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An integrated 'one-box' process for hydrogen production

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An integrated 'one-box' process for hydrogen production

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AN INTEGRATED ‘ONE-BOX’ PROCESS
FOR HYDROGEN PRODUCTION

by

Mitra Abdollahi

A Dissertation Presented to the
FACULTY OF THE USC GRADUATE SCHOOL
UNIVERSITY OF SOUTHERN CALIFORNIA
In Partial Fulfillment of the
Requirements for the Degree
DOCTOR OF PHILOSOPHY
(CHEMICAL ENGINEERING)

May 2011

Copyright 2011 Mitra Abdollahi

Acknowledgments

This thesis is dedicated to my parents, for teaching me the value of
education, for their continued support through the years, and their
unwavering faith in me.

I would like to express my sincere thanks to my advisors, professor
Tsotsis and professor Sahimi for giving me the opportunity to pursue a
PhD program, for their support and encouragement throughout the
course of my study, as well as their insightful guidance, comments and
suggestions.

I would also like to express my thanks to Professor Massoud Pirbazari
for serving on my dissertation committee, and for his helpful remarks
and advice.

I am eternally grateful for my husband, for his constant love, strength,
and patience, and his ability to raise my spirits.

I would also like to express thanks to my brother, for his kindness,
assistance and support.

In addition, I must thank my fellow graduate students for their help and
collaboration. I am indebted to all my friends for their great friendship
and encouragement throughout my life.

Last but not the least, I sincerely appreciate the help of Chemical
Engineering Department staff, and thank them for their kindness.

iii

Table of Contents

List of Tables v
List of Figures vi
Abstract ix
Chapter 1. Motivation and Background
1.1 Introduction 1
1.2 Hydrogen Production 3
1.2.1 Hydrocarbon-Based Routes 3
1.2.1.1 Catalytic Steam Reforming 4
1.2.1.2 Partial Oxidation 6
1.2.1.3 Auto-thermal Reforming 8
1.2.1.4 Coal Gasification
9
1.2.1.5 Water-Gas-Shift Reaction
10
1.2.2 Unconventional Methods for Hydrogen Production 11
1.2.2.1 Biological Methods 11
1.2.2.1.1 Hydrogen from Biomass 12
1.2.2.1.2 Direct Bio-processing Technologies 13
1.2.2.2 Hydrogen from Water 15

1.3 Scope of the Work 17
Chapter 2. Hydrogen Production from Coal-Derived Syngas
2.1 Introduction 23
2.2 Experimental Studies 30
2.3 Modeling and Data Analysis 34
2.4 Results and Discussion 37
2.4.1 Kinetic Studies 37
2.4.2 Membrane Reactor Studies 44
2.5 Reactor Design and Scale-up 58

2.6 Summary and Conclusion 67
Chapter 3. Hydrogen Production from Biomass-Derived Syngas
3.1 Introduction 69
3.2 Experimental, Modeling, and Data Analysis 72
3.3 Results and Discussion 78

iv
3.3.1 Membrane Stability 78
3.3.2 Membrane Reactor Studies 80

3.4 Summary and Conclusion 88
Chapter 4. Ultra Pure Hydrogen Production from Reformate
4.1 Introduction 91
4.2 Experimental Studies 99
4.3 Modeling and Data Analysis 102
4.4 Results and Discussion 104
4.4.1 Membrane Characterization 104
4.4.2 Reaction Studies 110
4.5 Reactor Design and Scale-up 115

4.6 Summary and Conclusion 121

Bibliography 123

v
List of Tables
2.1 Kinetic parameters for the power-law rate model 40
2.2 Physical and chemical properties of the Co-Mo/Al
2
O
3
catalyst

41
2.3 Mixed-gas permeation data for CMS#1 46
2.4 Single-gas permeation data for CMS#2 50
2.5 The base-case and the range of the experimental
conditions used in the simulations

59
3.1 Single-gas permeation data for the CMS membrane before
and after the MR experiments

83
4.1 Summary of published data for H
2
permeation through
palladium membranes

106
4.2 Measured single gas permeances and calculated
separation factors for the Pd membrane before and after
MR experiments

110
4.3 Hougan Watson rate parameters 111
4.4 CO content of the MR H
2
product (permeate side) at two
different sets of the experimental conditions

118
4.5 Measured and calculated CO conversion and H
2
recovery
for two sets of the experimental conditions
118

vi
List of Figures
1.1 Conventional reforming processes flow diagram 5
1.2 Pressure-swing adsorption reforming process flow diagram 6
1.3 Partial oxidation process flow diagram 7
1.4 Coal gasification process flow diagram 9
1.5 Amount of pollutants and carbon dioxide emissions from
hydrogen use in fuel cell, gasoline use in hybrid electric
and internal combustion engine vehicles

19
2.1 Experimental set-up used in H
2
production from coal-
derived syngas

31
2.2 Experimental vs. fitted CO conversion using the power law
rate expression at various packed-bed reactor experimental
conditions

43
2.3 Comparison of the experimental (a) conversion and (b)
recovery with the model predictions at T=250°C, P=5 atm
and SR=0.1 using CMS#1

47
2.4 Comparison of the experimental (a) conversion and (b)
recovery with the model predictions at T=250°C, P=3 atm
and SR=0.3 using CMS#1

48
2.5 Compositions of (a) reject and (b) permeate side at P=3 atm
and SR=0.3

51
2.6 Comparison of the experimental (a) conversion and (b)
recovery with the model predictions at T=300
o
C, P=3 atm
and SR=0.1 using CMS#2

53
2.7 Comparison of the experimental (a) conversion and (b)
recovery with the model predictions at T=300
o
C, P=3 atm
and SR=0.3 using CMS#2

54
2.8 Comparison of the experimental (a) conversion and (b)
recovery with the model predictions at T=300
o
C, P=5 atm
and SR=0.1 using CMS#2
56

vii
2.9 Comparison of the experimental (a) conversion and (b)
recovery with the model predictions at T=300
o
C, P=5 atm
and SR=0.3 using CMS#2

57
2.10 Effect of pressure on (a) conversion and (b) recovery, L=1
m, T=300°C and SR=0.3

61
2.11 Effect of sweep ratio on (a) conversion and (b) recovery, L=1
m, T=300°C and P=5 atm

62
2.12 Effect of the membrane area on (a) conversion and (b)
recovery, L=1 m, T=300°C, P=5 atm and SR=0.3

63
2.13 Effect of the H
2
permeance on (a) conversion and (b)
recovery, L=1 m, T=300°C, P=5 atm, SR=0.3

65
2.14 Effect of the H
2
permeance on CO concentration in the
product stream, L=1 m, T=300°C, P=5 atm, SR=0.3 and
H
2
/CO separation factor (S.F.)=146

66
2.15 Effect of the CO permeance on CO concentration in the
product stream, L=1 m, T=300°C, P=5 atm, SR=0.3 and H
2

permeance=0.54 m
3
/(m
2
.h.bar)

66
3.1 Schematic of the membrane testing unit used in the
membrane stability tests in the presence of simulated
biomass derived syngas with model tar and organic vapor
compounds

74
3.2 Experimental set-up used in H
2
production from biomass-
derived syngas

76
3.3 He permeance and He/N
2
selectivity of the CMS membrane
in the presence of naphthalene/toluene as model tar and
organic vapors, respectively, in the biomass gasifier off-gas
at T=250
o
C and P=2 atm

79
3.4 Testing of the fouling and regeneration of a CMS
membrane in the presence of naphthalene/toluene as
model tar and organic vapors, respectively, in the biomass
gasifier off-gas

81
3.5 Effect of the MR feed-side pressure on CO conversion at
T=300
o
C and SR=0.3
85

viii
3.6 Effect of the sweep ratio on CO conversion at T=300
o
C and
P=3 atm

85
3.7 Effect of the MR feed-side pressure on a) CO conversion
and b) H
2
recovery at T=300
o
C, SR=0.3, and W
c
=10 g

87
3.8 Effect of the sweep ratio on a) CO conversion and b) H
2

recovery at T=300
o
C, P =20 atm, and W
c
=10 g

89
4.1 SEM cross section of the Pd layer deposited on the ceramic
support

98
4.2 The experimental set-up used in H
2
production from
reformate
100
4.3
Flux of hydrogen as a function of () ( )
Fn P n
PP  at T=300
o
C
107
4.4 Effect of CO on membrane H
2
permeation 108
4.5 Measured vs. calculated CO conversion data using the
Hougen –Watson rate expression

112
4.6 Effect of pressure on a) CO conversion and b) H
2
recovery,
T=300
◦
C and SR=0.3

116
4.7 Effect of sweep ratio (SR) on a) CO conversion and b) H
2

recovery, T=300
◦
C and P=446 kPa

117
4.8 Effect of pressure on CO conversion, T=300
◦
C and SR=0.1 119
4.9 Effect of pressure on H
2
recovery, T=300
◦
C and SR=0.1 120
4.10 Effect of pressure on H
2
purity, T=300
◦
C and SR=0.1 121

ix
Abstract

Hydrogen is widely considered as a fuel for the future to address
the energy crisis and the environmental concerns. In this study, process
intensification in hydrogen production from natural gas, coal- and
biomass-derived syngas is investigated both experimentally and
theoretically. A novel reactor/separator system, termed the “one-box”
process is being employed. The heart of this system is a membrane
reactor (MR) that combines the water-gas shift (WGS) reaction with
hydrogen separation into a single unit, thus eliminating the need for the
commonly utilized two separate WGS reactors and a distinct
purification unit.
Impurity-resistant carbon molecular sieve (CMS) membranes and
sour-gas shift

catalysts are utilized in order to treat the gas streams
with simulated coal/biomass-derived syngas compositions. Both have
shown high tolerance for H
2
S and NH
3
, which are the main impurities
in the syngas (in fact, the sour-gas shift catalysts require sulfur to be
present in order to remain active). This adds another benefit to the
system by eliminating the need for gas clean-up upstream of the WGS
reactor, which saves energy, and significantly simplifies the process
design. The CMS membrane stability is further investigated in the
presence of model tar and organic vapor compounds. The membranes
proved to be stable at temperatures akin to the WGS environment.

x
Permeation loss occurred at lower temperatures; however, regeneration
was readily achieved by purging the membranes with inert gas at higher
temperatures. This indicates that at low temperatures only surface
coverage through condensation occurs, with little or no irreversible pore
plugging of the membrane.
The project focus has been on experimental investigations in
order to prove the feasibility of using the ‘one-box’ system for H
2
production and to validate a mathematical model developed for use for
scale-up investigations. Prior to their use in the MR experiments, the
membranes are characterized via single and multi-component gas
permeation measurements. Packed-bed experiments are also performed
in order to derive the kinetic rate expressions. The effect of different
experimental conditions on MR performance, in terms of CO conversion,
hydrogen product recovery and purity is experimentally studied, and
the results are compared with those from the mathematical model. The
model is further used for process scale-up, and in order predict the MR
performance under conditions akin to the industrial applications. It is
shown that the MR-based systems are capable of converting more CO
into H
2
when compared to the more traditional packed-bed reactors, in
some cases attaining almost complete CO conversions.
In addition to the CMS membrane-based MR, we have also
investigated palladium (Pd) membranes with infinite selectivity towards

xi
hydrogen, combined with low-temperature shift (LTS) catalysts are used
for ultra pure hydrogen production from gas streams with simulated
natural gas reformate compositions, to serve as a feed for fuel cell
applications. One key advantage of using the Pd membranes is that one
can produce a high-purity (> 99.9%) hydrogen product for fuel cell
applications.

1
Chapter 1
Motivation and Background
1.1 Introduction
Nowadays, 86% of the world energy needs are provided by fossil
fuels (Dorian et al., 2006). The United States is the 3rd largest oil
producer in the world. Yet, the amount of imported petroleum is about
60% of the total consumption. Given the political uncertainty in the
world (particularly in the Middle East), the increasing demand for
petroleum from developing countries, coupled with the fact that the
world’s recoverable oil reserves are inching closer to depletion, it is
becoming increasingly more important to develop appropriate and long-
term technologies based on alternative fuels, which progressively can
substitute for the declining crude oil reserves.
There are also significant environmental concerns with petroleum
usage. Currently, the atmospheric carbon dioxide concentration is the
highest ever recorded, exceeding pre-industrial levels by almost 100
ppm, and above concentrations measured in gas pockets of ice cores,
reaching back 400,000 years (Petit 1999). A recent analysis estimated
that increasing CO
2
and other emissions related to petroleum usage
may affect public health in areas where 50% of Americans live
(Yuzugullu, 2006).

2
Hydrogen (H
2
), as fuel, is an important example of a non-carbon
energy carrier which, if widely adopted, could potentially diminish the
prospect of global warming. As a result, in recent years moving towards
a hydrogen economy, where hydrogen is used as the primary energy
carrier, has attracted the attention of both business leaders and policy
makers. The potential of hydrogen as a clean source of energy has been
known for almost two centuries, since the first combustion engine
developed in 1805 by Isaac de Rivaz was fueled with hydrogen.
Hydrogen can be produced from different routes, making use of a
variety of raw materials. Although H
2
production from a number of
sources (e.g., natural gas and coal) produces carbon dioxide,
sequestration followed by capture and storage could be conducted at
the site of production, thus diminishing its environmental impact.
Hydrogen used in transportation and power generation burns clean,
with the exception of some NO
x
emissions, but without carbon
emissions. The U.S. Department of Energy has developed a long-term
plan for the development of hydrogen infrastructure and storage
technologies. To be competitive, the price of hydrogen per kg (energy
equivalent of 1 gallon of gasoline) should be in the range $ 2-4/kg (Petit,
1999).

3
1.2 Hydrogen Production
Hydrogen at standard temperature and pressure conditions is a
colorless, odorless, highly flammable gas. It constitutes 75% of the
universe's elemental mass. It burns according to the following reaction:
2 H
2
+ O
2
→ 2 H
2
O +572 kJ 1.1
∆H= -286 (kJ/mol)
As mentioned before, hydrogen can be produced through various
routes, and can be used to satisfy almost any energy need. At present,
hydrogen is produced mainly from fossil fuels, biomass and water.
Today, about 96% of the worldwide production of hydrogen depends on
fossil energy resources. The steam-reforming process alone produces
nearly 90% of all the hydrogen (Das et al., 2001). Methods utilized to
produce hydrogen from different raw materials are as follows.

1.2.1 Hydrocarbon-Based Routes
A variety of hydrogen-containing materials, such as methane,
ammonia, or methanol, can be converted into a hydrogen-rich stream.
Methane is the most common raw material in commercial use. The
main techniques used to produce hydrogen from hydrocarbon fuels are
steam reforming, partial oxidation (POX), and auto-thermal reforming
(ATR).

4
1.2.1.1 Catalytic Steam Reforming
Catalytic steam reforming is a well-known and commercially
available process for hydrogen production. The reforming reactions for
hydrocarbons and methanol are:

C
m
H
n
+ mH
2
O → mCO + (m + ½n) H
2
1.2

∆H = depends on the hydrocarbon
CH
3
OH + H
2
0 → CO
2
+ 3H
2
1.3
∆H = +49 (KJ/mol)

The catalysts used for the reaction are divided into two groups:
those with non-precious metal (typically nickel), and catalysts with
precious Group VIII metals. The product of the reforming reaction
consists primarily of hydrogen, carbon monoxide and carbon dioxide.
The reaction requires heat which comes from an external source. High
temperatures are utilized for reforming that may encourage carbon
formation, and steam-to-carbon ratios higher than the stoichiometric
are typically required to gasify coke when a nickel-based catalyst is
used. Coke formation is less severe over the noble Group VIII metals.
Alkaline components are usually also added to the catalyst support in
order to minimize the coke formation (Ross 1975).

5
Hydrocarbon reforming produces an off-gas with a high H
2
content, particularly if water gas shift reactors are used to further shift
the syngas from the reformer; see Figure 1.1. However, it also produces
high CO
2
emissions (since hydrocarbon combustion also provides the
energy for the endothermic reaction). Conventional scrubbers (Figure
1.1) or pressure- swing adsorption methods (Figure 1.2) are utilized to
remove the CO
2
, potentially for storage and disposal.

