KINETIC STUDY OF CATALYTIC PARTIAL OXIDATION OF …
Transcript of 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
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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
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(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.
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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.
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DEDICATION
This work is dedicated to my mother, Sadia Nurin, and my father, Md. Abdus
Samad Khan (R.I.P).
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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
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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
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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
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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
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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
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LIST OF FIGURES
3.1 Schematic of the experimental setup for the CPOX reforming of SD
using a packed bed tubular reactor (PBTR)
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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]
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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]
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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
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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
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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
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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
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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
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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
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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
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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
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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
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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
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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
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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
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‗‗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
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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
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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).
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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
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(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.
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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.
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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
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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).
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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
W/FSD,0 (kgcatalyst*s/kmolSD) 19008 28512 38052 47556
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
W/FSD,0 (kgcatalyst*s/kmolSD) 19008 28512 38052 47556
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
O2/SD = 8.0 (equivalent to O2/C = 0.62)
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
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
* 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)
52
Table 4.3c: SD conversion and H2 selectivity with W/FSD, 0 and T at O2/SD molar ratio of 9.3
O2/SD = 9.3 (equivalent to O2/C = 0.72)
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
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
* 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)
53
Table 4.3d: SD conversion and H2 selectivity with W/FSD, 0 and T at O2/SD molar ratio of 10.5
O2/SD = 10.5 (equivalent to O2/C = 0.82)
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
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
* 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)
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
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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.
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
Butyl cyclohexane 0.0653
1,2,3,4-Tetrahydronaphthalene 0.0824
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).
T (K) Cp (J/mol/K) T (K) Cp (J/mol/K)
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)
T (K) Cp (J/mol/K) T (K) Cp (J/mol/K)
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 cyclohexane 0.536085
1,2,3,4-Tetrahydronaphthalene 0.418278
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
rxnAcAseffmax,particle
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
rxnAcAseffmax,particle
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
; [Data taken from experimental run at T=1123 K,
O2/SD=8.0 and W/FA,0=19008 Kgcatalyst*s/kmolSD]
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
; [Data taken from experimental run at T=1123 K,
O2/SD=8.0 and W/FA,0=19008 Kgcatalyst*s/kmolSD]
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
controls resistance film if Rate
rate Observed p
cAb
obs,A ; ----------------------------------------------- (4.9)
Here, Catalyst bulk density = 3b
m
kg54.17
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
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
controls resistance film if Rate
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
; [Data taken from experimental run at T=1123 K,
O2/SD=8.0 and W/FA,0=19008 Kgcatalyst*s/kmolSD]
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
Run#2: T =1123K, O2/SD=8.0, W/FSD,0=28512kgcat*s/kmolSD
118
Run#3: T =1123K, O2/SD=8.0, W/FSD,0=38052kgcat*s/kmolSD
119
Run#4: T =1123K, O2/SD=8.0, W/FSD,0=47556kgcat*s/kmolSD
120
Run#5: T =1123K, O2/SD=9.3, W/FSD,0=19008kgcat*s/kmolSD
121
Run#6: T =1123K, O2/SD=9.3, W/FSD,0=28512kgcat*s/kmolSD
122
Run#7: T =1123K, O2/SD=9.3, W/FSD,0=38052kgcat*s/kmolSD
123
Run#8: T =1123K, O2/SD=9.3, W/FSD,0=47556kgcat*s/kmolSD
124
Run#9: T =1123K, O2/SD=10.5, W/FSD,0=19008kgcat*s/kmolSD
125
Run#10: T =1123K, O2/SD=10.5, W/FSD,0=28512kgcat*s/kmolSD
126
Run#11: T =1123K, O2/SD=10.5, W/FSD,0=38052kgcat*s/kmolSD
127
Run#12: T =1123K, O2/SD=10.5, W/FSD,0=47556kgcat*s/kmolSD
128
Run#13: T =1173K, O2/SD=6.7, W/FSD,0=19008kgcat*s/kmolSD
129
Run#14: T =1173K, O2/SD=6.7, W/FSD,0=28512kgcat*s/kmolSD
130
Run#15: T =1173K, O2/SD=6.7, W/FSD,0=38052kgcat*s/kmolSD
131
Run#16: T =1173K, O2/SD=6.7, W/FSD,0=47556kgcat*s/kmolSD
132
Run#17: T =1173K, O2/SD=8.0, W/FSD,0=19008kgcat*s/kmolSD
133
Run#18: T =1173K, O2/SD=8.0, W/FSD,0=28512kgcat*s/kmolSD
134
Run#19: T =1173K, O2/SD=8.0, W/FSD,0=38052kgcat*s/kmolSD
135
Run#20: T =1173K, O2/SD=8.0, W/FSD,0=47556kgcat*s/kmolSD
136
Run#21: T =1173K, O2/SD=9.3, W/FSD,0=19008kgcat*s/kmolSD
137
Run#22: T =1173K, O2/SD=9.3, W/FSD,0=28512kgcat*s/kmolSD
138
Run#23: T =1173K, O2/SD=9.3, W/FSD,0=38052kgcat*s/kmolSD
139
Run#24: T =1173K, O2/SD=9.3, W/FSD,0=47556kgcat*s/kmolSD
140
Run#25: T =1223K, O2/SD=6.7, W/FSD,0=19008kgcat*s/kmolSD
141
Run#26: T =1223K, O2/SD=6.7, W/FSD,0=28512kgcat*s/kmolSD
142
Run#27: T =1223K, O2/SD=6.7, W/FSD,0=38052kgcat*s/kmolSD
143
Run#28: T =1223K, O2/SD=6.7, W/FSD,0=47556kgcat*s/kmolSD
144
Run#29: T =1223K, O2/SD=8.0, W/FSD,0=19008kgcat*s/kmolSD
145
Run#30: T =1223K, O2/SD=8.0, W/FSD,0=28512kgcat*s/kmolSD
146
Run#31: T =1223K, O2/SD=8.0, W/FSD,0=38052kgcat*s/kmolSD
147
Run#32: T =1223K, O2/SD=8.0, W/FSD,0=47556kgcat*s/kmolSD
148
Run#33: T =1223K, O2/SD=9.3, W/FSD,0=19008kgcat*s/kmolSD
149
Run#34: T =1223K, O2/SD=9.3, W/FSD,0=28512kgcat*s/kmolSD
150
Run#35: T =1223K, O2/SD=9.3, W/FSD,0=38052kgcat*s/kmolSD
151
Run#36: T =1223K, O2/SD=9.3, W/FSD,0=47556kgcat*s/kmolSD
152
Extended TOS Stability Run: T =1173K, O2/SD=9.3, W/FSD,0=38052kgcat*s/kmolSD