.94 to 10.30 wt %, the selectivity to S.S.

increased by 150%, followed by another 53% as the S.A. concentration increased from

10.30 to 20.64 wt %.

4.5.2.3 Effect of Hydrogen Pressure

Table 4.3.1 shows that the rate of S.A. deoxygenation was slightly influenced by the change in PH2. Figure 4.3.1 reflects the slight effect of PH2 observed on the rate of S.A. deoxygenation. The figure shows that the reaction order of S.A. deoxygenation with respect to H2 was only 0.15.

Immer and Lamb reported that they observed a change in the rate of S.A. decarboxylation with the change in PH2. [11] However, the nature of the decarboxylation reaction indicates that it does not require H2, and in this study, I find no evidence that the rate of decarboxylation is H2 dependent. Therefore, in this study, the S.A. decarboxylation was treated as PH2 independent. This means that the slight S.A. deoxygenation dependence on

PH2 was connected to the hydro-deoxygenation step. The small dependence showed that only a small amount of H2 was required to open the first step in the S.A. hydro- deoxygenation route. Any further increase in PH2 would only have a minor effect on the rate of hydro-deoxygenation.

Table 4.3.2 shows that the change in PH2 had a negligible effect on the rate of n-C17 production. The reaction order of n-C17 with respect to H2 was -0.08, which confirms the small effect of PH2 on the rate of n-C17 production. For this reason, in this work, I assert that at 260 C, the rate of n-C17 production was independent of PH2 over the range of 127

- 931 psig.

In contrast, Table 4.3.2 shows that the rate of S.S. production was highly sensitive to PH2.

This really reflects the rate of alcohol production dependence on PH2. The rate of S.S. production increased 4.5 times as the PH2 increased from 127 to 931 psig. At 931 psig,

S.S. production came very close to the rate of n-C17 of production which means that the rate of aldehyde reduction came very close to or may even exceed the rate of aldehyde decarbonylation (when the ratio of CO/CO2 was low). Figure 4.3.3 shows that at

931 psig, the aldehyde reduction becomes highly competitive to the aldehyde decarbonylation, resulting in a pullback in the rate of n-C17 production. The figure shows that the reaction order for the S.S. production with respect to H2 was 0.68, which is again a reflection of the alcohol production dependence on PH2.

The rate of n-C18 production also increased, but with a weaker dependence on PH2. As the PH2 increased from 127 to 931 psig, the rate of n-C18 production increased by a factor of 2. Figure 4.3.3 shows that the reaction order of n-C18 with respect to H2 was 0.39.

This suggests that at high PH2, the alcohol was produced at high rate. The alcohol then condensed with S.A. forming S.S. at higher rates than reducing to n-C18. This was in agreement with my observations from a separate experiment when octadecanol was tested as the primary feed. When octadecanol was fed, the ratio of the rate n-C17 production to the rate of alcohol hydrogenation ranged from 8-10 which shows that the octadecanol reduction appeared to be an unfavorable pathway. Murzin’s group also reported a similar observation. [17]

With regard to the slower increase in the rate of n-C18 production, it is clear that as the

PH2 increases, the rate of the alcohol production increases. Once the alcohol is formed it 71

has two routes through which it can further react. The first route is to reduce and make n-C18, but this route is constrained by the availability of catalytic sites. The second route is to condense to form S.S., and this route is always available when the S.A. concentration in the feed is high. Therefore, the alcohol condenses at a much higher rate than it is reduced to n-C18, which is shown in the slower rate of n-C18 production.

In addition, the rate of n-C18 production was not affected by the strong increase in the amount of S.S. in the solution, therefore, S.S. hydrogenolysis seems not to be a major contributor to the rate of n-C18 production.

Table 4.3.2 shows that n-C17 had the highest selectivity at the lowest PH2, where the rate of S.S. production was the lowest. As the rate of S.S. production increased with the increase in PH2, the selectivity to n-C17 became smaller and smaller. The selectivity to

S.S. increased from 14 to 46% as the PH2 increased from 127 to 931 psig. However, the selectivity of n-C18 only increased from 5 to 7% for the same change in PH2.

4.5.2.4 Highest Temperature and Lowest Hydrogen Pressure

For the first time, at the highest temperature of 300 °C and lowest PH2 of 20 psig, the hydrocarbons were the only products that could be detected and quantified. No S.S. was detected. The ratio of the rate of n-C17 production to the rate of n-C18 production, shown in Table 4.4.2, was 117. However, the apparent amount of n-C18 was likely due to variations in the background, which means that all the S.A. would have converted to n-C17.

The significant increase in the rate of n-C17 production could be explained by the increase in both of the rates of S.A. decarboxylation and the aldehyde decarbonylation that yielded n-C17, while the aldehyde reduction to alcohol showed very weak competition due to the low PH2 of 20 psig.

