KINETIC STUDY OF CATALYTIC PARTIAL OXIDATION OF …

KINETIC STUDY OF CATALYTIC PARTIAL OXIDATION OF SYNTHETIC

DIESEL FOR HYDROGEN PRODUCTION

A Thesis

Submitted to the Faculty of Graduate Studies and Research

In Partial Fulfillment of the Requirements

For the Degree of

Master of Applied Science

In

Process Systems Engineering

University of Regina

By

Md. Faysal Ahamed Khan

Regina, Saskatchewan

January 2012

Copyright 2012: M. F. A. Khan

UNIVERSITY OF REGINA

FACULTY OF GRADUATE STUDIES AND RESEARCH

SUPERVISORY AND EXAMINING COMMITTEE

Md. Faysal Ahamed Khan, candidate for the degree of Master of Applied Science in Process Systems Engineering, has presented a thesis titled, Kinetic Study of Catalytic Partial Oxidation of Synthetic Diesel for Hydrogen Production, in an oral examination held on December 20, 2011. The following committee members have found the thesis acceptable in form and content, and that the candidate demonstrated satisfactory knowledge of the subject material. External Examiner: Dr. Fanhua Zeng, Petroleum Systems Engineering

Co-Supervisor: Dr. Raphael Idem, Process Systems Engineering

Co-Supervisor: Dr. Hussameldin Ibrahim, Process Systems Engineering

Committee Member: *Dr. Farshid Torabi, Petroleum Systems Engineering

Committee Member: Dr. Yongan Gu, Petroleum Systems Engineering

Chair of Defense: Dr. Nader Mobed, Department of Physics *Not present at defense

i

ABSTRACT

The focus of this research is to study the kinetics of the catalytic partial oxidation

(CPOX) of synthetic diesel (SD) for hydrogen production. The kinetic experiments were

done in a packed bed tubular reactor (PBTR) over a 5wt.%Ni/Ce0.5Zr0.33Ca0.085Y0.085

(5N/CZCaY) catalyst prepared by a surfactant-assisted route. The SD is composed of 75

vol.% saturated hydrocarbons and 25 vol.% aromatic hydrocarbons, with an average

chemical formula resembling commercial diesel C12.87H24.81. The kinetic experiments

were conducted at atmospheric pressure, in the temperature range of 1123-1223K (850-

950C), with oxygen/synthetic diesel (O2/SD) ratio in the range of 6.7-10.5 and W/FSD,0

(weight-time) in the range of 19008-47556 kgcatalyst*s/kmolSD. The experimental results

were used to derive an empirical power law rate model. This model was of the form:

n

O

m

SDRT

E

0SD 2NNekr

. Activation energy was found to be 16kJ/mol and the order of

reaction with respect to SD was 1.89 (≈2) and with respect to oxygen was found to be

0.41 (≈1/2). Estimation of the values of the model parameters was based on the

minimization of the sum of the residual squares of the reaction rates by Gauss-Newton

and Levenberg-Marquardt algorithm using non-linear regression (NLREG) software.

Excellent agreement between the experimental and predicted rate was established with an

absolute average deviation (AAD) of 8%. The 5N/CZCaY catalyst was tested for an

extended time on stream (TOS) operation in order to establish and demonstrate that the

catalyst is stable and also to ensure steady state performance. In addition, the effects of

reaction parameters such as reaction temperature, feed ratio (O2/SD), and weight-time

ii

(W/FSD,0) on the resultant catalytic activity of the chosen catalyst were also investigated

in order to obtain the optimal operating conditions for H2 production from CPOX of SD.

To the best of our knowledge, the current study is the first of its kind on the CPOX

reforming of SD.

iii

ACKNOWLEDGEMENTS

First, I would like to give thanks to the Almighty for giving me all the patience as

well as strength required to complete my degree.

I wish to thank Dr. Raphael Idem and Dr. Hussameldin Ibrahim for giving me the

opportunity to do a Masters of Applied Science at the University of Regina under their

supervision. They have been very helpful throughout the course of my research work,

providing expert advice, encouragement, and support. I would like to make special

reference to Dr. Ataullah Khan; the many hours spent in the lab trouble shooting the

experimental set up with him helped me a lot to understand reaction engineering, and his

knowledge of catalysis encouraged me to continue further study in this field. Without his

cooperation, it would not have been possible for me to complete this work on time.

I would like to thank all the members of the H2 production research group of the

University of Regina for their encouragement, support, and valuable suggestions

throughout my work and for all the brainstorming technical meetings with them every

two weeks.

Finally, I would like to thank the Faculty of Graduate Studies and Research of the

University of Regina for the financial support I have received thought out my research

work, and I acknowledge the International Test Centre for CO2 Capture (ITC) for giving

me the opportunity to do all my experimental work in their labs.

iv

DEDICATION

This work is dedicated to my mother, Sadia Nurin, and my father, Md. Abdus

Samad Khan (R.I.P).

v

TABLE OF CONTENTS

ABSTRACT i

ACKNOWLEDGMENTS iii

DEDICATION iv

TABLE OF CONTENTS v

LIST OF TABLES ix

LIST OF FIGURES x

LIST OF APPENDICES xi

NOMENCLATURE xiii

CHAPTER 1: INTRODUCTION 1

1.1 HYDROGEN ECONOMY 1

1.2 DIESEL REFORMING TECHNOLOGY 6

1.3 KNOWLEDGE GAP AND PROBLEM IDENTIFICATION 9

1.4 RESEARCH OBJECTIVES 11

vi

CHAPTER 2: LITERATURE REVIEW 12

2.1 PRECIS 12

2.2 HYDROGEN FROM FOSSIL FUEL REFORMING 13

2.2.1 Technologies 13

2.2.2 Diesel as Source of Fossil Fuel 14

2.3 REACTORS AND DELIVERY SYSTEMS USED FOR

DIESEL REFORMING 16

2.4 CATALYST USED FOR DIESEL REFORMING 17

2.5 KINETIC WORK DONE ON DIESEL REFORMING 20

CHAPTER 3: EXPERIMENTAL SECTION 22

3.1 PRECIS 22

3.2 CHEMICALS, GASES, EQUIPMENT, AND INSTRUMENTS 23

3.2.1 Chemicals 23

3.2.2 Gases 23

3.2.3 Equipments 23

3.2.4 Instruments 24

3.3 CATALYST PREPARATION 25

3.3.1 Support Preparation (Surfactant Assisted Method) 25

3.3.2 Catalyst Preparation (Wet Impregnation) 26

3.3.3 Pelletizing & Sieving 27

3.3.4 Activation 27

vii

3.4 CATALYST CHARACTERIZATION 28

3.4.1 N2 Physisorption 28

3.4.2 H2 Chemisorption 28

3.4.3 Temperature Programmed Reduction (TPR) 29

3.5 REACTION FEEDSTOCK 30

3.5.1 SD Composition (Feed#1) 30

3.5.2 Air Composition (Feed#2) 32

3.6 OPERATION CONDITIONS AND VARIABLES 33

3.7 EXPERIMENTAL SETUP AND PROCEDURE 36

3.7.1 Diagram of Schematic Setup 36

3.7.2 Description of Reaction Process 37

3.8 EQUATIONS USED TO CALCULATE CONVERSION, YIELD, AND

SELECTIVITY 39

CHAPTER 4: RESULTS & DISCUSSION 40

4.1 PRECIS 40

4.2 CATALYST CHARACTERIZATION RESULTS 41

4.2.1 N2 Physisorption 41

4.2.2 H2 Chemisorption 42

4.2.3 Temperature Programmed Reduction (TPR) 44

4.3 EXTENDED TOS STABILITY STUDY 46

4.4 KINETIC EXPERIMENTS RESULTS 48

4.4.1 Effect of W/FSD,0 55

viii

4.4.2 Effect of Feed Ratio (O2/SD) 55

4.4.3 Effect of Temperature 57

4.5 KINETIC MODELING STUDY 59

4.5.1 Empirical Rate Model 59

4.5.2 Investigation of Heat and Mass Transfer Limitation 61

4.5.2.1 Heat Limitation 61

4.5.2.2 Mass Limitation 63

4.5.3 Experimental Rates of Reaction 66

4.5.4 Estimation of Parameters of Rate Model 69

CHAPTER 5: CONCLUSIONS AND RECOMMENDATIONS 72

REFERENCES 74

APPENDICES 79

ix

LIST OF TABLES

3.1 Physical properties of fuels and their percentage composition in SD 31

3.2 Design of kinetics experiments 35

4.1 N2 physisorption result 42

4.2 H2 chemisorption result 43

4.3a SD conversion and H2 selectivity with W/FSD,0 and T at O2/SD molar

ratio of 6.7

50

4.3b SD conversion and H2 selectivity with W/FSD,0 and T at O2/SD molar

ratio of 8.0

51

4.3c SD conversion and H2 selectivity with W/FSD,0 and T at O2/SD molar

ratio of 9.3

52

4.3d SD conversion and H2 selectivity with W/FSD,0 and T at O2/SD molar

ratio of 10.5

53

4.4 Experimental kinetic data table 1 67

4.5 Experimental kinetic data table 2 68

4.6 Estimation of the values of the parameters of the models 69

x

LIST OF FIGURES

3.1 Schematic of the experimental setup for the CPOX reforming of SD

using a packed bed tubular reactor (PBTR)

36

4.1 TPR patterns of titled Ce0.5Zr0.33M0.085N0.085O2 support and 5wt.%Ni/

Ce0.5Zr0.33M0.085N0.085O2 catalyst

45

4.2 Extended TOS stability study over 5wt.%Ni/Ce0.5Zr0.33Ca0.085Y0.085O2

catalyst for CPOX of SD reforming reaction. [T=1173K; Feed Ratio:

O2/SD=9.3; W/FSD,0 = 38052 kgcat*s/kmolSD]

47

4.3 A representative GC data sheet of a SD reforming reaction experiment.

[T=1173K; O2/SD=6.7; W/FSD,0=19008kgcat*s/kmolSD]

49

4.4 Variation of SD conversion with weight-time (W/FSD, 0) at different

temperature and different O2/SD.

54

4.5 Effect of feed ratio on the SD conversion and product distribution at

W/FSD,0=38,052 kgcat*s/kmolSD and (a) T=1123K, (b) T=1173K, (c)

T=1223K

56

4.6 Effect of reaction temperature on the SD conversion and product

distribution at W/FSD,0=38052 kgcat*s/kmolSD and (a) O2/SD=8.0, (b)

O2/SD=9.3

58

4.7 Diffusion regime showed in terms of Wagnar modulus 64

4.8 Parity plot of predicted rate vs. experimental rate for PLM 1 70

4.9 Parity plot of predicted rate vs. experimental rate for PLM 2 70

xi

LIST OF APPENDICES

APPENDIX A: Calculation of molecular formula for synthetic diesel

(SD)

79

APPENDIX B: Calculation of Diffusion coefficient of SD in Air

(DAB) and Deff

83

APPENDIX C: Calculation of feed mixture (SD and Air) vapor

density (ρmix)

87

APPENDIX D: Calculation of feed mixture (SD & Air) vapor

viscosity (µmix)

88

APPENDIX E: Calculation of mass transfer coefficient (kc) 92

APPENDIX F: Calculation of heat capacity (Cp) feed stream (SD &

Air) at T = 1123K

94

APPENDIX G: Calculation of thermal conductivity of SD (λA) 97

APPENDIX H: Calculation of standard heat of reaction ( 0

rxnH ) 99

APPENDIX I: Calculation of heat transfer coefficient (h) 100

APPENDIX J: Calculation of internal pore heat transfer resistance

maxparticle,T

102

APPENDIX K: Calculation of external film heat transfer resistance

maxfilm,T

103

xii

APPENDIX L: Calculation of Mears‘ criteria for heat transport

limitation

104

APPENDIX M: Calculation of Weisz-Prater criterion for internal

mass diffusion

105

APPENDIX N: Calculation of external film diffusion limitation

(Levenspiel, 1999)

106

APPENDIX O: Calculation of Mears‘ criterion for external film

diffusion limitation

107

APPENDIX P: MSDS of compressed air, PRAXAIR 108

APPENDIX Q: NLREG code with results for PLM 1 & 2 109

APPENDIX R: Mole balance of Run#13 112

APPENDIX S: GC Datasheets of each experiment 116

xiii

NOMENCLATURE

Notation

Å Angstrom

Ac Reactor cross-sectional area

CAb Concentration of A in the bulk, kg/m3

Ca Calcium

CaO Calcium oxide

Ce Ceria

CH4 Methane

CO2 Carbon di-oxide

CO Carbon monoxide

Cp Specific heat capacity

ipd,wpC Wiesz-Praster Criterion for pore diffusion

d Diameter of reactor

dp Diameter of catalyst particle

D Diffusivity

xiv

DAB Diffusion of SD (A) in Air (B), m2/s

DNi Ni dispersion

E Activation energy

h Heat transfer coefficient

H2 Hydrogen

H2O Water

He Helium

kc Mass transfer coefficient, m/s

k0 Pre-exponential factor or collision factor

k Reaction rate constant for the experimental runs performed at feed

ratio 10.5 only

L Length of catalyst bed

Lc Characteristic length

Mi Molecular weight of i species

MW Wagner Modulus

N2 Nitrogen molecule

Ni Nickel

NiO Nickel oxide

xv

Ni Molar flow rate of i species

O2 Oxygen

P Pressure

R Molar gas constant, 8.314 J/mol/K

Rc Catalyst particle radius, m

Ar Rate of reaction with respect to SD, kmol/m3/s

SNi Ni surface area

SH2 Selectivity of H2

T Temperature

W/FSD,0 Weight-time

xi Mol fraction of compound (i) in the mixture

XSD Conversion of SD

Y Yttrium

Y2O3 Yttrium oxide

YH2 Yield of H2

Zr Zirconium

xvi

Abbreviations

APU Auxiliary power unit

ATR Autothermal reforming

BET Brunauer, Emmett and Teller

BEV Battery electric vehicle

CPOX Catalytic partial oxidation

CTAB Cetyl-trimethylammonium bromide

DI De-ionized

DOE Department of Energy

ER Eley-Rideal

FCV Fuel cell vehicle

FT-SD Fischer-Tropsch synthetic diesel

GTL Gas-to-liquids

HD Hexadecane

IC Internal Combustion

LH Langmuir-Hinshelwood

LHHW Langmuir-Hinshelwood Hougen-Watson

http://en.wikipedia.org/w/index.php?title=Stephen_Brunauer&action=edit&redlink=1

http://en.wikipedia.org/wiki/Paul_H._Emmett

http://en.wikipedia.org/wiki/Edward_Teller










xvii

LPG Liquefied petroleum gas

MSDS Material safety data sheets

NLREG Non-linear regression

NSc Schmidt number

NSh Sherwood number

NRe Raynolds Number

NPr Prandlt Number

PBTR Packed bed tubular reactor

POX Partial Oxidation

ppmv parts-per-million by volume

PV Pore volume

SA Surface area

sccm Standard cubic centimeter per minute

SD Synthetic diesel

SR Steam reforming

TCD Thermal conductivity detector

TPR Temperature programmed reduction

xviii

TOS Time on stream

ULSD Ultra low sulfur diesel

UHP Ultra high pure

Greek letter

p Void fraction

λ Thermal conductivity

Tortuosity factor

Φ Porosity

Ψ Shape factor

μ Vapour viscosity

ν Velocity

ρ Density

H Enthalpy of reaction

0

rxnH Heat of reaction at standard temperature and pressure (25°C and 1

atm)

Superscripts

m order with respect to A

xix

n order with respect to B

o order with respect to C

p order with respect to D

Subscripts

b bulk

cat catalyst

eff effective

g gas

ipd internal pore diffusion

max maximum

A synthetic diesel

B oxygen

C carbon monoxide

D hydrogen

1

CHAPTER 1

INTRODUCTION

1.1 HYDROGEN ECONOMY

Hydrogen is the lightest element in the periodic table of elements and is the most

abundant element in the universe. At standard temperature and pressure, hydrogen exists

as a gas. It is colourless, odorless, tasteless, and lighter than air. Hydrogen gas is a

diatomic molecule (H2); each molecule has two atoms of hydrogen. Like electricity,

hydrogen is an energy carrier (not an energy source), meaning it can store and deliver

energy in an easily usable form. Although abundant on earth as an element, hydrogen

combines readily with other elements and is almost always found as part of some other

substance, such as water (H2O), or hydrocarbons (Balat, 2008). That is why it needs to be

produced from these compounds in order to use it.

The fact, that hydrogen reacts with oxygen to produce energy and water (Eq. 1) is

the basis of so-called hydrogen energy.

;C25atmol

kJ8.285H;OHO

2

1H 0

222 ------------------------------------ (1.1)

Hydrogen has a high energy yield of 122kJ/g, which is 2.75 times greater than

hydrocarbon fuels, but not energy per unit volume—of any fuel. Its relatively low

volumetric energy content poses a significant challenge for storage. However,

combustion of hydrogen for energy produces only water, so it holds great promise as a

2

clean source. However, regardless of its source (water or other hydrogen-containing

compounds), hydrogen itself is considered to be the cleanest energy carrier. For that

reason hydrogen holds the promise of an ideal future fuel, with many social, economic,

and environmental benefits to its credit. Moreover, it also has the long-term potential to

reduce dependence on foreign oil and lower the carbon emissions from the transportation

and energy sectors.

Given the continued growth in the world‘s population, as well as the progressive

industrialization of developing nations, the global demand for energy is expected to

continue to escalate in the coming decades – by more than 50% until 2030, according to

the International Energy Agency (IEA) – with fossil fuels continuing to dominate global

energy use. At the same time, there is a growing global consensus that greenhouse gas

(GHG) emissions, which keep rising, need to be managed in order to prevent dangerous

anthropogenic interference with the climate system. These concerns over energy supply

security, climate change, as well as local air pollution and the increasing prices of energy

services are having a growing impact on policy development throughout the world (Ball

and Wietschel, 2009).

Today‘s energy and transport systems, which are based mainly on fossil energy

carriers, can in no way be considered sustainable in the long-term. The transportation

sector today accounts for some 18% of primary energy use and some 17% of global CO2

emissions (Ball and Wietschel, 2009).

Many countries around the world are seriously considering the implications of a

shift towards a hydrogen economy. There is no universally accepted definition of the

3

‗‗hydrogen economy,‘‘ but it is generally viewed as the replacement of the vast majority

of petroleum fuels used by transportation vehicles of all kinds (automobiles, trucks,

trains, and aircraft) with hydrogen being burned in IC engines, external-combustion (jet)

engines, or preferably, used in fuel cells to more efficiently generate power for

transportation (Balat, 2008). Although modern cars emit far less toxic pollutants,

including hydrocarbons, nitrogen oxides, carbon monoxide, and particulates, their

increasing number is resulting in increasing automobile pollution. Here, fuel cells will

help diminish poisonous emissions into the atmosphere and also have higher electrical

efficiency compared to internal combustion engines. Alternative, long-term visions of a

hydrogen economy have been articulated based on large-scale use of renewable fuels, a

mixture of clean coal and fossil fuels (with carbon sequestration), and/or nuclear power

(Balat, 2008).

On the other hand, alternative fuels are not available everywhere—one location

might prefer ethanol and another might be dominated by biodiesel, or GTL fuel, or

methane. Most of these fuels require a different engine technology for efficient operation.