Figure 1.1 Conventional reforming process flow diagram (Kothari et al.,
2004

1
p
C
∆

C
∆
Figure 1
1.2.1.2 Par
Hydr
partial oxid

C
m
H
n
+ ½m
∆H = depen
CH
3
OH + ½
∆H = -193.2
.2 Pressur
rtial Oxid
rocarbons
dation:
mO
2
→ mC
nds on the
½02 → CO
2 (KJ/mol
re-swing a
diagram
dation
or alcoho
CO + ½H
2

e hydrocar
O
2
+ 2H
2

l)

adsorption
(Kothari e
ols are als

rbon, exoth

n-based ref
et al. 2004
so converte

hermic

forming pr
4)
ed into CO

rocess flow
O and H
2

1
1
6

w
by
1.4
1.5

p
N
to
o
to
(F
re
d
F
a
F
In co
principle, d
Ni or Rh) a
o reduce
xidation r
o 2, which
FT) synthe
eactions t
difficult to
Figure 1.3
a partial ox

Figure 1.3
ontrast to
does not n
are utilized
the prope
reaction is
h is favora
esis reacto
that resul
control th
shows a f
xidation pr
3 Partial o
o reformin
need a cata
d in order
ensity to
that it pr
able for its
ors. A disa
ts in coke
he reactor
flow diagra
rocess.
xidation p

ng, the r
alyst, thou
r to lower
form coke
oduces a s
s downstre
advantage
e and hot
temperatu
am of hydr

process flow
reaction is
ugh cataly
the operat
e. An adv
syngas wit
eam use i
is the exo
t spot form
ure (Krumm
rogen prod
w diagram
s exotherm
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ting tempe
vantage o
th a H
2
/C
in the Fisc
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mation, a
menacher
duction fro
m (Kothari
mic and,
ally based
eratures a
f the part
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and makes
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et al., 200
7
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and
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ose
sch
the
s it
03).
ing

04)

8
1.2.1.3 Autothermal Reforming
An alternative process is autothermal reforming that uses a
combination of steam reforming for hydrogen production and of partial
oxidation in order to provide the heat of the reaction, resulting in an
overall thermally-neutral process. The overall reactions for
hydrocarbons and methanol fuels are:

C
m
H
n
+ ½mH
2
O + ¼mO2 → mCO + (½m + ½n)H
2
1.6
∆H = depends on the hydrocarbon, thermally natural

4CH
3
OH + 3H
2
O + ½02 → 4CO
2
+ 11H
2
1.7
∆H = 0

Auto-thermal reforming operates properly provided that the
steam-to-carbon and the oxygen-to-fuel ratios are properly controlled at
all times. In this case, one can control the reaction temperature and the
composition of the product gas and, at the same time, prevent coke
formation (Krumpelt et al., 2002). Autothermal reforming and partial
oxidation processes do not need an external heat source. On the other
hand, they require almost pure oxygen, or the product gas is diluted
with nitrogen. As a result, because of the cost of producing pure oxygen,

s
p
1

g
m
W
s
g

F
team refo
production

1.2.1.4 Co
In th
gasifier. Us
mixture con
WGS react
implified f
greater det
Figure 1.4
orming is
n in the ind
oal Gasific
his proces
sing heat
ntaining C
tion in ord
flow sheet
ail later in
4 Coal gas
s typically
dustry.
cation
ss coal is
and high
CO, CO
2,
C
der to prod
t of the p
n Chapter
sification p

y the pr
reacted w
pressure,
CH
4
and H
2
duce more
process. Th
2.
process flow

referred p
with oxyge
coal is co
2
. The syn
e hydrogen
he proces
w diagram
process fo
en and/or
onverted in
ngas is sub
n. Figure
s will be
m (Kothari
or hydrog
r steam in
nto a syng
bjected to
1.4 show
explained
et al., 200
9
gen
n a
gas
the
s a
d in

04)

10
1.2.1.5 Water-Gas-Shift Reaction (WGS)
First reported in 1888, the WGS reaction has became one of the
most important catalytic reactions in 1915, when the first coal-based
ammonia synthesis plant started operating (Kochloefl et al., 1997). As
noted previously, since all the above processes to produce hydrogen
also produce large amounts of carbon monoxide, one or more WGS
reactors, typically a combination of a high-temperature and a low-
temperature reactor, are being used to convert most of the CO into
additional hydrogen. The WGS reaction is given by:

CO + H
2
O → CO
2
+ H
2
1.8
∆H = -41.1 (KJ/mol)

The high-temperature WGS reaction (performed at temperatures
higher than 350
o
C) has fast kinetics, but is limited by the
thermodynamics for the amount of CO that can be shifted. As a result,
a low-temperature WGS reaction (performed at temperatures from 210
to 330
o
C) is used to further convert the carbon monoxide. The high-
and low-temperature WGS reactions use iron- and copper-based
catalysts, respectively (Nielsen et al., 2003).
Most hydrocarbon fuels contain some amount of sulfur that
poisons the conventional reforming and the WGS reaction catalysts.

11
This represents perhaps the biggest technical challenge facing these
technologies.

1.2.2 Unconventional Methods for Hydrogen Production
Hydrogen is produced by several other methods other than
reforming. A brief description of some of these techniques follows.

1.2.2.1 Biological Methods
Biological hydrogen production encompasses a broad group of
processes that utilize biomass as a raw material. The key advantage of
such processes is that, biomass is a renewable energy source that has,
in principle, an unlimited supply. Moreover, since the carbon in the
biomass results from the photosynthesis of atmospheric CO
2
, any
emissions resulting from the processing of such raw materials are CO
2
-
emission neutral. In fact, if the CO
2
is captured and stored, use of
biomass for fuels production has the potential to result in overall
negative CO
2
emissions. The use of waste biomass is of particular
interest here, as it currently represents a significantly underutilized
energy resource. A number of the biological hydrogen production
processes are further described below.

12
1.2.2.1.1 Hydrogen from Biomass
Biomass is the most likely future renewable organic substitute to
petroleum. The term biomass encompasses a wide range of different
materials, including animal and various other agricultural wastes, solid
municipal waste, woody crops and crop residues, aquatic plants, waste
paper, corn, and many others. Biomass can be used to produce
hydrogen via several methods including gasification, pyrolysis,
liquefaction, and hydrolysis. Of these, gasification is the technology that
is closer to commercial fruition.
Gasification technology is commercially practiced for a variety of
raw materials. It involves the partial oxidation of these materials into a
mixture of hydrogen, methane, carbon monoxide, carbon dioxide, and
nitrogen (Demirbas, 2006). Biomass gasification suffers from low
thermal efficiency, since the moisture contained in the biomass must
also be vaporized during the gasification process. It can be carried out
with or without a catalyst and in a fixed, entrained-flow or a fluidized-
bed reactor. Adding steam (in addition to oxygen) to the gasifier process
results in steam reforming and produces a hydrogen-rich syngas stream,
which can then be used as the feed to a WGS reactor for additional
hydrogen production (Demirbas, 2004). A separation process must be
used downstream of the WGS reactor in order to produce pure hydrogen.
Biomass gasification is further discussed later in Chapter 3.

13
1.2.2.1.2 Direct Bio-processing Technologies
Bio-processing technologies for hydrogen production have gained
attention over the last several years. The main technologies used for
bio-hydrogen production include direct photolysis, dark fermentative,
photo-fermentative processes, and hydrogen production by biological
WGS (that uses enzymes, instead of metal, to catalyze the process).
Some of these are further described below.
Direct Photolysis uses the same process that plants use to grow,
with the difference that it generates hydrogen instead of carbon-
containing biomass. Photosynthesis uses solar energy to generate the
reductant that is used for carbon dioxide reduction. In this process, two
photons are used for each electron that is removed from water and used
subsequently in carbon dioxide reduction or hydrogen formation. The
enzymes that catalyze hydrogen formation are typically absent in green
plants, so only CO
2
reduction happens in them. Microalgae, on the
other hand, have hydrogenase enzymes that are capable of producing
hydrogen, using solar energy.
One advantage of this technology is that it uses water as a source
of hydrogen, which is inexpensive and abundantly available. However, it
requires a considerable surface area to collect enough sunlight. Another
drawback is that the microorganisms produce oxygen along with
hydrogen that finally causes it to completely stop hydrogen production

14
(Kapdan et al., 2006). Moreover, producing hydrogen and oxygen
together poses safety concerns.
Dark fermentation uses anaerobic bacteria or some types of
algae that can grow on carbohydrate-rich substrates in the dark. For
fermentative processes, the biomass used needs to be biodegradable,
available in high quantities, inexpensive, and to have a high
carbohydrate content (Kapdan et al., 2006).
Photo-fermentative processes use the nitrogenase functionality
of purple non-sulfur bacteria to produce hydrogen. The sunlight breaks
water into protons, electrons, and oxygen. The nitrogenase catalyst
helps performing the reaction between protons and electrons with
nitrogen and ATP (adenosine triphosphate) in order to produce
ammonia, hydrogen and ADP (adenosine diphosphate).
In recent years, hydrogen production by fermentation has
attracted less attention than by photolysis. However, this method has
several advantages for industrial production. First, fermentative
bacteria have a very high hydrogen evolution rate. Second, they can
produce hydrogen from organic substrates continuously, both during
day and night. Third, they have a growth rate that is high enough for
the supply of the microorganisms to the production system.
Enzymatic WGS: hydrogen production by enzymatic WGS uses
photoheterotrophic bacteria that grow in the dark by consuming carbon

15
monoxide. The oxidation of the carbon monoxide to produce carbon
dioxide in the presence of water produces CO and H
2
, equivalent to the
WGS reaction, but through the use of enzymes rather than metals as a
catalyst. The process happens at relatively low temperatures and
pressures, and high conversion of carbon monoxide into carbon dioxide
and hydrogen is obtained (Levin et al., 2004). The conversion rate of the
process is higher than those of other biological processes.

1.2.2.2 Hydrogen from Water
Hydrogen is the most plentiful element in the form of water on
earth. Therefore, with water being a relatively cheap and abundantly
available source, production of hydrogen from water seems promising.
Water can be split to make hydrogen and oxygen, but the process
requires energy that must be obtained using solar energy, electricity or
heat.
Solar energy can be used in hydrogen production from water in a
variety of processes, including photolysis, photovoltaic-electrolysis, and
photo-electrochemical processes. In photolysis, photo-catalysts are
used that are able to adsorb light in the desired wavelength range. The
photolysis reaction is given below in which X represent the photo-
catalyst:

16
H
2
O + X + Light → X + H
2
+ ½ O
2
1.9

In photovoltaic-electrolysis, a photo-electrolyser, which is the
combination of a photovoltaic cell with an electrolyser upon exposing to
sunlight, converts water into hydrogen and oxygen.
In electrolysis, an electrical current passes through two
electrodes to break water into hydrogen and oxygen. The
electrochemical reaction is given by:

Cathode: 4H
2
O + 4e → 2H
2
+ 4OH
-

Anode: 4OH
-
→ 2H
2
O + O
2
+ 4e
Overall:
H
2
O+ Electricity → H
2
+ ½ O
2
+ Heat 1.10

Alkaline electrolysis processes are most commonly utilized.
Proton-exchange membrane electrolysis (PEM) and solid oxide
electrolysis cells (SOEC) are also being used (Grigoriev et al., 2006).
Electrolysis is a more expensive way for hydrogen production than the
conventional techniques. If non-renewable energy is used to produce
the required electricity, it may even result in higher CO
2
emissions than
the natural gas reforming.

17
In thermolysis, heat is used directly to break water into
hydrogen and oxygen. The problem is that water decomposes at 2500
o
C
and the number of materials stable at this temperature is severely
limited. Besides, viable heat sources are not always available.
Thermochemical processes help to lower the operating temperature
for thermolysis through the aid of chemical reactions. A challenge here
is that the various chemicals used in this method as intermediates
should be recovered for reuse.
Though, as noted above, unconventional methods to produce
hydrogen exhibit good potential and are generating great excitement,
most of the hydrogen production for the foreseeable future will come
from conventional reforming and gasification approaches. For the
United States in particular, given its abundant coal reserves and the
immense quantities of waste biomass that are generated, coal and
biomass gasification will remain a key focus.

1.3 Scope of the Work
The United States is becoming progressively more dependent on
imported oil in order to satisfy its energy needs. On the other hand, the
U.S. coal reserves are almost equal to the total world oil reserves. A
potential way towards U.S. energy independence, therefore, is the use of
these abundant coal resources that can supply its energy needs for up

18
to 250 years at the current domestic production rates. Coal can be
converted into liquid fuels, chemicals, and hydrogen using gasification
technology, as noted above.
Abundantly available everywhere in the world, biomass is another
attractive and comparatively inexpensive renewable energy source for
H
2
production. It can be derived from agricultural and forestry residues,
metropolitan and industrial wastes, and terrestrial and marine crops
grown only for energy purposes.
Global warming from increasing concentration of greenhouse
gases, mainly CO
2
, has become a key environmental issue over the past
several years. Figure 1.5, for example, compares the amount of
pollutant and carbon dioxide emissions from hydrogen used in fuel cells,
and from gasoline used in hybrid electric and in internal combustion
engine vehicles. From Figure 1.5 it is clear that gasification technology
which has the ability to produce clean synthesis gas and hydrogen from
coal and/or biomass has the potential to significantly reduce CO
2
and
other pollutant emissions.
Currently, the process of hydrogen production from syngas includes
the following steps. First, syngas needs to be cleaned in order to remove
its impurities. To do so, it is cooled down from temperatures >1,800 ºF
(as it leaves the gasifier) to ambient temperature, and is then reheated
again to 700
o
F in order to reach the temperature required for the next

s
re
te
to
fu
s
a
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in
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s
Figur
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tep that c
eactors, o
emperatur
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urther pur
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.
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GS reactor
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19
uch
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ane
t is
with
ess
gen
in
ogen
ngine

20
In this Thesis, the development of such a process, termed the “one-
box” process is described. It integrates the WGS reaction and hydrogen
separation step, together with carbon dioxide capture, and contaminant
removal into a single unit. The heart of the process is a novel membrane
reactor system, which combines the WGS reaction with hydrogen
separation. The WGS-MR exhibits increased yield and hydrogen purity.
Its development eliminates the need for two separate WGS reactors and
a separate purification section, thus reducing the capital costs.
Impurities in the synthesis gas (e.g. H
2
S, NH
3
, tar and organic
vapors) act as poisons for most of the conventional membrane and
catalyst materials. This mandates substantial clean-up upstream,
which is costly. The novelty of the WGS-MR under study in this project
is that, it uses carbon molecular-sieve membranes and Co-Mo/Al
2
O
3
catalysts, both of which show particularly high tolerance for such
impurities. In fact, the Co-Mo-based catalysts require sulfur to remain
active. This eliminates the need for gas clean-up upstream of the WGS
reactor which saves a lot of energy, and more importantly, significantly
simplifies the process design. In this project, the focus is on studying
the behavior of the WGS-MR and on validating the ability of these
membranes to function stably in the WGS environment.
In order to use H
2
in fuel cell application, purity is a very important
factor. Palladium membranes have almost infinite selectivity towards H
2
,

21
therefore, in this study, the possibility of H
2
production from clean
reformate through the use of the “one-box” system consisting of Pd
membranes and low-temperature Cu-Zn/Al
2
O
3
catalysts is also
investigated.
The rest of the Thesis is organized as follows:
Chapter 2 summarizes the results of the “one-box” process
performance investigation in H
2
production from coal-derived syngas.
The membranes are characterized through single and muti-component
gas permeation experiments. Packed-bed experiments are performed in
order to determine the reaction kinetics and rate parameters of the
commercial, sulfur-tolerant Co-Mo/Al
2
O
3
catalyst. The membrane
reactor performance is investigated for a range of applied pressures and
sweep ratios, and the results are compared with those from an
isothermal MR model. Finally, the same model is used to further
investigate the design features of the proposed process.
In Chapter 3, experimental results of the WGS-MR, using a feed with
simulated biomass-derived syngas composition are reported. The
membrane and catalyst stability are investigated in the presence of
common impurities, including H
2
S, NH
3
, tar and organic vapors. The
system performance is investigated under various experimental
conditions and compared with the model predictions. The model is used
for further studying the design aspects of the system.