CHAPTER 5

INDIVIDUAL REACTION PATHWAY CONTRIBUTION TO THE

RATE OF n-ALKANE PRODUCTION

5.1 Decarboxylation Contribution to the Total Rate of n-C17 Production

5.1.1 Introduction

As mentioned in section 4.1.2 and shown in Figure 5.1.1, the odd-numbered product, n-

C17, observed in the liquid phase was the combined product of S.A. decarboxylation and the decarbonylation of the aldehyde, octadecanal. The only way to quantify the contribution of each reaction pathway to the observed rate of n-C17 production is to determine the rate of CO2 and CO production, the co-products of the decarboxylation and decarbonylation, respectively. The rate of n-C17 production from decarboxylation is equal to the rate of CO2 production and the rate of CO production is equal to the rate of the aldehyde decarbonylation. This section discusses the study conducted to estimate the individual reaction contribution to the total rate of n-C17 production.

Hydrogenolysis C18H38 Decarboxylation

C17H35-COOH C17H36 Decarbonylation Pd/C

Pd/C + H2

C17H35-CHO Hydrodeoxygenation

Pd/C -/+ H2

C17H35-COOH Condensation C17H35-COO-C18H37 C17H35-CH2OH Reduction

Heat H2O Pd/C + H2

C18H38 Reduction

Figure 5.1.1 The observed odd-numbered product n-C17 was the combined product of the decarboxylation of S.A. and the decarbonylation of octadecanal. CO2 and CO are the co-products of the decarboxylation and decarbonylation respectively

5.1.2 Experiment

In order to determine the contribution of S.A. decarboxylation and the aldehyde decarbonylation to the total rate of n-C17 production, I designed two sets of experiments from which I could either directly estimate the rates of CO and CO2 production or estimate the ratio of CO/CO2.

The experiments were carried out at 260 and 280 °C and 100 psig total system pressure.

Measurements were taken by using 2.50 g of 3 wt % Pd/C and 2.5 wt % S.A. in the C24 solvent. A high WHSV of 20/hr was used in order to increase the concentration of CO

and CO2 in the gas phase while keeping the overall conversion under 10%. The gas flow rate was set to 25 mL/min.

The first set of experiments was conducted under pure H2 atmosphere, while the second set was in an H2/Ar mixture with low PH2. Although no H2 was flowing into the system in the second set, the minor amount of H2 generated from the dehydrogenation of the C24 solvent on the catalyst surface is the cause of the low PH2. The H2 produced from the solvent dehydrogenation was approximately 2 mole percent of the gas phase composition.

Given that the total system pressure was 100 psig, the PH2 was approximately 2 psig.

A new charge of catalyst was used for each data point. The experiments were typically conducted over a 9 hour period. The gas phase data were monitored by FTIR2. In order to account for all the CO and CO2 produced during the run and after the liquid ran out, a spectrum was collected every half hour for 14 hours.

At the end of the run, the CO and CO2 concentrations were converted to molar fluxes and plotted as a function of time. The area under the curve of each species was numerically integrated yielding the total number of moles of that species produced during the run.

The sum of the total moles of CO and CO2 produced was compared to the total number of moles of n-C17 produced during the run to confirm the accuracy of the gas phase results.

The ratio of n-C17/ (CO+CO2) was within +/- 15%.

2 A complete description of the FTIR and the analytical approach are discussed in Appendix A

5.1.3 Results

Under all conditions, the total amount of CO produced was higher than CO2. However, determining the number of moles of each gaseous product with the level of accuracy required was difficult, particularly, due to the fact that I was required to operate within a narrow temperature range, in order to limit the overall conversion. The challenge was quantifying the low concentrations of CO2. Any fluctuation in the parameters influenced the gas phase product distribution. Many observations had led to repeating experiments for more accuracy. One of the observations was that the time allowed between the end of the catalyst activation and the start of liquid flow had to be consistent in every run. The results varied widely when this rule was not followed.

Due to the sensitivity of the measurements, some data points were repeated multiple times to ensure the accuracy of the results. The overall CO/CO2 ratio was used to understand how the rates were influenced by the change in operating conditions.

Generally, the data showed that the ratio of CO to CO2 was sensitive to both the temperature and PH2. The ratio of CO/CO2 decreased as the temperature increased. At

280 °C and 100 psig of pure H2, the results of 3 runs showed that the ratio of CO/CO2 was 2.9, 3.3, and 3.5.

When the experiment was conducted at the same temperature but at the low PH2 of 2 psig, the ratio of CO/CO2 was 1.5.

At 260 °C in pure H2, the ratio was 7, 7.3, and 9.5. Finally, at 260 °C and PH2 of 2 psig the ratio of CO/CO2 was about 6.5.

5.1.4 Discussion

Under the higher PH2 of 100 psig, the contribution of the decarboxylation to the overall rate of n-C17 production was approximately 12% at 260 °C and increased to 25% as the temperature increased to 280 °C. While under the lower PH2 of 2 psig, the contribution of the decarboxylation to the overall rate of n-C17 production was approximately 17% at

260 °C, and increased to 40% as the temperature increased to 280 °C.