However, hydrogen can be produced using diverse, domestic resources, including fossil

fuels, coal, and natural gas; nuclear; and biomass and other renewable energy

technologies, such as wind, solar, geothermal, and hydroelectric power. This great

potential for diversity of supply is an important reason why hydrogen is such a promising

energy carrier (DOE, November 2008). In addition, hydrogen can be produced at large

central plants as far as several hundred miles from the point of end-use; semi-centrally,

25 to 100 miles from the point of end-use; or in small, distributed units located at or very

4

near the point of end-use, such as at refueling stations or stationary power sites (DOE,

November 2008).

At present, hydrogen, representing a market of roughly fifty billion US$ for 40 Mt

annual production, is mainly used as a chemical substance rather than a fuel. The most of

its current uses are found as processing agents in oil refineries (e.g., for desulphurization

and upgrading conventional petroleum) and in chemical production processes (e.g.,

methanol, ammonia, and pharmaceuticals) (Dincer, 2011).

For the passenger car market, fuel cells offer the benefits of zero-emissions

operation without the range and charging limitations of pure battery electric vehicles

(BEVs). While the market for fuel cell vehicles (FCVs) has been slower to develop than

many anticipated a few years ago, major automakers, including Toyota, Daimler, GM,

Honda, and Hyundai, have all publicly stated that fuel cells are a critical piece of a

complete clean vehicle portfolio. Commercialization is expected to accelerate beginning

in 2015. According to a recent report from Pike Research, cumulative commercial sales

of FCVs will surpass 1 million by the end of this decade, generating $16.9 billion in

annual revenue by 2020 (Pike research website).

U.S. Department of Energy has developed a multiyear plan with aggressive

milestones and targets for the development of hydrogen infrastructure, fuel cells, and

storage technologies. The targeted hydrogen cost is $2-4/kg (the energy equivalent of 1

gallon of gasoline) delivered as cited in Holladay et al. (2009). The United States stands

to profit from hydrogen technologies. A recent study projected global annual demand for

stationary and transportation fuel cell products to reach $46 billion by 2011 and more

5

than $2.5 trillion by 2021. Government and industry investment in hydrogen and fuel cell

technologies has positioned the United States as a leader in this rapidly growing market

(US DOE 2011). Global investment in hydrogen has accelerated dramatically over the

past few years and is now in the range of several US billion dollars. Japan also recently

announced plans to introduce around 4000 hydrogen filling stations by 2020. Perhaps the

best-known example of a ‗hydrogen economy‘ is Iceland, which has set a goal for a

complete transition to hydrogen by 2030 (DOE, November 2008).

6

1.2 DIESEL REFORMING TECHNOLOGY

Although hydrogen is the most abundant element in the universe, it does not

naturally exist in its elemental form on Earth. Pure hydrogen must be produced from

other hydrogen-containing compounds, such as fossil fuels, biomass, or water. Each

production method requires a source of energy (i.e., thermal (heat), electrolytic

(electricity), or photolytic (light) energy).

However, the lack of infrastructure for H2 production, distribution, and delivery,

as well as the current cost of hydrogen production from nuclear and renewable energy,

has led to the consideration of hydrogen generation from various hydrocarbons. Steam

reforming of natural gas (methane (CH4)) is already a well establish process to produce

hydrogen in the fertilizer industry. Various hydrocarbon fuels other than methane have

been investigated for H2 production by reforming technology, including, for example,

propane, LPG, butane (Ayabe et. al., 2002), alchohols (Ahmed et. al., 1998), gasoline

(Moon et. al., 2001, & Cheekatamarla and Thomson, 2005), diesel (Pereira et. al., 2000,

& Cheekatamarla and Lane, 2005a), and JP8 (Cheekatamarla and Lane, 2005a, &

Cheekatamarla and Lane, 2005b). Selection of fuel for hydrogen production depends on

technical and/or economic and political factors. However, for diesel, a well-developed

distribution network is already in place, making it an ideal source for H2 production

onboard or in stationary facilities supplying refueling stations with H2. Moreover,

gasoline and diesel both have a higher energy density (compared to natural gas) and

larger hydrogen content as compared with oxygenated hydrocarbons such as methanol

and ethanol (Ibrahim and Idem, 2007). The increasing demand of electric power in

vehicles for either controls or comfort and safety features has made auxiliary power unit

7

(APU) systems more attractive because of the low energy storage capacity of

conventional lead-acid batteries. The APU system converts chemical energy of fuel such

as gasoline or diesel directly into electric power without combustion. APU consists of

two technologies: fuel reforming to produce hydrogen and fuel cell technology to

produce electric energy from hydrogen. The high efficiency of APU units makes them the

technology of choice for an engine-independent supply of electrical power in all kinds of

vehicles for both driving and stationary modes (Kim et. al., 2011 & Lindermeir et. al.,

2007).

Steam reforming (SR), partial oxidation (POX), and auto-thermal reforming

(ATR) are the major technologies for producing hydrogen (Kim et al., 2011, and Yoon et

al., 2008). SR provides the highest reforming efficiency and H2 yield, but the reaction is

highly endothermic and, therefore, consumes a large amount of heat energy, so the

reactor needs to be designed to promote heat transfer from external heat sources.

Therefore, SR is less attractive for on-board production of H2 for transportation fuel cell

systems that require rapid start up and dynamic response (Yoon et al., 2008) but is

attractive for large- and small-scale industrial H2 generation. In contrast, POX systems

are highly exothermic and sustain the process heat once initiated. They also have the

advantage of rapid start up and fast response to load change due to fast reaction rates.

Though POX shows relatively lower hydrogen production compared to SR and ATR,

POX reforming requires a smaller reactor to achieve high conversions and eliminates the

need of steam as a feed, which makes POX systems attractive for on-board application

for APU. On the other hand, ATR combines POX and SR, and has the advantages of both

POX and SR, such as rapid start up, dynamic response, and relatively high H2 yield.

8

Diesel and JP8 ATR is reported as having higher reformer efficiency and fuel conversion

compared to SR and POX (Kang, I. and Bae, J., 2006, & Ahmed, S., and Krumpelt, M.,

2001). However, ATR requires a proper combination of POX and SR; otherwise,

formation of hot spots in the reactor might occur, and in such a situation, degradation of

catalyst will be inevitable. Mixing enhancement of reactants and carbon deposition on

catalyst are the two major challenges for diesel ATR. The other disadvantages associated

with the ATR process are the need for steam generation and larger reactor dimensions,

which makes ATR unsuitable for transportation and APU applications.

9

1.3 KNOWLEDGE GAP AND PROBLEM IDENTIFICATION

The greatest technical challenge to hydrogen production is cost reduction. For

transportation, a key driver for energy independence, hydrogen must be cost-competitive

with conventional fuels and technologies on a per-mile basis. This means that the cost of

hydrogen — regardless of the production technology, and including the cost of delivery

— must be in the range of $2.00 to $3.00 per gallon gasoline equivalent (untaxed) (DOE,

November 2008). Note that transportation fuels are often compared based on their

equivalency to gasoline. The amount of fuel with the energy content of one gallon of

gasoline is referred to as a gallon gasoline equivalent (gge).

In addition, the major challenge of diesel reforming is the deactivation of catalyst

due to coke formation, plugging of the catalyst bed, and hot spot formation during POX

reforming (Boon et al., 2011; Krummenacher et al., 2003; Parmar et al., 2009). Heavier

hydrocarbons in diesel with low H/C ratio are more likely to produce coke, as it is

thermodynamically favourable even under POX conditions (Parmar et al., 2009). To

overcome this, researchers have used precious metals as active components and complex

reactor (reformer) designs with spray nozzles or separate vapourization chamber in the

experimental set up. Performance tests, as well as kinetic study experiments, on synthetic

diesel, which consists of seven hydrocarbon components, have never been attempted in

the literature, to the best of our knowledge, due to the complicated reaction pathways

involved in the process. Studying reaction kinetics is very important for of three reasons:

(1) understanding the reaction mechanism, (2) designing a suitable reactor, and (3)

specifying operating conditions, control methods, and auxiliary equipment to meet the

technological and economic needs of the reaction process. The novelty of the current

10

work is that, in the current study, a base metal catalyst was employed; a syringe pump

was used to deliver the SD feed; a conventional packed bed tubular reactor (PBTR) was

used; and a mixture of seven different aliphatic, and aromatic, olefinic hydrocarbons,

which are predominately found in diesel, were used as synthetic diesel feed for all the

experimental runs performed in the present study.

11

1.4 RESEARCH OBJECTIVES

The objectives of the research presented in this thesis were to:

(1) Evaluate the performance of Ni-based catalyst for hydrogen production by POX

of SD feed.

(2) Test a realistic and complex synthetic diesel feed (mixture of paraffinic, olefinic,

and aromatic hydrocarbons) instead of a single hydrocarbon as a diesel surrogate

for POX reaction

(3) Use a simplified reactor setup (PBTR) and a conventional delivery system for

feeding the SD.

(4) Carry out experiments with a wide range of process parameters (temperature,

contact time, and O2/C ratio)

(5) Carry out a kinetic modeling study in order to establish an empirical rate

expression representing the system globally (i.e., to fit the experimental data).

12

CHAPTER 2

LITERATURE REVIEW

2.1 PRECIS

An extensive literature survey was conducted to identify the current developments

in the area of hydrogen production by diesel fuel reforming. As indicated earlier, the

contributions of the current research to the advancement of scientific knowledge can be

classified into four core areas, which are as follows: (1) introduction of a complex SD

feed; (2) simplified reaction process, delivery system, and reactor design; (3) use of Ni-

based, inexpensive catalyst; (4) studying kinetics of the diesel POX reforming reaction.

The current chapter reviews the scientific literature in the above stated four core areas.

13

2.2 HYDROGEN FROM FOSSIL FUEL REFORMING

2.2.1 Technologies

An outline of significant current and developing hydrogen production

technologies is given below (adopted from Holladay et al., 2009).

Hydrogen production technologies:

A. Fuel Processing

Hydrocarbon reforming

Steam reforming (SR)

Partial oxidation (POX)

Auto thermal reforming (ATR)

Pyrolysis

Plasma reforming

Aqueous phase reforming

Ammonia reforming

B. Non reforming hydrogen production

Hydrogen from biomass

Photobiological & Photoelectrochemical

Hydrogen from water

Electrolysis

High temperature thermochemical water splitting

14

Nuclear high temperature electrolysis

Renewable electrolysis (electricity from wind, solar,

geothermal, hydroelectric power, etc., to split water)

Steam reforming (SR), partial oxidation (POX), and auto-thermal reforming

(ATR) are the major technologies for producing hydrogen from hydrocarbon fuels. Their

pros and cons have been discussed in brief in section 1.2. Significant work on POX

reforming of hydrocarbon fuels has been reported in the literature (Kim et al., 2011;

Mundschau et al., 2008; Krummenacher et al., 2003; Haynes et al., 2010a, 2010b, 2008;

Shekhawat et al., 2006; Ibrahim and Idem, 2006 & 2007; Elghawi et al., 2008).

Meanwhile, Boon et al. (2011), Gawade et al. (2010), and Thormann et al. (2009) have

worked on SR, and Cheekatamarla and Lane (2005, 2006), Karatzas et al. (2010),

Alvarez-Galvan et al. (2008), Mota et al. (2010) have worked on ATR of hydrocarbon

fuels.

2.2.2 Diesel as Source of Fossil Fuel

The current study used a synthetic diesel free of sulfur, composed of 75 vol.%

saturated hydrocarbons and 25 vol.% aromatics, and possessing a hydrogen-to-carbon

ratio (H/C) of < 2 (1.94), which is far less than natural gas (CH4), where H/C is 4. It is a

well-known fact in the literature that it is difficult to reform higher hydrocarbon feeds, as

lower H/C ratios promote coking and catalyst deactivation. From an extensive literature

survey, it was noted that most of the research reported to date employed feeds with a

higher H/C ratio (>2.0) composed of either one single paraffin (Kim et al., 2011, Gawade

15

et al., 2010, Thormann et al., 2009, Alvarez-Galvin et al., 2008), a combination of two

paraffins, or one aromatic component along with paraffin (Pitz and Mueller, 2011) as

surrogates for diesel. In addition, some researcher have used ultra low sulfur diesel

(ULSD) (15 ppmv S) (Krummenacher et al., 2003) and diesel surrogates containing

sulfur as feeds but have used expensive precious metal catalysts and/or special reactor

designs or have introduced a sulfur purifier (scrubber) as an additional unit to the

experimental setup.

Krummenacher et al. (2003) used ULSD to produce H2 by POX, but their H/C

ratio was ~2.14 (>2.0). They also used a Rh-based catalyst and further employed a spray

nozzle to introduce the feed along with a vapourization chamber. Mundschau et al.

(2008) used pump grade diesel for POX. Haynes et al. (2010a, 2010b, 2008) and

Shekhawat et al. (2006), from the same group, used n-tetradecane as a source of paraffin,

1-methylnaphthalene (1-MN) as a source of aromatic component, and dibenzothiophene

(DBT) as a source of sulphur (50 ppmv) to formulate a synthetic diesel feed. Elghawi et

al. (2008) have also used ULSD (≤ 15 ppmv) for POX over precious metal catalyst.

Boon et al. (2011) used commercial diesel (Aral Ultimate and BP Ultimate) to

investigate the SR process. Cheekatamarla and Lane (2005a, 2006) used SD (with 10

ppm sulfur) and JP8 (1000 ppm sulfur) in an ATR process. Likewise, Karatzas et al.

(2010) also worked with low sulfur FT diesel in an ATR process.

16

2.3 REACTORS AND DELIVERY SYSTEMS USED FOR DIESEL

REFORMING

Kim et al. (2011) used a fixed bed reactor, but their reformer used a specially

designed nozzle for feeding hexadecane (diesel surrogate) and an electrically-heated cell

as a fuel vapourizer/ignition chamber prior to processing in a catalyst bed. Likewise, an

automotive fuel injector and vapourization chamber was also designed by Krummenacher

et al. (2003) for heavy hydrocarbon POX. Mundschau et al. (2008) proposed a very

uniquely-designed membrane reactor. Haynes et al. (2010a, 2010b, 2008), and from the

same group, Shekhawat et al. (2006), used a very simple fixed bed reactor with a furnace

around it. Their set up resembles the one employed in the current study. Furthermore, the

above-mentioned groups worked with sulfur-containing diesel fuels. Elghawi (2008)

worked with a millisecond fixed bed reactor for ULSD, FT-SD, and biodiesel POX

reforming.

Boon et al. (2011) and Gawade et al. (2010) also used a specially-designed nozzle

and vapourization chamber for SR of fuel reforming. Cheekatamarla et al. (2006),

Karatzas et al. (2010), and Alvarez-Galvan et al. (2008) all used either a spray nozzle or

vapourization chamber or both for ATR hydrocarbon fuel (diesel/surrogates of diesel)

reforming.

17

2.4 CATALYST USED FOR DIESEL REFORMING

Catalysts have been used for over 2000 years (Fogler, 1999). The impact of

catalysis and catalysts is substantial. Today, over 90% of all industrial chemicals are

produced with the aid of catalysts. Catalysts impact a sizable fraction of any nation‘s

gross domestic product. In 1991, it was estimated that the total value of fuels and

chemicals derived from catalysts exceeded $900 billion/year. World catalyst demand is

forecast to grow to $16.3 billion through 2012 according to the Freedonia Group (Armor,

2010). In 2003, global sales of catalysts exceeded 12 billion dollars, which was up from

9.3 billion dollars in 1998 (Armor, 2010). This section provides a summary of the types

and nature of catalysts that have been used for diesel or synthetic diesel or surrogates of

diesel reforming to produce H2.

Kim et al. (2011) have reported POX of n-Hexadecane (n-HD) as a surrogate for

diesel over a Pd -based monolith-type catalyst (PdO/CeO2/BaO/SrO/Al2O3). According to

their study, Pd was used because noble metals are coke tolerant and Ni is not. On the

other hand, monolith-type catalysts prevent hot spots from occurring on the catalyst

surface, which decreases catalyst fouling and sintering. Krummenacher et al. (2003)

reported POX of higher hydrocarbon (decane, hexadecane, and ULSD) over

~5wt%Rh/Al2O3 monolith-type catalyst to produce syngas with 80% H2 selectivity and

>98% fuel conversion with several hours of stable operation.

Haynes et al. (2008) attempted POX of n-tetradecane (TD) as a surrogate for

diesel over 1wt%Pt/γ-Al2O3 and Co0.4Mo0.6Cx catalyst, resulting in 98% conversion and

>76% H2 yield for seven hours of stable operation, establishing no need for a noble

18

metal-based catalyst. However, the process failed when they add 1-methylnaphthalene (1-

MN) as a source of aromatics and dibenzothiophene (DBT) as a source of sulfur (50

ppmw) to simulate diesel. Haynes et al. (2010a) came up with an improved catalyst, a

Ru-substitute pyrochlore catalyst (La1.5Sr0.5Ru0.05Zr1.95O7-δ), and in comparison with

traditional 5%Ru/ γ-Al2O3 for POX of TD, 1-MN, and DBT mixture at 900°C, both gave

>99% conversion and >75% H2 yield and was stable for 5 hours of experimental run.

Later on, the same group came with a Ni-based pyrochlore catalyst (La1.5Sr0.5Zr1.72

Ni0.15O7-δ) for the same feed and same operating conditions, resulting in >50% H2 yield

and stable performance with no deactivation of catalyst (Haynes et al., 2010b).

Elghawi et al. (2008) reported CPOX of ULSD, rapeseed methyl ester (RME) –

biodiesel, and Fischer-Tropsch synthetic diesel fuels over a prototype monolith catalyst

(1%Rh/CeO2-ZrO2) in a millisecond reactor. Mundschhau et al. (2008) used perovskite

catalyst (La0.5Sr0.5CoO3-δ and La0.5Sr0.5FeO3-δ) in a membrane reactor for POX of pump

grade, low-sulfur diesel fuel (<9 ppm S by mass) at 950ºC. They found the activity of

perovskite catalyst to be similar to noble metal and to be sulfur tolerant.

SR and ATR of diesel were attempted more often by researchers than POX. Boon

et al. (2011) have reported SR of commercial diesel (Aral Ultimate and BP Ultimate)

over commercial Ni-based and precious metal-based catalysts for 118 hours and 1190

hours, respectively, without any deactivation of catalyst and with 100% conversion and

40-50% H2 yield. Thormann et al. (2009) using Rh/CeO2 catalyst and Gawade et al.

(2010) using 0.5wt% Rh10wt%Ni/γ-Al2O3 have also reported successful POX of

hexadecane as a surrogate for diesel. Cheekatamarla and Lane (2005a) used Pt/CeO2

catalyst for ATR of SD (with 10 ppm S) and JP8 (1000 ppm S). They presented 50 hours

19

worth of results where they have shown stable performance for SD conversion, but

deactivation occurred with JP8 due to higher sulfur content. The same group,

Cheekatamarla and Lane (2006), presented a modified catalyst (Ni-Pt/CeO2) for ATR of

JP8 with stable H2 production for 50 hours with >75% H2 yield. As JP8 has about 1000

ppm sulfur, this is a great achievement for them. Karatzas et al. (2010) used Rh-based

monolithic catalyst (3wt% Rh supported on alumina doped with Ce/La (Rh3.0Ce10La10/δ-

Al2O3) for ATR of n-tetradecane, low sulfur, and Fischer-Tropsch diesel for 40 hours

with >99% conversion and stable performance.