22
In Chapter 4, the experimental results of the WGS-MR, using a feed
with simulated reformate composition are reported. The system
performance is investigated under various experimental conditions and
compared with the model predictions. The model is used for predicting
the system performance under conditions akin to the industrial
applications.

23
Chapter 2
Hydrogen Production from Coal-Derived Syngas
The research presented in this Chapter has already been published
(Abdollahi et al., 2010).

2.1 Introduction
Being comparatively inexpensive and relatively plentiful coal is an
attractive energy source (Shoko et al., 2006). Integrated gas combined
cycle (IGCC) power plants, in particular, show promise for
environmentally-benign power generation. In such plants, coal is
gasified into synthesis gas, and is then processed in water-gas shift
(WGS) reactors to further enhance its hydrogen content for clean-power
generation. As noted in Chapter 1, the current process involves, as the
first step, reacting coal with steam and/or oxygen in a gasifier to
produce syngas. The syngas must then be cooled down in order to
remove its contaminants, especially H
2
S, and then reheated to be
further reacted with steam in WGS reactors to maximize its hydrogen
content. The WGS reaction is exothermic and its equilibrium conversion
decreases with temperature. Therefore, typically two reactors, one
operating at high temperature (HTS) and another at lower temperature
(LTS) are used in order to overcome equilibrium limitations, and to
increase CO conversion of the feed at space velocities in the range of

24
(400-2500 hr
-1
) (Ladebeck et al., 2003). The gas stream exiting the
WGS reactors must be treated further in separation units to produce
nearly pure hydrogen. The total process, as it now stands, is complex
and energy intensive.
In order to avoid using the dual WGS reactor system, membrane
reactors (MR) have been considered for this application. Through the
use of the MR, and by removing H
2
in situ from the WGS reaction
mixture, the equilibrium is shifted towards the products, ultimately
resulting in higher conversions in a single-stage reactor. Dense
palladium (Pd) (Basile et al., 1996; Tosti et al., 2003; Iyoha et al., 2007;
Barbieri et al., 2008; Bi et al., 2009; Brunetti et al., 2009), microporous
silica (Giessler et al., 2003; Brunetti et al., 2007; Battersby et al., 2008;
Battersby et al., 2009), and carbon molecular sieve (CMS) membranes
(Harale et al., 2007; Sa et al., 2009; Harale et al., 2010) have been
studied for use in the MR applications for the WGS reaction. For
example, Bi et al. (2009) used a Pt/Ce
0.6
Zr
0.4
O
2
catalyst with a
hydrogen-selective Pd membrane and attained an improved
performance using feeds with compositions that match those at the exit
of industrial reformers. For a reactor temperature of 350
o
C, a feed-side
pressure of 1200 kPa, and a steam to CO ratio equal to 3, the
conversion remained above the thermodynamic equilibrium value for
feed hourly space velocities up to 9100 l kg
-1
h
-1
. However, H
2
recovery

25
decreased rapidly with increasing feed space velocity. The highest
hydrogen purity attained was 99.7%. Brunetti et al. (2009) used a
syngas mixture as the feed for the WGS reactor and upgraded it in an
one-stage MR using a Pd–Ag membrane. For a reactor temperature of
325
o
C, feed-side pressure of 600 kPa, and an hourly space velocity of
2600 h
-1
, they obtained a CO conversion of 90% and a H
2
recovery of
80%. Iyoha et al. (2007) used an MR containing multiple tubular Pd and
80 wt% Pd –20 wt% Cu membranes to perform the WGS reaction in the
absence of heterogeneous catalyst particles at 900
o
C, which is the
temperature an MR positioned just downstream of a coal gasifier would
typically encounter. At an hourly space velocity of 720 h
-1
, they attained
a CO conversion of 93% and a H
2
recovery of 90% using the Pd
membrane, and a CO conversion of 66% and a H
2
recovery of 85% using
the Pd-Cu membrane.
Despite being able to attain high CO conversions and to deliver
high hydrogen purity, Pd membranes have also drawbacks which have
limited their widespread industrial applications. Palladium is an
expensive material, which makes membranes made out of the metal
typically more expensive than other inorganic membranes. H
2
O and CO
block dissociation sites on the membrane surface, resulting in a
reduction in the hydrogen recovery (Li et al., 2007), especially at
temperatures below 300
o
C. H
2
S is known to adversely affect Pd

26
membrane characteristics, even at single-digit ppm concentration levels.
Exposure of the Pd to H
2
S has been shown, for example, to not only
reduce the permeability, but to also result in the formation of an
irreversible grey surface scale of palladium sulfide (Hurlbert et al.,
1961). Pd membrane exposure to H
2
S has also been shown to result in
dramatic pitting of the membrane surface (Kulprathipanja et al.. 2005).
The drawbacks associated with Pd membranes have motivated
the use of other types of high-temperature resistant membranes.
Giessler et al. (2003), for example, compared the performance of
molecular sieve silica membranes in the WGS reaction to that of Pd-
composite membranes. Their results suggest that the molecular sieve
silica membranes perform better than the Pd-composite ones. Since
silica membranes are known to be sensitive to steam, they proposed the
functionalization of the membranes using surfactants (e.g.,
triethylhexylammonium bromide) in order to form a hydrophobic silica
surface. They reported that the steam presence resulted in a decrease in
the H
2
and CO
2
permeances. However, the H
2
/CO
2
selectivity remained
constant between prior to and after steam exposure. Brunetti et al.
(2007) studied the performance of a traditional packed-bed reactor (PB)
and a membrane reactor making use of a silica membrane on a porous
stainless steel (SS) support for the WGS reaction. They observed that
increasing the temperature, especially at lower pressures, increased the

27
conversion difference between the MR and the PB. At a space velocity of
2070 h
-1
, a temperature of 280
o
C, and a pressure of 400 kPa, the MR
yielded a CO conversion of 95%, which represents an increase of 8%
over the conversion in the PB. The CO concentration in the permeate
side ranged from 1 to 10% depending on the operating conditions. The
membrane exhibited higher permeances, but no change in selectivity
(H
2
/CO, H
2
/CO
2
) before and after the reaction. The membrane flux (H
2
,
CO and CO
2
)

was shown to be a linear function of the driving force, and
no inhibition effect of other gases on the hydrogen flux was observed.
CMS membranes, which have attracted recent attention for gas
separation applications (Ismail et al., 2001), have also used in WGS-MR
applications. Harale et al. (2007. 2010) studied a hybrid adsorbent-
membrane reactor (HAMR) system to carry out the WGS reaction. They
used a hydrotalcite sorbent for CO
2
adsorption and nanoporous
hydrogen-selective CMS membranes. The HAMR system attained CO
conversions that are significantly higher than the corresponding
equilibrium conversions and proved to be more efficient than the PB as
well as the conventional MR. Sa et al. (2009), in a modeling study,
compared the performance of an MR containing Pd membranes with an
MR containing CMS membranes for H
2
production by methanol steam
reforming. They concluded that the major difference in the performance
of these MR occurs when the amount of hydrogen produced is low. At

28
low contact time, the Pd-MR presents a higher driving force for H
2
than
the CMS-MR, and delivers higher H
2
recovery. For the Pd-MR, low
hydrogen production rate means lower driving force since hydrogen is
the only permeating species. Changing the contact time has very little
effect on the Pd-MR recovery of H
2
, and only becomes noticeable at very
high contact times. On the other hand, all species permeate in the
CMS-MR, therefore, H
2
partial pressure increases with the permeation
of the other species (mainly water) towards the permeate side which
results in a higher driving force for H
2
and in increasing the CMS-MR
hydrogen recovery. The Pd-MR performance was further enhanced by
higher retentate and lower permeate pressures, while the CMS-MR
performed better for intermediate trans-membrane pressure gradients.
To our knowledge, all the WGS-MR studies, so far, have been
performed with pure feeds in the absence of impurities that one may
typically encounter in coal-derived syngas, particularly H
2
S, which has
proven detrimental to both conventional WGS (HTS and LTS) catalysts
(Newsome 1980) and Pd membranes (Hurlbert et al., 1961;
Kulprathipanja et al., 2005). We study, instead, a WGS-MR treating a
feed which contains substantial quantities of H
2
S (several thousand
ppm) typical of what may be encountered in the off-gas of a coal gasifier.
The idea is to make this MR the “heart ” of a “one-box ” process in which the
gasifier syngas is fed directly into the WGS reactor, which then

29
effectively converts the CO into hydrogen in the presence of H
2
S and
other impurities, and delivers a substantially contaminant-free
hydrogen product. For the MR, we have chosen to use CMS membranes
which are prepared by the deposition of polymeric precursors on
tubular alumina substrates commercially available by Media and
Process Technology, Inc. (for further details about the preparation
technique, see Sedigh et al. 1998; 1999; 2000). In ongoing field studies
by our team these membranes have already proven stable in the
treatment of commercial refinery streams containing high levels of
contaminants, such as H
2
S and NH
3
. CMS membranes have also been
shown previously (Harale et al., 2007, 2010) to be highly stable in the
presence of steam in the WGS reaction environment. Since H
2
S in the
syngas poisons the common WGS catalysts (Newsome, 1980), in our
study we are making use of a so-called sour-shift catalyst (Kochloefl et
al., 1997), in order to overcome the problem of catalyst poisoning. These
sulfur-resistant WGS catalysts, containing sulfided Co-Mo or Ni-Mo
supported on alumina, and on various other supports, such as zeolites
(Laniecki et al., 1993), titania and zirconia (Laniecki et al., 2000), were
first proposed almost twenty years ago; they have since been shown to
exhibit good performance (at relatively low temperatures, 250-350
o
C)
with syngas feeds containing high H
2
S concentrations (if adequately

30
pre-sulfided prior to use, they perform satisfactorily even in feed
streams that contain low level of H
2
S).
In what follows, we first discuss the experimental studies to
determine the reaction kinetics and rate parameters of the commercial
sulfur-tolerant Co-Mo/Al
2
O
3
catalyst that we utilize. Then, we discuss
the experimental membrane reactor performance for a range of
pressures and sweep ratios, and compare it with results from a simple
isothermal MR model. Finally, the same model is used to further
investigate the design features of the proposed process.

2.2 Experimental Studies
A schematic of the MR system used in this study is shown in
Figure 2.1. The tubular CMS membrane is sealed inside the tubular SS
reactor using graphite O-rings and compression fittings. The catalyst
particles are first thoroughly mixed with ground quartz particles, and
are then loaded into the annular space in between the membrane and
the reactor body. We dilute the catalyst with inert quartz particles in
order to completely fill the annular reactor volume, and to be able to
operate the reactor bed under isothermal conditions. The experimental
system consists of three sections: (i) the feed section, which consists of
gas cylinders, mass flow controllers (MFC), syringe pumps, and the
steam generating units; (ii) the reactor section, which consists of the

31
MR, a furnace for heating the reactor, pressure gauges for measuring
the pressure, two condensers and two moisture traps to remove the
water from the reject- and permeate-side streams of the reactor, and
two traps to remove the H
2
S from the same streams; (iii) the analysis
section, that consists of a gas chromatograph to analyze the
concentration of the exit gas streams, two bubble flow-meters for
measuring the total flow-rates, and Drager tubes to measure the H
2
S
concentration.

Figure 2.1 Experimental set-up used in H
2
production from coal-
derived syngas

32
For the experiments, the reactor is maintained isothermal by
placing it in a six-zone furnace, the temperature in each zone controlled
with the aid of six temperature controllers and thermocouples installed
in six different locations in the bed. An additional thermo-well is
installed in the bed in order to monitor the temperature along the
length of the bed using a sliding thermocouple. The feed and sweep gas
streams flow at specified rates controlled by mass-flow controllers.
Pressures are controlled by adjusting the needle valves at the exit of the
reactor and sweep sides. Pressure gauges are installed at the inlet and
outlet of both the feed and permeate sides in order to monitor the
pressure.
Two syringe pumps are used in order to supply a controlled flow
of water into two steam-generating units (one connected to the feed line,
and the other to the sweep stream line). The steam generators are well-
insulated SS vessels which are packed with quartz beads, in order to
accelerate the water evaporation and to dampen out any fluctuations in
the flow of the steam that is generated. They are heated by heating
tapes wrapped around them, and their temperature is controlled with
the aid of temperature controllers. All the stream lines, including feed,
sweep, permeate, and reject lines are insulated and heat-traced using
heating tapes. Their temperature is also controlled with temperature
controllers. In particular, the feed and sweep gas flows are preheated to

33
the reaction temperature before entering the reactor.
The above experimental system is also utilized to carry-out
permeation studies for characterizing the membrane properties. For
such experiments, the sweep gas (permeate-side) inlet is closed, gas
flows into the feed-side and the flow rates and the compositions of the
permeate and reject streams are measured. For calculating the water
permeance, the permeated stream by-passes the condenser and goes
directly into the adsorbent bed where the water is captured. The
amount of water that permeates is calculated by measuring the weight
of the adsorbent before and after water permeation. During the MR
experiments, the gas streams exiting the reject and permeate sides flow
first through condensers and then through moisture-traps in order to
capture the water. The flow rates of the water-free stream are then
measured by a bubble flow-meter. A small slip-stream from both the
reject and permeate sides is intermittently removed to measure the H
2
S
content through the use of Drager tubes. These are graduated tubes
that contain a Cu compound that reacts with the H
2
S and produces
CuS, which results into a color change from blue to black. The degree of
color change, read on a linear scale on the colorimetric detection tube,
is translated into an accurate measurement of the level of H
2
S (as low
as 0.2 ppm) present in the gas stream. Another small slip-stream from
both the reject and permeate sides is allowed to pass through an

34
adsorption bed (in order to remove its H
2
S content), and is then used to
measure the composition with an online gas chromatograph.
To carry-out the packed-bed reactor experiments (to compare its
performance with that of the MR) and for measuring the catalytic
reaction kinetics, the same procedure is followed, except that the inlet
and exit valves for the sweep gas are closed.

2.3 Modeling and Data Analysis
To analyze the data, we use an isothermal co-current flow (feed to
permeate) MR model utilized by our group for describing such reactors
(Hwang et al., 2008). Several assumptions are made in order to simplify
the mathematical analysis. Briefly, it is assumed that the reactor
operates isothermally (this has been validated experimentally) under
ideal gas law conditions, and that the external mass-transfer
resistances are negligible for the catalyst and the membrane. Catalyst
internal diffusional limitations are included in the overall rate
coefficients. Mass transfer through the membrane is described by the
following empirical equation:

2.1

35
where F
j
is the molar flux for component j (mol/m
2
. h), P
j
F
(bar) the
partial pressure for component j in the feed side, P
j
P
(bar) the partial
pressure for component j in the permeate side, and U
j
(mol/m
2
. h. bar)
the permeance for component j.
Mass balances for each component in the feed and permeate sides
are described by Equations 2.2 and 2.3, respectively.

1
2.2

2.3

where
F
j
n is the molar flow rate (mol/h) for component j in the feed side,
P
j
n the corresponding molar flow rate (mol/h) in the permeate side, V the
reactor volume variable (m
3
),
m
 the surface area of the membrane per
unit reactor volume (m
2
/m
3
), υ
j
the stoichiometric coefficient for
component j (negative for reactants and positive for products),
v
 the
bed porosity in the feed side,
c
 the fraction of the solid volume
occupied by the catalysts,
c
 the catalyst density (g/m
3
), and
F
r the
WGS reaction rate (mol/g. h)
The pressure drop in the packed-bed is calculated using the Ergun
equation (Equations 2.4-6, below).

36

10

2.4

1.75

2.5

2.6

where
F
P is the feed-side pressure (bar),
F
A the cross-sectional area
available to flow for the reactor feed side (m
2
),
F
 the viscosity (g/m.h),
P
d the particle diameter in the feed side (m),
F FF
Gu   the superficial
mass flow velocity in the feed side (g/m
2
·
h),
F
u the average velocity of
the fluid (m/h),
F
 the average fluid density (g/m
3
), and
c
g the gravity
conversion factor. The above set of equations are solved numerically
together with the following boundary conditions:

at V = 0 n
j
= n
j0
and P
F

=
0
F
P

where
0
F
P is the inlet feed-side pressure (bar) and n
j0
is the inlet molar
flow rate for component j (mol/h). CO conversion is defined by Eq. 2.7.