Therefore, it could be concluded that S.A. decarboxylation was more sensitive to temperature than the S.A. decarbonylation. At constant PH2, both the S.A. decarboxylation and the aldehyde decarbonylation rates increased as the temperature increased. However, the relative increase in the rate of S.A. decarboxylation as the temperature increased was higher than the relative increase in the rate of aldehyde decarbonylation over the same change in temperature, forcing the ratio of CO to CO2 to become smaller.

In contrast, the increase in PH2 yields an increase in the ratio of CO/CO2. As mentioned earlier, S.A. decarboxylation is assumed to be independent of the PH2, while the rate of hydro-deoxygenation is slightly dependent on the PH2. The rate of hydro-deoxygenation slightly increases as the PH2 increases and ultimately yields a minor increase in the rate of aldehyde production. Since, the aldehyde prefers decarbonylation, it decarbonylates to yield n-C17. Since the rate of decarboxylation is unchanged, the ratio of the rate of CO production to the rate of CO2 production becomes larger.

However, earlier observations with regard to the effect of P H2 on the rate of n-C17 production suggested that the overall rate of n-C17 production was independent of PH2. I believe that as the rate of aldehyde decarbonylation increases in the presence of H2, the rate of S.A decarboxylation decreases by nearly the same amount.

5.2 Stearyl Stearate Hydrogenolysis Contribution to the Total Rate of n-

C18 Production

5.2.1 Introduction

Similar to n-C17, n-C18 was also a combined product of two reactions. Figure 5.2.1 shows that in addition to the alcohol reduction, n-C18 could be produced by the hydrogenolysis

S.S. Since S.S. appeared as a final product in the deoxygenation of S.A. in relatively large amounts, the S.S. hydrogenolysis could be a potential source of n-C18 and could play an important role in the overall product distribution. For this reason, I decided to estimate the S.S. contribution to the overall rate of n-C18 production during S.A. deoxygenation.

Furthermore, since the condensation prevents the alcohol reduction, I believe that the condensation could potentially limit the overall rate of n-C18 production. For these reasons, I decided to investigate the effect of operating conditions on the rate of S.S. hydrogenolysis.

Hydrogenolysis C18H38 Decarboxylation

C17H35-COOH C17H36 Decarbonylation Pd/C

Pd/C + H2

C17H35-CHO Hydrodeoxygenation

Pd/C -/+ H2

C17H35-COOH Condensation C17H35-COO-C18H37 C17H35-CH2OH Reduction

Heat H2O Pd/C + H2

C18H38 Reduction

Figure 5.2.1 The observed even-numbered product n-C18 is the combined product of the alcohol reduction and the S.S. hydrogenolysis

Another critical parameter that needed to be taken into consideration was the difference in the ester residence time on the catalyst bed. In order to compare the residence time of the S.S. formed during the deoxygenation of S.A. and the residence time of the ester fed directly as primary feed the following approach was used: during the S.A. deoxygenation, the alcohol could potentially form anywhere between the top to the bottom of the catalyst bed. Once the alcohol is formed, given the high concentration of S.A. and the temperature of the reactor, the alcohol can instantly condense forming S.S.

Therefore, in order to estimate an average residence time for the S.S. produced in the catalyst bed, I assumed that on average the S.S. resides ½ of the residence time of the

S.A. Therefore, when the S.S. is the primary feed, it could be assumed that the residence time of the S.S. on the catalyst bed is approximately twice the residence time of S.S. produced during the deoxygenation of S.A.

It was necessary to study the kinetics of S.S. hydrogenolysis as primary feed on the 3 wt

% Pd/C. But in order to be able distinguish between the hydrogenolysis and the alcohol reduction products, palmityl stearate (P.S.) was used instead of S.S. In which case, the n-C16 was the hydrogenolysis product, while n-C18 was still the alcohol reduction product.

This study was conducted in the same reactor system described in section 3.4.3 by using

1 gram of 3 wt% Pd/C. The effect of temperature, P.S. initial concentration, and hydrogen pressure on the rate of n-C16 production is presented and a kinetic framework is developed based on the rate of hydrogenolysis, which is equal to the rate of n-C16 production.

5.2.2 Feed Materials

A four liter batch of P.S. was prepared by blending equimolar amounts of S.A. and hexadecanol in a paraffinic mixture of C26/C28 solvent to yield a solution of 20wt % P.S. upon completion. The reason for using this alternative carrier solvent is that the 1- hexadecanol elutes from the GC column within the distribution of C24 isomers which makes its precise quantification difficult.

The composition of the paraffinic mixture n-C26/n-C28was as follows: 45.6 wt % n-C26;

30.0 wt % n-C28; 2.0 wt % n-C30; 5.6 wt % n-C24 and the balance was a mixture of C26 and C28 isomers.