A close examination of the literature shows that most of the researchers have tried

precious metal-based catalyst, but because of high costs and non-availability, the precious

metal catalysts are not suitable for commercial applications. Ni-based catalysts can be

employed to resolve the above issues, as their catalytic activity is similar to that of noble

metals; however, the Ni-based catalysts suffer from deactivation on account of carbon

deposition. The coking can be overcome by employing an easily reducible redox support

(Sukonket et al., 2011).

In the current work, Ni-based, low cost catalyst was successfully tested for SD POX for

an extended period of operation with desirable results.

20

2.5 KINETIC WORK DONE ON DIESEL REFORMING

Kinetic study of diesel or synthetic diesel reforming, either by SR or POX or

ATR, is yet to be published, to the best of our knowledge. Most of the published works in

the area of diesel reforming have concentrated on catalyst development.

However, Thormann et al. (2009) used hexadecane (HD) as a surrogate for diesel

and developed a kinetic model for steam reforming of HD over Rh/CeO2 catalyst in a

microchannel reactor. They reported 100% conversion for HD, reaction order, with

respect to HD, of ≈ 0.5, and activation energy of 71.0kJ/mol for dissociative adsorption

of HD on Rh sites. According to Thormann et al. (2009), Patel et al. (2007) has also done

HD steam reforming in millimeter channels coated with Rh/Al2O3 catalyst and proposed

a power law model, an Eley-Rideal (ER) model, and a Langmuir-Hinshelwood (LH)

model. While their LH model failed mechanistically and the power law failed to predict

the experimental results, the ER model was a success. Gawade et al. (2010) have studied

combustion and reforming kinetics of n-HD over Rh//Ni catalyst supported on alumina

and designed a flexible fuel reformer where combustion and reforming can be done

simultaneously. They have developed three mechanistic models (ER, LH bimolecular

adsorption, and LH dual site models) and found the ER model produced a good fit and

fulfilled the thermodynamic criteria. They also found that the power law model produced

the best fit among the other models. They reported the reaction order with respect to HD

to be 0.31 and activation energy to be 44.1±12.74kJ/mol for combustion of HD from the

power law model and compared this result with Sawyer‘s (1995) Ph.D. work, where

Sawyer reported 37-54 kJ/mol for dodecane combustion. Gawade et al. (2010) also

reports 99.3±0.4 kJ/mol activation energy for steam reforming of HD from ER model.

21

Praharso et al. (2004) presented a dual site LH model that was a good fit for steam

reforming of iso-Octane over Ni-based catalyst. Praharso et al. (2004) reported the

reaction order with respect to iso-octane to be ≈ 0.2 and ≈ 0.5 with respect to steam and

activation energy of 44±2.2 kJ/mol. Praharso‘s work also demonstrated a trend of lower

activation energy for higher HC steam reforming.

Ibrahim and Idem (2007) reported kinetic study of POX of synthetic gasoline, and

a dual site mechanism of LHHW formulation was found to best fit the experimental data,

and they estimated the apparent activation energy to be 19kJ/mol, while the reaction

order with respect to gasoline was 3.1. Their study is of great interest, as they have used a

mixture of five hydrocarbons to represent gasoline. Finally, after an exhaustive literature

search, it can be confidently stated that the current work on kinetics of SD POX is the

first of its kind in the scientific literature.

22

CHAPTER 3

EXPERIMENTAL SECTION

3.1 PRECIS

A total of thirty six (36) experiments and an extended time on stream stability

experiment were conducted to study the kinetics of the catalytic partial oxidation reaction

of synthetic diesel. This section provides the details of those experimental runs. First, the

catalyst preparation method is described, along with catalyst characterization procedures.

Then, the composition of feed materials, operating conditions, and variables for the

reaction are presented. Finally, a brief description of the reaction process is added for the

reader.

23

3.2 CHEMICALS, GASES, EQUIPMENTS & INSTRUMENTS

3.2.1 Chemicals

Nickel (II) Nitrate Hexahydrate [Ni(NO3)2.6H2O] salt (99.999% purity, Aldrich)

Cerium (III) Nitrate Hexahydrate [Ce(NO3)3.6H2O] (99% purity, Aldrich)

Zirconyl(IV) Nitrate Dihydrate [ZrO(NO3)2.2H2O] (99.9% purity, Alfa Aesar)

Calcium (II) Nitrate Tetrahydrate [Ca(NO3)2.4H2O] (99% purity, Aldrich)

Yttrium (III) Nitrate Hexahydrate [Y(NO3)3.6H2O] (99.8% purity, Aldrich)

Cetyl Trimethyl Ammonium Bromide [C19H42NBr], (CTAB; g98% purity, Sigma

Aldrich)

Ammonium hydroxide reagent grade, 28-30% (w/w); ACS-Pure: Fisher

3.2.2 Gases

Helium gas: Praxair (Ultra High Purity Grade, 99.999%)

Nitrogen gas: Praxair (99.999% purity)

5%H2/bal. N2: Praxair

Air (19.5-23.5%O2/bal. N2): Praxair

3.2.3 Equipment

Tube furnace: ZCP 386, Zesta Engineering, Ltd.

Digital mass flow controller: GFC 171S, 0-500 mL/min, Aalborg

24

Mass flow meter (GFM): GFM 171S, 0-500 mL/min, Aalborg

Hydraulic press, model 3912, Carver

U.S.A. standard test sieve, ASTM, E-11 specification, Fisher Scientific Company

Isotemp Muffle Furnace, Model 550-126, Fisher Scientific Company

Sensor Thermocouples, Type-K, AFEU0AQ180UK05X-2, Zesta Engineering

Limited

3.2.4 Instrument

Surface area and porosity: Micromeritics ASAP 2010

Metallic surface area and dispersion: Micromeritics ASAP 2010

Temperature programmed reduction (TPR): ChemBET 3000 TPR/TPD,

Quantachrome

25

3.3 CATALYST PREPARATION

A ceria-based mixed oxide catalyst of nominal composition

5wt%Ni/Ce0.5Zr0.33Ca0.085Y0.085 was employed in the current study, and it is abbreviated

as 5N/CZCaY. The selection of 5N/CZCaY catalyst for this study was the result of a

screening study taking n-Hexadacane as a surrogate for SD by our group (Idem et al.,

2010). Cerium (Ce) and zirconium (Zr) combinations have drawn much attention

recently, as they could improve textural properties, enhance activity, improve active

metal dispersion, eliminate or reduce coke formation, and prevent thermal sintering when

used as support material in catalyst (Ibrahim and Idem, 2008). Here, CaO and Y2O3 were

added as promoters for the catalyst (Idem et al., 2010). The 5% nickel loading was

selected on the basis of Ibrahim and Idem‘s (2008) studies on a similar hydrocarbon feed

(i.e. gasoline), where they have shown that lower nickel loadings, typically 3-5wt%,

exhibit better degrees of dispersion and high levels of stable conversion and selectivity.

An earlier publication from our group states that surfactant-assisted method yields better

catalyst than the conventional co-precipitation method, and a higher surfactant/metal ratio

is desirable for better stability (Sukonket et al., 2011). In accordance with the above

reference, the CZCaY support was prepared by surfactant assisted method, a

surfactant/metal ratio of 1.25 was found to be optimum. The details of the preparation

steps are given below.

3.3.1 Support Preparation (Surfactant Assisted Method)

In order to prepare the CZCaY (Ce0.5Zr0.33Ca0.085Y0.085) mixed oxide support by

the surfactant-assisted method, appropriate quantities of Cerium (III) nitrate hexahydrate,

26

Zirconyl(IV) nitrate dihydrate, Calcium nitrate tetrahydrate, and Yttrium (III) nitrate

hexahydrate precursor salts were dissolved in deionized (DI) water in different beakers.

Separately, a calculated amount of cetyl trimethyl ammonium bromide (CTAB) was

dissolved in DI water at 60ºC. All the metal nitrate solutions were then added to the

surfactant solution to obtain a mixture solution. The molar ratio of surfactant/metal

[CTAB]/[Ce+Zr+Ca+Y] was maintained at 1.25. Aqueous ammonia was gradually added

to the aforementioned mixture solution under vigorous stirring until precipitation was

complete (pH 11.8). The addition of ammonia induced the precipitation of gelatinous

yellow-brown colloidal slurry. The slurry was stirred for 60 min in a glass reactor, and

subsequently transferred into Pyrex glass bottles, sealed, and aged ―hydrothermally‖ in

autogenous pressure conditions for 5 days at 90ºC. After this timeframe, the bottles were

cooled, and the resulting precipitate was filtered and washed repeatedly with warm DI

water. The resulting cakes were oven-dried at 120oC overnight and finally calcined at

650oC for 3 hours in flowing air.

3.3.2 Wet Impregnation

A nominal 5 wt % Ni was loaded over the above-prepared support by a standard

wet impregnation method. In a typical impregnation, about 14.25 g of the catalyst support

was immersed in 127.8 mL of a 0.1 M Ni(NO3)2 solution. The mixture was subjected to

slow heating under constant stirring in a hot oil bath so as to remove the excess water.

The temperature of the oil bath was kept constant at 85oC, and in approximately 8 hours,

dried powder of Ni impregnated catalyst was formed. The dried powder thus obtained

was calcined at 650ºC in a furnace in air for 3 hours.

27

3.3.3 Pelletizing & Sieving

The catalyst powder was pelletized into thin wafers using a 4 cm i.d. die set

pressed under a hydraulic press. Thus obtained wafers were gently crushed and sieved to

desirable size (0.8 mm) by passing through the appropriate sieves.

3.3.4 Activation

Prior to each test run, the catalyst was activated in situ by reducing it (from NiO

to Ni) at 700C for 2 hours duration under flowing 5% H2 bal. N2 (100 sccm).

28

3.4 CATALYST CHARACTERIZATION

3.4.1 N2 Physisorption

The BET surface area, pore volume, and average pore size measurements for

5N/CZCaY catalyst and its CZCaY support were obtained by N2 physisorption analyses

at liquid N2 temperature using a Micromeritics ASAP 2010 instrument. Prior to analysis,

the samples were degassed at 180C under vacuum for 4 hours. Average pore size and

pore volume was analyzed using the desorption branch of the N2-isotherm. Each sample

was analyzed using N2 physisorption at least twice in order to establish repeatability. The

deviation in these measurements was ≤ ±5%.

3.4.2 H2 Chemisorption

The metallic surface area and metal dispersion in the 5N/CZCaY catalyst sample

were estimated using hydrogen chemisorption at 35C using a Micromeritics ASAP 2010

instrument. Prior to analysis, the catalyst sample was dried at 120C and then reduced in

situ in flowing H2 gas (UHP grade) at 700C for 2 hours (in order to mimic the reduced

state formed during the course of a typical catalytic run) followed by evacuation at 700C

for 1 hour before cooling down to 35C. The metallic surface area (SNi) was calculated

with the help of the following expression:

SNi = 13.58×10-20

NM (m2/gcat.) ---------------------------------------------------------------- (3.1)

Where, NM is the number of hydrogen molecules adsorbed in the monolayer per gram of

catalyst. The above expression was derived by considering the surface occupied per atom

of nickel as 6.79 Å2 per atom (considering the density of nickel as 8.91 g/cm

3 and a face-

29

centred cubic lattice) and the adsorption stoichiometry as 2 surface nickel atoms per

hydrogen molecule. The nickel dispersion (D %) was then calculated as the percentage of

surface nickel atoms with respect to total nickel atoms in the catalysts (Iglesia and

Boudart, 1983; Tsay and Chang, 2000). The H2 chemisorption analysis was repeated

twice for the sample in order to check reproducibility. The deviation in these

measurements was < ±1%.

3.4.3 Temperature Programmed Reduction (TPR)

The H2-TPR of 5N/CZCaY catalyst and CZCaY support samples were performed

on a Quantachrome ChemBET 3000 unit equipped with a thermal conductivity detector

(TCD). For both samples (except pristine NiO) investigated by TPR, exactly the same

amount was analyzed so as to make comparison possible. Prior to TPR measurements,

the samples were outgassed at 180C in an inert atmosphere (N2 UHP grade) for 2 hours.

The reducibility of the support, as well as that of the catalyst prepared, were studied using

the TPR technique in a temperature range from ambient to 105C at a heating rate of

15C/min using 5%H2/bal.N2 as the reactive gas (flow rate = 45 sccm). The total reactive

gas consumed during TPR analysis was measured. The H2 uptake as a function of TCD

response vs. temperature was used to plot the TPR profile. For reference purposes, the

TPR profiles of pristine NiO and CeO2 were also studied. Each sample was analyzed

using TPR at least twice in order to establish reproducibility. The deviation in Tmax values

was found to be less than ± 4C.

30

3.5 REACTION FEEDSTOCK

The main reactants of catalytic partial oxidation of synthetic diesel (SD) are SD

and oxygen (O2). Details of these two reactant feeds are provided in the following

sections.

3.5.1 SD Composition (Feed#1)

Firstly, synthetic diesel fuel is not a commercial grade diesel. Commercial diesel

fuel consists of approximately 75 vol.% saturated hydrocarbons and 25 vol.% aromatic

hydrocarbons, with an average chemical formula ranging approximately from C10H20 to

C15H28 (U.S. Air force, 1989, and Parmar et. al., 2009). In our lab, we mixed various

compounds that are predominantly found in commercial diesel to prepare a synthetic

diesel (SD) mixture while maintaining the above-mentioned saturated and aromatic

hydrocarbon percentages. The individual components were obtained from Sigma-Aldrich

(Canada) and had purity higher than 99.0%. The average molecular formula, based on the

weighted average of these components, was C12.87H24.81, and the average molecular

weight was 179.54 g/mol. Calculations for finding the average molecular formula,

weight, and density of SD can be found in Appendix A. Table 3.1 shows the physical

properties of fuels, as well as their percentage contribution, in the making of the SD.

31

Table 3.1 Physical properties of fuels and their percentage composition in SD

Chemical Compound Volm

Frac

Mass

Frac

Density

(g/mL) at

250C (lit.)

Mol. Wt.

(g/mol)

Purity

(%)

Chemical

Formula

Hexadecane 0.500 0.482 0.773 226.44 99.0 C16H34

Dodecane 0.250 0.234 0.750 170.34 99.0 C12H26

Decahydro-napthelene 0.050 0.056 0.896 138.25 99.0 C10H18

Butyl cyclohexane 0.050 0.051 0.818 140.27 99.0 C10H20

1,2,3,4-

Tetrahydronaphthalene

0.050 0.061 0.973 132.21 99.0 C10H12

Butyle Benzene 0.050 0.054 0.860 134.22 99.0 C10H14

1-methyl naphthalene 0.050 0.062 1.001 142.20 95.0 C11H10

Total 1.000 1.000 0.8078 179.54 C12.87H24.81

32

3.5.2 Air Composition (Feed#2)

Air was introduced in the reaction as a source of oxygen, which is the other

reactant besides SD. A compressed air cylinder was obtained from Praxair (Canada).

According to their MSDS, compressed air is a mixture of 19.5-23.5% oxygen with

balance being nitrogen (Appendix P). For the sake of calculation, it is assumed to be 21%

oxygen and the balance nitrogen.

33

3.6 OPERATION CONDITIONS & VARIABLES

A total of thirty six (36) experiments were conducted to collect kinetic data for the

CPOX of SD. Besides these runs, an extended time on stream (TOS) stability study was

conducted for seventeen (17) hours under the best operating conditions. Isothermal

conditions were maintained for all the above experimental runs. A fixed catalyst particle

size (0.8 mm) was used for all runs. A fixed amount of α-Al2O3 (7.6 gm) of the same

particle size as the catalyst (0.8 mm) was used as a diluent for all the experimental runs.

Also, for all the experimental runs, the reactor was kept at atmospheric pressure.

In order to collect kinetic data from the experimental runs, three (3) parameters

were varied: temperature (T), feed ratio (O2/SD), and weight-time (W/FSD,0). According

to Parmar et al. (2009), carbon-free operation for partial oxidation is possible for the

entire temperature range of interest (600-1000°C) if O2/C ratio is 1.1 or greater. However,

at higher temperature, lower O2/C could allow carbon free operation. Parmar et al. (2009)

proposed a diagram showing the carbon formation zone after a thermodynamics analysis

of the diesel reforming process and thermodynamics equilibrium gas phase product

distribution. From the economics stand point, it is desirable to operate the reactor at

reasonably lower operating temperatures and close to the stoichiometric O2/C ratio

needed for partial oxidation (~0.5) in order to maximize the syngas (CO + H2) yields and

at the same time to avoid coking. From work on thermodynamics analysis of equilibrium

product distribution by Parmar et al.(2009), it is observed that at 1000°C and an O2/C of

1, a maximum of 20% CO2 can be produced along with ≥ 20% H2 and < 30% CO and

without any carbon formation. At a lower temperature than 1000°C and an O2/C of 1,

CO2 production will increase and H2 production will decrease. Accordingly, the

34

important factors that were considered in designing the boundary operating conditions of

the current study are (1) carbon formation free operation, (2) low temperature, (3) O2/SD

ratio close to partial oxidation regime, and (4) enhanced syngas (H2 + CO) yields. On the

whole, the 36 experimental runs were planned and successfully executed by taking the

three process variables into consideration, that is temperature (1123, 1173 and 1223K),

O2/SD feed ratio (6.7, 8.0,9.3, and 10.5) and W/FSD,0 (19,008, 28,512, 38,052, and 47,556

kgcatalyst*s/kmolSD). In the current work, the weight-time (W/FSD, 0) was varied by varying

the catalyst amount (mass) employed in a given run (W) while keeping the molar feed

rate of SD (FSD, 0) constant. The catalyst amount ‗W‘ was varied as 100, 150, 200, and

250 mg to obtain the respective W/FSD, 0 values of 19,008, 28,512, 38,052, and 47,556

kgcatalyst*s/kmolSD.

Fixed operating parameters:

a) Reactor Pressure, P = 1 atm

b) Initial Feed (SD) Flow Rate, FSD, 0 =4.5 ml/h =5.2577×10-9

kmol/s

c) Catalyst Particle Diameter, dp = 0.8 mm

d) Inert material used (α-Al2O3) = 7.6 g

Variable operating parameters:

1) Temperature = 1123, 1173 and 1223 K

2) O2/SD molar ratio = 6.7, 8.0,9.3, and 10.5

3) Weight-time (W/FSD,0) = 19,008, 28,512, 38,052, and 47,556

kgcatalyst*s/kmolSD

35

Table 3.2 Design of kinetic experiments

1123 K

W/FSD,0 (kgcatalyst*s/kmolSD) 19008 28512 38052 47556

O2/SD

8.0 Run 1 Run 2 Run 3 Run 4

9.3 Run 5 Run 6 Run 7 Run 8

10.5 Run 9 Run 10 Run 11 Run 12

1173 K


O2/SD

6.7 Run 13 Run 14 Run 15 Run 16

8.0 Run 17 Run 18 Run 19 Run 20

9.3 Run 21 Run 22 Run 23 Run 24

1223 K


O2/SD

6.7 Run 25 Run 26 Run 27 Run 28

8.0 Run 29 Run 30 Run 31 Run 32

9.3 Run 33 Run 34 Run 35 Run 36

36

3.7 EXPERIMENTAL SETUP AND PROCEDURE

3.7.1 Schematic of Experimental Setup

Figure 3.1: Schematic diagram of the experimental setup used for the POX reforming of

SD.