,

,

2.7

37
where
is the CO molar flow rate at the inlet (mol/h), ,
is the
CO molar flow rate at the exit of the reactor’s feed side (mol/h), and
,
is the CO molar flow rates at the exit of the reactor’s permeate
side (mol/h). Hydrogen recovery (
2
Re
H
) is given by Equations 2.8.

,

,
, 2.8

where ,
is the hydrogen molar flow rate at the exit of the reactor’s
feed side (mol/h) and , is the hydrogen molar flow rate at the exit
of the reactor’s permeate side (mol/h).

2.4 Results and Discussion
In what follows, we present the results, and describe and discuss
their implications.

2.4.1 Kinetic Studies
The kinetics of the WGS reaction has received substantial
attention in recent years (Twigg, 1990; Knozinger et al., 1997; Rase,
2000). Several researchers, in particular, have also studied the WGS
kinetics over sour WGS catalysts (Cimino et al., 1975; Hou et al., 1983;
Srivatsa, 1987; Srivatsa et al., 1988; Spillman, 1988; Ramaswamy,
1990; Lund, 1996). Various rate expressions have been reported and

38
different mechanisms (Twigg, 1990) have been proposed to explain the
observed reaction rate equations. Most researchers, however, make use
of an empirical, power-law rate expression, without reference to any
specific reaction mechanism. Weller and coworkers (Srivatsa, 1987;
Srivatsa et al., 1988; Spillman, 1988; Ramaswamy, 1990) have used the
following empirical power law rate expression for the WGS reaction over
a sour-shift catalyst:

1 2.9

1

.

.

where P
j
is the partial pressure for component j, and K
eq
is

the overall
reaction equilibrium constant

(Choi et al., 2003). They studied the
reaction over a sulfided Mo/Al
2
O
3
they prepared in the laboratory

(Spillman, 1988), as well as a Co-Mo/Al
2
O
3
catalyst they prepared in
the laboratory and a commercial Co-Mo/Al
2
O
3
catalyst

(Srivatsa, 1987;
Srivatsa et al., 1988; Ramaswamy, 1990). For the WGS reaction over
the Mo/Al
2
O
3
catalyst they reported two sets of reaction rate
parameters, corresponding to the experiments at two different pressures,
namely 5 and 15 atm (these parameters are shown in Table 2.1). For

39
the data with the Co-Mo/Al
2
O
3
catalysts they reported one set of rate
parameters also shown in Table 2.1. Lund

(1996) analyzed the same
data using a microkinetic model, which provides a better insight into
the behavior of the catalyst (including the presence of a maximum in
the rate as a function of the CO/H
2
O ratio, and the effect of H
2
S), but
which also makes use of a much larger set of rate parameters that must
be determined experimentally.
In our study, a commercial Co-Mo/Al
2
O
3
sour shift catalyst is
utilized to perform the WGS reaction (the physical properties of the
catalyst are shown in Table 2.2).
The catalyst and quartz particles are crushed separately into
smaller particles and their sizes are sorted with the aid of mesh-screens
in the range of 600-800 m. Prior to loading into the reactor, the
catalyst is mixed and diluted with the quartz particles in order to
completely fill the reactor space and to be able to conveniently operate
the reactor bed under isothermal conditions. The catalyst particles are
irregular in shape, but are roughly considered spherical for the
estimation of the bed properties. Since the Co and Mo metal
components of the fresh catalyst, as received, are in the oxidized form,
they must be sulfided prior to the reaction.

40

Srivatsa (1987)
219
38826
0.52
0.21
-0.1
-

Table 2.1 Kinetic parameters for the power-law rate model

Spillman (1988)
P=5 atm
6.3
22064
0.7
0.14
-
-
Spillman (1988)
P=15 atm

6.0

24911
0.8
0.29
-0.07
-
This work

4.9

28145
0.61
(±0.0605)
0.37
(±0.0752)
-0.38
(±0.0103)
-0.54
(±0.0079)

k
0

mol/(atm
(a+b+c+d)
·h·g)
E
(J/mol)
a
b
c
d

41

Catalyst form Extrudates
Catalyst size 0.003 m

Chemical
Composition

CoO: 3-4 wt%
MoO
3
: 13-15 wt%
Al
2
O
3
:

80-85 wt%
Bulk Density 592.68*10
3
g/m
3

Surface Area
160-220

m
2
/g
Pore Volume 0.55-0.65*10
-6
m
3
/g

Table 2.2 Physical and chemical properties of the Co-Mo/Al
2
O
3
catalyst

The activation procedure involves the in-situ reduction of the metals
using a gas mixture containing H
2
and H
2
S using the temperature and
pressure protocol as specified by the catalyst manufacturer. Since the
catalyst manufacturer did not provide any reaction rate information on
the catalyst, a series of kinetic experiments have been carried out in a
packed-bed reactor system. We also utilized the above empirical power
law rate expression (2.9) to fit our experimental packed-bed reactor data
(including the packed-bed experiments carried out in tandem with the
MR experiments), using nonlinear regression analysis. The rate
parameters (including their 95% confidence limits) are also shown in

42
Table 2.1. The goodness of fit can be seen in Figure 2.2, in which we
compare the experimental and fitted CO conversions for various
packed-bed temperatures (220-300
o
C) and pressures (1-5 atm) using
two different feed compositions namely (H
2
: CO: CO
2
: CH
4
: H
2
S
=2.6:1:2.1:0.8:0.05), typical of the composition of a coal/oxygen-blown
gasifier off-gas and (H
2
:CO:CO
2
:N
2
:CH
4
:NH
3
:H
2
S
=0.67:1.00:1.00:2.67:0.2:0.00067:0.0006), typical of the composition of
an air-blown biomass gasifier off-gas. We utilize in these experiments
W
C
/F
CO
values in the range of 70-320, where W
C
(10-12 g) is the weight
of undiluted catalyst, and F
CO
is the molar flow rate of CO (mol/h). We
also tested all four sets of rate parameters in Table 2.1 in simulating the
conversion from an isothermal packed-bed reactor operating at 250
o
C
with feeds similar to those used by Weller and coworkers (the studies
were carried out using feeds not containing H
2
or CO
2
), and there was
satisfactory agreement between the values obtained. However, as the
concentration of the H
2
/CO
2
in the feed increased differences in the
calculated conversions were amplified, indicative of the common
problem with such empirical rate expressions in that they are often
unable to predict data far removed from the set of experimental
conditions utilized to derive them. Our empirical rate expression,
however, adequately describes all the experimental packed-bed and MR
data generated in our study.

43

Figure 2.2 Experimental vs. fitted CO conversion using the power law
rate expression at various packed-bed reactor experimental conditions

44
2.4.2 Membrane Reactor Studies
The purpose of these experiments was to show that both the
membrane and catalyst perform stably under the WGS reaction
environment. In all the MR experiments reported here, we used a feed
with composition (on a dry basis) of H
2
:CO:CO
2
:CH
4
:H
2
S
=2.6:1:2.13:0.8:0.05 (intended to simulate a coal gasifier’s exit
composition), and a near stoichiometric H
2
O/CO ratio in the feed of 1.2.
Commercial WGS reactors generally operate in the presence of
substantial excess steam. One potential advantage of WGS-MR is that
they give the same conversion, at or near stoichiometric H
2
O/CO ratios,
thus the choice of the low H
2
O/CO ratio in our experiments. Two
different CMS membranes of the exact same dimensions (L=254 mm,
ID=3.5 mm, OD=5.7 mm) were utilized in the isothermal experiments
reported here. The first membrane (CMS#1) with relatively high flux but
moderate selectivity was used for a preliminary series of membrane
reactor experiments, and also for an extended series of membrane
characterization studies, with the entire series of experiments lasting
for over one month. Table 2.3 presents three sets of mixed-gas
permeation data. The first composition is that of the dry feed for the MR
experiments. The second composition is the exit (on a dry basis) reactor
composition corresponding to 70% CO conversion, while the third
composition is the same as the second one, but with water being

45
present (the lower H
2
S content in the second and third compositions is
because H
2
S is purchased premixed in the CO gas cylinder). As can be
seen in Table 2.3, varying the mixed-gas composition (including the H
2
S
concentration) has little effect on the permeance of most of the gases
(other than hydrogen for which the permeance varies by <15%). Figure
2.3 shows the CO conversion and H
2
recovery at three different c
CO
W
F
for
the MR experiments at P=5 atm and permeate steam sweep gas ratio
SR=

n
j0
P

n
j0
F
  
= 0.1 (the error bars reflect the carbon and hydrogen
loss or gain due to the experimental errors in measuring the flow rates
and compositions). In the experiments we used 15 g of the catalyst
diluted with 80 g of ground quartz glass, packed inside the MR in the
annular space between the reactor wall and the membrane (membrane
shell-side). Shown on the same figure are the simulated conversion and
recovery lines using the model, utilizing the experimental rate
expression, and the last set of mixed-gas permeances in Table 2.3.
Figure 2.4 shows the conversion and hydrogen recovery at 250
o
C for a
different set of experimental conditions as shown in the figure caption.

46

Gas Mixture Composition
(1)
H
2
:CO:CO
2
:CH
4
:H
2
S =
39.5%:15.2%:32.4%:12.2%:0.7%

Gas
Permeance
m
3
/(m
2
.h.bar)
Separation
Factor (S.F)
H
2
1.37 1.0
CO 0.02 68.5
CO
2
0.05 27.4
CH
4
0.01 137.0
H
2
S 0.01 137.0

(2)
H
2
:CO: CO
2
:CH
4
:H
2
S=
45.36%:4.52%:38.91%:11%:0.21%

Gas
Permeance
m
3
/(m
2
.h.bar)
Separation
Factor (S.F)
H
2
1.40 1.0
CO 0.02 70.0
CO
2
0.04 35.0
CH
4
0.01 140.0
H
2
S 0.01 140.0
(3)
H
2
:CO:CO
2
:CH
4
:H
2
O:H
2
S =
42.45%:4.233%:36.4%:10.29%:6.43%:0.197%

Gas
Permeance
m
3
/(m
2
.h.bar)
Separation
Factor (S.F)
H
2
1.56 1.0
CO 0.02 78.0
CO
2
0.05 31.2
CH
4
0.01 156.0
H
2
O 1.1 1.4
H
2
S 0.01 156.0

Table 2.3 Mixed-gas permeation data for CMS#1

(a
(b
a)
b)
Figure
recovery w
2.3 Comp
with the m
parison of t
model pred
u

the experi
dictions at
using CMS
mental (a)
T=250°C,
S#1
) conversio
P=5 atm a
on and (b)
and SR=0.
47

.1

(a
(b
CO C i %
a)
b)
Figure
recovery w
0
25
50
75
100
0
CO Conversion%
2.4 Comp
with the m
parison of t
model pred
u
100
W

the experi
dictions at
using CMS
200
Wc/Fco (g-cat
mental (a)
T=250°C,
S#1
0
t.h/mol-CO)
) conversio
P=3 atm a
300
on and (b)
and SR=0.
MR EXP
MR SIM
48

.3
400

49
In order to further explore the range of appropriate operating conditions
for the CMS-MR, and to further validate the ability of membranes and
catalysts to function stably, a second series of MR experiments for the
WGS reaction were carried out at 300
o
C, a temperature which pushes
the limits of application for both the sour-shift catalyst and the CMS
membranes. In these experiments, we again used the same feed gas
composition and H
2
O/CO ratio, with the experiments being carried out
at two different feed pressures (3 and 5 atm, with the permeate side
pressure under atmospheric conditions), and two different permeate
steam sweep gas ratios of 0.1 and 0.3. In this series of experiments, in
tandem with the MR experiments, we also carried out under identical
conditions PB experiments (every MR experiment was followed by a PB
experiment during which the permeate side is kept closed, as previously
described). A different membrane (CMS#2), this time with high
selectivity but relatively low permeability was utilized, and 10 g of the
catalyst diluted with 80 g of ground quartz glass was packed in the
membrane shell-side (the reactor shell-side volume was slightly smaller
in these series of experiments, thus the use of a smaller amount of
catalyst).
Since the emphasis in these experiments was on the MR
performance, we carried out only a limited number of permeation
studies using single gases before and after the MR experiments. Prior

50
studies by our group with several of these membranes indicate that the
mixed-gas permeances of the various gases generally remain relatively
close to the values measured during the single-gas experiments (Harale
et al., 2007). Table 2.4 indicates the single gas permeances for this
membrane measured prior to the initiation of the reactor experiments.

Pure Gas

Permeance
m
3
/(m
2
.h.bar)
Separation Factor
(S.F)
H
2
0.5354 1
CO 0.0037 145.88
CO
2
0.0107 50.03
CH
4
0.0014 385.18
H
2
O 0.0922 5.8

Table 2.4 Single-gas permeation data for CMS#2

51
(a)

(b)

Figure 2.5 Compositions of (a) reject and (b) permeate side at P=3 atm
and SR=0.3

52
For H
2
S, the MR experiments indicated that it does not permeate
through the membrane (within the detection limit of the Drager tube
utilized) and, hence, its permeance was taken to be zero, since it has no
impact on the modeling results. (In extensive studies in which both the
surface of the membrane module and the plumping were specifically
coated to avoid potential wall adsorption, the H
2
S permeance was
always found to lie in between the permeance of CO/N
2
and CH
4
). The
permeance of water was deduced by fitting all the compositional data
available for both the reject and permeate sides of the MR (e.g., see
Figure 2.5 for the fit for one set of such data). The MR experiments
lasted more than one month during which period membrane gas
permeances changed less than 7% before and after the MR experiments,
indicative of the good stability of the membranes under the WGS-MR
environment. The model discussed earlier was again used to simulate
the experimental results, together with the experimental empirical
power-law rate expression and the experimental single-gas permeances
(Table 2.4) as discussed above.
Figure 2.6 shows the CO conversion and H
2
recovery at three
different c
CO
W
F
for the MR and PB experiments at 3 atm and a steam
sweep ratio of 0.1. Figure 2.7 presents the CO conversion and H
2

recovery at the same conditions mentioned above, but at a steam sweep
ratio equal to 0.3.

(a
(b
a)
b)
Figure
recovery
0
25
50
75
100
100
CO Conversion%
e 2.6 Comp
with the m
0
parison of
model pred
u
200
W

f the exper
dictions at
using CMS
0
Wc/Fco (g-cat
rimental (a
t T=300
o
C
S#2
300
t.h/mol-CO)
M
P
a) conversi
, P=3 atm
0
MR EXP
PB EXP
ion and (b)
and SR=0
400
MR SIM
PB SIM
53

)
0.1

(a
(b
a)
b)
Figure
recovery w
0
25
50
75
100
10
CO Conversion%
2.7 Comp
with the m
0
parison of t
model pred
u
200
W

the experi
dictions at
using CMS
0
Wc/Fco (g-cat
mental (a)
T=300
o
C,
S#2
300
.h/mol-CO)
M
P
) conversio
P=3 atm a
0
MR EXP
PB EXP
on and (b)
and SR=0.
400
MR SIM
PB SIM
54

.3

55
Figures 2.8 and 2.9 show the CO conversion and H
2
recovery for the MR
and the PB experiments at 5 atm and steam sweep ratios equal to 0.1
and 0.3, respectively (the solid lines in the figures represent the
modeling results).
These figures make it clear that the model does, generally, a good
job in predicting the experimental MR as well as the PB behavior. The
MR attains conversions which are higher than those for the packed-bed.
The relatively low conversions are due to limitations with the size of our
laboratory system, which accommodates only one small CMS
membrane, and the limited amount of catalyst that can be filled inside
the small reactor. The pressure drop measured in the laboratory (and
also the one calculated using the Ergun equation) was also negligible
due to the same reasons. Higher conversions can be attained for higher
c
CO
W
F
, as the simulation results also indicate (see below). We are limited,
however, in our laboratory system by the amount of catalyst that can be
utilized. Once the amount of catalyst is fixed, the maximum value of
c
CO
W
F
attained is determined by the minimum flow rate one is able to
provide, reflecting the lower limits of the MFC in our experimental
system (it should be noted, however, that high c
CO
W
F
values are also not
interesting for practical applications).

