IGEC-1
Proceedings of the International Green Energy Conference
12-16 June 2005, Waterloo, Ontario, Canada
Paper No. IGEC-1-Keynote-Elnashaie
EFFICIENT PRODUCTION AND ECONOMICS OF THE CLEAN FUEL HYDROGEN
Said Elnashaie
Chemical Engineering Department, Auburn University, Auburn, Alabama 36849, USA,
E-mail: nashaie@eng.auburn.edu
&
Chemical and Biological Engineering Department, University of British Columbia, Vancouver, British
Columbia, Canada. E-mail: nashaie@chml.ubc.ca
ABSTRACT
This paper/plenary lecture to this green energy
conference briefly discusses six main issues:
1) The future of hydrogen economy.
2) Thermo-chemistry of hydrogen production for
different techniques of autothermic operation
using different feedstocks.
3) Improvement of the hydrogen yield and
minimization of reformer size through
combining fast fluidization with hydrogen and
oxygen membranes together with CO2
sequestration.
4) Efficient production of hydrogen using novel
Autothermal
Circulating
Fluidized
Bed
Membrane Reformer (ACFBMR).
5) Economics of hydrogen production.
6) Novel gasification process for hydrogen
production from biomass.
It is shown that hydrogen economy is not a Myth as
some people advocate, and that with well-directed
research it will represent a bright future for humanity
utilizing such a clean, everlasting fuel, which is also
free of deadly conflicts for the control of energy
sources. It is shown that autothermic production of
hydrogen using novel reformers configurations and
wide range of feedstocks is a very promising route
towards achieving a successful hydrogen economy.
A novel process for the production of hydrogen from
different renewable biomass sources is presented
and discussed. The process combines the
principles of pyrolysis with the simultaneous use of
catalyst, membranes and CO2 sequestration to
produce pure hydrogen directly from the unit. Some
of the novel processes presented are essential
components of modern bio-refineries.
INTRODUCTION
It is expected that by 2050, the world will double its
energy demand. Due to environmental awareness
the world also needs low-emission and low-carbon
energy since people are very unlikely to tolerate
increased pollutions and their possible effects on
the climate. It will take more than a decade before
alternatives can compete effectively with fossil fuels.
It is certain that there is a viable exciting future for
renewables sources of energy. Many alternatives
are competing and most probably the final outcome
will be an optimal blind of all of them. The most
promising sources today are: Solar energy, Wind
energy, Biomass, Biofuels, Geothermal energy,
Hydro-electricity and Hydrogen. Our concentration
in this paper and the plenary lecture is on
Hydrogen, which offers one of the most challenging
prospects for this century: sustainable and
emission-free energy. In order to generate energy
from hydrogen, a “fuel cell” is usually used; direct
efficient clean combustion of hydrogen is also
possible. The different views regarding “Hydrogen
Economy” are briefly discussed, while most of the
paper and the lecture are devoted to the efficient
production of clean hydrogen from different
feedstocks aiming not only at the optimization of
specific processes and designs but also at
conceptual
optimization
challenging
the
configurations and basic fundamentals of the
processes. A sequential de-bottlenecking approach
is utilized to reach the most efficient processes,
including the breaking of traditionally established
barriers against the increase of efficiency and
productivity. Some of the processes developed and
discussed are essential parts for modern biorefineries. The use of hydrogen as an energy carrier
or major fuel requires development in several
industrial segments, including:
1) Production of hydrogen from fossil fuels,
biomass, or water involves thermal, electrolytic,
and photolytic processes.
2) Distribution of hydrogen from production and
storage sites involves pipelines, trucks, barges,
and fueling stations
3) Storage of hydrogen for delivery, conversion,
and use involves tanks for both gases and
liquids at ambient and high pressures and
reversible/irreversible metal hydride systems
3) Conversion for making of electricity and/or
thermal energy involves combustion turbines,
reciprocating engines, and fuel cells.
Each industrial segment is an integral part to
the building of a hydrogen-based economy, and
the development of one segment relies on
corresponding developments of all others.
The End-Use of hydrogen includes: portable
power in different devices, energy applications and
computers, transportation systems: such as fuel
additives, fuel-cell vehicles, internal combustion
engines, and in propulsion systems for space
shuttles. It also includes the use for stationary
energy generation systems.
necessitating very high steam to carbon ratio.
This can be overcome by special reactorregenerator configuration.