37

3.7.2 Description of Reaction Process

Experimental runs to collect intrinsic kinetic data for CPOX of SD were

performed in a tubular fixed-bed reactor, as shown in Fig. 3.1. The reactor (ID = 12.7 mm

and length ≈ 550mm), made of Inconel 625 high temperature alloy, was housed vertically

in an electric tubular furnace (Zesta Engineering Ltd.) and was controlled by a

programmable temperature controller (Zesta). All the gases were regulated through pre-

calibrated mass (gas) flow controllers with digital readout units (Aalborg Instruments).

SD was fed by means of a syringe pump (KD scientific). The catalyst bed temperature

was recorded and controlled by sliding a k-type thermocouple into the center of the

catalyst bed.

In a typical experiment, a known amount (100, 150, 200, and 250 mg) of 0.8 mm-

sized catalyst particles were mixed with 7.6 g of 0.8 mm-sized α-Al2O3 to obtain a

catalyst bed height of ~ 4.5 cm. Prior to each run, the catalyst was activated in situ by

reducing it at 700C for 2 hours using a gas mixture of 5 vol. % H2 in N2 (flow rate = 100

sccm). Air (avg. 21% O2 bal. N2) was fed at a calculated flow rate in order to obtain the

desired O2/SD ratio (6.7, 8.0, 9.3, and 10.5). The kinetic experiments were performed at

three different temperatures, namely 1123, 1173, and 1223K. After attaining steady state

at the designated temperature, the reformate gas samples were analyzed at regular

intervals of 20~30 min. The product reformate stream coming from the reactor was

passed through a water-cooled condenser and then passed through an ice-cold knockout

trap to separate permanent gases from condensate (mainly water). The composition of the

product gas was monitored by an online gas chromatograph equipped with a thermal

conductivity detector (GC-TCD) (HP 6890, Agilent Technologies) using the molecular

38

sieve and Haysep columns (Alltech Associates) and with helium as the carrier gas. Some

tests were repeated in order to check for reproducibility. The maximum error in the SD

conversion and H2 selectivity data was < 1%.

Absence of heat and mass transfer limitations (details can be found in section

4.4.2) was confirmed by theoretical calculation to guarantee intrinsic kinetic data

collection from the experiment, as was done previously by Ibrahim and Idem (2007). In

addition, in order to ensure plug flow conditions (i.e., the velocity and temperature profile

in the reactor is radially uniform) and absence of back mixing and channelling, certain

criteria reported by Froment and Bischoff (1990) and Rase (1987) in their work were

implemented. These criteria were followed in similar kinetics studies conducted by

Ibrahim and Idem (2007). These criteria are: (1) ratio of catalyst bed height to catalyst

particle size (L/dp) ≥ 50 and (2) ratio of catalyst bed diameter to particle size (d/dp) ≥ 10.

In the current work, both the above-mentioned criterion was met. L/dp and d/dp of 56.2

and 15.9, respectively, were used, thereby satisfying the requirements for plug flow

conditions.

39

3.8 EQUATIONS USED TO CALCULATE CONVERSION, SELECTIVITY

AND STABILITY

The overall reaction used for the development of the kinetic model for SD partial

oxidation reforming is given as:

C12.87H24.81 + 6.435 O2 12.87 CO + 12.41 H2; ----------------------------------------- (3.2)

In this work, the conversion of diesel is defined as the part, carbon based, of the

fuel converted in C1-2 species in reformate/product stream. Higher hydrocarbons (CxHy)

in reformate (x>2) stream were considered unconverted and it was assumed that liquid

product stream was only water and no hydrocarbon. Hence, synthetic diesel conversion

was calculated using the moles of species in gaseous reformate via

10087.12)N(

2)N(NN%)mol(X

inlet

SD

outlet

HC

outlet

CO

outlet

CO

SD622

; ----------------------------------------- (3.3)

The hydrogen selectivity (SH2) is defined as actual moles of hydrogen produced

divided by the theoretical expected moles of hydrogen in SD feed divided by conversion

and calculated according to

100

X2

81.24N

N%)mol(S

SD

inlet

SD

outlet

H

2H2

; --------------------------------------------------- (3.4)

The equation used for calculating reforming reaction stability is shown below:

hours 1st two of avg.SD

hours 1st two of avg.SDhours last two of avg.SD

)X(

)X()X(Stability

; -------------------------------------- (3.5)

40

CHAPTER 4

RESULTS & DISCUSSION

4.1 PRECIS

The first part of this chapter discusses the results of the catalyst characterization

studies, and the second part deals with the results of the extended time on stream (TOS)

stability study of the 5N/CZCaY catalyst used for CPOX of SD at optimal operating

conditions. The third part discusses the effects of reaction parameters (operating

variables) such as reaction temperature (T), feed ratio (O2/SD), and weight-time

(W/FSD,0) on the resultant catalytic activity of the chosen catalyst in order to obtain the

optimal operating conditions to maximize the hydrogen yield from SD by CPOX. The

last part discusses the kinetic modeling studies of CPOX of SD.

41

4.2 CATALYST CHARACTERIZATION RESULTS

4.2.1 N2 Physisorption

The surface area, pore volume, and average pore diameter measurements of the

ceria-based quaternary oxide support ‗CZCaY‘ and corresponding catalyst ‗5N/CZCaY‘

developed in this study are presented in Table 4.1. The support sample exhibited

reasonably high surface area >129 m2/g, average pore diameter >87 Å (mesopore), and

cumulative pore volume 0.38 cc/gcat. The development of higher surface area can be

attributed to the method of preparation adopted in the current work (Terribile et al.,

1998). Upon impregnation of a nominal 5 wt. % Ni over the surface of the support, the

surface area and cumulative pore volume decreased. This is a general phenomenon

observed in the case of supported catalysts when an active component is impregnated

over its surface. The observed decrease is mainly due to penetration of the dispersed

nickel oxide into the pores of the support.

The measurements of pore volume per unit surface area (PV/SA) can also be

found in Table 4.1. Sukonket et al. (2011) showed in their work that the supports and

corresponding catalysts possessing higher pore volume per surface area > 1.7 x 10-9

m

exhibit exceptionally good performance. In this regard, it is important to mention that the

current support and catalyst were prepared by analogous routes (Sukonket et al., 2011)

and exhibit a PV/SA > 2.6×10-9

m. The addition of Ca and Y in the CZ support improved

PV/SA considerably; for CZ, PV/SA was 1.5×10-9

, and for CZCaY, it is 2.89×10-9

, more

than CZCa and CZY. PV/SA ratio was used to compare characteristics among the

catalysts. PV/SA means pore volume of catalyst per unit surface area of that catalyst.

42

More pore volume will have more active sites in a catalyst, which will ensure better

performance meaning better conversion.

Table 4.1: N2 physisorption results

M

BET SA

(m2 g

-1)

Pore

Volume

(cc g-1

)

Avg. Pore

Diameter

(Å)

Pore Vol./

BET SA

(10-9

m)

Ref.

Ce0.6Zr0.4O2 201 0.30 41 1.50 Sukonket et. al.,2011

5%Ni/Ce0.6Zr0.4O2 184 0.20 41 1.10 Sukonket et. al.,2011

Ce1-(x+y)ZrxMyO2 Supports - CZM

Ca 127.5 0.28 60.3 2.17 Sengupta, 2011

Y 188.2 0.50 83.1 2.66 Sengupta, 2011

CaY 129.83 0.38 87.3 2.89 This work

5Ni/Ce1-(x+y)ZrxMyO2 Catalysts - 5Ni/CZM

Ca 103.7 0.19 57.7 1.83 Idem, 2011

Y 187 0.36 68 1.92 Idem, 2011

CaY 120.3 0.32 81.9 2.65 This work

4.2.2 H2 Chemisorption

H2 chemisorption technique was employed to estimate the metallic surface area

and metal dispersion of active component (Ni). Table 2 compares the Ni surface area and

Ni dispersion data of the various kinds of ceria-based binary, ternary, and quaternary

oxide catalysts prepared by analogous surfactant assisted route. As observed from Table

4.2, the Ni surface area of the 5N/CZCaY quaternary catalyst is greater than that of the

43

5N/CZCa or 5N/CZY ternary oxide catalysts. However, it is low compared to that of the

5N/CZ binary oxide catalyst. More Ni dispersion ensures that Ni as active sites are well

distributed and structured over the surface of catalyst instead of Ni agglomerate, which is

not catalytically active. Similarly, low Ni surface area also indicates the presence of Ni as

agglomerates. Agglomerates, which are lumps of active component (Ni, in this case)

particles, are more prone to coking.

Table 4.2: H2 chemisorption results

M

Ni Dispersion (%)

DNi

Ni Surface Area

(m2 g

-1 cat.)

SNi

Ref.

5%Ni/Ce0.6Zr0.4O2 7.4 2.5 Sukonket et. al.,2011

5Ni/Ce1-(x+y)ZrxMyO2 Catalysts - 5Ni/CZM

Ca 6.2 2.1 Idem, 2011

Y 5.8 1.9 Idem, 2011

CaY 6.6 2.2 This work

44

4.2.3 Temperature Programmed Reduction (TPR)

Representative TPR patterns of the CZCaY support and corresponding Ni-

supported catalyst are shown as a function of temperature in Figure 4.1. As observed, the

TPR profile of the pure support (blue line) exhibits two broad H2 consumption peaks in

the temperature range of 650 – 750C and 850 – 900C. These two peaks can be

attributed to the reduction of surface and bulk oxygen anions. The reduction profile

observed here is very comparable with that of the pristine ceria sample, which shows two

characteristic reduction regimes, surface shell reduction (485C) and bulk reduction

(850C) (Khan and Smirniotis, 2008). According to the literature, the TPR trace for ceria

is not controlled by the rate of diffusion of the oxygen vacancies; instead, a surface

reduction process and the difference of both thermodynamic and kinetic properties

existing in the mixed oxide micro crystals are critical factors that control this rate

(Giordano et al., 2000). The TPR profiles of NiO impregnated supports exhibit a low

temperature H2 uptake peak at ~ 360 – 480C (Tmax of NiO Ni) denoting the reduction

of ‗NiO‘ species to metallic ‗Ni‘ species (Saha, 2011). The shift in the TPR peaks of the

support component upon impregnation of Ni can be attributed to the interaction between

the ceria component and the nickel component, which lowers the reduction temperature.

45

Figure 4.1: TPR patterns of titled Ce0.5Zr0.33Ca0.085Y0.085O2 support and

5wt.%Ni/Ce0.5Zr0.33Ca0.085Y0.085O2 catalyst. [Original in color]

-5

0

5

10

15

20

25

30

0 100 200 300 400 500 600 700 800 900 1000

Sig

na

l, m

V

Temperature, C

Temperature VS Signal

CZCY- 1.25

5Ni-CZCY-1.25

46

4.3 EXTENDED TOS STABILITY STUDY

An extended TOS stability study was performed to demonstrate the high level of

stability of the chosen 5N/CZCaY catalyst. The experiment was done at 1173K with a

feed ratio of 9.3 (O2/SD) and with W/FSD,0 of 38052 kgcat*s/kmolSD. To maintain the

desired feed ratio, SD feed flow rate was fixed at 4.5 ml/h and air flow rate at 360cc/min

(≡ 21600 ml/h). To obtain the desired W/FSD, 0, 200 mg catalyst of 0.8 mm particle size

was used. Before the reaction, the catalyst was activated by reducing it at 700C in

flowing 5%H2/bal. N2. Within the period of operation, no deactivation was observed, and

the conversion and selectivity were observed to be steady and stable, as shown in Figure

4.2 (a). Reaction stability was also calculated using equation (3.5) and found to be 0.009,

which indicates that the reaction process was very stable for the experiment time period.

Product distribution over the 17 hours TOS operation was also found to be steady and

stable, as shown in Figure 4.2 (b). The experiment was intentionally shut down after 17

hours.

47

(a)

(b)

Figure 4.2: Extended TOS stability study over 5wt.%Ni/Ce0.5Zr0.33Ca0.085Y0.085O2

catalyst for CPOX of SD reforming reaction. [T=1173K; Feed Ratio: O2/SD=9.3; W/FSD,0

= 38052 kgcat*s/kmolSD]

0

10

20

30

40

50

60

70

80

90

100

0 2 4 6 8 10 12 14 16 18

(mol%

)

Time on Stream (hr)

SD conversion H2 selectivity H2 yield

0

5

10

15

20

25

30

35

40

45

50

55

60

0 2 4 6 8 10 12 14 16 18

Pro

du

ct D

istr

ibu

tion

(m

ol%

)

Time on Stream (hr)

H2 CO C2H6 CO2

48

4.4 KINETICS EXPERIMENT RESULTS

Experimental section 3.5 followed by Table 3.2 describe how the 36 kinetic

experiments were designed. From each experiment, I have collected average product

distribution of gaseous product outlet stream data (online GC data sheet). A simple

chemical engineering mole balance calculation around the reactor using this outlet GC

data and known inlet feed condition will result solving the process for each experiment

for SD conversion (XSD) and H2 selectivity (SH2). Now, for the readers to get an idea, a

representative online GC data sheet of an experiment that was conducted at 1173 K

temperature, 6.7 O2/SD ratio, and 19008 kgcat*s/kmolSD W/FSD,0 was presented in figure

(4.3), and a detail chemical engineering mole balance around the reactor for this

experiment was presented in Appendix R. However, a representative GC data sheet of

each of all the thirty six (36) experiments and one (1) extended TOS stability experiment

have been included in Appendix S.

Finally, the results obtained in terms of the SD conversion (XSD), calculated using

equation (3.3), and H2 selectivity (SH2), calculated using equation (3.4), at different

W/FSD,0 and at different reaction temperatures are tabulated in Table 4.3a-d, with each

table presenting the results at a fixed feed ratio (O2/SD); the results are also presented in

graphical form in Figure 4.4.

49

Figure 4.3: A representative online GC data sheet of a SD reforming reaction

experiment. [T=1173K, O2/SD=6.7, and W/FSD,0=19008 kgcat*s/kmolSD]

50

Table 4.3a: SD conversion and H2 selectivity with W/FSD, 0 and T at O2/SD molar ratio of 6.7

O2/SD = 6.7 (equivalent to O2/C = 0.52)

Run# T W/FSD,0 Results Output

* Results Input

**

XSD SH2 inlet

SDN outlet

H2N outlet

CON outlet

CO2N

outlet

HC 62N

(K) (kgcat*s/kmolSD) (mol%) (mol%) kmol/s kmol/s kmol/s kmol/s kmol/s

13 1173 19008 73.83 57.47 5.57E-09 2.932E-08 3.743E-08 1.474E-08 3.806E-10

14 1173 28512 81.10 62.84 5.57E-09 3.521E-08 4.528E-08 1.173E-08 5.651E-10

15 1173 38052 82.59 67.38 5.57E-09 3.845E-08 4.819E-08 1.002E-08 4.956E-10

16 1173 47556 83.52 67.23 5.57E-09 3.879E-08 4.912E-08 1.018E-08 2.867E-10

25 1223 19008 75.11 64.53 5.57E-09 3.349E-08 3.917E-08 1.443E-08 1.219E-10

26 1223 28512 83.59 72.33 5.57E-09 4.178E-08 4.964E-08 9.926E-09 1.800E-10

27 1223 38052 85.28 70.29 5.57E-09 4.142E-08 5.023E-08 1.070E-08 1.036E-10

28 1223 47556 83.62 70.86 5.57E-09 4.094E-08 4.745E-08 1.215E-08 1.748E-10

* XSD and SH2 was calculated by Equation (3.3) and (3.4) respectivily

** These inputs have been used in Equation (3.3) and (3.4)

51

Table 4.3b: SD conversion and H2 selectivity with W/FSD, 0 and T at O2/SD molar ratio of 8.0



* Results Input

**

XSD SH2 inlet

SDN outlet

H2N outlet

CON outlet

CO2N

outlet

HC 62N


1 1123 19008 73.81 38.60 5.57E-09 1.968E-08 2.407E-08 2.755E-08 6.435E-10

2 1123 28512 79.23 46.26 5.57E-09 2.533E-08 3.251E-08 2.217E-08 1.056E-09

3 1123 38052 83.58 49.42 5.57E-09 2.854E-08 3.568E-08 2.137E-08 1.434E-09

4 1123 47556 84.59 50.77 5.57E-09 2.967E-08 3.634E-08 2.156E-08 1.367E-09

17 1173 19008 78.01 47.64 5.57E-09 2.568E-08 3.235E-08 2.162E-08 9.751E-10

18 1173 28512 83.45 52.35 5.57E-09 3.018E-08 3.999E-08 1.853E-08 6.521E-10

19 1173 38052 84.47 55.64 5.57E-09 3.248E-08 4.181E-08 1.729E-08 7.266E-10

20 1173 47556 85.39 56.05 5.57E-09 3.307E-08 4.201E-08 1.819E-08 5.033E-10

29 1223 19008 84.31 53.01 5.57E-09 3.088E-08 4.216E-08 1.764E-08 3.211E-10

30 1223 28512 85.69 54.74 5.57E-09 3.241E-08 4.467E-08 1.649E-08 1.345E-10

31 1223 38052 86.24 58.82 5.57E-09 3.505E-08 4.202E-08 1.966E-08 7.004E-11

32 1223 47556 89.99 65.98 5.57E-09 4.103E-08 4.798E-08 1.642E-08 5.528E-11



52

Table 4.3c: SD conversion and H2 selectivity with W/FSD, 0 and T at O2/SD molar ratio of 9.3



* Results Input

**

XSD SH2 inlet

SDN outlet

H2N outlet

CON outlet

CO2N

outlet

HC 62N


5 1123 19008 81.85 40.23 5.57E-09 2.276E-08 2.681E-08 3.032E-08 7.700E-10

6 1123 28512 83.90 44.47 5.57E-09 2.578E-08 3.179E-08 2.680E-08 7.756E-10

7 1123 38052 86.78 44.47 5.57E-09 2.667E-08 3.225E-08 2.817E-08 8.910E-10

8 1123 47556 88.21 44.98 5.57E-09 2.741E-08 3.323E-08 2.839E-08 8.071E-10

21 1173 19008 86.98 42.12 5.57E-09 2.531E-08 3.425E-08 2.655E-08 7.735E-10

22 1173 28512 88.35 44.76 5.57E-09 2.732E-08 3.593E-08 2.602E-08 6.919E-10

23 1173 38052 88.20 46.35 5.57E-09 2.824E-08 3.688E-08 2.565E-08 3.500E-10

24 1173 47556 92.69 54.82 5.57E-09 3.511E-08 4.049E-08 2.534E-08 3.077E-10

33 1223 19008 88.15 47.12 5.57E-09 2.870E-08 3.802E-08 2.441E-08 3.773E-10

34 1223 28512 89.86 50.27 5.57E-09 3.121E-08 4.077E-08 2.328E-08 1.826E-10

35 1223 38052 91.04 52.03 5.57E-09 3.272E-08 4.226E-08 2.273E-08 1.344E-10

36 1223 47556 92.01 51.43 5.57E-09 3.270E-08 4.368E-08 2.202E-08 1.305E-10



53

Table 4.3d: SD conversion and H2 selectivity with W/FSD, 0 and T at O2/SD molar ratio of 10.5



* Results Input**

XSD

SH2 inlet

SDN outlet

H2N outlet

CON outlet

CO2N

outlet

HC 62N


9 1123 19008 87.21 39.62 5.57E-09 2.387E-08 2.908E-08 3.231E-08 5.666E-10

10 1123 28512 88.83 36.47 5.57E-09 2.239E-08 2.718E-08 3.542E-08 5.412E-10

11 1123 38052 90.71 38.71 5.57E-09 2.426E-08 2.927E-08 3.454E-08 6.074E-10

12 1123 47556 97.51 44.99 5.57E-09 3.031E-08 3.681E-08 3.230E-08 3.937E-10



54

Figure 4.4: Variation of SD conversion with weight-time (W/FSD, 0) at different

temperature and different O2/SD ratio.[Original in Color]

55

4.4.1 Effect of W/FSD, 0

The W/FSD,0 was varied by changing the mass of the catalyst while keeping the

molar flow rate of SD (FSD, 0) constant at 4.5 mL/h. Consequently, the higher the value of

W/FSD, 0, the more catalyst was used for that run. The effect of W/FSD, 0 on SD conversion

and H2 selectivity at different temperatures is shown in Table 4.3a-d. The effect is well

described by Figure 4.4, as it shows that for all cases, SD conversion increased with

increasing W/FSD, 0, as expected. However, if we observe Figure 4.4 closely, after W/FSD,0

= 30000kgcat*s/kmolSD, SD conversion was almost constant. This might be because the

reaction has reached its thermodynamics limit.