(a
(b
a)
b)
Figure
recovery w
0
25
50
75
100
100
CO Conversion%
2.8 Comp
with the m
0
parison of t
model pred
u
200
W

the experi
dictions at
using CMS
0
Wc/Fco(g-cat.
mental (a)
T=300
o
C,
S#2
300
h/mol-CO)
M
P
) conversio
P=5 atm a
MR EXP
PB EXP
on and (b)
and SR=0.
400
MR SIM
PB SIM
56

.1

(a
(b
a)
b)
Figure
recovery w
0
20
40
60
80
100
10
CO Conversion%
2.9 Comp
with the m
00
parison of t
model pred
u
200
W

the experi
dictions at
using CMS
0
Wc/Fco (g-cat
mental (a)
T=300
o
C,
S#2
300
.h/mol-CO)
M
PB
) conversio
P=5 atm a
0
R EXP M
B EXP P
on and (b)
and SR=0.
400
MR SIM
PB SIM
57

.3

58
A key conclusion from this series of experiments is that the membrane
and the catalyst exhibited robust behavior and remained stable
throughout the series of experiments which lasted almost one month in
the presence of hydrogen sulfide under the harsh WGS environment.

2.5 Reactor Design and Scale-up
Since the model performs reasonably well in describing the
experimental results, it can be used to further study the effect of
various parameters on WGS-MR performance in terms of reactor
conversion, hydrogen recovery, and purity. The target here is to choose
appropriate conditions which maximize both the CO conversion and H
2

recovery, and minimize the CO content of the hydrogen product. In the
simulations that follow the experimental power-law reaction rate
expression together with the experimental single-gas permeances for the
CMS#2 membrane were used. The membrane length is increased to 1 m
and it is assumed that the catalyst and quartz are packed along the
entire length of the membrane. The conditions utilized, unless otherwise
noted in the figures’ captions, are listed in Table 2.5.

59

Parameter Base case
Applied
range
Feed side pressure (P
F
) 5 atm 5-30 atm
Permeate side pressure (P
P
) 1 atm -
Reactor temperature (T
R
) 300
o
C -
Steam sweep to feed ratio (SR) 0.3 0.3-2
Number of membranes 1 1-4
Length of the membrane (L) 1 m -
Inner diameter of the membrane (ID) 0.0035 m -
Outer diameter of the membrane (OD) 0.0057 m
Inner diameter of the reactor 0.0318 m
Weight of the catalyst (W
c
) 10 g -
Hydrogen permeance
0.54
m
3
/(m
2
.h.bar)
0.5-3
m
3
/(m
2
.h.bar)
H
2
/CO separation factor (S.F) 146 100-300

Table 2.5 The base-case and the range of the experimental conditions
used in the simulations

60
Figure 2.10 shows the effect of pressure on the WGS-MR performance.
Since the coal gasifier typically operates in the pressure range of 20-30
atm, the advantages of operating the CMS-MR at high pressures is
obvious. Increasing the pressure helps increase both the conversion
and the hydrogen recovery by increasing the partial pressure difference
of hydrogen across the membrane. The pressure effect is more
prominent at lower c
CO
W
F
, and for a constant weight of catalyst that
means higher feed flow rates. As expected, even at the highest
pressures the reactor does not attain complete conversion due to the
omnipresent loss of CO, which indicates that a more appropriate type of
reactor may be a hybrid system consisting of a packed-bed, followed by
an MR (Barbieri et al., 2008). Figure 2.11 shows the effect of the sweep
ratio. As the figure indicates, increasing the SR increases both the
conversion and recovery, as expected, since sweeping helps maintain
the permeate-side partial pressures low. Increasing the sweep ratio does
not affect the CO transport through the membrane as much as
increasing the reactor side pressure, and as a result the impact of CO
loss is not as severe and CO conversion continues to increase as the
sweep ratio increases. Figure 2.12 shows the effect of the membrane
area on the MR performance. In these simulations, we have kept the
amount of catalyst constant and have increased the number of
membranes that are packed into the reactor. The amount of quartz

u
m
u
(a
(b
F
utilized ha
membrane
up to four m
a)
b)
Figure 2.1
as been a
and the r
membrane
0 Effect of
adjusted
reactor wa
es inside t
f pressure
T=30

to fill th
ll (we estim
the reactor

e on (a) con
00°C and S
he annula
mate that
r).
nversion a
SR=0.3
ar space b
we can co
and (b) reco
between t
omfortably
overy, L=1
61
the
y fit

1 m,

(a
(b
F
a)
b)
Figure 2.11 Effect of f sweep ra
m, T=3

atio on (a) c
300°C and
conversion
P=5 atm
n and (b) r recovery, L
62

L=1

(a
(b
a)
b)
Figure 2 2.12 Effec
recovery
ct of the m
y, L=1 m, T

membrane a
T=300°C, P
area on (a
P=5 atm an
) conversio
nd SR=0.3
on and (b)
3
63

64
As Figure 2.12 indicates, increasing the number of membranes (n),
which translates into increasing the membrane area per unit catalyst
weight, increases both CO conversion and H
2
recovery, due to the more
rapid transfer of products to the permeate side. This favorably shifts the
reaction equilibrium towards the product side. Similarly to the
pressure effect, the effect of increasing the membrane area is stronger
at lower c
CO
W
F
, which at constant weight of the catalyst corresponds to
higher feed flow rates. Figure 2.13 shows the effect of varying the H
2

permeance on the MR performance (while keeping the separation
factors towards the other species constant, as indicated in Table 2.4).
As expected, increasing the H
2
permeance helps increasing both CO
conversion and H
2
recovery, the effect being more pronounced for the
hydrogen recovery. Figure 2.14 shows the effect of varying hydrogen
permeance (while maintaining the separation factors constant) on CO
concentration (on a dry basis) in the product stream. Note that while
higher permeances have a positive impact on CO conversion and
hydrogen recovery, they have, on the other hand, a negative impact on
hydrogen purity. Finally, Figure 2.15 shows the effect of varying the
H
2
/CO separation factor by varying the CO permeance, while
maintaining the permeances of all other species constant (see Table 2.4),
on the CO concentration (on a dry basis) in the product stream.
Decreasing the CO permeance impacts the conversion (not shown here)

65
by decreasing the inadvertent CO loss to the permeate side, but as
expected, its most significant impact is lowering the CO concentration
in the product stream.
(a)

(b)

Figure 2.13 Effect of the H
2
permeance on (a) conversion and (b)
recovery, L=1 m, T=300°C, P=5 atm and SR=0.3

66

Figure 2.14 Effect of the H
2
permeance on CO concentration in the
product stream, L=1 m, T=300°C, P=5 atm, SR=0.3 and H
2
/CO
separation factor (S.F.)=146

Figure 2.15 Effect of the CO permeance on CO concentration in the
product stream, L=1 m, T=300°C, P=5 atm, SR=0.3 and H
2

permeance=0.54 m
3
/(m
2
.h.bar)

67
As noted previously, the proposed “one-box” approach is being
studied in the context of IGCC power plants, where the goal is to carry
out the WGS step without the need to cool the gasifier off-gas stream to
remove its various contaminants (e.g., H
2
S) and then having to reheat it
back to the WGS reaction temperature. Turbines, internal combustion
(IC) engines and proton exchange membrane (PEM) fuel cells have been
studied for power generation using the hydrogen-enriched syngas.
Turbines and IC engines are significantly more tolerant to low levels of
sulfur, CO and other contaminant than PEM fuel cells. If PEM fuel cells
are the option of choice for power generation, however, then as the
simulations above indicate an additional polishing step
(absorption/adsorption for the H
2
S and other contaminants and
membrane separation/preferential oxidation for CO) may be required.
However, the energy needed for such a step, when treating the permeate
stream of the proposed “one-step” process, is likely to be a fraction of
what would be needed to treat the off-gas of a conventional sour-shift
reactor.

2.6 Summary and Conclusions
In this Chapter, the “one-box” process which combines reaction
and membrane separation in the same unit was experimentally
evaluated for the WGS reaction. The kinetics of the same reaction over

68
sulfided Co-Mo/Al
2
O
3
catalyst was investigated and a data-validated
rate expression and kinetic parameters were obtained. Nanoporous
carbon molecular sieve membranes were used for the in-situ hydrogen
separation. The membranes’ performance was investigated under the
operating conditions and their transport properties were used for the
model predictions. The modeling studies indicated good agreement with
the experimental data. The MR performance was investigated for a
range of pressures and sweep ratios, and showed higher CO
conversions and H
2
purity compared with those of the traditional
packed-bed reactor. The effect of the membrane properties and
experimental conditions on the performance of the system was also
investigated. The “one-box ” process proved to possess several advantages
over the traditional systems including increasing CO conversion,
decreasing the amount of steam required for the reaction, and being
able to deliver a product with significantly lower CO content. Using
impurity-resistant catalyst adds another advantage to this system by
allowing one to perform the reaction in the presence of hydrogen sulfide;
for the IGCC power plants this would result in considerable energy
savings. The catalyst and the CMS membranes have demonstrated good
stability in the presence of hydrogen sulfide in continuous reactor
experiments lasting over a month.

69
Chapter 3
Hydrogen Production from Biomass-Derived Syngas
The research presented in this Chapter has already been published
(Abdollahi et al., 2010). The work was carried out collaboratively with
researchers at Media and Process Technology, Inc. Specifically, the
testing of the CMS membranes in the presence of tar and toluene
vapors was carried out by M&P but is included here for completeness.

3.1 Introduction
The abundant availability of coal and biomass in the USA makes
them both attractive and comparatively inexpensive energy sources for
H
2
production. The possibility of H
2
production from coal-derived
syngas in the presence of its impurities, utilizing the proposed “one-box ”
system is investigated and the results were presented above. Biomass,
however, has several advantages over coal; it is a renewable energy
source, abundantly available in the world. It comes in various forms,
ranging from agricultural and forestry residues and municipal and
industrial waste, to terrestrial and marine crops grown specifically for
energy production. It consumes atmospheric CO
2
during growth;
therefore, producing hydrogen from biomass has the potential to result
in substantially lower net CO
2
emissions than when using coal.

70
Though biological processes have been proposed for H
2

production from biomass, they suffer from low production rates and
from process inefficiencies

(Levin, 2009). As a result, as with coal,
gasification remains today the main method for H
2
production from
biomass. It converts the biomass, in the presence of air (oxygen) and/or
steam, into a syngas consisting mainly of H
2
, CO, CO
2
, N
2
(when air is
used), and lower molecular weight (MW) hydrocarbons, principally CH
4
,
as well as smaller concentrations of organic vapors, and high MW
compounds known collectively as tars

(Bridgwater et al., 2003;
Schiefelbein et al., 1989). The latter are a major issue in biomass
gasification that must be dealt with during process development.
Appropriate modifications in gasifier design and operating conditions
along with using catalysts and additives, as well as novel technologies
such as supercritical water gasification

(Yu et al., 1993; Sutton et al.,
2001; Devi et al., 2003; Guo et al., 2010) help to minimize tar formation,
and also to convert the hydrocarbon vapors in the syngas into H
2
and
CO via steam reforming. However, their presence in the syngas cannot
be completely eliminated and they must, therefore, be taken into
account during the design of downstream processes for further syngas
clean-up and processing.
As for coal derived syngas, maximizing the hydrogen content in
the biomass derived syngas requires cooling it down to remove its

71
contaminants, especially H
2
S and NH
3
, and then heat it up again to
further react it with steam in WGS reactors. The gas stream exiting the
WGS reactors must be treated further in additional separation units to
produce pure hydrogen. Aznar et al. (2006), for example, used WGS
reactors downstream from a fluidized-bed biomass gasifier and a steam-
reforming catalytic bed. They obtained CO conversions higher than 90%
and H
2
content as high as 73 vol% on a dry basis. The CO conversion
and the increase in H
2
content correlated with the steam/CO ratio in
the syngas at the inlet of the WGS reactor. Effendi et al. (2005) studied
the H
2
production from a contaminant-free model biomass-derived
syngas through a combination of steam reforming (in excess steam) in a
fluidized-bed reactor over a Ni/Al
2
O
3
catalyst, followed by two fixed-bed
WGS reactors. The steam reforming step resulted in more than 98 %
CH
4
conversion, while the WGS reactors raised the H
2
content to 68%.
The total process of H
2
production from biomass-derived syngas,
as it is now envisioned, is very energy-intensive. As with coal, “process
intensification, ” in which gas clean-up, the WGS reaction, and hydrogen
separation are all integrated into one unit, offers potential advantages
in scaling-up, and provides for better efficiency and lower costs. In this
study, we investigated the use of the “one-box” process (containing CMS
membrane and Co-Mo/Al
2
O
3
catalyst) for this purpose and the results
are presented in this Chapter. We first present results of the stability

72
tests of the CMS membranes under the expected operating conditions in
the presence of simulated biomass gasifier off-gas containing model
organic vapor and tar compounds. We then present WGS packed-bed
and MR experiments with feeds containing H
2
S and NH
3
at levels typical
of those encountered in biomass-derived syngas. The experimental MR
performance for a range of pressures and sweep ratios is discussed and
compared with the results obtained from the isothermal MR model.
Lastly, the same model is used to further investigate the design features
of the proposed process.

3.2 Experimental, Modeling, and Data Analysis
Our CMS membranes have proved to be very stable at the operating
temperatures in the range of 200 to 300 ºC, under aggressive
hydrothermal conditions (Harale et al 2007; Harale et al. 2010;
Abdollahi et al. 2010)

for extended periods of time (more than a month).
Pilot-plant tests (lasting as long as one week) have also been conducted
with these CMS membranes with both refinery hydrocracker and coal
gasifier off-gases laden with various contaminant gases, including
organic vapors, H
2
S, NH
3
, etc., and the membranes proved remarkably
stable. As discussed in Chapter 2, the membranes also proved to be
stable for more than a month under the WGS environment using a feed
with high levels of H
2
S, with no apparent change in the membrane

73
permeation characteristics (less than 7% before and after the MR
experiments), indicative of the good stability of the membranes in the
presence of the impurities. Tar-like species, expected in the biomass
gasifier off-gas, represented a new potential challenge to these
membranes prior to the start of our project, however. A key goal of our
studies, therefore, was to validate the ability of CMS membranes to also
function stably in the presence of these added contaminants. Based on
previous literature work (Coll et al., 2001), in this study naphthalene
was chosen as a model tar-like species, while toluene was used as an
organic vapor simulant. Helium was used as an inert carrier gas during
the testing, in order to generate a simulated off-gas stream containing
0.8 vol% naphthalene and 6.4 vol% toluene with the aid of a vaporizer
(see below). Most of the membrane stability tests were conducted at 250
ºC and total system pressure of ~2 atm. However, a number of tests
were also carried out in which the membrane was purposefully “fouled”
by operating at low temperatures (which result in the condensation of
naphthalene and toluene within the membrane structure) in order to
verify whether these membranes could be regenerated following such
operational upsets.
The experimental system used in these tests is shown in Figure
3.1. It consists of two parts, namely (i) the feed subsystem, and ii) the
membrane subsystem. The feed subsystem is used to prepare and

74
deliver simulated tar-laden biomass gasifier off-gas to the membrane. It
consists of an HPLC pump and check-valves and shut-off valves used to
control the liquid flow to the unit. During the tests, a solution of
naphthalene in toluene is first prepared and then delivered by the HPLC
pump into a constant temperature vaporization chamber at a constant
flow rate. Gas is also delivered to the unit using a separate line, its flow
controlled via a mass flow controller.