HYDROGEN ECONOMY, REALITY OR MYTH?
The hydrogen economy, reality or a myth?, this is a
very difficult question to answer, for it is very clear
that extremely varying views are strongly
expressed. If it were only views it would have been
similar to any philosophical or ideological problems
or even religious or political problems we face, but
actually the amount of capital invested in order to
make hydrogen economy a reality is tremendous.
Billions of dollars are spent by companies as well as
funding agencies in USA, Canada, Europe, Japan,
etc. Many Universities around the world have
established centers of research excellence for
hydrogen research and established distinguished
Professorship for the same purpose.
PRODUCTION OF HYDROGEN
Produced hydrogen is either consumed on site
(“captive” hydrogen) or distributed via pipelines or
trucks (“merchant” hydrogen). Hydrogen does not
naturally exist in its elemental form on Earth, it must
be produced from other compounds. Each method
of production requires energy in some form, such as
heat, light, or electricity, to initiate the process. The
energy content of hydrogen is much higher than the
energy necessary to release it as shown later. In the
United States, approximately 95 percent of
hydrogen is currently produced via Catalytic Steam
Reforming (CSR) using Nickel catalyst, extracting
hydrogen from both hydrocarbons and steam, the
net reactions are highly endothermic. The reactions
are very fast but are limited by thermodynamic
equilibrium and in conventional steam reformers it
requires very high temperatures to increase the
equilibrium
conversion.
Cracking
of
the
hydrocarbons at these high temperatures causes
carbon formation, which necessitates the use of
high steam to hydrocarbon ratio to avoid catalyst
deactivation(Elnashaie and Elshishini, 1993). The
carbon formation increases with the use of higher
hydrocarbons and renewable feedstocks such as
biomass and bio-oils. This inefficient CSR process
for hydrogen production suffers from four main
bottlenecks to be overcome through well-directed
multidisciplinary research (Elnashaie and Garhyan,
2003):
1) Diffusional limitations: associated with the
catalyst and expressed as the effectiveness
factors, necessitating the use of small catalyst
particles and special reactor configuration to
remove these severe limitations (Sammels,
et.al., 2000;Sahd and Drnevich, 2000; Chen
and Elnashaie,2002,2003c,2004; Chen et.al.,
2003a,b;2004a,b; Prasad and Elnashaie,2002).
2) Thermodynamic limitations: due to the
reversibility of the reactions, necessitating the
use of high temperatures to increase the values
of the equilibrium constants. The other
alternative is to “break the thermodynamic
limitations”
using
hydrogen
selective
membranes and/or CO2 sequestration (BrunTsekhovoi, et.al, 1988; Prasad and Elnashaie,
2004)
3) Thermal
limitations:
due
to
high
endothermicity of the reactions. This can be
overcome by using “autothermic” operation
(Chen and Elnashaie, 2005a; Prasad and
Elnashaie, 2005)
4) Catalyst deactivation: due to carbon formation
associated with hydrocarbons cracking and
Some views are very positive regarding the
future of hydrogen economy and billions of
dollars are spent to achieve that. The basic
arguments in favor of hydrogen economy are :
The elimination of pollution caused by fossil
fuels, the elimination of greenhouse gases, the
elimination of economic dependence
and
distributed production. They also emphasize that
the problems with the fossil fuel economy are so
great, and the environmental advantages of the
hydrogen economy so significant, that the push
toward the hydrogen economy is very strong. For
example Senator Dominici, one of the strong
advocates of hydrogen economy and a strong
supporter for President Bush’s hydrogen initiative,
sees the development of hydrogen technology as
an essential component of establishing American
energy independence in the 21st Century. He says:
“Hydrogen technology may truly revolutionize our
transportation technology for years to come,” and
also adds: “It offers the promise of perfectly clean
transportation”. However, he also notes the serious
research and technological hurdles currently facing
government and industry, noting that the $2–billion
investment in the energy bill for hydrogen research
will be just a fraction of what it would take to bring
this idea to fruition. Another strong supporter of
hydrogen economy is Senator Dorgan, he says:
“Shouldn’t
we
together,
Republicans
and
Democrats…understand the peril to our economy
with respect to the direction we are headed in our
reliance on oil from troubled parts of the world?”
Noting we still fuel our cars the same way we did
100 years ago, he expressed his support for
establishing timelines for development of hydrogen
technology.