4.4.2 Effect of Feed Ratio (O2/SD)

The molar ratio of reactants, O2/SD (alternatively O2/C), was varied from 6.7 to

10.5 (O2/C=0.52 to 0.82) by changing the flow rate of air while keeping the SD flow rate

constant at 4.5 mL/h. The corresponding results are given in Figure 4.5. The figure shows

that the conversion of SD increases with an increase in the feed ratio, O2/SD.

Furthermore, the H2 and CO production decreased with the increase of feed molar ratio

(O2/SD) while CO2 production increased for all cases (i.e., all the temperatures

investigated). The above observations can be attributed to the presence of excess oxygen

over the stoichiometric value of O2/C = 0.5 needed for a partial oxidation regime

(O2/SD= 6.435) and leaning towards the combustion regime.

56

Figure 4.5: Effect of feed ratio on the SD conversion and product distribution at

W/FSD,0=38,052 kgcat*s/kmolSD and (a) T=1123K, (b) T=1173K, (c) T=1223K

(a)

(b)

(c)

57

4.4.3 Effect of Temperature

Overall, as shown in Figure 4.6b, at high O2/SD of 9.3, with the increase of

temperature, SD conversion stays almost the same and H2 production also does not vary

significantly, while the decrease in CO2 production is accompanied by a consequent

increase in CO. However, at low O2/SD, shown in Figure 4.6a, a small peak in product

distribution is observed. In general, H2 and CO were increasing with the increase of

temperature, and the trend was reversed for CO2.

58

Figure 4.6: Effect of reaction temperature on the SD conversion and product distribution

at W/FSD,0=38052 kgcat*s/kmolSD and (a) O2/SD=8.0, (b) O2/SD=9.3

(a)

(b)

59

4.5 KINETIC MODELING STUDY

4.5.1 Empirical Rate Model

The overall reaction used for the development of the kinetic model for CPOX of

SD reforming is given as:

C12.87H24.81 + 6.435 O2 12.87 CO + 12.41 H2; ----------------------------------------- (3.2)

;mol

kJ1192H0

rxn (The calculation is shown in Appendix H).

An empirical, reversible power law rate model can be written as:

p

D

o

C

n

B

m

A

)RT

E(

0A NNNNekr

; ------------------------------------------------------------------- (4.1)

Here, A = SD; B = O2; C = CO; D = H2;

Ar = rate of reaction with respect to SD, kmol/m3/s

k0 = pre exponential factor or collision factor

E = activation energy, J/mol

T = reaction temperature, K

R = molar gas constant, 8.314 J/mol/K

NA = molar flow rate of A, kmol/s

NB = molar flow rate of B, kmol/s

NC = molar flow rate of C, kmol/s

ND = molar flow rate of D, kmol/s

m = order of reaction with respect to A

n = order of reaction with respect to B

o = order of reaction with respect to C

p = order of reaction with respect to D

60

As CO and H2 were not part of the feed in the current work, the power law

equation (4.1) can be further simplified to:

n

B

m

A

)RT

E(

0A NNekr

; -------------------------------------------------------------------------- (4.2)

From this point, the simplified power law rate equation (eq 4.2) will be used to fit the

experimental data.

As experimental runs #9-12 with a feed ratio of O2/SD = 10.5 were attempted

only at a single temperature, 1123K (850⁰C), these runs were, therefore, excluded from

fitting into the simplified power law rate equation (4.2). For only this set of runs, a new

power law equation can be proposed as:

m

AA Nkr ;-------------------------------------------------------------------------------------- (4.3)

Here, k = reaction rate constant for the experimental runs performed at feed ratio = 10.5

61

4.5.2 Investigation of Heat and Mass Transfer Limitation

Kinetic data collection in any experiment can only be considered intrinsic in the

absence of heat and mass transport limitations. Since catalytic POX reactions at high

temperatures (above 1123K) are very fast and tend to be mass-transfer limited, it is very

important to determine to what extent, if at all, these transport resistances affect the rate

of reaction. Several correlations are available in the literature to determine the effects that

interparticle and intraparticle heat and mass transport limitations could have on the rate of

reaction. In this work, these effects were investigated at a temperature of 1123K, the

lowest temperature used for the reaction.

4.5.2.1 Heat Limitation

The internal pore heat transfer resistance was estimated using the Prater analysis,

adopted from Ibrahim and Idem (2007), is given by:

eff

rxnAcAseffmax,particle

H)CC(DT

; --------------------------------------------------------- (4.4)

where ΔTparticle,max is the upper limit of temperature variation between the pellet centre

and its surface, ΔHrxn is the heat of reaction, CAs and CAc are, respectively, the

concentrations at the pellet surface and centre (assumed, respectively, to be the same as

bulk concentration and zero, as suggested by Levenspiel, 1999), and Deff is the effective

mass diffusivity obtained from

pAB

eff

DD (Fogler, 1999), where DAB is the bulk

diffusivity of component A in B (i.e., SD in air), which in turn, is estimated using the

Brokaw equation (Perry and Green, 1997).

62

The value for DAB was found, at the temperature of 1123K, to be 1.521×10-6

m2/s.

The effective diffusivity Deff was estimated to be 9.67×10-8

m2/s. εp is the void fraction

(estimated as the ratio of the volume occupied by voids to the total bed volume = 0.5) and

calculated using the formula ])(d/d

)2(d/d1[073.038.0ε

2

p

2

p

p

(Geankoplis, 2003), where d

and dp are the internal diameter of the reactor and the diameter of the particle,

respectively. η is the tortuosity factor, taken as 8 (Ibrahim and Idem, 2007; Akpan et al.,

2007; Fogler, 1999). λeff is the effective thermal conductivity obtained using the

correlation, λeff /λ=5.5+0.05NRe (Walas, 1990) for packed bed tubular reactors. λ is the

molecular thermal conductivity calculated using the Wassiljewa correlation to be

3.4525×10-2

W/m/K (Perry and Green, 1997). The detailed calculation is shown in

Appendix G. The effective thermal conductivity, λeff, was found to be 1.908×10-4

kW/m/K, is shown in Appendix J. A value of 0.613K was obtained for ΔTparticle,max. The

detailed calculation of internal pore heat transfer resistance, ΔTparticle,max, is shown in

Appendix J.

The heat transfer limitation across the gas film was determined using the

following correlation adopted from Ibrahim and Idem, 2007:

h

H)r(LT

rxnobs,Ac

max,film

; ---------------------------------------------------------------- (4.5)

where ΔTfilm,max is the upper limit of the temperature difference between the gas bulk and

the pellet surface, Lc is the characteristic length, rA,obs is the observed rate of reaction, h is

the heat transfer coefficient (estimated from the correlation 3/2

Pr

p

DH NuC

hJJ

,

63

where JH is the heat transfer J factor,

p

Pr

CN , and λ is the molecular thermal

conductivity). The detailed calculation is shown in Appendix I. The JD factor is given by

the following correlations: 4069.0

Re

p

D N4548.0

J

(Geankoplis, 2003);

)1(

udN

p

p

Re

. kc

is the mass transfer coefficient obtained as, 1.051×10-2

m/s. The detailed calculation is

shown in Appendix E. The heat transfer coefficient, h, was determined to be 1.635×10-2

kJ/m2/s/K. A value of 1.344K was obtained for ΔTfilm,max. The calculation of external film

heat transfer resistance, ΔTfilm,max, is shown in Appendix K.

Additionally, a more rigorous criterion for determining the onset of the heat

transport limitation during reaction, which was developed by Mears (1971), was also

used to further ascertain the insignificance of heat transfer resistance in the rate of

reaction:

15.0RhT

HERr2

rxncbobs,A

;---------------------------------------------------------------------- (4.6)

In substituting the numerical values for the terms on the left hand side (LHS) of Eq. (4.6),

a value of 6.152×10-3

is obtained, which is much less than 0.15. Hence, the heat transport

limitation did not occur. The detailed calculation of Mears‘ criteria is shown in Appendix

L.

4.5.2.2 Mass Limitation

The internal pore mass transfer resistance was calculated using the Weisz–Prater

criterion, adopted from Ibrahim and Idem, 2007, as given by:

64

Aseff

2

CCobs,A

ipd,wpCD

RrC

;-------------------------------------------------------------------------- (4.7)

Where, Cwp,ipd is the Weisz–Prater criterion for internal pore diffusion, ρC is the pellet

density, and Rc is the catalyst radius. The estimated value for Cwp,ipd was 0.444. This

value is much less than 1. Thus, this result indicates that the concentration of reactant on

the catalyst surface is more or less the same as the concentration within its pores.

According to Fogler (1999), this result is obtained as a consequence of the absence of

internal pore diffusion limitations. A detailed calculation is shown in Appendix M.

Wagnar modulus can also be calculated to see if there is any internal pore mass

transfer resistance by using the following equation, adopted from Lavenspiel, 1999:

e

obsAA2

WD

)C/r(LM

; ------------------------------------------------------------------------ (4.8)

If, MW < 0.15, then there is no internal diffusion resistance, and if MW > 0.15, then strong

pore diffusion exists.

Figure 4.7: Diffusion regime shown in terms of Wagnar modulus [Levenspiel, 1999]

Here, L = characteristic length; L=Rc/3, for spheres; Rc = radius of catalyst particle;

So, L = 1.333×10-4

m

65

All the parameters of the RHS of equation 4.8 were determined and Wagner modulus was

calculated to be 2.51×10-2

, which is much less than 0.15, which indicates that there is no

internal pore diffusion.

To determine whether film mass transfer resistance has any effect on the rate of

reaction, the ratio of observed rate to the rate if film resistance exists was examined, as

shown in Levenspiel, (1999). Eq. (4.9) illustrates this criterion:

6

d

kC

r

controls resistance film if Rate

rate Observed p

cAb

obs,A ;----------------------------------------------- (4.9)

The estimated value for the ratio in Eq. (4.9) was calculated to be 1.038×10-2

.

This result indicates that the observed rate is significantly less than the limiting film mass

transfer rate. Thus, the resistance to film mass transfer should not influence the rate of

reaction (Levenspiel, 1999). A detailed calculation for this is shown in Appendix N.

Mears‘ criterion (Fogler, 1999) is often considered a more rigorous criterion for

determining the onset of mass transport limitation in the film. Therefore, it was applied to

determine if there was any mass transfer limitation during the collection of the kinetic

data. This correlation is given as:

15.0Ck

nRr

Ac

cbobs,A

;-------------------------------------------------------------------------- (4.10)

The value of the LHS of the equation is 1.297×10-2

which is far less than the RHS.

Therefore, it can be concluded that there was no mass transport limitation in the film. A

detailed calculation is shown in Appendix O.

66

4.5.3 Experimental Rates of Reaction

To determine the rate of reaction of CPOX of SD experimentally with an integral

(plug flow) reactor, a differential method of analysis was used. Details of this method of

analysis can be found in Levenspiel, 1999, in chapters 3 and 18. For the integral (plug

flow) reactor, the differential analysis gave the following equation:

)F/W(d

dX

)F/W(d

dXr

0,SD

SD

0,A

AA ;----------------------------------------------------------- (4.11)

From here, we can see that the slope of XSD vs. W/FSD,0 will give the experimental

reaction rate (mol/kgcat/s) at a W/FSD,0. XSD vs. W/FSD,0 graphs were drawn for all the

temperatures and feed ratios and are given in Figure 4.3. The experimental rates of the

reaction were obtained from Figure 4.3 as the derivatives of the SD conversion vs.

W/FSD,0 curves, as given in Tables 4.4 and 4.5. It should be noted that the curves were

been generated using data analysis and graphing software Origin Pro 8.

Here, NA, NB, NC, and ND were calculated using the following equations:

NA=NA,0 - XANA,0;----------------------------------------------------------------------------- (4.12)

NB=NB,0 - 6.435 × (XANA,0);------------------------------------------------------------------ (4.13)

NC=12.87 × (XANA,0);------------------------------------------------------------------------- (4.14)

ND=12.41 × (XANA,0);------------------------------------------------------------------------- (4.15)

NA,0 = initial molar flow rate of SD (A), kmol/s; NB,0 = initial molar flow rate of O2 (B),

kmol/s; XA = SD conversion, calculated using equation (3.3) from the experimental data.

67

Table 4.4: Experimental kinetic data table 1

Run T rA × 10⁸ NA × 10⁸ NB × 10⁸ NC × 10⁸ ND × 10⁸ I (O₂/SD)

K kmol/(kgcat*s) kmol/s kmol/s kmol/s kmol/s

1 1123 788.20 0.147 1.886 5.132 5.151 8.0

2 1123 504.69 0.116 1.674 5.556 5.577 8.0

3 1123 367.81 0.094 1.528 5.848 5.870 8.0

4 1123 288.21 0.078 1.418 6.068 6.091 8.0

5 1123 386.18 0.094 2.239 5.853 5.874 9.3

6 1123 247.28 0.078 2.135 6.061 6.083 9.3

7 1123 180.21 0.068 2.063 6.204 6.226 9.3

8 1123 141.21 0.060 2.009 6.311 6.335 9.3

13 1173 558.97 0.129 1.050 5.381 5.401 6.7

14 1173 357.92 0.106 0.899 5.682 5.703 6.7

15 1173 260.85 0.091 0.796 5.889 5.911 6.7

16 1173 204.40 0.080 0.718 6.045 6.067 6.7

17 1173 443.29 0.109 1.631 5.642 5.663 8.0

18 1173 283.84 0.092 1.512 5.881 5.903 8.0

19 1173 206.86 0.080 1.430 6.045 6.067 8.0

20 1173 162.09 0.071 1.368 6.169 6.192 8.0

21 1173 246.06 0.068 2.062 6.207 6.230 9.3

22 1173 157.55 0.058 1.995 6.339 6.363 9.3

23 1173 114.82 0.051 1.950 6.430 6.454 9.3

24 1173 89.97 0.046 1.915 6.499 6.523 9.3

25 1223 488.50 0.120 0.993 5.494 5.514 6.7

26 1223 312.80 0.101 0.862 5.756 5.778 6.7

27 1223 227.96 0.088 0.772 5.937 5.959 6.7

28 1223 178.63 0.078 0.703 6.074 6.096 6.7

29 1223 302.65 0.083 1.452 6.001 6.023 8.0

30 1223 193.79 0.071 1.371 6.163 6.186 8.0

31 1223 141.23 0.063 1.315 6.275 6.299 8.0

32 1223 110.67 0.056 1.273 6.360 6.383 8.0

33 1223 217.46 0.060 2.007 6.317 6.340 9.3

34 1223 139.24 0.051 1.948 6.434 6.457 9.3

35 1223 101.48 0.045 1.908 6.514 6.538 9.3

36 1223 79.52 0.041 1.878 6.575 6.599 9.3

68

Table 4.5: Experimental kinetic data table 2

Run T rA × 10⁸ NA × 10⁸ NB × 10⁸ NC × 10⁸ ND × 10⁸ I (O₂/SD)

K kmol/(kgcat*s) kmol/s kmol/s kmol/s kmol/s

9 1123 589.52 0.072 2.806 6.143 6.165 10.5

10 1123 377.48 0.049 2.647 6.460 6.484 10.5

11 1123 275.10 0.033 2.538 6.678 6.703 10.5

12 1123 215.57 0.021 2.456 6.843 6.868 10.5

69

4.5.4 Estimation of Parameters of Rate Model

Estimation of the power law model parameters was based on the minimization of

the sum of the residual squares of the reaction rates by the Gauss-Newton and Levenberg-

Marquardt algorithms using non-linear regression software (NLREG). The values

obtained for the parameters presented are in Table 4.6. The NLREG code is given in

Appendix Q.

The validation of the models was based on the determination of percentage

average absolute deviation (AAD%) between the predicted rate using the proposed

kinetic model and experimentally obtained rate. According to the calculations, both the

models had an AAD <15%, which is acceptable. The formula for calculating AAD% is:

%100rate alExperiment

rate Predicted rate alExperiment%AAD

;----------------------------------- (4.16)

Table 4.6: Estimation of the values of the parameters of the models

Parameter PLM 1 PLM 2 (only O2/SD=10.5

runs)

k0 or k 3.22×1015

6.02×102

E (J/mol) 1.6×104

Not applicable

m 1.89 0.88

n 0.41 Not applicable

AAD (%) 7.9 4.9

Besides calculating AAD%, a parity plot (Figure 4.7 for power law model 1 and

Figure 4.8 for power law model 2) of experimental rate vs. predicted rate was also drawn

to depict how well the models fit the experimental data.

70

Figure 4.8: Parity plot of predicted rate vs. experimental rate for PLM 1

Figure 4.9: Parity plot of predicted rate vs. experimental rate for PLM 2

0

3E-06

6E-06

9E-06

0 3E-06 6E-06 9E-06

0

3E-06

6E-06

9E-06

0 3E-06 6E-06 9E-06

Pre

dic

ted

ra

te, k

mol/

(kg

ca

t*s)

Experimental rate, kmol/(kgcat*s)

Pre

dic

ted

ra

te, k

mol/

(kg

ca

t*s)

Experimental rate, kmol/(kgcat*s)

71

From the parity plot, it can be concluded that the experiment rates and model

predicted rates matched extremely well, and the models are acceptable. Finally, power

law model 1 for the temperature range of 1123-1223 K and feed ratio (O2/SD) in the

range of 6.7-9.3 (except 6.7 O2/SD for 1123 K) can be written as:

41.0

B

89.1

A

RT

16000

15

A NNe1022.3r

;------------------------------------------------------------ (4.17)

Power law model 2 for temperature 1123 K and O2/SD of 10.5 can be written as:

88.0

AA N602r ;---------------------------------------------------------------------------------- (4.18)

72

CHAPTER 5

CONCLUSIONS & RECOMMANDATIONS

The following conclusions can be drawn from the work presented in this thesis:

Ni-based catalyst (5N/CZCaY) performed well for the CPOX of SD under a wide

range of operating conditions. Interestingly, no catalyst deactivation was

observed.

A realistic synthetic diesel feed was used in the investigation. The SD was made

by mixing seven different paraffinic, olefinic, and aromatic hydrocarbons. The

average chemical composition of the SD was C12.87H24.81, where the H/C ratio was

< 2 (1.93).