Figure 3.1 Schematic of the membrane testing unit used in the
membrane stability tests in the presence of simulated biomass derived
syngas with model tar and organic vapor compounds
The liquid solution is fed to the top of the mixing chamber and
vaporized in a hot zone. This design assured that a uniform gas/vapor
stream was delivered to the membrane and prevented the naphthalene

Porous Membrane
He
P
HPLC
Pump
Naphthalene -
Toluene Feed
Vessel
MFC
P
Bubble
Flow
Meter
Carbon
Trap
Back Pressure
Regulator
Shut-off
Valve
Check
Valve
Hot
Zone
Bubble
Flow
Meter

75
from crystallizing in the feed lines prior to vaporization. The membrane
subsystem consists of a stainless steel (SS) module placed in a
temperature-controlled furnace, in which the CMS membrane (with
dimensions of 5.7 mm OD, 3.5 mm ID, and 254 mm in length) is sealed
with the aid of graphite packing. During the experiments, the retentate
stream in the membrane module passes through a carbon-trap to
remove the naphthalene and toluene, and its flow rate is then measured
by a bubble-flow meter, as is the flow rate of the permeate-side stream.
(There is no need to use a carbon trap for the permeate stream, as for
these highly selective membranes very little naphthalene or toluene
escapes through). The pressure of the retentate chamber was controlled
by a back-pressure regulator, while the permeate-side pressure was
always kept atmospheric.
In addition to testing whether the CMS membranes can perform
satisfactorily in the presence of tar-like species, another primary
objective of this project was to experimentally verify that the CMS
membranes and Co-Mo/Al
2
O
3
catalyst can also function stably as
components of an MR system treating biomass-derived syngas for H
2

production through the WGS reaction. The experimental apparatus
used for these experiments is shown in Figure 3.2. The set-up is the
same used in the MR experiments for H
2
production form coal derived
syngas (explained in Chapter 2) with the only difference that we added

76
another trap before the gas chromatograph for removing NH
3
from the
exit streams before they enter the GC, and added Drager tubes to the
analysis section for measuring the NH
3
concentration.

Figure 3.2 Experimental set-up used in H
2
production from biomass-
derived syngas

During the MR experiments, the gas streams exiting the reject and
permeate sides flow through condensers and then through the
moisture-traps in order to capture the water. The flow rates of the

77
water-free streams are then measured by a bubble flow-meter. A small
slip-stream from both the reject and permeate sides is passed through
the Drager tubes to measure its H
2
S and NH
3
content. As noted
previously, the graduated tube used for measuring H
2
S contains a Cu
compound. In the presence of H
2
S, the Cu compound reacts with it and
produces CuS, which results into a color change from blue to black.
The graduated tube used for measuring NH
3
contains a pH indicator.
Changing the pH, because of NH
3
being present in the gas stream,
changes the color of the indicator from yellow to blue. The degree of the
color change in the Drager tubes can be translated into the ppm
concentration level of the H
2
S and NH
3
in the stream. A small slip-
stream from both the reject and permeate sides is passed through two
different commercial adsorption beds in order to remove its H
2
S and
NH
3
content respectively, and is then used to measure the composition
with an online gas chromatograph.
Another key goal of the project was to validate, by WGS-MR
experiments, the isothermal MR model previously discussed in Chapter
2, so that simulations can be performed to identify optimized reactor
operating conditions (in terms of residence time, reactor pressure and
temperature, and steam purge rate) for the various end-use applications.
The results of the experiments and the simulations are as follows.

78
3.3 Results and Discussion
3.3.1 Membrane Stability
During the testing of membrane performance in the presence of
tar (naphthalene) and organic vapor (toluene) surrogate compounds, the
membrane performed well. Figure 3.3, for example, shows the
permeance and selectivity for one of the CMS membranes tested as a
function of time at an operating temperature of 250 ºC and a pressure
of ~2.0 atm. After an initial decay in the membrane permeance, the
performance was stable throughout the remainder of the run. The
initial ~10 to 15% decline in membrane permeance is attributed to
naphthalene/toluene molecules blocking the largest pores in the
membrane. However, pores below a certain size are inaccessible to these
large molecules, and remain free for He (and H
2
) permeation. Very little
impact was observed on the membrane selectivity during the test.
Condensation and accumulation of tar-like species on the membrane
surface leading to significant or even complete blocking of the
membrane surface was, prior to the initiation of our study, a key
concern for potential fouling of the membrane during biomass gasifier
off-gas conditioning and H
2
recovery. Though the membrane tests
above indicated that the membrane is stable under temperature and
pressure conditions akin to the WGS environment, one cannot exclude
situations where, due to various process upsets, the membrane surface

79
will foul due to tar deposition. It is important then that the CMS
membrane can be regenerated under field-implementable conditions
consisting of temperatures below 300 ºC, and in the presence of inert
purge gases such as steam, N
2
, etc.

Figure 3.3 He permeance and He/N
2
selectivity of the CMS membrane
in the presence of naphthalene/toluene as model tar and organic
vapors, respectively, in the biomass gasifier off-gas at T=250
o
C and P=2
atm
During this project, fouling and subsequent regeneration were tested by
exposing the membrane to naphthalene and toluene at low
temperatures to promote surface fouling, followed by regeneration by He

80
purging at 250 ºC. Figure 3.4 shows that considerable permeation loss
occurs when a CMS membrane is operated in the presence of these
model tar and organic vapor compounds at 150ºC. However,
regeneration is readily achieved by purging of the membrane in the
presence of inert gas at 250 ºC. This indicates that at low temperatures
only surface coverage through condensation occurs with little or no
irreversible pore-plugging of the membrane. These results are also
consistent with our previous experience with these membranes in pilot-
plant tests conducted in the presence of VGO hydrocracker and coal
gasifier off-gas.

3.3.2 Membrane Reactor Studies
As noted above, the key goal in this study is to validate the ability of the
WGS-MR, (which is to function as the “heart” of the proposed “one-box”
process) to effectively convert CO in the biomass-derived syngas in the
presence of its common impurities, and to deliver a contaminant-free
hydrogen product.

81

Figure 3.4 Testing of the fouling and regeneration of a CMS membrane
in the presence of naphthalene/toluene as model tar and organic
vapors, respectively, in the biomass gasifier off-gas

All MR experiments reported here were carried out with a simulated
syngas feed with the following composition (H
2
: CO: CO
2
: N
2
: CH
4
: NH
3
:
H
2
S = 0.67: 1.00: 1.00: 2.67: 0.2: 0.00067: 0.0006) which is typical of
the composition of an air-blown biomass gasifier off-gas (Corella et al.
1998). A near stoichiometric H
2
O/CO ratio in the feed of 1.1 is used.
In order to study the H
2
S/NH
3
effect on the MR performance, in this
series of experiments we did not add any tar or organic vapors into the
syngas. Nevertheless, we do not expect these added contaminants to

82
alter any of the conclusions, as the membranes have already proven
robust in their presence.
MR experiments were carried out using a CMS membrane (with
dimensions of 5.7 mm OD, 3.5 mm ID, and 254 mm in length). For
these experiments 10 g of the commercial Co-Mo/Al
2
O
3
catalyst were
diluted with 80 g of ground quartz glass and were packed in the
membrane shell-side. The experiments were carried out at different feed
pressures, with the permeate side pressure under atmospheric
conditions, and with different permeate steam sweep gas ratios (SR),
defined as the ratio of inlet molar flow rate in the permeate side to the
inlet molar flow rate in the feed side. Packed-bed reactor experiments
(as previously described, in these experiments the permeate side is kept
closed) have also been carried out under identical conditions, using the
same feed composition in order to compare the results with those of the
MR experiments.
Prior to the initiation of the MR experiments, the membrane was
characterized through single-gas permeation experiments. The
permeances of the various species are shown in Table 3.1. Previous
studies by our group with CMS membranes indicate that the mixed-gas
permeances of the various gases generally remain fairly close to the
values measured during the single-gas experiments

(Harale et al., 2007).
For the H
2
S and NH
3
syngas contaminants, the MR experiments

83
indicate that they do not substantially permeate through the membrane,
and their permeance was set equal to zero for the simulations (In
extensive studies in which both the surface of the membrane module
and the plumping were specifically coated to avoid potential wall
adsorption, the H
2
S permeance was always found to lie in between the
permeances of N
2
and CH
4
).

Pure Gas
(Before MR experiments)
Pure Gas
(After MR experiments)
Gas
Permeance
m
3
/(m
2
.h.bar)
Separation
Factor
S. F.
Permeance
m
3
/(m
2
.h.bar)
Separation
Factor
S. F.
H
2
1.121 1.00 0.828 1.00
CO 0.024 46.71 0.026 31.85
CO
2
0.062 18.08 0.072 11.5
N
2
0.014 80.07 0.015 55.2
CH
4
0.009 124.56
H
2
O 0.374 3.00

Table 3.1 Single-gas permeation data for the CMS membrane before
and after the MR experiments

84
Figure 3.5 shows the effect of the feed-side pressure on the CO
conversion at different

for the MR and the packed-bed experiments
at 300
o
C and an SR equal to 0.3. Figure 3.6 shows the effect of the
steam SR on the CO conversion for the MR experiments at different

at 300
o
C and a feed pressure of 3 atm (the error bars in both figures
reflect the carbon loss or gain due to experimental errors in measuring
the flow rates and compositions). The lines represent the simulated
conversions using our power-law rate expression (Table 2.1) and the
single gas permeances shown in Table 3.1. In general, the CO
conversion increases with increasing

, feed pressure, and permeate
side steam SR. This is consistent with the general expectation, given the
higher driving force for H
2
permeation and removal from the reaction
zone, and the resultant shift in the production of additional H
2
product.
From these figures, it is clear that there is generally good
agreement between the model predictions and the experimental results.
The MR always delivers higher CO conversions than those of the
packed-bed reactor, ranging from 5%-12% depending on the
experimental conditions used. A key conclusion from these experiments,
lasting for more than two months, is that the catalyst and membrane
exhibited fairly robust behavior. We observed, for example, no notable
changes in catalyst activity.

85

Figure 3.5 Effect of the MR feed-side pressure on CO conversion at
T=300
o
C and SR=0.3

Figure 3.6 Effect of the sweep ratio on CO conversion at
T=300
o
C and P=3 atm
0
25
50
75
100
0 90 180 270 360
CO Conversion%
Wc/Fco (g-cat.h/mol-CO)
MR EXP, P=5 atm
MR EXP, P=3 atm
PB EXP
MR SIM, P=5 atm
MR SIM, P=3 atm
PB FIT
0
25
50
75
100
0 90 180 270 360
CO Conversion%
Wc/Fco (g-cat.h/mol-CO)
MR EXP, SR=0.3
MR EXP, SR=0.1
PB EXP
MR SIM, SR=0.3
MR SIM, SR=0.1
PB FIT

86
The single gas permeances of various species were also measured after
all the experiments were completed (see Table 3.1). The permeance of
the less permeable species (CO and N
2
) changed very little, but the
permeance of H
2
for this particular membrane decreased by 25%
.
(though it is still high for this type of membranes).
The good agreement between the simulations and experiments
makes it possible to use the model to further study the effect of various
parameters on system performance in terms of CO conversion and H
2

recovery. Doing so, allows one to identify the appropriate conditions
which maximize both conversion and product recovery. Figure 3.7
shows, for example, the effect of feed pressure on MR performance.
Increasing pressure helps by increasing both CO conversion and H
2

recovery by increasing the partial pressure difference of H
2
across the
membrane. The pressure effect is more prominent at lower

, where it
compensates for the decrease in contact time between the gas species
and the catalyst due to the high flow rate (in these simulations W
c
is
kept constant). As shown in the figure, increasing pressure increases H
2

recovery to close to 100%. Though these pressures are significantly
higher than the operating range of our bench-scale experimental system,
they are very much in line with the range of pressures biomass gasifiers
usually operate under, and the advantage of using the “one-box ” process
to produce a contaminant-free H
2
is, therefore, obvious.

87
a)

b)

Figure 3.7 Effect of the MR feed-side pressure on a) CO conversion and
b) H
2
recovery at T=300
o
C, SR=0.3, and W
c
=10 g

0
25
50
75
100
0 90 180 270 360
CO Conversion%
Wc/Fco (g-cat.h/mol-CO)
P=5 atm
P=10 atm
P=20 atm
PB
0
25
50
75
100
0 90 180 270 360
Recovery%
Wc/Fco (g-cat.h/mol-CO)
P=5 atm
P=10 atm
P=20 atm

88
Figure 3.8 shows the effect of steam SR on MR performance at P=20
atm. With our laboratory system we were limited to feed pressures up to
~ 6 atm, and thus we used steam as a sweep stream. Under commercial
conditions, however, the reactor will be operated at high pressures (>20
atm), and as Figure 3.8 indicates, the use of sweep is not at all
necessary to obtain the desired conversion and recovery (it should be
noted, however, that low-pressure steam, to be used as sweep, is
plentifully available in the Chemical Industry, and it does not entail the
expense of high-pressure steam to be used as a reactant in high-
pressure WGS reactors).

3.4 Summary and Conclusions
The proposed “one-box” process, in which reaction and separation
are combined in the same unit was successfully utilized for producing
hydrogen from a feed with simulated biomass-derived syngas
composition containing common impurities such as H
2
S and NH
3
. A
CMS membrane was used for the in-situ hydrogen separation. The
membrane was characterized in terms of its gas permeances which were
used for the model predictions. The CMS membrane stability was also
investigated in the presence of naphthalene and toluene as model tar
and organic vapor compounds, and the membrane proved to be stable
at the experimental conditions akin to the WGS reaction environment.

89
a)

b)

Figure 3.8 Effect of the sweep ratio on a) CO conversion and b) H
2

recovery at T=300
o
C, P=20 atm, and W
c
=10 g

0
25
50
75
100
0 90 180 270 360
CO Conversion%
Wc/Fco (g‐cat.h/mol‐CO)
SR=0.3
SR=0.1
SR=0.01
SR=0.001
PB
0
25
50
75
100
0 90 180 270 360
Recovery%
Wc/Fco (g‐cat.h/mol‐CO)
SR=0.3
SR=0.1
SR=0.01
SR=0.001

90
At lower temperatures (e.g., 150
o
C), permeation loss occurred; however,
regeneration was readily achieved by purging of the membrane in the
presence of inert gas at 250 ºC. This indicates that at low temperatures
only surface coverage through condensation occurs, with little or no
irreversible pore plugging of the membrane.
The performance of the MR (the “one-box” process) using such
membranes and catalysts was investigated experimentally for a range of
pressures and sweep ratios; the MR showed higher conversions
compared with those of the traditional packed-bed reactor. Parallel
modeling investigations indicated good agreement with the experimental
data. The performance of the system under different experimental
conditions was further investigated using the model.
The “one-box” process shows several advantages over the
traditional packed-bed system, including improvements in CO
conversion and H
2
purity, while allowing one to perform the reaction in
the presence of hydrogen sulfide and ammonia and being able to deliver
a contaminant-free hydrogen product. Use of the process in hydrogen
production from biomass-derived syngas should, therefore, result in
considerable energy savings.

91
Chapter 4
Ultra Pure Hydrogen Production from Reformate
4.1 Introduction
Fuel cells are widely considered as a clean and energy efficient
technology that shows promise as future potential replacements for the
internal-combustion (IC) engines in vehicles, IC engines and gas
turbines in stationary power generation, and for batteries in portable
power applications. Proton exchange membrane (also known as polymer
electrolyte membrane or PEM) fuel cells are particularly promising as
they are compact and lightweight units that work at comparatively
lower temperatures and pressures, which makes them good candidates
for use in mobile and small-scale, distributed stationary power
generation applications. PEM fuel cells make use of a thin, permeable
polymeric membrane as the electrolyte, and platinum (Pt) electrodes are
used on either side of the membrane in order to catalyze the reactions.
Hydrogen is typically utilized as the fuel fed to the anode, where it
catalytically splits into protons (which transport through the membrane
to the cathode) and electrons that travel along an external load circuit
to the cathode, thus producing electricity. Oxygen (as air) is
simultaneously supplied to the cathode, where it combines with the
hydrogen ions to produce water. The hydrogen, that PEM fuel cells need
to operate can be produced from different renewable (e.g., water,

92
biomass) and non-renewable (e.g., fossil fuels) feed-stocks through
various processes. Due to its highest efficiency, currently steam
methane reforming (SMR) is the primary industrial process that is used
for H
2
production. CO that is produced as an undesirable by-product of
SMR is known to be particularly detrimental for PEM fuel cell operation.
Even at trace concentrations CO being present in the H
2
stream is
reported to degrade PEM fuel cell performance as it poisons severely the
electro-active Pt surface in the anodes, thus preventing H
2
oxidation.