Some views are balanced, seeing benefits
associated with the hydrogen economy and at
2
the same time seeing great difficulties and
expecting research to solve these problems
(e.g., Dr. Robert Hirsch, who is a senior energy
program advisor at SAIC, and who has held
senior positions in industry, government and
the non-profit sector and who also served in
numerous
advisory committees
and
is
immediate past chairman of the Board on
Energy and Environmental Systems of the NST
of the United States). His view as briefly
expressed in his article in Chemical Engineering
Progress (CEP) on November 2004 can be
summarized in the following points:
All negative sides:
 An H2 “mirage” collapsed 30 years ago after
wasting $7 billion for the US and another $6
billion for Europe and Japan.
 H2 is an inefficient energy carrier not an energy
source.
 Conversion of fossil fuels to hydrogen
inherently involves a large irreversible energy
loss.
 Converting solar or nuclear energy to H2 is a
thermodynamic crime and technical insanity.
 Any use of H2 as a substitute for a fissile fuel
will increase our( meaning US) dependence on
imported oil and gas as well as increase global
emission.
 H2-powered cars would increase both
greenhouse and smog-forming emission by a
factor of 5.
 H2 is the most dangerous fuel.
Positive sides:
1. H2 supply on earth is infinite
2. In internal combustion engines H2 burns without
emission of CO2, particulates and SOx .
3. In fuel cells, H2 can be converted to electric
powers at attractive efficiencies with negligible
emissions.
Suggested alternative solutions:
 Reduce fuel consumption, e.g.: eliminate gas
guzzling SUVs and forcing everybody to buy a
hybrid car.
 Build thermal solar and nuclear plants.
Professor Shinnar also adds that the only purpose
of the H2 car is to divert attention from the above
two solutions he is suggesting. But why is this
“Conspiracy” against these two solutions, Professor
Shinnar only tell us about a deep psychological
need in the US society!!, but why also in Canada,
Europe and Japan we have no answer to that. He
also considers the fuel cells a fantasy, and a nonstarter. In the January 2005 issue of CEP, we find
in pages 4-6, seven letters to the Editor supporting
and adding to the views of Professor Shinnar.
However in the same issue of the same Journal we
find an entire page devoted to the promotion of
hydrogen economy under the title: “Hydrogen Fuel
Cell Vehicles: Driving Towards Reality”, which is
referring to the successful efforts of Dave Austgen
the vice-president of Technology for Shell
Hydrogen, the entire page is packed with prohydrogen economy and fuel cells arguments. They
also refer the reader to page 116 of the same issue
of CEP, where we find Shell celebrating its first new
hydrogen filling station in Washington, DC.
Negative sides:
1. H2 is an energy carrier and not a fuel that
occurs in nature, implying that primary energy
from an external source must be invested to
produce hydrogen.
2. H2-fueled systems have large economic
barriers to overcome.
3. H2 systems are also still facing problems with
regard to safety, and public acceptance.
Some Preliminary Conclusions:
The problems facing H2 as a fuel today are not
eternal and innovative research can solve them.
The world oil production peak followed by decline,
this decade or the next decade will change the
situation in favor of hydrogen as a fuel. The higher
emphasis on the environment will push forward
hydrogen as the most clean and efficient fuel.
Research should continue for improving technically
and economically the hydrogen production
processes.
Some Views are negative regarding the future of
hydrogen economy and are calling the hydrogen
economy a Myth. One of the most
“antihydrogen” views is by Dr. Reuel Shinnar , who
is a distinguished Professor of Chemical
Engineering at City College. He has taught and
did research for 40 years and published more
than a 100 papers and introduced powerful
design and economic evaluation techniques. His
view as expressed in his article in Chemical
Engineering Progress (CEP) on November 2004
can be summarized in the following points:
Some of the other important facts contradicting
the views of Shinnar are available in the same
issue of CEP :
Under the title: “Making Fuel Cells More Affordable”
there is a description of the efforts by PolyFuel, Inc.
(Mountain View, CA) in co-operation with other
companies to make the hydrogen fuel cell more
economically competitive. Also Dow Chemical
Co.(Midland, MI) and GM (Detroit; MI) launch
Phase II of their collaborative effort to prove the
3
viability of hydrogen fuel cells for vehicles and
distributed power generation. In addition BOC, Inc.