The S.A. and hexadecanol in C26/C28 solvent was stirred while heated to 65°C. About 6 mL of concentrated H2SO4 were added slowly to the solution. The solution was left for three days to achieve the highest conversion possible. In order to monitor the reaction

progress, samples were taken twice a day during the three day period. The reaction was considered complete when the conversion to P.S. exceeded 97%. At this point, approximately 150 mL of distilled H2O were added to the solution to separate the H2SO4 and excess reactants from the P.S. solution. The P.S. solution was sent for analysis to confirm the composition. The solution had 21.40 wt % P.S. but since the target P.S. solution concentration was 20 wt %, some of the 21.40 wt% solution was slightly diluted by adding n-C26/n-C28 to yield 19.65 wt % P.S. solution.

The more diluted P.S. solutions were obtained by sequential dilution of the 21.40 wt % solution by adding more of the C26/C28 solvent. Samples of the diluted solutions were also sent for analysis to verify their compositions. The P.S. in C26/C28 solution was solid at room temperature. For this reason, the solution was kept stirred and heated at 70 °C just until before it was ready to be fed, so that it stayed liquid.

5.2.3 Effect of Temperature on P.S. Hydrogenolysis

5.2.3.1 Experiment

In order to understand the effect of temperature on the P.S. hydrogenolysis, this part of the study was conducted under pure H2, which entered the system at 26 mL/min. A temperature range of 225 - 250 °C was selected for the study. The experiment was conducted by using 1.00 g of the 3 wt % Pd/C. The system pressure was set to 110 (+/-

3) psig. The 19.65 wt % P.S. was fed to the reactor at WHSV = 1/hr. At these conditions, the P.S. overall conversion ranged from 2.80 to 10.76%.

In order to examine the temperature impact on the rate of P.S. hydrogenolysis, and the rate of product formation, the temperature was varied by 5 °C increments while the system pressure, liquid flow, and H2 flow remained unchanged. This part of the study section (5.2.3) was conducted on one charge of catalyst.

Typically, before the end of the day, an operating temperature would be established and the system was allowed to reach steady state overnight. The following morning, the transient liquid processed overnight was collected and discarded. The experiment ran for another 8 hours at steady conditions, after which, the sample was collected and sent for analysis. The cycle started again by setting the next desired operating temperature.

5.2.3.2 Results

Table 5.2.1 shows the effect of temperature on the rate of P.S. hydrogenolysis.

Table 5.2.1 The effect of temperature on the rate of P.S. hydrogenolysis

T Rate of P.S. [˚C] Hydrogenolysis -1 -1 [mol∙gcat ∙min ] 225 1.98E-07 230 2.64E-07 235 3.30E-07 240 4.09E-07 250 6.23E-07

1.00 g of the 3 wt % Pd/C was used. The 19.65 wt % palmityl stearate was fed to the reactor at WHSV = 1/hr. The system pressure was set to 110 (+/- 3) psig. The study was conducted under pure H2 that entered the reactor at 26 mL/min

Figure 5.2.2 shows the effect of temperature on the rate of P.S. hydrogenolysis in

Arrhenius form. The linear fit to the data appears excellent over the entire range. The activation energy derived from the slope of the line was 98kJ/mol.

-14.2

-14.4

]) 1

- -14.6

.min

1 -

cat -14.8

∙g [mol

-15.0 y = -11783x + 8.2547 R² = 0.997

P.S. disapp. P.S. -15.2 ln(r

-15.4

-15.6 0.0019 0.00192 0.00194 0.00196 0.00198 0.002 0.00202 1/T (1/K)

Figure 5.2.2 Effect of temperature on the rate P.S hydrogenolysis in Arrhenius form

Measurements were taken over temperature range of 225 - 250 ˚C, by using 1.00g of 3 wt % Pd/C. The 19.65 wt % P.S.in C26/C28 solution was fed to the reactor at WHSV = 1/hr. The system pressure was set to 110 psig. The study was conducted under pure H2 that entered the reactor at 26 mL/min.

5.2.4 Effect of Initial Palmityl Stearate Concentration

5.2.4.1 Experiment

In order to understand the effect of the P.S. initial concentration on the rate of P.S. hydrogenolysis, the rate of P.S. hydrogenolysis was monitored as the initial concentration of P.S. varied from 2.15 - 21.40 wt % in the C26/C28 paraffinic mixture. The measurements were taken at 235 °C, by using 1.00 g of 3 wt % Pd/C. The P.S. solution was fed to the reactor at a rate corresponding to WHSV = 1/hr. The system pressure was set to 108 +/- 3 psig. The study was conducted under pure H2, which entered the reactor at 26 mL/min. The data presented in this section (5.2.4) were all collected on one charge of 3 wt% Pd/C catalyst.