The uniqueness of the current work is that a conventional (standard) reactor setup

and typical liquid delivery system were used, as opposed to the use of highly

sophisticated expensive spray nozzles, ignition chambers, vapourization

chambers, or complex reactor assemblies used by other research groups.

To the best of our knowledge, the current study is the first of its kind on CPOX

reforming of SD.

Equations 4.16 and 4.17 are acceptable from AAD% calculations and parity plot

fits. The activation energy of CPOX of SD over 5N/CZCY was found to be

16kJ/mol, and the reaction order with respect to SD was 1.89 and 0.41 with

respect to oxygen.

73

Based on the scope of the current work, the following recommendations can be

made for future research:

Introduce sulfur into the SD mixture and eventually use pump grade diesel to

evaluate the performance of the 5N/CZCaY catalyst.

In depth characterization of the fresh and used catalysts should be conducted in

order to understand the catalyst further.

A long-term stability and performance study of 50-400 h should be conducted.

The reaction mechanism of CPOX of SD should be determined, based on a

mechanistic approach of LHHW and ER formulation, and a rate expression

should be generated to support the mechanism.

Finally, a reactor modeling study should be performed on the system to design a

suitable reactor for the current application.

74

REFERENCES

Ahmed, S., & Krumpelt, M. (2001). Hydrogen from hydrocarbon fuels for fuel cells.

International Journal of Hydrogen Energy 26(4), 291–301.

Ahmed, S., Krumpelt, M., Kumar, R., Lee, S. H. D., Carter, J. D., Wilkenhoener, R., &

Marshall, C. (1998). Catalytic partial oxidation reforming of hydrocarbon

fuels, in: Presented at Fuel Cell Seminar, Palm Springs, CA.

Akpan, E., Sun, Y., Kumar, P., Ibrahim, H., Aboudheir, A., & Idem, R. (2007). Kinetics,

experimental and reactor modeling studies of the carbon dioxide reforming of

methane (CDRM) over a new Ni/CeO2–ZrO2 catalyst in a packed bed tubular

reactor. Chemical Engineering Science 62, 4012–4024.

Akande, A., Aboudheir, A., Idem, R., & Dalai, A. (2006). Kinetic modeling of hydrogen

production by the catalytic reforming of crude ethanol over a co-precipitated

Ni–Al2O3 catalyst in a packed bed tubular reactor. International Journal of

Hydrogen Energy 31, 1707–1715.

Alvarez-Galvan, M. C., Navarro, R. M., Rosa, F., Briceno, Y., Ridao, M. A., & Fierro, J.

L. G. (2008). Hydrogen production for fuel cell by oxidative reforming of

diesel surrogate: Influence of ceria and/or lanthana over the activity of

Pt/Al2O3 catalysts. Fuel 87, 2502-2511.

Armor, J. N. (2010). A History of Industrial Catalysis. Catalysis Today, In Press.

Ayabe, S., Omoto, H., Utaka,T., Kikuchi, R., Sasaki, K., Teraoka, Y., & Eguchi, E.

(2003). Catalytic autothermal reforming of methane and propane over

supported metal catalysts. Applied Catalysis A: General 241, 261-269.

Balat, M. (2008). Potential importance of hydrogen as a future solution to environment

and transportation problems. International Journal of Hydrogen Energy, 33,

4013–4029.

Ball, M., & Wietschel, M. (2009). The future of hydrogen-opportunities and challenges.

International Journal of Hydrogen Energy 34 (2), 615-627.

Boon, J., Dijk, E. v., Munck, S. d., & Brink, R. v. d. (2011). Steam reforming of

commercial ultra-low sulphur diesel. Journal of Power Sources, in press.

Cheekatamarla, P. K., & Lane, A. M. (2005a). Catalytic autothermal reforming of diesel

fuel for hydrogen generation in fuel cells I. Activity tests and sulfur poisoning.

Journal of Power Sources 152, 256-263.

75

Cheekatamarla, P. K., & Lane, A. M. (2005b). Efficient bimetallic catalysts for hydrogen

generation from diesel fuel. International Journal of Hydrogen Energy. 30

(11), 1277-1285.

Cheekatamarla, P. K., & Thomson, W. J. (2005). Poisoning effect of thiophene on the

catalytic activity of molybdenum carbide during tri-methyl pentane reforming

for hydrogen generation. Applied Catalysis A: General. 287 (2), 176-182.

Cheekatamarla, P. K., & Lane, A. M. (2006). Catalytic autothermal reforming of diesel

fuel for hydrogen generation in fuel cells II. Catalyst poisoning and

characterization studies. Journal of Power Sources 154, 223-231.

Dincer, I. (2011). Green methods for hydrogen production. International Journal of

Hydrogen Energy, in press.

Elghawi, U., Theinnoi, K., Sitshebo, S., Tsolakis, A., Wyszynski, M. L., Xu, H. M., . . .

Mightouf, A. (2008). GC-MS determination of low hydrocarbon species (C1–

C6) from a diesel partial oxidation reformer. International Journal of

Hydrogen Energy 33, 7074-7083.

Felder, R. M., & Rousseau, R. W. (1999). Elementary principles of chemical processes.

3rd

ed. Singapore: John Wiley & Sons, Inc.

Fogler, H. S. (1999). Elements of Chemical Reaction Engineering. 3rd

ed. Englewood

Cliffs, NJ: Prentice-Hall.

Froment, G. F., & Bischoff, K. B. (1990). Chemical Reactor Analysis and Design. 2nd

ed. New York: Wiley.

Gawade, P. V., Patel, D., Lipscomb, G. G., & Abraham, M. A. (2010). Kinetics and

modeling of the flexible fuel reformer: n-Hexadecane steam reforming and

combustion. Industrial & Engineering Chemistry Research 49, 6931-6940.

Geankoplis, C. H. (2003). Transport Processes and Separation Process Principles. 4th

ed. New York: Wiley.

Giordano, F., Trovarelli, A., Leitenburg, C. D., & Giona, M. (2000). A model for the

temperature-programmed reduction of low and high surface area ceria.

Journal of Catalysis, 193, 273-282.

Haynes, D. J., Berry, D. A., Shekhawat, D., Xiao, T. C., Green, M. L. H., & Spivey, J. J.

(2008). Partial Oxidation of n-Tetradecane over 1 wt % Pt/γ-Al2O3 and

Co0.4Mo0.6Cx Carbide Catalysts: A Comparative Study. Industrial &

Engineering Chemistry Research 47, 7663-7671.

Haynes, D. J., Campos, A., Berry, D. A., Shekhawat, D., Roy, A., & Spivey, J. J.

(2010a). Catalytic partial oxidation of a diesel surrogate fuel using an Ru-

substituted pyrochlore. Catalysis Today 155, 84-91.

http://www.sciencedirect.com/science/article/pii/S0360319905000595?_alid=1838700805&_rdoc=1&_fmt=high&_origin=search&_docanchor=&_ct=3&_zone=rslt_list_item&md5=9cd16c3b63643cacf0d3e269d33c2339



http://www.sciencedirect.com/science/article/pii/S0926860X0500205X?_alid=1838699919&_rdoc=1&_fmt=high&_origin=search&_docanchor=&_ct=1&_zone=rslt_list_item&md5=60f1253cf1e3a886e9ee638b75aa2dae




76

Haynes, D. J., Campos, A., Smith, M. W., Berry, D. A., Shekhawat, D., & Spivey, J. J.

(2010b). Reducing the deactivation of Ni-metal during the catalytic partial

oxidation of a surrogate diesel fuel mixture. Catalysis Today 154, 210-216.

Holladay, J. D., Hu, J., King, D. L., & Wang, Y. (2009). An overview of hydrogen

production technologies. Catalysis Today, 139, 244–260.

Idem, R. O., Khan, A., Ibrahim, H., Tontiwachwuthikul, P., Kumar, P., & Srisuwan, T.

(2010). Catalyst for feedstock and process flexible hydrogen production. US

Provisional Patent Application, 61/311, 055, (File # 1001).

Ibrahim, H., & Idem, R. (2006). Kinetic studies of the partial oxidation of isooctane for

hydrogen production over a nickel-alumina catalyst. Chemical Engineering

Science 61 (17), 5912-5918.

Ibrahim, H., & Idem, R. (2007). Kinetic studies of the partial oxidation of gasoline

(POXG) over a Ni–CeO2 catalyst in a fixed bed flow reactor. Chemical

Engineering Science 62, 6582–6594.

Ibrahim, H., & Idem, R. (2008). Single and mixed oxide-supported nickel catalysts for

the catalytic partial oxidation reforming of gasoline. Energy & Fuels 22, 878–

891.

Iglesia, E., & Boudart, M. (1983). Decomposition of formic acid on copper, nickel, and

copper-nickel alloys. Journal of Catalysis, 81, 204-213.

Kang, I., & Bae, J. (2006). Autothermal reforming study of diesel for fuel cell

application. Journal of Power Sources 159 (2), 1283–1290.

Karatzas, X., Creaser, D., Grant, A., Dawody, J., & Pettersson, L. J. (2010). Hydrogen

generation from n-tetradecane, low-sulfur and Fischer-Tropsch diesel over Rh

supported on alumina doped with ceria/lanthana. Catalysis Today. In press.

Khan, A., & Smirniotis, P. G. (2008). Relationship between temperature-programmed

reduction profile and activity of modified ferrite-based catalysts for WGS

reaction Journal of Molecular Catalysis A: Chemical, 280, 43-51.

Kim, H., Yang, J., & Jung, H. (2011). Partial oxidation of n-hexadecane into synthesis

gas over a Pd-based metal monolith catalyst for an auxiliary power unit (APU)

system of SOFC. Applied catalysis B: Environmental 101, 348-354.

Krummenacher, J. J., West, K. N., & Schmidt, L. D. (2003). Catalytic partial oxidation of

higher hydrocarbons at millisecond contact times: decane, hexadecane, and

diesel fuel. Journal of Catalysis 215, 332-343.

Levenspiel, O. (1999), Chemical Reaction Engineering. 3rd

ed.. New York: Wiley.

Lindermeir, A., Kah, S., Kavurucu, S., & Muhlner, M. (2007). On-board diesel fuel

processing for an SOFC-APU-Technical challenges for catalysis and reactor

design. Applied Catalysis B: Environmental 70, 488-497.

77

Mears, D.E. (1971). Tests for transport limitations in experimental catalytic reactors.

Industrial & Engineering Chemistry Process Design and Development 10,

541–547.

Mundschau, M. V., Burk, C. G., & Gribble Jr., D. A. (2008). Diesel fuel reforming using

catalytic membrane reactors. Catalysis Today 136, 190-205.

Moon, D. J., Sreekumar, K., Lee, S. D., Lee, B. G., & Kim, H. S. (2001). Studies on

gasoline fuel processor system for fuel-cell powered vehicles application.

Applied Catalysis A: General 215(1-2), 1-9.

Mota, N., Navarro, R. M., Alvarez-Galvan, M. C., Al-Zahrani, S. M., & Fierro, J. L. G.

(2010). Hydrogen production by reforming of diesel fuel over catalysts

derived from LaCo1-xRuxO3 perovskites: Effect of the partial substitution of

Co by Ru (x=0.01-0.1). Journal of Power Sources. In press.

NIST web chemistry website, (Oct 20, 2011): (http://webbook.nist.gov/chemistry/cas-

ser.html)

Pacheco, M., Sira, J., & Kopasz, J. (2003). Reaction kinetics and reactor modeling for

fuel processing of liquid hydrocarbons to produce hydrogen: isooctane

reforming. Applied Catalysis A: General 250, 161–175.

Parmar, R. D., Kundu, A., & Karan, K. (2009). Thermodynamic analysis of diesel

reforming process: Mapping of carbon formationa boundary and

representative independent reactions. Journal of Power Sources 194, 1007-

1020.

Patel, D., Gawade, P., Goud, S., Abraham, M. A., & Lipscomb, G. G. (2007). Steam

reforming and combustion kinetics of n-hexadecane for modeling and

simulation of a novel catalytic reformer. In: AIChE annual meeting 2007, Salt

Lake City.

Perry, R. H., & Green, D.W. (1997). Perry’s Chemical Engineer’s Handbook. 7th

ed.

New York: McGraw-Hill.

Pereira, C., Bae, J.M., Ahmed, S., & Krumplet, M. (2000). Liquid Fuel reformer

development: ATR of diesel fuel, in: Presented at the US Department of

Energy, Hydrogen Program technical review, San Ramon, CA.

Pike research website, (Oct 20, 2011): (http://www.pikeresearch.com/newsroom/fuel-

cell-vehicle-market-to-reach-16-9-billion-by-2020)

Pitz, W. J., & Mueller, C. J. (2011). Recent progress in the development of diesel

surrogate fuels. Progress in Energy and Combustion Science 37, 330-350.

Praharso, P., Adesina, A. A., Trimm, D. L., & Cant, N. W. (2004). Kinetic study of iso-

octane steam reforming over a nickel-based catalyst. Chemical Engineering

Journal 99, 131–136.

http://www.sciencedirect.com/science/article/pii/S0926860X01005269?_alid=1838702643&_rdoc=3&_fmt=high&_origin=search&_docanchor=&_ct=3&_zone=rslt_list_item&md5=61692f2262db2c6a5b7f865a6b00b12a

http://www.sciencedirect.com/science/article/pii/S0926860X01005269?_alid=1838702643&_rdoc=3&_fmt=high&_origin=search&_docanchor=&_ct=3&_zone=rslt_list_item&md5=61692f2262db2c6a5b7f865a6b00b12a

http://webbook.nist.gov/chemistry/cas-ser.html

http://webbook.nist.gov/chemistry/cas-ser.html

http://www.pikeresearch.com/newsroom/fuel-cell-vehicle-market-to-reach-16-9-billion-by-2020

http://www.pikeresearch.com/newsroom/fuel-cell-vehicle-market-to-reach-16-9-billion-by-2020

78

Rase, H. F. (1987). Chemical Reactor Design for Process Plants. New York: Wiley.

Saha, B. (2011). Sulfur tolerant catalyst for hydrogen production by the reforming of

biogas. (MASc. Dissertation). University of Regina, SK, Canada.

Sawyer, J. (1995). The oxidation of volatile organic compounds on Pt/Al2O3 catalyst.

(PhD. Dissertation). University of Tulsa, OK.

Sengupta, P. (2011). Catalysts for hydrogen production by the steam reforming of

mixtures of oxygenated hydrocarbons. (MASc. Dissertation). University of

Regina, SK, Canada.

Shekhawat, D., Gardner, T. H., Berry, D. A., Salazar, M., Haynes, D. J., & Spivey, J. J.

(2006). Catalytic partial oxidation of n-tetradecane in the presence of sulfur or

polynuclear aromatics: Effects of support and metal. Applied Catalysis A:

General 311, 8–16.

Sukonket, T., Khan, A., Saha, B., Ibrahim, H., Tantayanon, S., Kumar, P., & Idem, R.

(2011). Influence of the catalyst preparation method, surfactant amount, and

steam on CO2 reforming of CH4 over 5Ni/Ce0.6Zr0.4O2 catalysts. Energy &

Fuel, (ASAP).

Terribile, D., Trovarelli, A., Llorca, J., Leitenbrug, C., & Dolcetti, G. (1998). The

preparation of high surface area CeO2-ZrO2 mixed oxides by surfactant-

assisted approach. Catalysis Today, 43, 79-88.

Thormann, J., Maier, L., Pfeifer, P., Kunz, U. Deutschmann, O., & Schubert, K. (2009).

Steam reforming of hexadecane over a Rh/CeO2 catalyst in microchannels:

Experimental and numerical investigation. International Journal of Hydrogen

Energy 34, 5108-5120.

Tsay, M. T., & Chang, F.W. (2000). Characterization of rice husk ash-supported nickel

catalysts prepared by ion exchange. Applied Catalysis A: General, 203, 15-22.

U.S. Air force, (1989). The Installation Restoration Toxicology Guide. Volume (1-5).

Wright-Patterson Air Force Base, OH.

U. S. DOE, (2008). Website: (http://www.hydrogen.energy.gov/index.html). (Data

retrieved on Dec 12, 2010).

U. S. DOE, (2011). Website: (http://www.hydrogen.energy.gov/index.html). (Data

retrieved on Oct 20, 2011).

Walas, S. M. (1990). Chemical Process Equipment—Selection and Design. MA:

Butterworth-Heinemann.

Yoon, S., Kang, I., & Bae, J. (2008). Effects of ethylene on carbon formation in diesel

autothermal reforming. International Journal of Hydrogen Energy 33, 4780-

4788.

http://www.hydrogen.energy.gov/index.html

http://www.hydrogen.energy.gov/index.html

79

APPENDIX A

Calculation of molecular formula for synthetic diesel (SD)

Table 3.1 has been reproduced here:

Chemical Compound Volume

Fraction

Density

(g/mL) at

250C (lit.)

Mol. Wt.

(g/mol)

Chemical

Formula

Hexadecane 0.500 0.773 226.44 C16H34

Dodecane 0.250 0.750 170.34 C12H26

Decahydro-napthelene 0.050 0.896 138.25 C10H18

Butyl cyclohexane 0.050 0.818 140.27 C10H20

1,2,3,4-

Tetrahydronaphthalene

0.050 0.973 132.21 C10H12

Butyle Benzene 0.050 0.860 134.22 C10H14

1-methyl naphthalene 0.050 1.001 142.20 C11H10

Here, the density, molecular weight, and molecular formula of each compound have been

collected from NIST web chemistry websites.

Assume, SD Mixture = 1mL

So, the volume amount needed for each compound of SD = Volume fraction × 1mL

Calculation for Hexadecane:

Hexadecane Volume = 0.50 × 1mL = 0.50 mL

Amount of Hexadecane (mass, g) = Volume × Density = 0.50 × 0.773 = 0.3865 g

80

Similarly, the mass has been calculated for all six (6) other compounds of SD.

Chemical Compound Mass amount (g)

Hexadecane 0.3865

Dodecane 0.1875

Decahydro-napthelene 0.0448

Butyl cyclohexane 0.0409

1,2,3,4-Tetrahydronaphthalene 0.0487

Butyle Benzene 0.0430

1-methyl naphthalene 0.0501

For Hexadecane:

Mole amount = mass/molecular weight = 0.3865 g/226.44 g/mol = 1.707×10-3

mol

Similarly, mole amounts have been calculated for all other compounds in the SD.

Chemical Compound Mole amount (mol)

Hexadecane 1.707×10-3

Dodecane 1.101×10-3

Decahydro-napthelene 3.241×10-4

Butyl cyclohexane 2.916×10-4

1,2,3,4-Tetrahydronaphthalene 3.680×10-4

Butyle Benzene 3.204×10-4

1-methyl naphthalene 3.520×10-4

So, total mol of SD = (1.707×10-3

+ 1.101×10-3

+ 3.241×10-4

+ 2.916×10-4

+ 3.680×10-4

+

3.204×10-4

+3.520×10-4

) = 4.464×10-3

mol

Mol fraction of each compounds = mol of compound/total mol of SD

Chemical Compound Mole fraction

Hexadecane 0.3824

Dodecane 0.2466

Decahydro-napthelene 0.0726



Butyle Benzene 0.0718

1-methyl naphthalene 0.0789

81

If we assume mass is additive, then, the average molecular weight =

7n

1i

ii )xMW( ;

(Felder & Rousseau, 1999)

Here, xi= mol fraction of compound (i) in the mixture

So, Avg. molecular weight of SD = 179.54 g/mol

Now, assume the volume of the solution is proportional to the mass; then:

Density of liquid mixture of SD compounds =

7n

1i

iimix )x( ; (Felder & Rousseau,

1999)

Here, xi= mol fraction of compound (i) in the mixture.