(Murray Hill, NJ) is partnering with Membrane
Reactor Technologies, Inc. (MRT; Vancouver,
British Columbia, Canada;) and HERA Hydrogen
Storage Systems Inc. (HERA; Longueuil, Montreal,
Canada) to develop a low-cost hydrogen generation
and delivery systems. The US Department of
Energy (DOE, Washington, DC) is contributing $3.5
millions over 3-yr period to fund the project.
On response to this situation, and before
starting to write this paper for the plenary
lecture to this International Green Energy
Conference (IGEC) I contacted a large number of
academicians and companies and federal
agencies involved in R&D for hydrogen
economy to get their feed back. From the large
number of pro-hydrogen economy responses I
got , I give below a summary of the most
interesting answers, the first view is from an
international hydrogen organization and the
second is from an academician very active in
hydrogen R&D:
Prof. Shinnar failed is to look at a bigger picture.
Energy system of the future is something that is
going to replace the fossil fuels hopefully before we
run out of them or before the environment runs out
of capacity to absorb pollution created by burning
fossil fuels. It is therefore a wrong comparison,
comparing energy system of today with the future
energy system, and concludes that this one is
better. If this one was good and could run forever
we would not need to look for another one”.
Second View: “Hydrogen is already a major
industrial intermediate used for many purposes.
Much of the research on it would be useful and
should continue whatever happens with respect to
fuels or fuel cells. Yes of course one must consider
the
CO2
production
and
thermodynamic
inefficiencies with respect to H2 production. But it
must be considered in the context of the overall life
cycle analysis, e.g. if starting with natural gas from
well to ultimate use in stationary or mobile
applications. When the most favorable hydrogen
production technology is combined with the most
favorable fuel cell technology, there are indications
that some thermodynamic gains, admittedly small,
may be realized. The advantages of engines, which
emit essentially only H2O in the urban environment,
could be substantial for health, aesthetics, buildings,
etc. Other benefits include reductions in noise. The
technologies being developed now can in principle
be extended readily to biomass and various wastes,
reducing greenhouse gas emissions further. I would
personally ban SUVs and we can certainly make
better use of alternative forms of energy. But a
multi-faceted approach is needed, and hydrogen
provides some options where it may well be
important as an energy carrier”.
First View: “Ever since hydrogen has come up in
the midst of attention, there are critics, and most of
the time they are right. However, they are right in
criticizing their own wrong perception of hydrogen
economy. Their arguments are unbeatable. For
example, Prof. Shinnar says “hydrogen is not an
energy source” and he is absolutely right. No
scientist could have ever said that hydrogen is an
energy source. This may have come up from some
politician who did not understand the concept of
hydrogen economy. Prof. Shinnar also says that
production of hydrogen from fossil fuels does not
make sense. I agree with that one too. He also says
that there should be more emphasis on energy
efficiency. I could not agree more. Hydrogen
economy does not mean continuation of wasteful
practices implemented when plentiful and cheap
fossil fuels were available. On the contrary it means
very frugal use of valuable energy. Production of
hydrogen from electricity does not make sense from
thermodynamic point of view. Correct again. But a
very same argument could be used on generation of
electricity from fossil fuels - we waste 2/3 of energy
in fuel to produce electricity. From thermodynamic
point of view that is very inefficient. Yet we are
doing it and it is economical because electricity can
do many things coal cannot and we are willing to
pay for it. No scientist proposes to convert electricity
to hydrogen and then hydrogen to electricity if that
electricity could be used directly. Hydrogen comes
into the picture when and where electricity cannot
be used directly. The correct title of the future
energy system therefore should be Electric and
Hydrogen Economy. Such a system can be made to
work with any energy sources of the future. Where
Provisional Conclusion: From the brief discussion
above, it will be safe to conclude that hydrogen is a
very promising clean fuel and that the obstacles facing
hydrogen economy can all be solved through extensive
well direct multidisciplinary research covering all
aspects of production, storage, transportation,
conversion, etc. In this paper and plenary lecture we
will concentrate on production, but before that it will be
interesting to have a definition for the hydrogen
economy. It seems that the Shell Hydrogen definition
is very suitable: “The hydrogen economy is a
world fundamentally different than the world
we know now. In the hydrogen economy
hydrogen
is
available
to
everyone,
everywhere—from the corner fueling station
to the large industrial facility on the outskirts
of town. Countries will not be dependant
anymore on a single source of fuel. Hydrogen
is produced, cleanly and cost-effectively,
from a variety of sources like renewables,
4
other words: the right model is the one with the
optimum degree of sophistication (Rutherford
Aris).