In order to examine the effect of the of P.S. initial concentration on the rate of P.S. hydrogenolysis and the rate of product formation, the sequence of P.S. concentrations was fed in a descending order. On the first day, 21.40 wt% solution was fed and the system was allowed to reach steady state overnight. The following morning, the transient liquid processed overnight was collected and discarded. The system ran for 8 hours, after which, the sample was collected for analysis.

In preparation for the next concentration, the liquid left in the syringe pump was emptied into the feed reservoir, from which it was pipetted out. About 50 mL of the 10 wt % solution was used to rinse the syringe pump and feed reservoir by repeatedly filling and emptying the pump into the feed reservoir. The liquid was collected from the feed

reservoir and discarded. The process was repeated 3 times. The same process was repeated before introducing each new P.S. concentration.

5.2.4.2 Results

Table 5.2.2 shows the rate of P.S. hydrogenolysis as P.S. initial concentration changed.

Measurements were taken at 235 °C, 1.00 g of 3 wt % Pd/C. In order to determine the effect of P.S. initial concentration on the rate of P.S. hydrogenolysis, ln (rp.s.) was plotted versus ln [P.S.]as shown in Figure 5.2.3. The slope of the line which is the reaction order with respect to P.S. was found to be 0.7.

Table 5.2.2 Effect of P.S. initial concentration on the rate of P.S. hydrogenolysis

Wt % of P.S. in the Rate of P.S. hydrogenolysis -1 -1 feed [mol∙ gcat . min ]

2.15 7.84E-08 5.21 1.18E-07 10.07 2.29E-07 19.65 3.32E-07 21.40 3.79E-07

The P.S. concentration varied from 2.15 - 21.40 wt % in the C26/C28 solvent. Measurements were taken at 235 ˚C. The P.S. solution was fed to the reactor at a rate that corresponds to WHSV = 1/hr. The system pressure was set to 108 +/- 3 psig. The study was conducted in pure H2 that entered the reactor at 26 mL /min.

-14.6

-14.8

]) 1 - -15.0

.min -15.2 1

- y = 0.7 x - 16.95

cat -15.4 R² = 0.9803

-15.6 [mol∙g -15.8

-16.0 P.S. disapp. P.S.

-16.2 ln(r -16.4

-16.6 0.5 1.0 1.5 2.0 2.5 3.0 3.5 ln[P.S.[wt %]]

Figure 5.2.3 Effect of P.S. initial concentration on the rate of P.S. hydrogenolysis

5.2.5 Effect of Hydrogen Pressure

5.2.5.1 Experiment

In order to determine the effect of PH2 on the rate of P.S. hydrogenolysis, the rate of P.S. disappearance was examined over 59 to 920 psig range of PH2. The measurements were taken at 235 °C by using 1.00 g of 3 wt % Pd/C. In this section (5.2.5.1), a 20.60 wt%

P.S solution was used. The 20.60 wt % P.S. solution was fed to the reactor at a rate that corresponded to WHSV = 1/hr. The overall P.S. conversion ranged from 5.00 - 7.50% over the range of PH2. The target PH2 was achieved by controlling the Ar/H2 mole ratio, while keeping the system pressure constant at 920 psig. The H2 mole fraction in the gas phase was confirmed by GC. The rate of hydrogenolysis was monitored at 5 different

PH2. The sequence started with the highest PH2 of 920 psig and went in a descending order. The data presented in this section (5.2.5.1) were all collected on one charge of catalyst.

5.2.5.2 Results

Table 5.2.3 shows the rate of P.S hydrogenolysis as the PH2 ranged from 59 - 917 psig.

Temperature was set to 235 °C; 1.00g of 3 wt % Pd/C and 20.60 wt% of P.S solution were used. Figure 5.2.4 shows the P.S. hydrogenolysis rate dependence on PH2 which was determined by plotting ln(rn-C16) versus ln (PH2). The reaction order for the P.S. hydrogenolysis with respect to H2 was found to be 0.25.

Table 5.2.3 Effect of hydrogen pressure on the rate of P.S. hydrogenolysis

PH2 Rate of n-C16 production -1 -1 [psig] [mol∙ gcat ∙min ] 917 6.2E-07 454 4.9E-07 216 4.2E-07 117 3.6E-07 59 3.1E-07

Measurements were taken by using 1.00 g of 3 wt % Pd/C and at 235 ˚C. The 20.60 wt % P.S. solution was fed to the reactor at a rate that corresponded to WHSV = 1/hr. The PH2 was varied by changing the H2/Ar molar ratio, while the total system pressure was set to 920 +/- 1% psig. The PH2 varied from 59 - 917 psig.