The density of liquid mixture of SD compounds =ρmix= 0.8078 g/mL.

For Hexadecane:

Mol of C in hexadecane (HD) = mol of HD × mol of C in 1 mol of HD

From the molecular formula, mol of C in 1 mol of HD = 16

So, Mol of C in hexadecane (HD) = 1.707×10-3

mol HD × 16 mol of C in 1 mol of HD =

2.731×10-2

mol C

Mol of H in hexadecane (HD) = mol of HD × mol of H in 1 mol of HD

From the molecular formula, mol of C in 1 mol of HD = 34

So, Mol of H in hexadecane (HD) = 1.707×10-3

mol HD × 34 mol of H in 1 mol of HD =

5.804×10-2

mol H

Similarly, mol of C and H has been calculated for each compound and are given in the

table below:

Chemical Compound Mol of C (mol) Mol of H (mol)

Hexadecane 2.731×10-2

5.803×10-2

Dodecane 1.321×10-2

2.862×10-2

Decahydro-napthelene 3.241×10-3

5.833×10-3

Butyl cyclohexane 2.916×10-3

5.832×10-3

1,2,3,4-Tetrahydronaphthalene 3.680×10-3

4.416×10-3

Butyle Benzene 3.204×10-3

4.485×10-3

1-methyl naphthalene 3.872×10-3

3.520×10-3

Total 5.743×10-2

1.107×10-1

82

So, Total mol of C = 5.743×10-2

mol, and

Total mol of H = 1.107×10-1

mol

H/C ratio = 1.107×10-1

mol/5.743×10-2

mol = 1.928

So, the empirical formula of SD = CH1.928

To determine the molecular formula, we can write the following:

Molecular weight from empirical formula × n = Molecular weight of SD mixture

We have, then, calculated:

Molecular weight of SD mixture = 179.54 g/mol

So, Molecular weight from the empirical formula (CH1.9282) = 12.0107×1+1.00794×1.928

= 13.95 g/mol

Here, atomic weight of C = 12.0107 g/mol and atomic weight of H = 1.00794g/mol

So, 13.95 g/mol × n = 179.54 g/mol

n = 12.87

The molecular formula of SD can be written as = (CH1.928)n = (CH1.9282)12.87 = C12.87H24.81

83

APPENDIX B

Calculation of Diffusion coefficient of SD in Air (DAB) and Deff

Assumptions:

Dipole moment (i ) of SD =

A = 0.1 Debyes

Dipole moment ( i ) of Air = B = 0 Debyes

Here, A = SD and B = Air.

For a binary mixture of hydrocarbon, polar components in a low-pressure system, the

Brokaw equation has been referenced to calculate DAB in Perry and Green (1997).

The Borkaw equation is as follows:

D

2

AB

2/1

AB

2/3

ABP

MT001858.0D

; ------------------------------------------------------------------- (B.1)

Here, P = 1 atm; T = 850⁰C = 1123 K

Now, MA = 179.5 g/mol; MB = 29 g/mol

We know, mol/g10005.4295.179

295.179

MM

MMM 2

BA

BAAB

So, now, we need to calculate ζAB & ΩD. The correlations needed to calculate these two

parameters are given below, adopted from Perry and Green (1997):

2

BAAB

; -------------------------------------------------------------------------------- (B.2)

3/1

i,bi V18.1 ; So, 3/1

A,bA V18.1 and 3/1

B,bB V18.1

10.0575.1*909.4*

D T911.1T54.44 ; ------------------------------------------------------ (B.3)

i,bi,b

2

i3

iTV

1094.1

; So,

A,bA,b

2

A3

ATV

1094.1

and

B,bB,b

2

B3

BTV

1094.1

i,b

2

i

*

i T)3.11(18.1k

;

So, A,b

2

A

*

A T)3.11(18.1k

and B,b

2

B

*

B T)3.11(18.1k

84

2/1*

B

*

A

*

AB

*

AB

* kTT

We know, k = boltzman constant = 1.38×10-16

erg/K

For the case of SD,

Tb, A = boiling point (Tb) = 560 K

Vb,A = molar volume at boiling point (Vbi) can be calculate from the ideal gas law

mol

cm10595.4K560

atm1

mol*K

atm*cm05746.82

P

RTV

34

3

A,b

A,b

For the case of AIR,

Tb, B = boiling point (Tb) = 90 K

Vb,B = molar volume at boiling point (Vbi) can be calculate from the ideal gas law

mol

cm10385.7K90

atm1

mol*K

atm*cm05746.82

P

RTV

33

3

B,b

B,b

Calculation of ζAB:

26.42)10595.4(18.1V18.1 3/143/1

A,bA

98.22)10385.7(18.1V18.1 3/133/1

B,bB

62.322

98.2226.42

2

BAAB

Calculation of ΩD:

7

4

2

3

A,bA,b

2

A3

A 1054.756010595.4

1.01094.1

TV1094.1

0

9010385.7

01094.1

TV1094.1

3

2

3

B,bB,b

2

B3

B

8.660560))1054.7(3.11(18.1T)3.11(18.1k

27

A,b

2

A

*

A

85

14*

A 10119.9k8.660

2.10690))0(3.11(18.1T)3.11(18.1k

2

B,b

2

B

*

B

14*

B 10466.1k2.106

142/114142/1*

B

*

A

*

AB 10656.3)10466.110119.9(

239.410656.3

11231038.1kTT

14

16

*

AB

*

8647.0)239.4911.1239.454.44()T911.1T54.44( 10.0575.1909.410.0575.1*909.4*

D

Calculation of DAB:

s

m10521.1

s

cm10521.1

8647.0)62.32(1

)10005.4()1123(001858.0

P

MT001858.0D

26

22

2

2/122/3

D

2

AB

2/1

AB

2/3

AB

Calculation of Deff:

Deff = Effective mass diffusivity

pAB

eff

DD ; (Fogler, 1999)

εp = void fraction (the ratio of the volume occupied by voids to the total catalyst bed

volume)

εp calculated using the following formula (Geankopolis, 2003):

])(d/d

)2(d/d1[073.038.0ε

2

p

2

p

p

; ------------------------------------------------------------- (B.4)

Here, d = ID of reactor = 12.7 mm

dp = diameter of particle = 0.8 mm

8points twoosebetween th distanceshortest

points obetween tw travelsmolecule a distance actualtortuosity ; (Akpan et al.,

2007; Fogler, 1999; Ibrahim and Idem, 2007)

86

So,

5087.0])8.0/7.12(

)28.0/7.12(1[073.038.0]

)(d/d

)2(d/d1[073.038.0ε

2

2

2

p

2

p

p

s

m1067.9

8

5087.010521.1DD

28

6pAB

eff

87

APPENDIX C

Calculation of feed mixture (SD and Air) vapour density (ρmix)

Let, SD = A & Air = B

MA = 179.5 g/mol

MB = 29 g/mol

Mmix = ∑ yi * Mi

Here, yA = 0.0242 & yB = 0.9758

So, Mmix = yA * MA + yB * MB = 0.0242*179.5 + 0.9758*29 = 32.64 g/mol

P = 1 atm

T = 850⁰C = 1123 K

R = 8.205746E-5 m3.atm/K/mol

3

mixmix

m

kg3542.0

TR

MP

88

APPENDIX D

Calculation of feed mixture (SD & Air) vapour viscosity (µmix)

To calculate the vapour viscosity of the feed mixture of SD and air, first, we have to

calculate the vapour viscosity of SD (µA), then the vapour viscosity of air (µB), and then

the viscosity of their mixture (µmix).

Calculating µA is very challenging because SD itself is a mixture of seven (7) different

hydrocarbons. Here is the calculation for calculating µA, µB, and µmix.

Calculating µA:

The most accurate method (5 – 10% error) to predict vapour viscosity of pure

hydrocarbon is the Stiel and Thodos method as reported in Perry and Green (1997) at low

pressure. To use this method, only molecular weight, critical temperature, and critical

pressure are required. The equations used are as follows:

6/1

C

3/2

C

2/14

vT

PNM106.4 ; ------------------------------------------------------------------- (D.1)

Here, 5.1Tfor )67.1T58.4(0001778.0N r

625.0

r

The resultant viscosity is in centipoises (mPa.sec) if Tc and Pc are given in K and Pa,

respectively.

Now, we have modified the method by assuming that SD is a pure hydrocarbon (though

SD is a mixture of seven different hydrocarbons), and the molecular weight, and critical

temperature and pressure of SD have been calculated by the following formulas:

ni

1i

iiASD MyMM

ni

1i

i,ciPc TyT

ni

1i

i,ciPc PyP

89

Chemical Compound yi Mi yi*Mi Tc Pc yi*Tc yi*Pc

g/mol K Pa

Hexadecane 0.382 226.44 86.59 723.0 1.41E+06 276.47 5.39E+05

Dodecane 0.247 170.34 42.01 658.0 1.82E+06 162.27 4.49E+05

Decahydronapthalene 0.073 138.25 10.04 704.0 3.20E+06 51.11 2.32E+05

Butyl cyclohexane 0.065 140.27 9.16 650.0 2.56E+06 42.46 1.67E+05

1,2,3,4-

Tetrahydronaphthalene 0.082 132.21 10.90 720.0 3.70E+06 59.36 3.05E+05

Butyl Benzene 0.072 134.22 9.63 660.5 2.89E+06 47.41 2.07E+05

1-methyl nahthalene 0.079 142.20 11.21 772.0 3.60E+06 60.88 2.84E+05

Molecular weight of SD = MSD = MA = 179.5 g/mol

SD (mixture of 7 hydrocarbons) critical temperature = TPc = 699.95 K

SD (mixture of 7 hydrocarbons) critical pressure = PPc = 2.18×106 Pa

Here, Tr = T/TPc = 1123/699.95 = 1.6 > 1.5

So,

4625.0625.0

r 10264.5)67.16.158.4(0001778.0)67.1T58.4(0001778.0N

Modified vapour viscosity equation for SD:

s*m

kg10833.1s*mPa 0183286.0

95.699

1018.25.17910264.5106.4

T

PNM106.4

5

6/1

3/262/144

6/1

C

3/2

C

2/14

v

Calculating µB:

To calculate vapour viscosity of air at 850⁰C, we have used Sutherland‘s formula:

2/3

0

0T

T

b

a

; --------------------------------------------------------------------- (D.2)

Here, a = 0.555T0 + C and b = 0.555T + C

C = Sutherland constant

µ = vapour viscosity (in cP) at T (in ⁰R)

µ0 = vapour viscosity (in cP) at T0 (in ⁰R)

90

For air,

T0 = 524.07 ⁰R, µ0 = 0.01827 cP, and C = 120

T = 850 ⁰C =1123 K = 2021 ⁰R

a = 0.555T0 + C = 0.555×524.07 + 120 = 410.9

b = 0.555T + C = 0.555×2021 + 120 = 1242

So, air vapour viscosity at 850⁰C will be:

s*m

kg10579.4

cP 0457912.007.524

2021

1242

9.41001827.0

T

T

b

a

5

2/32/3

0

0

Calculating µmix:

To calculate the gaseous mixture of SD and air at low pressure, the method used by

Bromley and Wilke (error of only about 3% ) was used (Perry and Green, 1997):

n

1in

1j,1i i

j

ij

imix

y

yQ1

; ----------------------------------------------------------------- (D.3)

j

i

24/1

i

j

2/1

j

i

ij

M

M18

M

M1

Q

; [Note: Perry and Green, 1997, had a typing error in this

equation as published, so this is corrected from their original paper.]

For a SD-air mixture, the equations can be written as:

B

A

24/1

A

B

2/1

B

A

AB

M

M18

M

M1

Q

&

A

B

24/1

B

A

2/1

A

B

BA

M

M18

M

M1

Q

B

ABA

B

A

BAB

Amix

y

yQ1

y

yQ1

91

Here, the feed stream is at T =1123K, O2/SD=8.0, W/FA,0=19008kgcatalyst*s/kmolSD. From

these data, mol fraction of SD and air was calculated and found to be as follows:

yA= mol fraction of SD = 0.0242

yB=mol fraction of Air = 0.9758

2588.0

M

M18

M

M1

Q

B

A

24/1

A

B

2/1

B

A

AB

003.4

M

M18

M

M1

Q

A

B

24/1

B

A

2/1

A

B

BA

s*m

kg10326.4

y

yQ1

y

yQ1

5

B

ABA

B

A

BAB

Amix

92

APPENDIX E

Calculation of mass transfer coefficient (kc)

Mass transfer coefficient (kc or kg or k‘) has been calculated using the correlation for

packed bed for gas, as reported in Perry and Green, 1997, Table 5-27, situation [A].

3/1

Sc

49.0

ReSh )N()N(91.0N ; ------------------------------------------------------------- (E.1)

Here,

a

vN s

Re

and

ABG

GSc

DN

and

AB

sSh

D

dkN

Here,

vs = superficial velocity; needed to calculate

ρ = ρG = ρmix = 0.3542 kg/m3; calculated in Appendix C

µ = µG = µmix = 4.326×10-5

kg/m/s; calculated in Appendix D

Ψ = shape factor = 1; [for particle]

DAB =Diffusion co-efficient of SD in Air = 1.521×10-6

m2/s

ds = dp = catalyst particle diameter = 0.8 mm

a = effective interfacial mass transfer area per unit volume = 6(1-ε)/dp

Here, ε is defined as the void fraction available for gas flow, which is as εp.

So, ε = εp = 0.5087; calculated in Appendix A

And, dp is defined as the diameter of sphere, but we have catalyst particles in a packed

bed, so we needed to find an equivalent spherical diameter for a catalyst particle of same

volume, so that, it can used to find ―a‖. The equivalent spherical diameter of the particle

is denoted as ―dp*‖ and the calculation is shown below:

Volume of catalyst bed = bed

2

bed H2

dV

; d = diameter of reactor = 12.7 mm; Hbed

= height of catalyst bed = 45 mm

So, Vbed = 5.7×10-6

m3

Catalyst bulk density = 336

6

bed

catb

m

kg54.17

m107.5

kg10100

V

W

93

Catalyst particle density = 3

3

p

bc

m

kg49.34

5087.0

mkg54.17

Volume of catalyst particles = 36

3

6

c

catcat m109.2

mkg49.34

kg10100WV

So, now,

m10769.14

V32d

2

d

3

4V 2

3/1

cat*

p

3*

p

cat

So, 1

*

p

p

p

m6.166d

)1(6

d

)1(6a

Calculation of vs (superficial velocity):

Q = volumetric flow rate of feed stream (SD & Air) = 5.183×10-6

m3/s; [data taken from

experimental run where T =850⁰C, O2/C=0.62 and W/FA,0=5.0]

Ac = cross sectional area of catalyst bed = 1.267×10-4

m2

So, s/m10091.4A

Qv 2

c

s

011.2

m6.1661s*m

kg10326.4

m

kg3542.0

s

m10091.4

a

v

a

vN

15

3

2

mix

mixssRe

3.80

s

m10521.1

m

kg3542.0

s*m

kg10326.4

DN

26

3

5

ABG

GSc

53.53.80011.2191.0NN91.0N3/149.03/1

Sc

49.0

ReSh

We know that,

s

m10051.1

m108.0

s/m10521.153.5

d

DNk

D

dkN 2

3

26

s

ABSh

AB

sSh

94

APPENDIX F

Calculation of heat capacity (Cp) of feed stream (SD & Air) at T =1123K

First, heat capacity in ideal the gas state (Cp) at 1123 K for all the seven pure

hydrocarbon components was calculated using the following procedure:

For, hexadecane and dodecane, we used the following equation to calculate Cp, as

presented in Perry and Green (1997):

22

P)T/5Ccosh(

T/5C4C

)T/3Csinh(

T/3C2C1CC

; ----------------------------------- (F.1)

Data for C1 to C5 were collected from Perry and Green, 1997, Table 2-198.

At, T =1123 K,

Cp, Hexadecane=0.8884337 kJ/mol/K

Cp, Dodecane=0.6729175 kJ/mol/K

For decahydronapthalene, we could not use eqn (F.1) due the absence of C1 to C5

parameters. Instead, we have found Cp with different temperature data from the NIST

web chemistry website. So, we have fit those data to a third order polynomial to get an

equation for Cp as a function of temperature. The data are as follows:

T (K) Cp (J/mol/K) T (K) Cp (J/mol/K)

50 39.75 600 349.48

100 59.89 700 392.97

150 82.98 800 429.58

200 108.15 900 460.57

273.15 151.73 1000 486.93

298.15 168.1 1100 509.43

300 169.37 1200 528.68

400 236.42 1300 545.21

500 297.57

The third order polynomial equation fitted to these data is:

Cp,Decahydronapthalene = -2.42781+0.59628*T+3.36×10-5

*T2 - 1.33×10

-7* T

3; ------------ (F.2)

95

Like decahydronapthalene, the third order polynomial equation was fitted for 1,2,3,4-

tetrahydronaphthalene and 1-methyl naphthalene to get Cp=f(T).


50 39.51 600 293.63

100 55.34 700 325.91

150 75.22 800 352.66

200 98.28 900 375.08

273.15 136.97 1000 394

298.15 150.9 1100 410.07

300 151.98 1200 423.77

400 206.65 1300 435.51

500 254.31 Cp, 1,2,3,4-Tetrahydronapthalene = 0.08272+0.55495*T-9.75×10

-5*T

2 - 5.79×10

-8* T

3; ------- (F.3)


50 39.1 700 320.7

100 59.1 800 344.5

150 81.1 900 364.3

200 105.9 1000 381

273.15 145.6 1100 395

298.15 159.3 1200 407

300 160.4 1300 418

400 212.3 1400 426 500 256.2 1500 434 600 291.8

Cp, 1-methyl napthalene= -4.64709+0.65369*T-3.12×10-4

*T2 + 4.66×10

-8* T

3; -------------- (F.4)

For butyl cyclohexane and butyl benzene, the Harrison and Seaton equation was used to

calculate Cp at T = 1123 K.

The Harrison and Seaton equation, adopted from Perry and Green, (1997), is:

Cp = a1 + a2 C + a3 H; -------------------------------------------------------------------------- (F.5)

Here, C = number of carbon atoms in the compound, H = number of hydrogen atoms in

the compound, and a1, a2, a3 = constant parameter given in Table 2-387 of Perry and

Green (1997).

Cp was calculated at the designated temperature.

Then, a third order polynomial was fitted for butyl cyclohexane and butyl benzene as

follows:

Cp, Butyl cyclohexane= -46.20469+1.01532*T-5.858×10-4

*T2 + 1.277×10

-7* T

3; ----------- (F.6)

Cp, Butyl Benzene= -46.92655+0.89776*T-5.66807×10-4

*T2 + 1.35667×10

-7* T

3; -------- (F.7)

96

SD components Cp (kJ/mol/K)

Hexadecane 0.888434

Dodecane 0.672917

Decahydronapthalene 0.520994



Butyl Benzene 0.438588

1-methyl nahthalene 0.401524

The Cp of SD, a mixture of the seven components, was calculated using the following

equation adopted from Felder and Rousseau, 1999, equation 8.3-13.

components mixture all i,PimixP )T(Cy)T()C( ; --------------------------------------------------- (F.8)

So, K*mol

kJ676151.0CC)T()C( A,PSD,PmixP

Like hexadecane and dodecane, for air, we have used equation (F.1) to calculate Cp, as

presented in Perry and Green (1997).