Thermal Demand of the Highly Endothermic
Reforming Reactions and Autothermicity
Reforming reactions are the best for hydrogen
production, mainly because they extract the
hydrogen of both the hydrocarbon and the water, as
an example for methane:
such as biomass and water, fossil fuels or
even nuclear energy, using advanced
technologies to ensure that any carbon
released in the process does not escape into
the atmosphere. Hydrogen is delivered and
stored routinely and safely. Hydrogenpowered fuel cells and engines are as
common as the gasoline and diesel engines
of the late 20th century—they power our cars,
trucks, buses and other vehicles, as well as
our homes, offices and factories”.
CH4+ H2O(g)
CO + 3H2 (H0298 = 206 kJ/mol),
Energy Cost /mole H2 = 68.7 kJ/mole H2
(1)
CH4 + 2H2O(g)
CO2 + 4H2 (H0298 =165 kJ/mol),
Energy Cost /mole H2 = 41.25 kJ/mole H2
(2)
EFFICIENT HYDROGEN PRODUCTION USING
NOVEL REFORMERS
Reaction (1) uses ~ 67% more energy per mole H2
produced than reaction (2). Reaction (2) is the most
efficient for hydrogen production (4moles H2 per
mole CH4) and energy efficiency. Breaking the
thermodynamic equilibrium by membranes and/or
sequestration push towards this reaction (2) by
pushing the exothermic shift reaction forward:
CO(g) + H2O(g)
CO2(g) + H2(g)
0
(H 298 = - 41 kJ/mol, exothermic)
(3)
Main Approach and Methodology
The approach used for R&D depends upon a
number of principles, which can be very briefly
summarized in the following:
- Optimal utilization of sequential de-bottlenecking to overcome in a sequential manner
the bottlenecks facing the maximization of the
process performance, whether this process only
involves optimization of a given configuration
and the processes taking place within its
boundaries or whether it involves change of
configuration and processes involved.
- Maximum utilization of optimal configuration
and optimal design/operation.
- Combined fundamental and practical research
and maximum utilization of fundamental
findings to formulate novel ideas and generalize
successful ones.
- Multidisciplinary teaming and research is a
must for serious research and development of
novel competitive processes.
- Optimal
combined
use
of
rigorous
mathematical/computer
models,
experimentation and pilot plant verification.
- Maximum utilization of available industrial data.
- Verification of developed models/design
equations against industrial data
- Iterative use of mathematical modeling,
optimization
and
nonlinear
dynamics
techniques with experimental, pilot plant and
industrial data.
- Utilizing both process optimization techniques
together
with
conceptual
optimization
techniques in order to reach the best
configuration with its optimum design and
operating parameters.
- Developing suitable parameters for fair
comparisons between the optimums of
competing processes and designs.
- Models/design equations should be: As simple
as possible but not simpler (Albert Einstein), in
Elementary Calculations for Autothermicity
For the highly endothermic steam reforming
reactions the process can be autothermic if there
are other exothermic reactions balancing it. There
are different techniques to achieve that. These
techniques are analyzed here using very simple
material/ energy balances to find out their basic
characteristics with special emphasize on the
maximum hydrogen yield.
A- Methane ( Natural Gas)
1-The Oxidative reforming/Steam Reforming:
The oxidative reforming can be represented by the
following two reactions:
CH4(g) + 0.5O2(g)
CO(g) + 2H2(g)
(H0298 = - 36 kJ/mol)
(4)
and,
CO(g) + H2O(g)
CO2(g) + H2(g)
(H0298 = - 41 kJ/mol)
(5)
The two reactions produce 3 moles of hydrogen per
mole of CH4, and produces 77kJ/mole reacted of
CH4
Steam Reforming is: CH4 + 2H2O
CO2 + 4H2
(H0298 = 165 kJ/mol)
(6)
For autothermal operation it is almost trivial to
compute, through very simple mole and energy
balances, the maximum H2 yield to be equal to ~
3.15 mol H2/mol CH4 (78.8 % of the theoretical
yield of 4 moles H2 per mole CH4).
5
2-The Oxidative reforming (without shift
reaction)/Steam Reforming: If the oxidative
reforming catalyst and operating conditions tend not
to promote reaction (5) and is dominated by
reaction (4), the process becomes very inefficient
and also it produces more CO which is more
harmful than CO2. The use of hydrogen membranes
prevents this disastrous situation to develop by
favoring the shift reaction (5). The situation without
reaction (5) gives for autothermic operation:
maximum Hydrogen Yield = ~ 2.55 mol H2/mol
CH4, representing 63.8% of the theoretical yield,
which is very low.