-14.2

-14.3

-14.4 y = 0.25x - 16.019

R² = 0.995

])

1 -

-14.5

.min

1 -

cat -14.6

]) [mol∙g -14.7

P.S. disapp. P.S. -14.8 ln(r -14.9

-15.0

-15.1 4.00 4.50 5.00 5.50 6.00 6.50 7.00

ln(PH2 [psig])

Figure 5.2.4 Effect of hydrogen pressure on the rate of P.S. hydrogenolysis

5.2.6 Discussion

The hydrogenolysis rate dependence on temperature is shown in Table 5.2.1. The rate of hydrogenolysis tripled as the temperature increased from 225 to 250 °C. The actual activation energy of P.S. hydrogenolysis was estimated at 98 kJ/mol.

The value of the reaction order, 0.7, for the P.S. hydrogenolysis with respect to P.S. shows that the P.S. concentration has a considerable impact on the rate of P.S. hydrogenolysis. On the other hand, Table 5.2.3 shows that the PH2 has a lower impact on the rate of P.S. hydrogenolysis. The rate of hydrogenolysis increased by only a factor of two when PH2 increased by a factor of 15. That was reflected in the low value of the apparent hydrogen reaction order of 0.25. Combining the above parameters yields:

, Equation 5.2-1 where, is the rate of P.S. disappearance, which is also equal to the rate of P.S. hydrogenolysis.

Comparing the apparent hydrogen reaction order of 0.25 obtained for the P.S. hydrogenolysis to the apparent reaction order of n-C18 production with respect to H2 which was estimated at 0.68 in section 4.3.2, shows that the high dependence of the rate of n-C18 estimated in section 4.3.2 is linked to the rate of n-C18 produced from alcohol reduction, not the n-C18 produced from the hydrogenolysis step.

In order to estimate the hydrogenolysis contribution to the rate of n-C18 production during

-7 -1 -1 the deoxygenation of S.A., the rate of P.S. hydrogenolysis of 6.23∙10 mol∙min ∙gcat

observed at 250 C, 110 psig, and 19.6 wt. % P.S. was selected; then, adjusted to the conditions at which the highest rate of S.S. production was observed during the S.A. deoxygenation. The highest rate of S.S. production observed during the entire S.A.

-7 -1 -1 deoxygenation was 7.53∙10 mol∙min ∙gcat (from Table 4.4.2), which is equivalent to approximately 2.35 wt % of the total mass of liquid leaving the reactor. The 2.35 wt % of S.S. concentration was observed at 260 °C and 930 psig.

-7 -1 -1 Therefore, the rate of P.S hydrogenolysis of 6.23∙10 mol∙min ∙gcat was adjusted to 260

°C, 931psig, and 2.35 wt % of ester concentration. The adjustment was performed by using the P.S. hydrogenolysis apparent activation energy of 98 kJ/mol calculated in section 5.2.3.2, the reaction order with respect to the ester concentration estimated at 0.7 in section 5.2.4.2, and the PH2 reaction order with respect to H2 estimated at 0.25 in

-7 -1 -1 section 5.2.5.3. The adjusted rate obtained was 1.24∙10 mol∙min ∙gcat

The residence time of P.S. in the catalyst bed is twice the residence time of S.S. formed

-7 -1 -1 during the S.A. deoxygenation. Therefore, dividing 1.24∙10 mol∙min ∙gcat by 2 would

-8 -1 -1 give 6.20∙10 mol∙min ∙gcat , which was the estimated rate of n-C18 production via hydrogenolysis during S.A. deoxygenation at 260 °C, 930 psig, and 2.35 wt % S.S. concentration.

-8 -1 -1 -7 Comparing the rate of 6.20∙10 mol∙min ∙gcat to the rate of n-C18 production of 1.22∙10

-1 -1 mol∙min ∙gcat presented in Table 4.3.2 shows that the S.S. hydrogenolysis contributed approximately 50 percent of the total rate of n-C18 produced at these conditions.

In the above estimation, the S.S. concentration of 2.35 wt % that corresponded to the highest rate of S.S. production during the entire S.A. study was used. For this reason, and due to the strong dependence of the rate of S.S. hydrogenolysis on the S.S. concentration, the estimated S.S. hydrogenolysis contribution of 50% is the highest contribution that the

S.S. hydrogenolysis made to the overall rate of n-C18 during this study. However, in general, as the concentration of S.S. increases, the contribution of S.S hydrogenolysis to the overall rate of n-C18 production becomes larger.

With regard to the effect of S.A. initial concentration on S.S. hydrogenolysis contribution to the overall rate of n-C18 production, since the rate of S.S. production is highly dependent on S.A., and the rate of hydrogenolysis is highly dependent on the concentration of S.S., it can be concluded that the contribution of S.S. generally increases with increase in S.A. initial concentration.

In connection to the effect of PH2 on S.S. hydrogenolysis contribution to the overall rate of n-C18 production, since the rate of S.S. production is highly dependent on PH2, and the rate of S.S. hydrogenolysis is highly dependent on the concentration of S.S., it can be concluded that the contribution of S.S. generally increases with the increase in PH2.