Data for C1 to C5 were collected from Perry and Green, 1997, Table 2-198.

At T =1123 K,

Cp, Air=Cp,B =3.3521×10-2

kJ/mol/K

Again, we have used eqn (F.8) to calculate the Cp of the mixture of SD and air, that is, the

feed stream at 1123 K. Here, the mole fraction of SD is 0.0242 and air is 0.9758 for an

experimental run at O2/C =0.62 and W/FA,0=5.0.

Finally, Cp = (Cp)feed= 0.049073 kJ/mol/K

97

APPENDIX G

Calculation of thermal conductivity of SD (λA)

As the SD used in this study is a mixture of the seven different hydrocarbons in Table

3.1, we have used the Misic and Thodos equations, as presented in Perry and Green,

(1997), to calculate the molecular thermal conductivity of each of the pure components of

the SD using the correlations below. Then, we used the Wassiljewa correlation to

calculate the gas mixture of those seven components of SD to calculate the thermal

conductivity of SD.

Misic and Thodos equations to calculate molecular thermal conductivity (kG):

P3/2

r

7

G

C)14.5T52.14(10k ; For, Tr > 1; --------------------------------------- (G.1)

And, here,

3/2

c

2/16/1

cP

325.101MT

Where, kG = vapour thermal conductivity, W/m/K

Tr = reduced temperature, T/Tc

T = temperature, K=1123 K

Tc = critical temperature, K; [Data collected from Perry and green, 1997]

Cp = heat capacity at constant pressure and 1123K, J/kmol/K; [see Appendix F]

M = molecular weight

Pc = critical pressure, kPa; [Data collected from Perry and green, 1997]

Using these equations, the kG for each component is presented below:

Components of SD kG, W/m/K

Hexadecane 3.17879E-02

Dodecane 3.62989E-02

Decahydronapthalene 3.29202E-02

Butyl cyclohexane 4.92402E-02

1,2,3,4-Tetrahydronaphthalene 3.25314E-02

Butyl Benzene 4.77502E-02

1-methyl nahthalene 3.02438E-02

98

The Wassiljewa correlation (Perry and Green, 1997): low pressure (1 atm or less) gas

mixture

n

1i jj

iim

Ay

kyk ; --------------------------------------------------------------------------------- (G.1)

Where, km = mixture thermal conductivity,W/m/K = λA

n = number of components = 7

yi,j = mole fraction of component I or j in the vapour mixture

ki = thermal conductivity of pure component I at 1123 K temperature = kG

Aj = binary interaction parameter

Aj is obtained by the method of Lindsay and Bromley as reported in Perry and Green,

(1997):

i

ij

22/1

i

j

4/3

i

j

j

ij

ST

ST

ST

ST

M

M1

4

1A

Where, µi,j = vapour viscosity of pure component i or j at temperature T, Pa.s; [calculated

by the Stiel and Thodos equation as described in Perry and Green, 1997]

T = 1123 K

Sij=Sji=C(SiSj)1/2

Si,j=1.5Tbi,j

Tbi,j=normal boiling temperature of pure component i or , K; [data collected from

Perry and Green, 1997]

C=1.0

Mi,j=molecular weight of i or j

Using this Wassiljewa correlation, the vapour thermal conductivity at 1123 K of SD was

calculated and found to be 3.4525×10-2

W/m/K.

So, λA= 3.4525×10-2

W/m/K

99

APPENDIX H

Calculation of standard heat of reaction ( 0

rxnH )

To calculate 0

rxnH , we have to calculate 0

SD,fH first. We have collected the heat of

formation at standard temperature (25ºC) and standard pressure (1 atm) for all seven

components of SD from the NIST website. These are given in the following table:

SD components yi (kJ/mol) (kJ/mol)

Hexadecane 0.382 -374.9 -143.361

Dodecane 0.247 -288.1 -71.047 Decahydronapthalene 0.073 -169.2 -12.284 Butyl cyclohexane 0.065 -213.2 -13.927 1,2,3,4-Tetrahydronaphthalene 0.082 26.0 2.143 Butyl Benzene 0.072 -12.8 -0.919 1-methyl nahthalene 0.079 116.9 9.218

Then, we used the following formula to calculate the0

SD,fH .

mol

kJ2.230Hy)atm1,C25(H 0

i,fi

00

SD,f

Also, from NIST, we found that 0

CO,fH = -110.52 kJ/mol

The standard heat of reaction can be calculated from the standard heat of formation of the

product and reactant using the following equation adopted from Felder and Rousseau,

1999, equation 9.3-1:

tstanreac

0

i,fiproducts

0

i,fii

0

i,fi

0

rxn HHHH ; ------------------------------ (H.1)

The partial oxidation reaction of SD is:

2281.2487.12 H41.12CO87.12O775.6HC

So, according to eqn (H.1), we can write,

0

SD,f

0

O,f

0

H,f

0

CO,f

0

rxn HH775.6H41.12H87.12H22

2.2300775.6041.1252.11087.12H0

rxn

mol

kJ1192H0

rxn

100

APPENDIX I

Calculation of heat transfer coefficient (h)

Heat transfer coefficient is estimated from the following correlation (adopted from

Ibrahim and Idem, 2007):

3/2

Pr

P

DH NuC

hJJ

; ------------------------------------------------------------------------ (I.1)

where, JH=heat transfer J factor

Cp=heat capacity of feed stream at T(850ºC) = (Cp)feed=0.049073kJ/mol/K;

[calculated in Appendix F]

u =superficial velocity; [calculated in Appendix E]

ρ=vapour density of the reactant feed = ρmix; [calculated in Appendix C]

NPr=Prandlt number = CPμ/λ

µ=vapour viscosity of feed=µmix; [Calculated in Appendix D]

λ=thermal conductivity of SD =λA =3.4525×10-2

W/m/K; [calculated in Appendix

G]

Here, JD factor can be calculated by the following correlation (adopted from Ibrahim and

Idem, 2007):

4069.0

Re

p

D N4548.0

J

; ------------------------------------------------------------------------- (I.2)

Here, NRe is defined by the following equation:

p

p

Re1

udN

Here, dp= catalyst particle diameter = 0.8 mm

u=superficial velocity=4.091×10-2

m/s

ρ=vapour density of the reactant feed = ρmix=0.3542 kg/m3; [calculated in

Appendix C]

µ=vapour viscosity of feed=µmix=4.326×10-5

kg/m/s; [calculated in Appendix D]

εp=void fraction of catalyst bed=0.5087

101

So,

5455.01

udN

p

p

Re

Again, 144.1N4548.0

J 4069.0

Re

p

D

Now, to convert (Cp)feed from kJ/mol/K to kJ/kg/K, we can divide it by the molecular

weight of the feed mixture, which we have previously calculated to be 32.6 kg/kmol:

K*kg

kJ505.1

kmol

mol1000

kg6.32

kmol

K*mol

kJ049073.0C

feedP

And,

886.1

K*m

W104525.3

s*m

kg10326.4

K*kg

kJ505.1

CCN

2

5

A

mixmixPPPr

Now,

K*s*m

kJ10635.1uC

N

JhN

uC

hJJ

2

2

P3/2

Pr

H3/2

Pr

P

DH

102

APPENDIX J

Calculation of internal pore heat transfer resistance (ΔTparticle,max)

Calculate particle heat transfer:

eff


H)CC(DT

; --------------------------------------------------------- (J.1)

Here, CAc=0, and CAs=CAb=1.014 mol/m3

ΔHrxn=-1192 kJ/mol; [Calculated in Appendix H]

Deff=9.67×10-8

m2/s; [Calculated in Appendix A]

λeff= effective thermal conductivity; [needs to be calculated]

Here, λeff can be calculated using the correlation, λeff/λ=5.5+0.05NRe (Walas, 1990) for

packed bed tubular reactor.

Again, NRe was calculated in Appendix I, using the following correlation:

5455.0)1(

udN

p

p

Re

So,K*m

kW10908.1N05.05.5 4

effReeff

Now,

K613.0T

K*m

kW10908.1

mol

kJ1192

m

mol)0014.1(

s

m1067.9

H)CC(DT

max,particle

4

3

28

eff


103

APPENDIX K

Calculation of external film heat transfer resistance (ΔTfilm,max)

Calculate film heat transfer:

h

H)r(LT

rxnobs,A

max,film

; ------------------------------------------------------------------ (K.1)

Here, L = Rc/3 = 0.8mm/3 = 0.1333mm

ΔHrxn=-1192 kJ/mol; [calculated in Appendix H]

s*kg

kmol102.788r

catalyst

8

obs,A

; [Data taken from experimental run at T=1123 K,

O2/SD=8.0 and W/FA,0=19008 Kgcatalyst*s/kmolSD]

So,

s*m

mol1383.0rr

3bobs,Aobs,A

Here, Catalyst bulk density = 3b

m

kg54.17

K*s*m

kJ10635.1h

2

2 ; [calculated in Appendix I]

K 344.1T

K*s*m

kJ10635.1

mol

kJ1192

s*m

mol1383.0m101333.0

h

H)r(LT

max,film

2

2

3

3

rxnobs,A

max,film

104

APPENDIX L

Calculation of Mears’ criteria for heat transport limitation

Mears‘ criteria for heat transfer limitation:

15.0RhT

HERr2

rxncbobs,A

; --------------------------------------------------------------------- (L.1)

Here, s*kg

kmol102.788r

catalyst

8

obs,A



Catalyst bulk density = 3b

m

kg54.17

Rc=radius of catalyst particle =0.4mm

E=activation energy=1.6×104kJ/kmol

ΔHrxn=-1192 kJ/mol; [Calculated in Appendix H]

h=heat transfer coefficient=1.635×10-2

kJ/m2/s/K

T=850ºC=1123 K

R=molar gas constant=8.314 J/mol/K

15.010152.6

kJ

J1000

K*mol

J314.8K1123

K*s*m

kJ10635.1

mol

kJ1192

kmol

kJ106.1m104.0

m

kg54.17

s*kg

kmol102.788

RhT

HERr

3

2

2

2

43

3

catalyst

8

2

rxncbobs,A

105

APPENDIX M

Calculation of Weisz-Prater criterion for internal mass diffusion

The Weisz-Prater criterion for calculating the internal diffusion limitation is given as:

Aseff

2

ccobs,A

ipd,wpCD

RrC

; ------------------------------------------------------------------------

(M.1)

Here, s*kg

kmol102.788r

catalyst

8

obs,A



Catalyst particle density =3

3

p

bc

m

kg49.34

5087.0

m/kg54.17

; [Appendix E]

Rc = catalyst particle radius, m = 0.4×10-3

m

s

m1067.9D

28

eff

; [calculation is showed in Appendix B]

CAs =Concentration of A in the catalyst surface = CAb = Concentration of A in the bulk,

kg/m3 = 1.014 mol/m

3 [because, there is no external (film resistance, so concentration of

A in the bulk and concentration of A in the surface can be assumed same]

So,

1444.0C

kmol1

mol1000

m

mol014.1

s

m1067.9

104.0m

kg49.34

s*kg

kmol102.788

CD

RrC

ipd,wp

3

23

23

3

8

Aseff

2

ccobs,A

ipd,wp

After calculating ipd,wpC , I found that ipd,wpC < 1. So, we can claim that there are no

internal diffusion limitations and, consequently, no concentration gradient exists within

the pellet.

106

APPENDIX N

Calculation of external film diffusion limitation (Lavenspiel, 1999)

Recall equation 4.9:

6

d

kC

r


rate Observed p

cAb

obs,A ; ----------------------------------------------- (4.9)

Here, Catalyst bulk density = 3b

m

kg54.17

s*kg

kmol102.788r

catalyst

8

obs,A



So, s*m

mol1383.0rr

3bobs,Aobs,A

CAb = Concentration of A in the bulk = CA0 = 1.014 mol/m3

dp= catalyst particle diameter = 0.8 mm

kc = mass transfer coefficient = 1.051×10-2

m/s [calculation is shown in Appendix E]

2

3

2

3

3p

cAb

obs,A

10038.1controls resistance film if Rate

rate Observed

6

m108.0

s

m10051.1

m

mol014.1

s*m

mol1383.0

6

d

kC

r


rate Observed

The estimated value for the ratio in Eq. (3.10) is found to be much less than 1, so, then,

the observed rate will be significantly less than the limiting film mass transfer rate. Thus,

we can conclude that the resistance to film mass transfer should not influence the rate of

reaction for the reforming reaction in this study (Levenspiel, 1999).

107

APPENDIX O

Calculation of Mears’ criterion for external film diffusion limitation

Mears‘ criterion:

15.0Ck

nRr

Ac

cbobs,A

Here, s*kg

kmol102.788r

catalyst

8

obs,A



Catalyst bulk density = 3b

m

kg54.17

Rc = catalyst particle radius, m = 0.4×10-3

m

n = overall reaction order ≈ 2.5

kc = mass transfer coefficient = 1.051×10-2

m/s [calculation is shown in Appendix E]

CA = Concentration of A in the bulk = 1.014 mol/m

3

15.010297.1kmolmol

mol1000kmol10297.1

m

mol014.1

s

m10051.1

5.2m104.0m

kg54.17

skg

kmol102.788

Ck

nRr

25

3

2

3

3

8

Ac

cbobs,A

As, the LHS of Mears‘ criterion is less than 0.15, we can claim that there will be no

external (film) resistance to mass transfer during the kinetic study experiments in this

research.

108

APPENDIX P

MSDS of compressed air, PRAXAIR

A partial section is reprinted here to show the composition of Praxair compressed air:

109

APPENDIX Q

NLREG code with results for PLM 1 & 2

The two codes for the two power law models (PLM 1 & PLM 2) with results obtained

with NLREG software:

NLREG Code for Power law model 1:

110

Results of regression (Power law model 1):

111

NLREG Code for Power law model 2:

Results of regression (Power law model 2):

112

APPENDIX R

Mole balance of Run#13

Here, as an example, a detail chemical engineering mole balance around the reactor for

the experiment at 1173K temperature, 6.7 O2/SD ratio, and 19008 kgcat*s/kmolSD W/FSD,0

of Run#13 was presented to calculate the SD conversion (XSD) and H2 selectivity (SH2).

The process block diagram of the Run#1 can be drawn as:

Converting SD flow rate to molar SD flow rate:

From Appendix A, The density of liquid mixture of SD compounds =ρSD= 0.8 g/mL

And, the molecular weight of SD = MSD=179.5 g/mol

Molar flow rate of SD = (4.5 mL/h)*(0.8 g/mL)/(179.5 g/mol)

= 2.0251×10-2

mol/h

= 5.57×10-9

kmol/s

Converting Air flow rate to molar O2 flow rate and molar N2 flow rate:

From Ideal gas law, we know,

PV = nRT

Here,

P = pressure of system = 1 atm

V = Volume of the system = Air flow rate = 263 mL/min

R = molar gas constant = 0.08205746 L*atm/K/mol

Packed Bed

Tubular

Reactor

Reactant 1

SD = 4.5 mL/h

Reactant 2

Air = 263

mL/min

Product 1 (mol%)

13.173% H2

16.815% CO

6.623% CO2

0.171% C2H6

63.208% N2

Product 2

H2O

113

T = Air inlet temperature (room temperature) = 250C = 298 K

n = moles of Air

So,

s

kmol107925.1

mol1000

kmol

s60

min

mL1000

L

K298mol*K

atm*L08205746.0

min

mL263atm1

RT

PVn 7

Molar flow rate of Air = 1.7925×10-7

kmol/s

From Appendix P, we found that, our air cylinder from PRAXAIR contains 19.5-23.5%

O2 and balanced N2.

Let‘s assume, we have 21% O2 and 79% N2 in the Air.

So,

O2 molar flow rate =0.21 × 1.7925×10-7

kmol/s = 3.7643×10-8

kmol/s

N2 molar flow rate =0.79 × 1.7925×10-7

kmol/s = 1.4161×10-7

kmol/s

The block Diagram can be redrawn as:

Now, if we do mole balance around the reactor, we will get the molar flow rate of

product stream 1. The molar composition of the product stream was collected from the

online GC data sheet. Once we calculate the molar flow rate of product stream 1, using

these compositions, we can get the molar flow rate of each product components.

Product 1 (mol%)

13.17% H2

16.81% CO

6.62% CO2

0.17% C2H6

63.2% N2

0.01% O2

Reactant 2

O2 = 3.7643×10-8

kmol/s

N2 = 1.4161×10-7

kmol/s

SD = 5.57×10-9

kmol/s

Packed Bed

Tubular

Reactor

Reactant 1

114

N2 Mole balance around the reactor: (Basis = 1 s)

N2 in = 1.4161×10-7

kmol

N2 out = (Product1 flow rate) × 0.63208

Now, N2 in = N2 out

Product1 flow rate = (1.4161×10-7

)/0.63208

Product1 flow rate = 2.24×10-7

kmol

So,

H2 molar flow rate = (Product1 flow rate) × 0.131 = 2.93×10-8

kmol

CO molar flow rate = (Product1 flow rate) × 0.168 = 3.76×10-8

kmol

CO2 molar flow rate = (Product1 flow rate) × 0.066 = 1.47×10-8

kmol

C2H6 molar flow rate = (Product1 flow rate) × 0.0017 = 3.80×10-10

kmol

So, inlet

SDN = 5.57×10-9

kmol/s

outlet

H2N = 2.93×10

-8 kmol/s

outlet

CON = 3.76×10-8

kmol/s

outlet

CO2N = 1.47×10

-8 kmol/s

outlet

HC 62N = 3.80×10

-10 kmol/s

Calculating XSD:

Recall Equation (3.3),

10087.12)N(

2)N(NN%)mol(X

inlet

SD

outlet

HC

outlet

CO

outlet

CO

SD622

So,

%02.74%)mol(X

10087.12)1057.5(

2)1080.3(1047.11076.3%)mol(X

10087.12)N(

2)N(NN%)mol(X

SD

9

1088

SD

inlet

SD

outlet

HC

outlet

CO

outlet

CO

SD622

115

Calculating SH2:

Here, inlet

SDN = 5.57×10-9

kmol/s

outlet

H2N = 2.93×10

-8 kmol/s

%02.74XSD

So,

%3.57%)mol(S

100

74.02

81.241057.5

1093.2%)mol(S

100

X2

81.24N

N%)mol(S

2H

9

8

2H

SD

inlet

SD

outlet

H

2H2

116

APPENDIX S GC Datasheets of each experiment

Here, representative GC data sheets of each of all the thirty six (36) experiments and one

(1) extended TOS stability experiment have been included:

Run#1: T =1123K, O2/SD=8.0, W/FSD,0=19008kgcat*s/kmolSD

117


118


119


120


121


122


123


124


125


126


127


128


129


130


131


132


133


134


135


136


137


138


139


140


141


142


143


144


145


146


147


148


149


150


151


152

Extended TOS Stability Run: T =1173K, O2/SD=9.3, W/FSD,0=38052kgcat*s/kmolSD

KINETIC STUDY OF CATALYTIC PARTIAL OXIDATION OF …

Documents

Transcript of KINETIC STUDY OF CATALYTIC PARTIAL OXIDATION OF …