CH4, representing 100 % of the theoretical yield ,
however this is only a local(riser reformer)
autothermicity, but not a global( entire CFB reactorregenerator), because of
heat consumed for
calcinations.
B-Higher Hydrocarbons
We will present here only the complete oxidation
/steam reforming case because it is the most
efficient and the carbon formation/carbon burning
case is thermally very similar.
1) n –Butane
Complete Oxidation is as follows:
C4H10 + 6.5 O2 → 4 CO2 + 5 H2O
(H0298 = - 2658.5 kJ/mol).
The reaction produces no hydrogen but produces
2658.5 kJ/mole reacted of n-C4H10.
Steam Reforming is as follows:
n-C4H10 + 8 H2O → 4 CO2 + 13 H2
(H0298 = 649.9 kJ/mol) .
For autothermal operation maximum Hydrogen yield
is ~ 10.45 mol H2/mol n-C4H10 (80.4 % of the
maximum theoretical yield of 13).
3-Complete Oxidation /Steam Reforming :It is
interesting to notice that for autothermic operation
the complete oxidation/ steam reforming coupling is
more efficient than the oxidative reforming/steam
reforming coupling (even the efficient one with shift
conversion), full oxidation is:
CH4 + O2 → CO2 + 2H2O(g),
(H0298 = - 802 kJ/mol)
(7)
The reaction produces no hydrogen but produces
802 kJ/mole of CH4. Steam Reforming is reaction
(6) above and for autothermal operation maximum
H2 Yield is ~ 3.317 mol H2/mol CH4( 82.9 % of the
theoretical yield, which the highest).
2) Heptane
Complete Oxidation is as follows: C7H16 + 11 O2
→ 7 CO2+ 8 H2O (H0298 = - 4501.48 kJ/mol).
The reaction produces no hydrogen but produces
4501.48 KJ/mole reacted of heptane.
Steam Reforming is as follows:
C7H16 + 14 H2O → 7 CO2 + 22 H2 (H0298 =
819.6.9 kJ/mol).
For autothermal operation maximum Hydrogen
Yield is ~ 18.6 mol H2/mol C7H16 (84.5 % of the
maximum theoretical yield of 22).
4-Carbon formation / steam reforming in the
reformer and carbon combustion in the
regenerator:
It is almost as efficient as case 3 above, and relies
on cracking part of the methane in the reformer
according to the reaction:
CH4
C + 2H2 (H0298 = 75 kJ/mol) (8)
Coupled to steam reforming according to reaction
(6) also in the reformer, followed by carbon burning
for catalyst regeneration and heat production in the
regenerator according to the reaction:
C + O2 → CO2 (H0298 = - 393.5kJ/mol)
(9)
Autothermal operation for this process gives
maximum Hydrogen Yield of ~ 3.3 mol H2/mol CH4,
representing 82.5 of the theoretical yield of 4 moles
hydrogen per mole methane. Which is almost the
highest. An important question remains: Is it
Possible to achieve autothermicity using CO2
sequestration alone? The answer is : Yes without
carbonate regeneration, i.e.: Supply in-situ heat but
needs heat to get regenerated and recycled.
Energy Cost Per mole of Hydrogen from
Different Hydrocarbons (kJ/mole H2)
Methane: 41.25,
Butane: 37.33,
Heptane: 37.25.
Heat of Combustion of Hydrogen is ~ 240 kJ/mole
H2. Thus we spend very low % of the energy in
hydrogen to obtain it: 17.2% from methane, 15.6%
from Butane and 15.5% from Heptane.
Sequestration Load (mole CO2/mole hydrogen)
from Different Hydrocarbons
Sequestration reaction using CaO is:
CaO + CO2
CaCO3
(H0298 = -178 kJ/mol, exothermic)
(10)
If the amount of CaO used is enough to sequester
all CO2 then the total is exothermic: 165 –178 = -13
kJ/mol. Maximum Hydrogen Yield= 4 mol H2/mol
Methane: 0.250,
Butane: 0.308,
Heptane: 0.318.