CHAPTER 6

CONCLUSION AND FUTURE STUDIES RECOMMENDATIONS

6.1 Conclusion

The deoxygenation of S.A. over Pd/C catalyst, in the presence of hydrogen yielded three liquid products: n-heptadecane, n-octadecane, and S.S.,the condensation product of octadecanol and unreacted S.A. A comprehensive study was conducted to understand the effect of temperature, initial stearic acid concentration, and hydrogen pressure on the rate of stearic acid deoxygenation and the rate of liquid product formation. Rate expressions were derived for the S.A. deoxygenation and the liquid products formed. With regard to the hydrocarbons, the derived rate expressions were based on the observed total rate of production. These rate expressions will help estimating the rate of liquid species and product selectivity under any set of operating conditions. The rate of S.A. deoxygenation can be estimated from

Equation 6.1-1

The pre-exponential factor (A0) has units.

While, the rates of n-alkane production can be estimated from the following equations

, Equation 6.1-2

And, Equation 6.1-3

Since, n-C17, n-C18, and S.S. were the only 3 products of S.A. deoxygenation, the rate of

S.A. disappearance can be written as:

Equation 6.1-4

So, the rate of S.S. production can be written as:

Equation 6.1-5

With regard to the rate of S.S. hydrogenolysis during the deoxygenation of S.A., the rate can be estimated from

Equation 6.1-6

Although the number of quantifiable products was only three, no less than seven individual reactions were taking place within the reactor. The observed rate for each n- alkane production was found to be the combined rate of two different reactions. The rate of n-C17 production was the combined product of S.A. decarboxylation and octadecanal decarbonylation. In order to estimate the contribution of each reaction to the total rate of

n-C17 production, I took a closer look at the gaseous co-products CO and CO2. The rate of CO production, the octadecanal decarbonylation product, was compared to the rate of

CO2 production. Under all conditions, the rate of CO production was higher than the rate of CO2 production, which means that under all conditions the octadecanal decarbonylation was the primary contributor to the rate of n-C17 production. Examining the influence of temperature and PH2 on the ratio of CO to CO2 showed that as the temperature increased at constant PH2, the ratio of CO to CO2 becomes smaller, which means that decarboxylation increases at a faster pace than the octadecanal decarbonylation. On the other hand, the rate of S.A. decarbonylation increased with increase in PH2, while the rate of S.A. decarboxylation showed a small decrease to compensate for the increase in the rate of decarbonylation, because the sum of the rates of decarbonylation and decarboxylation are equal to the rate of n-C17 production which is independent of PH2. The lowest CO/CO2 ratio was observed at the highest temperature and lowest pressure, where decarboxylation contributed about 40% of the overall rate of n-C17 production and was the highest level observed during the study.

Similar to n-C17, n-C18 was also a combined product of two reactions, alcohol reduction and S.S. hydrogenolysis. In order to understand whether the S.S. was the major source of n-C18, the rate of n-C16 produced by P.S. hydrogenolysis was compared to the rates of n-

C18 produced during the deoxygenation of S.A. At 260 °C, 930 psig, and 2.35 wt % S.S. concentration, the S.S. hydrogenolysis contributed approximately 50% of the total rate of n-C18 production.

6.2 Future Work Recommendation

First, as it was earlier discussed, I believe that H2 is an essential element to the process feasability, as it increases the overall conversion of the fatty acid, influences the product selectivity and extends the catalyst lifetime. Generally, I recommend future studies to include H2. I also believe that operating at steady state in continuous mode is most useful to thoroughly understand the process in a production-like operation. Operating in continuous mode would minimize issues that may arise during scale-up.

Further studies of the kinetics of the deoxygenation of fatty acids in the presence of H2 in continuous mode operation are recommended. However, in order to study the effect of temperature and PH2, the concentration of the fatty acid must be constant across the catalyst bed. Keeping a constant concentration across the catalyst bed, in a fixed bed reactor, can only be achieved by keeping the S.A. conversion as low as possible.

In order to be able to study the process at lower ranges of PH2, using a H2 flow controller that can control the slower H2 flow rate is recommended.

In future work similar to ours, I recommend using a lighter hydrocarbon solvent with available vapor-liquid thermodynamic data. The reason is that the light solvent would have a lower dehydrogenation equilibrium constant value, which consequently would yield a smaller amount of H2 liberated from solvent dehydrogenation relative to a heavy hydrocarbon. At the end, this would reduce the PH2 of the system, particularly, when measurements at low PH2 are needed.

With regard to the catalyst, although Murzin’s group studied several metal/support combinations to determine the catalyst that yielded the highest S.A. conversion, their investigation was conducted in the absence of H2. Therefore, studying other metal/support combinations in the presence of H2 is also recommended. Pt/C and Rh/C would be a good place to start.

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104

APPENDIX

VARIAN 660 FTIR SPECTROMETER GAS PHASE

ANALYSIS

Approach

The FTIR unit was set up approximately 5 feet away from the reactor exit. The gas exiting the reactor went through a back pressure regulator, then directed to the FTIR by a switching valve. The FTIR cylindrical sampling cell was 10 cm long and had a 3 cm ID.