The tolerance of PEM full cells to CO depends on the material used for
preparing the anode. Pt–Ni/C, for example, is tolerant to 50 ppm CO
(Okada et al. 2007), while Pt – Ru on a defect-free carbon nano-tube has
been reported capable to tolerate up to 100 ppm of CO (Yoo et al. 2007).
For the hydrogen produced from SMR to be utilized in PEM fuel cells,
care must, therefore, be paid towards decreasing its CO content prior to
entering the fuel cell stack. Typically, the first step involves utilizing the
water gas shift (WGS) reaction in tandem with SMR to react the CO in
the reformate mixture and to produce additional H
2
.

Conventional WGS
reactors can attain high CO conversions, but not to the level required to
prevent CO-induced PEM fuel cell stack performance degradation. This
then means that following the H
2
/CO
2
separation step (e.g., via
pressure swing adsorption (PSA)) an additional CO removal step must
be implemented involving either the catalytic oxidation of CO (this is

93
known as Preferential Oxidation or PROX) (4-1) or methanation (4-2) in
order to reduce the CO level into the acceptable fuel cell operational
range.
2CO + O
2
→ 2CO
2
4.1
CO + 3H
2
→ CH
4
+ H
2
O 4.2
Both approaches end-up consuming part of the H
2
(since in the
presence of oxygen some catalytic oxidation of the hydrogen is
unavoidable). Methanation offers more simplicity than PROX, since no
oxygen needs to be added to the reactor. On the other hand, it typically
results in higher hydrogen consumption for the same CO conversion.
However, for either process, the lower the CO content

is, the less
hydrogen is consumed and the more efficient the process is. Therefore,
it is highly desirable to be able to produce pure H
2
or H
2
with an
acceptable CO content, without scarifying part of it and increasing the
process complexity.
The above conventional SMR+WGS+PSA+PROX (or methanation)
system is one of significant complexity; to add to this complexity, since
WGS is an equilibrium-limited reaction, usually two reactors are
utilized, one operating at high temperatures (high-temperature shift or
HTS) and the other at low temperatures (low-temperature shift or LTS)

94
in order to maximize the CO conversion. In order to improve on process
efficiency and to minimize system complexity membrane reactors are
employed for the WGS reaction. These reactors combine reaction and
separation in one unit, and in the context of their application for pure
hydrogen production provide many advantages. By removing hydrogen
from the reactor, conversion increases and less steam is required. CO
2

is separated in the reject side at high pressure, thus eliminating the
need for using a separate PSA unit. With highly permselective
membranes the need for using a separate PROX or methanation step
may be completely eliminated, but even for less permselective
membranes the overall system efficiency for the PROX (or the
methanation) step improves and hydrogen loss is minimized. Even the
use of inexpensive, low-pressure steam as a sweep for the permeate side
of these WGS-MR may provide an advantage, as it generates a pre-
humidified H
2
product as a feed, a requirement for many PEM fuel cell
stack applications.
Membrane reactors have been used for both the SMR (Shu et al.
1994; Chen et al. 2008) and WGS reactions (Basil et al. 2001; Iyoha et
al. 2007; Barbieri et al. 2008; Giessler et al. 2003; Harale et al. 2007;
Harale et al. 2010; Abdollahi et al. 2010) for H
2
production. Due to the
high process temperatures required, metal membranes like Pd and its
alloys (Basil et al. 2001; Iyoha et al. 2007; Barbieri et al. 2008)

95
inorganic membranes like microporous silica (Giessler et al. 2003), and
carbon molecular sieve (CMS) membranes (Harale et al. 2007; Harale et
al. 2010; Abdollahi et al. 2010) have been utilized. CMS membranes
provide a great advantage for the WGS application in the context of
Integrated Gas Combined Cycle (IGCC) applications for coal and
biomass, as they are very robust to the various impurities (e.g., H
2
S,
NH
3
, tar, etc.) one encounters typically in such systems, and which are
highly detrimental to the mechanical stability of the metal membranes.
Pd and Pd-alloy membranes are, however, well-suited for high-purity
hydrogen production via the use of SMR for fuel cell applications, due
to their very high (in principle infinite) selectivity towards hydrogen. As
result, a number of groups have investigated their application to the
WGS reaction. Janson et al. (2009), for example, performed a techno-
economic analysis of two process schemes involving the use of Pd
membrane reactors integrated in natural gas based combined power
generation cycles. These include (a) a process utilizing a WGS
membrane reactor, and (b) one that utilized a SMR membrane reactor.
The cycle involving the use of the WGS membrane reactor proved to
provide higher efficiency and lower capital cost. Janson et al. (2009)
attributed this to the fact that the typical feed to an SMR unit contains
methane and steam, with little or no H
2
being present. The WGS feed,
on the other hand, contains considerable amounts of H
2
, which are

96
produced during the SMR reaction. Thus, when using the Pd membrane
for the WGS step, since the driving force for H
2
transport across the
membrane is higher, less of a membrane area is required. They also
note that SMR reactors typically operate at much higher temperatures
when compared to their WGS counterparts. High temperature has a
negative impact on Pd membrane stability due to accelerated membrane
degradation; to counteract this impact one may have to utilize thicker
membranes, increasing the membrane cost, but more importantly also
reducing the H
2
flux.
A number of researchers have studied the use of Pd membrane
for the WGS reaction. Basile et al. (2001) used Pd and Pd/Ag
membranes for the WGS reaction and reported conversion values higher
than thermodynamic equilibrium. The feed to their reactor contained no
inerts or other reaction products one would expect to find in a typical
reformate mixture, and they utilized nitrogen as sweep. Iyoha et al.
(2007) performed the WGS reaction in the absence of any catalyst at a
temperature of 900
o
C and an operating pressure of 241 kPa making use
of Pd and Pd-alloy (80 wt% Pd –20 wt% Cu) tubular membranes. The high
rate of hydrogen extraction through the Pd-based membranes resulted
in increasing carbon monoxide conversions to 93%. CO conversion
decreased from 93% to 66% and hydrogen recovery from 90% to 85%
when the Pd membranes were replaced with Pd-Cu ones, which Iyoha et

97
al. (2007) attributed to the lower permeance of the Pd-Cu alloy. Though
the study of Iyoha et al. (2007) provides the advantage of not needing to
use a catalyst the long-term mechanical stability of these materials at
these extremely high temperatures raises concerns. Bi et al (2009) used
a Pt/Ce
0.6
Zr
0.4
O
2
catalyst with a highly H
2
permeable and selective
Pd/ceramic composite membrane in a WGS membrane reactor. The
system provided CO conversions above thermodynamic equilibrium at
350
o
C, P=1200 kPa and steam to CO ratio=3. The system showed a
decline in H
2
recovery with increasing gas hourly space velocity under
those conditions. The highest reported H
2
purity by the system was
99.7%.
Though some preliminary work has been presented so far on the
use of Pd membranes for the WGS reaction, most of the studies utilized
pure feeds and not typical reformate compositions which in addition to
CO contain substantial amounts of H
2
and CO
2
. In addition typically
inert gas is utilized as the sweep, which is unlikely to be the case in the
commercial case as it will entail substantial additional costs. The
membranes furthermore were of a small laboratory size. In this study
we investigate instead the use of “commercial-scale” Pd membrane,
prepared via electroless plating on 0.05µm pore size asymmetric tubular
ceramic substrate. Figure 4-1 shows a typical SEM cross section of the
Pd membranes provided by Media and Process Technology Inc.

R
to
p
a
th
u
H

F
Realistic fe
ogether w
performanc
and steam
he simulat
used to stu
H
2
recovery

Figure 4-1
eeds with
ith low pr
ce is inves
sweep ra
tion result
udy the ef
y and puri
1 SEM cro
a simulat
ressure st
stigated at
atios. The
ts from a m
ffect of ex
ity of the h
oss section

ted reform
team as a
t different
experime
mathemat
xperimenta
hydrogen p
n of the Pd
support

0.0
Inte
Cera
Pd
mate compo
a sweep. T
t feed pres
ental resul
tical mode
al conditio
product.
d layer dep
t
5µm Pore Size
ermediate Laye
amic Substrate
d Membrane
osition are
The memb
ssures an
lts are com
el. The mod
ons on CO
posited on
e
er
e being us
brane reac
d flow rat
mpared w
del is furth
O conversi
the ceram
98
sed
ctor
tes,
with
her
on,
mic

99
4.2 Experimental Studies
A schematic of the set-up used in this study is shown in Figure
4.2. The set-up is the same used in the MR experiments for H
2

production from coal derived syngas (explained in chapter 2) with the
only difference that there is no need for H
2
S trap before the GC since a
feed with simulated clean reformate composition is used.
A tubular Pd membrane (L = 762 mm, ID = 3.5mm, OD= 5.7mm)
provided by Media and Process Technology, Inc. (M&P), is utilized which
is prepared by depositing a thin Pd layer on a mesoporous alumina
support. The commercially available porous α-Al
2
O
3
tubes are cut into
length, rinsed and dried. The activation and deposition procedure
available in the literature is adopted for Pd deposition (Gade et al. 2008).
The tubes are activated by dipping into the Pd acetate solution for one
minute, then dried and calcined. The substrate thus activated is plated
at room temperature by immersing in the electroless Pd plating bath
with the composition described in the reference. The layer thickness is
dependent upon the deposition time. The thickness of ~4 µm is
prepared for the membranes used in this study. The Pd membranes
thus prepared are rinsed in distilled water overnight and dried, ready
for use. The supported Pd membrane is inserted in the middle of a
stainless steel reactor and sealed with the aid of graphite o-rings and

100
compression fittings. A commercial low-temperature Cu/Zn/Al
2
O
3
(44%
CuO, 44% ZnO, balance Al
2
O
3
) catalyst is utilized. The catalyst particles
are crushed into smaller pieces, and a fraction of the particles with their
size in the range of (600-800 μm) is selected using sieving trays; they
are then mixed with crushed quartz particles (in the same size range)
and loaded into the annular space in between the membrane and the
reactor wall. The reason for diluting the catalyst particles with inert
quartz particles is so that the reactor volume is completely filled (to
avoid flow channeling) but also in order to assure isothermality along
the reactor length.

Figure 4.2 The experimental set-up used in H
2
production from
reformate

101
Feed with a simulated reformate composition (containing H
2
, CO, CO
2

and CH
4
) is prepared (using the mass flow controllers) and flowed
through the feed line toward the membrane at a certain flow rate.
Controlled flow of water are provided by the syringe pumps and directed
toward the steam generating units. During the MR experiments, feed is
preheated to the reactor temperature and directed toward the bed of the
catalyst, where the WGS reaction happens. Steam, used as sweep, is
directed toward the membrane side. H
2
passes through the membrane
and exits the reactor through the permeate line. Remaining gas in the
reactor side is flowing through the reject line. Upon leaving the reactor,
both permeate and reject streams are passing through the condensers,
and then through the moisture traps to remove their water content. The
flow rates of the water-free streams are then measured using the bubble
flow meters. Small slip streams of both permeate and reject are
directing toward the Gas Chromatograph, where their composition is
precisely measured. Single gas permeance measurements are performed
before and after the MR experiments to characterize the membrane, and
also checking the membrane stability. In these experiments, sweep side
is close. Gas flows through the feed line, and the flow rate of the reject
and the permeate sides are measured using the bubble flow meters.

102
4.3 Modeling and Data Analysis
The isothermal co-current model previously discussed in chapter
2 is used to describe the MR performance. The reactor is assumed to be
working isothermally (this also validated experimentally), under ideal
gas law. If membrane is ideal, only hydrogen should pass through it.
Mass transfer for hydrogen through Pd membrane is described with the
following empirical equation (also known as Sieverts law):

4. 3

where
is the H
2
molar flux (mol/m
2
. h), (bar) the H
2
partial
pressure in the feed side, (bar) the H
2
partial pressure in the
permeate side, n the pressure exponent, and

(mol/m
2
. h. bar
n
) the
H
2
permeance.
Ideally, Pd membranes would allow only hydrogen to permeate
through. In reality, often small cracks and defects develop which permit
small amounts of other gases to permeate through. For porous
membranes a value of n=1 is mostly used in calculations, since the
main transport mechanisms are Knudsen diffusion for mesoporous
membranes, and molecular sieving for microporous membranes. Mass
transfer for species rather than H
2
then is the same as shown by Eq.
2.1.

103
H
2
Mass balances in the feed and permeate sides are described by
Eq. 4.4 and Eq. 4.5, respectively.

1

4.4

4.5

where is the H
2
molar flow rate (mol/h) in the feed side,
the
corresponding molar flow rate (mol/h) in the permeate side, V the
reactor volume variable (m
3
), the surface area of the membrane per
unit reactor volume (m
2
/m
3
), the bed porosity in the feed side, the
fraction of the solid volume occupied by the catalysts, the catalyst
density (g/m
3
), and the WGS reaction rate (mol/g. h).
Mass balances for components rather than H
2
in the feed and
permeate sides are the same as described by Eq. 2.2 and Eq. 2.3,
respectively.
The pressure drop in the packed-bed is calculated using the
Ergun equation (2.4-6).
The above mentioned set of equations is solved numerically together
with the boundary conditions of:

104
At V = 0:

and

where
is the inlet feed-side pressure (bar) and

is the inlet molar
flow rate for component j (mol/h).
finally, Eq. 2.7 and 2.8 are used to calculate CO conversion and H
2

recovery respectively.

4.4 Results and discussion
4.4.1 Membrane Characterization
Hydrogen transport through Pd and Pd alloy membranes is
commonly described by Sieverts’ law (Eq. 4.3). The law is consistent
with a mechanism involving the dissociative adsorption of hydrogen on
the Pd surface, diffusion of hydrogen through the metallic lattice, and
recombination of hydrogen atoms and desorption from the downstream
membrane surface. In practical situations external gas-phase mass
transfer limitations, the impact on surface adsorption/desorption of the
presence of various surface impurities and poisons, and Knudsen
diffusion and Poiseuille flows through the membrane also substantially
affect H
2
permeation through Pd membranes (Ward et al. 1999). For a
clean, defect free, thick Pd membrane at high temperatures (>600 K),
diffusion through the bulk metal is likely to be the rate limiting step
and and the Sievert coefficient n=0.5. However, for Pd films less than

105
10 m thick, particularly those prepared on porous ceramic supports,
external mass transfer through the porous layer can provide substantial
added resistance to transport. The presence of impurities on the
membrane surface and at the grain boundaries also have substantial
effect with the net result that the Sievert coefficient can vary depending
greatly on the membrane fabrication method. In such cases it is
common to use Eq. 4.3 for the flux of hydrogen, in which the value of
the pressure exponent (n) is regressed on the experimental data (it
usually varies between 0.5 and 1). Our group has done extensive
experiments in order to find the pressure exponent (n) in Eq. 4.3 for H
2

transport through the ultra long, ultra thin Pd membranes provided by
Media and process Technology Inc. Our results (not shown here)
indicates that the membrane obeys Eq. 4.3 with the exponent n = 0.75-
1. Table 4.1 summarizes the literature data found for the pressure-
dependence of hydrogen flow through palladium and composite
palladium membranes. It also includes the data generated by our group
(Harale 2008; Hwang 2009). Figure 4.3 shows H
2
permeation data at
T=300
o
C with one of the supported Pd membranes used in this study.
The membrane obeys well Eq. 4.3 with the exponent n = 0.96.

106
Membrane
Material
Temperature
(K)
Pressure
Exponent
Reference
Pd 623-723 0.68
Hulbert et al.
1961
Pd-Ag/Al
2
O
3
673 0.76
Uemiya et al.
1991
Pd/Al
2
O
3
573-773 1
Yan et al.
1994
Pd/Al
2
O
3
673 0.6-0.7
Xomeritakis
1998
Pd-Ag 623 1
Amandusson
et al.
2001
Pd 623-1173 0.5-0.62 Morreale et al.
2003
Pd-Ag/Al
2
O
3
673-773 0.5 Gue et al.
2003
Pd-Ag/ -Al
2
O
3
473-616 0.968
Chen et al.
2003
Pd-Cu/ceramic 623-723 0.8-1
Roa et al.
2003
Pd/Al
2
O
3
723 1
Harale
2008
Pd/Al
2
O
3
523-723 0.75
Hwang
2009

Table 4.1 Summary of published data for H
2
permeation through
palladium membranes

107

Figure 4.3 Flux of hydrogen as a function of () ( )
Fn P n
PP  at T=300
o
C

CO is known to affect H
2
permeance of the Pd membrane by adsorbing
on the membrane surface especially at lower temperatures (<250
o
C)
(Amandusson et al. 2001). For a feed consist of CO and H
2
, almost no
H
2
permeates the Pd membranes at temperatures below 150
◦
C. For
temperatures of 300
◦
C and above, however, CO has almost no effect on
the hydrogen permeation (Amandusson et al. 2001). We performed a
series of the experiment to study the effect of CO on H
2
permeance of
the Pd membrane under study. We exposed the membrane to CO for a
certain amount of time, and then suddenly exposed it to H
2
to see
whether we can attain the initial H
2
flux at the same pressure. Figure

108
4.4 shows the results at 300°C. As shown in the figure, for all the cases,
the H
2
flux almost instantly reaches the same value to that of
before CO exposure, no matter how long the membrane was exposed to
CO. No required regeneration time agrees with the literature data stated
above, that no CO adsorption occurs at 300
o
C.

Figure 4.4 Effect of CO on membrane H
2
permeation

109
As mentioned before, only atomic hydrogen can permeate through
the Pd metal lattice, with all other gases found in a reformate mixture
having negligible solubility and diffusivity; perfect Pd membranes,
therefore, should in principle have nearly infinite selectivity towards H
2
.
That they do not, is indicative of the fact that other gases permeate
through defects and metal lattice imperfections that are often
unavoidable during the preparation of ultra-thin layer Pd membranes.
The hydrogen permselectivity is then an important index of the quality
of the resulting Pd membranes. Table 4.2 shows the measured single
gas permeances and calculated separation factors for the Pd membrane
utilized in the MR experiments measured first before the experiments
were initiated and then again after the experiments were completed
(during which time the membrane stayed for more than one month on
stream). As Table 4.2 indicates, the membrane under study possesses
high hydrogen permeability and very large hydrogen permselectivity
towards the other gases found in the reformate mixture which is the
feed for the WGS reactor. Comparing the hydrogen permeation rates
measured before and after the MR experiments, there is a small
difference between the values measured (~6% decrease in H
2

permeance); however, the permeation rates of the other gases remained
virtually unchanged, indicating a satisfactory stability of the Pd
membranes under WGS environment.

110

Gas
Before MR Exp. After MR Exp.
Permenace
(m
3
/m
2
.h.bar)
Separation
Factor (S.F)
Permenace
(m
3
/m
2
.h.bar)
Separation
Factor (S.F)
H
2
16.8200 1 15.7420 1
CO 0.0071 2369 0.0070 2248
CO
2
0.0057 2950 0.0058 2714
Ar 0.0068 2473 0.0064 2459
N
2
0.0066 2548 - -
CH
4
- - 0.0069 2281

Table 4.2 Measured single gas permeances and calculated separation
factors for the Pd membrane before and after MR experiments

4.4.2 Reaction Studies
A commercial Cu/Zn-based supported LTS catalyst has been
used in this study. This type of catalyst has been extensively utilized for
the low temperature WGS reaction. Extensive kinetic investigations for
an expanded range of pressures, temperatures, and feed compositions
with a Cu-Zn supported catalyst of a very similar composition were
previously carried out by our group (Harale et al. 2010). A reaction rate
expression consistent with a Hougen–Watson type surface mechanism

111
was developed (4.6-7), where the various rate parameters are
summarized in Table 4.3.

.
.
.
.

.

.

^
4.6

.

.

4.7

This rate expression also satisfactorily describes the packed-bed
experimental data derived with the catalyst used in this study, as
Figure 4.5 indicates. Before its use in these experiments, the catalyst
was activated. The catalyst was reduced by passing N
2
as carrier gas
through the catalyst bed, and adding controlled amounts of hydrogen at
certain temperatures as instructed by the manufacturer. The hydrogen
reduces the copper oxide in the catalyst to copper metal. The activation
is complete when the inlet and outlet H
2
concentrations are almost the
same.
E 30.387 (kJ/mol)
k
0
48.16 (mol/g/s/atm
2
)

−12 0.0178

−28 0.0410

−0.86 0.0303

Table 4.3 Hougan Watson rate parameters

112

Figure 4.5 Measured vs. calculated CO conversion data using the
Hougen –Watson rate expression

Membrane reactor experiments were carried out using the
experimental system depicted in Figure 4.2, utilizing a “commercial size ”
M&P supported Pd membrane (L = 762 mm, ID = 3.5 mm, OD= 5.7 mm),
whose transport characteristics are shown Table 4.2.
As noted above, past studies, utilizing Pd membrane reactors for
the WGS reaction have used feed compositions very different from what
are used industrially. In this study, the feed gas consists of all species
potentially present in the stream exiting the SMR reformer upstream of

113
the WGS reactors. For the membrane reactor experimental data
presented here, in particular, the feed had the following composition (in
terms of molar ratios) H
2
: CO: CO
2
: CH
4
: H
2
O = 5.22:1:0.48:0.1:2.8,
which is chosen to match the calculated equilibrium conversion from
an SMR reformer operating at a temperature of 850
◦
C and a pressure of
1000 kPa. For the membrane reactor experiments, 30 g of the
commercial Cu/Zn/Al
2
O
3
catalyst were diluted with ground quartz
glass (enough to fill the reactor volume) and were packed in the
membrane shell-side (the annular volume between the membrane and
the reactor shell). The reactor was maintained under isothermal
conditions, and the permeate side was maintained under atmospheric
pressure conditions. Steam was utilized as the sweep gas stream.
In the experiments the effect of varying the feed-side (reactor)
pressure and permeate side sweep ratio (SR= ratio of inlet steam molar
flow rate in the permeate side to the inlet molar feed flow rate in the
feed side) on reactor conversion and recovery are studied.
Figures 4.6a and b show the CO conversion and H
2
recovery as a
function of W
c
/F
co
(weight of undiluted catalyst (g) over the molar flow
rate of CO (mol/h)), for a temperature of 300
◦
C and SR=0.3 for two
different operating pressures namely 308 and 446 Kpa. Shown on
Figure 4.6a is also the equilibrium conversion based on the feed
composition. Clearly the removal of hydrogen from the reaction mixture

114
has a substantially beneficial effect on reactor conversion. Increasing
the W
c
/F
co
and the reactor pressure significantly improves conversion
and hydrogen recovery, as expected, with almost complete conversions
and hydrogen recoveries in excess of 90% being attained. Figures 4.7a
and b show the CO conversion and H
2
recovery as a function of W
c
/F
co

for a temperature of 300 ◦C, a reactor pressure of 446 kPa and three
different sweep rates, namely SR=0 (no sweep), SR=0.1, SR=0.3 (shown
on the Figure 4.7a is also the calculated equilibrium conversion). As
shown in these two figures, increasing sweep ratio increase both
conversion (up to almost complete conversion) and recovery (up to
almost 90%), but the effect is stronger for the recovery. Increasing the
sweep ratio (as is the case with increasing the feed pressure – see figs
4.6a-b) increases the hydrogen partial pressure difference across the
membrane, thus providing higher driving force for H
2
to diffuse through
the membrane. The use of sweep is of particular value for the smaller,
and industrially more relevant W
c
/F
co
.
Another important indicator of good reactor performance is the
purity of the H
2
product, since as noted above, fuel cells are very
vulnerable to the presence of CO even in small quantities (ppm level).
Table 4.4 shows the CO content of H
2
product from the membrane
reactor (the stream exiting the permeate side) for a number of different
experiments. The CO concentration in the hydrogen product for all

115
experiments reported in Table 4.4 remains below 75 ppm, which is
acceptable for some types of PEM fuel cells, as noted previously (Yoo et
al. 2007). For other types of PEM fuel cells which require CO
concentrations below 50 ppm, this could be conveniently accomplished
by CO methanation, since the low (<75 ppm) CO content in the
permeate side stream implies that very little H
2
is going to get
consumed during the CO methanation.

4.5 Reactor Design and Scale-up
Due to limitations in terms of the membrane area and the range
of pressures that could be employed in our experimental system it was
not possible to completely explore via experimentation the full range of
capabilities of Pd membrane reactors for the WGS reaction. Instead, in
this study we have made use of the reaction rate expression (Eq. 4.6)
and the measured single gas permeances in order to investigate,
through simulations utilizing the mathematical model described above,
the behavior of the WGS-MR system for conditions more akin to the
industrial ones (e.g., in terms of pressure and feed flow rates). The
model, as Table 4.5 (which presents measured and calculated CO
conversions and H
2
recoveries for two sets of the experimental
conditions) indicates provides good agreement between measured and

116
calculated quantities, thus lending credence to the findings of the
simulations.
a)

b)

Figure 4.6 Effect of pressure on a) CO conversion and b) H
2
recovery,
T=300
◦
C and SR=0.3
85
90
95
100
100 200 300 400 500
CO Conversion %
Wc/Fco (gr‐cat.h/mol‐CO)
P=446 kPa
P=308 kPa
Xe
70
75
80
85
90
95
100
0 100 200 300 400 500
Recovery %
Wc/Fco (g‐cat.h/mol‐CO)
P=446 kPa
P=308 kPa

117
a)

b)

Figure 4.7 Effect of sweep ratio (SR) on a) CO conversion and b) H
2

recovery, T=300
◦
C and P=446 kPa
85
90
95
100
100 200 300 400 500
CO Conversion %
Wc/Fco (gr‐cat.h/mol‐CO)
SR=0.3
SR=0.1
no sweep
Xe
70
75
80
85
90
95
100
0 100 200 300 400 500
Recovery%
Wc/Fco (g‐cat.h/mol‐CO)
SR=0.3
SR=0.1
no sweep

118
T=300
o
C, P=446 kPa
no sweep

T=300
o
C, P=446 kPa
Wc/Fco=466 (gr-cat.h/mol-CO)

Wc/Fco
(gr-cat.h/mol-
CO)

CO
Concentration
(ppm)

Sweep Ratio
(SR)

CO
Concentration
(ppm)

466 73 0.3 59
266 55 0.1 68
186 41 None 73

Table 4.4 CO content of the MR H
2
product (permeate side) at two
different sets of the experimental conditions

W
c
/Fco
Conv.
(Exp.)
Conv.
(Sim.)
Rec.
(Exp.)
Rec.
(Sim.)
T=300
o
C, P= P=446 kPa, SR=0.3
266 99.7 98.6 88.9 90.3
T=300
o
C, P=446 kPa, SR=0.1
266 99.6 98.2 87.1 86.9

Table 4.5 Measured and calculated CO conversion and H
2
recovery for
two sets of the experimental conditions

119
Figure 4.8 shows the effect of feed pressure on membrane reactor
conversion. Increasing reactor pressure has a significant effect on
conversion especially at lower Wc/Fco, where it compensates for the
decrease in contact time between the gas species and the catalyst due
to the high flow rates (in these simulations W
c
is kept constant). The
effect levels off, however, at the higher Wc/Fco. These results are very
promising, considering the fact that high pressures (~ 20 atm and
higher) and low Wc/Fco (<10) are the conditions which are currently
applied industrially. Increasing pressure has also the additional
beneficial effect in that it increases H
2
recovery (Fig. 4.9) by increasing
the partial pressure difference of H
2
across the membrane.

Figure 4.8 Effect of pressure on CO conversion, T=300
◦
C and SR=0.1
0
20
40
60
80
100
02468 10
CO Conversion %
Wc/Fco (g‐cat.h/mol‐CO)
P=1000 kPa
P=1500 kPa
P=2000 kPa

120

Figure 4.9 Effect of pressure on H
2
recovery, T=300
◦
C and SR=0.1

Another very important factor determining good MR performance
is the H
2
purity, as previously noted. As Figure 4-10 shows, increasing
pressure decrease the H
2
purity considering the fact that it provides
higher driving force for other species rather than H
2
to diffuse through
the flaws of the membrane. The effect becomes stronger at higher
Wc/Fco where the flow rate is low and the contact time between the gas
and the membrane is higher allowing for a relatively larger fraction of
CO to transport through the membrane to the permeate side without
undergoing reaction. The purity of the hydrogen product for realistic
0
20
40
60
80
100
0 246 8 10
Recovery %
Wc/Fco (g‐cat.h/mol‐CO)
P=1000 kPa
P=1500 kPa
P=2000 kPa

121
pressures and Wc/Fco is very high, indicating that the use of Pd-MR for
industrially relevant conditions is indeed quite promising.

Figure 4.10 Effect of pressure on H
2
purity, T=300
◦
C and SR=0.1

4.6 Summary and Conclusion
In this study, an ultra-thin, long, high-performance (in terms of
H
2
permeance and selectivity) palladium membrane is used in a
membrane reactor system to produce pure hydrogen through the use of
the water-gas shift reaction from a gas stream with simulated reformate
composition. The membrane is characterized using the single gas
permeance measurements. A Cu-Zn/Al
2
O
3
catalyst is utilized for the
99.8
99.85
99.9
99.95
100
0 246 8 10
Purity %
Wc/Fco (g‐cat.h/mol‐CO)
P=1000 kPa
P=1500 kPa
P=2000 kPa

122
WGS reaction. The system performance is investigated under various
experimental conditions, namely, different pressures, feed flow rates
and sweep ratios. The best performance is obtained at T=300
o
C, P=446
kPa and the permeate sweep gas ratio=0.3 with almost complete CO
conversion and 90% hydrogen recovery. The product hydrogen purity is
always at more than 99.9% with CO concentration of less than 100 ppm.
A model is developed and its predictions are compared with the
experimental data, with good agreement between the two. The model is
used for further studying the design aspects of the system. It is shown
that the Pd membrane reactor system under study is capable of
delivering almost complete CO conversion and H
2
recovery at the
experimental conditions akin to the industrial applications. The
membrane exhibits significant stability with only 6% change in H
2

permeance and almost no change in the H
2
selectivity after using the
system for more than a month under WGS environment. Hence, it is a
very efficient system for hydrogen production for fuel cell applications.

123
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Abstract (if available)

Abstract Hydrogen is widely considered as a fuel for the future to address the energy crisis and the environmental concerns. In this study, process intensification in hydrogen production from natural gas, coal- and biomass-derived syngas is investigated both experimentally and theoretically. A novel reactor/separator system, termed the “one-box” process is being employed. The heart of this system is a membrane reactor (MR) that combines the water-gas shift (WGS) reaction with hydrogen separation into a single unit, thus eliminating the need for the commonly utilized two separate WGS reactors and a distinct purification unit.

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Asset Metadata

Creator Abdollahi, Mitra (author)

Core Title An integrated 'one-box' process for hydrogen production

School Viterbi School of Engineering

Degree Doctor of Philosophy

Degree Program Chemical Engineering

Degree Conferral Date 2011-05

Publication Date 11/03/2011

Defense Date 03/11/2011

Publisher University of Southern California (original), University of Southern California. Libraries (digital)

Tag carbon membrane,hydrogen production,membrane reactor,OAI-PMH Harvest,palladium membrane,water-gas shift reaction

Language English

Contributor Electronically uploaded by the author (provenance)

Advisor Tsotsis, Theodore T. (committee chair), Pirbazari, Massoud M. (committee member), Sahimi, Muhammad (committee member)

Creator Email abdollah@usc.edu,mitra_abd100@yahoo.com

Permanent Link (DOI) https://doi.org/10.25549/usctheses-m3883

Unique identifier UC1165007

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Dmrecord 471342

Document Type Dissertation

Rights Abdollahi, Mitra

Type texts

Source University of Southern California (contributing entity), University of Southern California Dissertations and Theses (collection)

Repository Name Libraries, University of Southern California

Repository Location Los Angeles, California

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