These dimensions yield a volume of approximately 70 ml. The sampling cell was positioned in the FTIR analysis chamber so that the beam passes through the cell without interference. The MCT detector is cooled with liquid N2. The FTIR inside compartment and the sampling cell were purged with two separate house N2 gas lines. Purging with N2 would start 2 hours before the start of the experiment. Just before the start of the experiment, the N2 line that purges the sampling cell is stopped. The reactor effluent line is set to flow through the cell. Purging the FTIR inside compartment with N2 continued throughout the entire experiment. The FTIR detector was cooled with liquid N2 every 4-5 hours.

105

Calibration

1. The calibration gas composition was: 20 mol % CO, 5% mol CO2, 7 mol % CH4,

62 mol % H2 and the balance of 6 mol % consisted of a mixture of light

hydrocarbons (C2-C6).

2. First, the GC was calibrated with the calibration gas described in 1.

3. Once the GC was calibrated, the calibration gas was run through the FTIR.

4. The CO peak appeared between 2240- 2030 wavenumbers [cm-1] (from left to

-1 right), while the CO2 peak area appeared between 2388-2282 [cm ]. Figure 1

presents an FTIR spectrum that includes CO and CO2 peaks.

Sample (21) CO2 CO Peak Areas 1.13199 1.81013

CO Peak

CO2 Peak

Figure 1 FTIR spectrum that includes CO and CO2 peaks.

106

5. In order to obtain different concentrations of CO and CO2, the calibration gas was

diluted in N2 to yield different CO and CO2 concentrations.

6. In order to dilute the calibration gas, the N2 Mass Flow Controller (MFC) was

used to deliver different flow rates of N2, which was mixed with the calibration

gas to yield the different concentrations of CO and CO2 tabulated in Tables 1 and

7. After each dilution the exact gas molar composition was run through the GC to

find the exact concentration.

8. Next the diluted gas was run through the FTIR and the areas of CO and CO2

peaks were obtained.

9. The different concentrations of CO and CO2 [PPM] were tabulated versus the

peak area obtained from the FTIR for the species as shown in Table 1 and 2.

10. In order to determine the correlation between the area of the peak to the actual

concentration of each of the CO and CO2, the area of each of the CO and CO2

peaks given by the FTIR was plotted as X values versus the corresponding

concentration of CO and CO2 obtained from the GC as Y values as shown in

Figure 1 and 2.

11. The line intercept in Figure 2 and 3 were forced to go through zero. The slope of

the line derived from each plot was used to convert peak area obtained from the

FTIR to concentration in PPM.

12. From Figure 2, in order to find the concentration of CO in PPM, the area of the

CO peak was multiplied by 711.91. Similarly, from Figure 3, in order to find the

107

concentration of CO2 in PPM, the area of the CO2 peak was multiplied by 136.49.

Table 3 is an example of how the area of each peak converted to concentration in

PPM.

Table 1 CO peak area obtained from FTIR versus the concentration in PPM given by the GC

CO peak Area CO PPM based on GC 0.39 272.87 0.42 273.96 0.52 425.65 0.56 425.65 0.61 477.34 0.65 483.26 0.95 684.16 0.97 684.16 1.12 861.64 1.13 860.52 1.69 1178.25 1.76 1178.25

108

1400 y = 711.91x 1200 R² = 0.9868

1000

800

600

400

CO Concentration in PPM (GC) 200

0 0 0.5 1 1.5 2 CO Integrated Area (FTIR)

Figure 2 CO integrated area obtained from FTIR plotted versus the concentration in PPM given by the GC the slope of the line was 711.91and the intercept was forced to zero.

109

Table 2 CO2 peak area obtained from FTIR versus the concentration in PPM given by the GC CO2 Peak Area CO Concentration in PPM 0.79 68.35 0.80 68.35 1.07 105.96 1.12 105.96 1.21 120.07 1.25 120.07 1.79 181.68 1.83 181.68 2.17 215.27 2.18 215.27 2.90 328.45 3.02 328.45 5.17 611.33 5.18 611.33 8.23 1082.97 8.23 1082.97 11.25 1670.13 11.27 1670.13

110

1800

1600

1400

1200

1000 y = 136.49x 800

600

400

CocentrationPPM in (GC)

200 CO 0 0 2 4 6 8 10 12

CO2 Integrated Area (FTIR)

Figure 3 CO2 integrated area obtained from FTIR plotted versus the concentration in PPM given by the GC the slope of the line was 136.49 and the intercept was forced to zero.

Table 3 An example of areas of CO and CO2 peaks converted to concentration in PPM

CO CO 2 CO CO Integrated Integrated 2 [PPM] [PPM] Area Area

Sample (21) 1.13 1.81 154.50 1288.63

111