From these numbers it is clear that CO2 production
per mole of hydrogen produced increases from
6
methane to higher hydrocarbons. The increase is
23.2% from methane to Butane , but only 1.4 %
from Butane to Heptane. This is important with
regard to global warming effects and also with
regard to the design of the simultaneous
sequestration and the dry reforming processes.
separated are used in a novel dry reforming
process (autothermic reactor-regenerator) to
produce syngas, this part of the process can be
combined with a biological process for the use of
CO2 to produce SCP for animal feed. There are
many biological processes for the use of CO2 to
produce SCP, one of the most suitable ones is
the use of the process of aerobic oxidation of
ferrous to ferric ions by microorganisms such as
Acidothiobacillus ferrooxidase (Thiobacillus
ferrooxidans) (Nemati, et.al., 1998) because in
addition to its utilization of CO2 to produce SCP
for animal feed it simultaneously oxidizes the
ferrous ions from the redox fuel cell and re-supply
its cathode by the ferric ions. There is a joint
NSERC project between UWO and UBC in
Canada utilizing this process with the redox fuel
cell.
Novel and Classical Reformers
Characteristics of Conventional Fixed Bed
Reformers
First Generation Reformers (FGRs)
FGRs are characterized by: hundreds of parallel
tubes, large catalyst pellets to minimize pressure
drop causing low effectiveness factors ( = 10 -3 10 –2, Elnashaie and Elshishini,1993), typical
temperatures of 800 - 1100 K and pressures of 20 25 bar, Huge furnace for supplying heat, high steam
to methane ratio to minimize carbon formation,
requirement of H2 purification steps (See Fig.1).
Three main Types of Membranes
Pd, Pd/alloy Porous Stainless Steel (PSS) and
other types for in-situ hydrogen removal in the
reformer, Zirconium based and Perovskite
membranes, and others for in-situ oxygen supply
in the reformer. Polyethersulfone (PES) thermally
stable polymers mixed matrix membranes for
hydrogen permeation to separate residual
hydrogen in exit gas from reformer.
Bubbling Fluidized Bed Reformer. Second
Generation Reformer (SGR)
Main Features: Powdered catalyst particles ( =
1), Hydrogen permselective membranes for
“breaking” thermodynamic equilibrium, oxygen feed
for efficient in-situ heat supply, aiming at
autothermicity. For more details see the site of
Membrane Reactor Technology, MRT (Elnashaie
and Adris, 1989; Adris, et.al.1991; Adris,
et.al.1994a, b; Adris et.al., 2002).
Methane Reforming in the Novel Reformer
The model for this configuration involves the
kinetics of all the possible reactions taking place in
the reformer and the regenerator, including carbon
formation in the reactor and burning for catalyst
regeneration in the regenerator, as well as the
sequestration of CO2 in the reactor and calcination
to regenerate the carbonates in the regenerator.
The model also includes the rate of hydrogen
permeation through the hydrogen membranes, and
the oxygen supply through the oxygen membranes.
We will present here process comparison between
the FGRs and the novel TGR concentrating on the
riser reformer, leaving the interaction between the
reformer and the regenerator to the next section.
The basis of the comparison will be some industrial
data of FGR, and the Fast Fluidization TGR
employing different techniques to further improve its
performance. The industrial FGR data for tubes and
furnace is given, and the same parameters( except
when it is intrinsically different) are used for the
TGR. The schematic representation of the TGR is
shown in Fig.2 where we concentrate in this section
on the left hand side part, which is the membrane
reformer.
Circulating Fluidized Bed(CFB) Reformer. Third
Generation Reformer(TGR)
The novel process consists (as schematically
shown in Fig. 2) of a riser catalytic reformer the
catalyst and the CO2 adsorbent traveling upwards
through it in the fast fluidization or the pneumatic
transport regions. The reformer is equipped with
hydrogen and oxygen selective membranes. A
sweep gas( usually steam)is used in the hydrogen
membranes. Pure hydrogen is produced from the
hydrogen membranes side after condensing the
steam and drying the pure hydrogen. The remaining
reformed gases flow out of the reformer together
with the solid(s). The traces of hydrogen remaining
with the exit gases can be recovered using special
membranes developed by our group. The solid is
regenerated using air to burn the carbon
deposited on the catalyst and to re-calcine the
carbonate formed in the reformer. The
regenerated hot solid(s) are recycled to the
reformer. The heat of the catalyst regeneration is
sufficient to supply the heat for the highly
endothermic reforming reactions. The gases
7
Industrial Reformer Data. Basis For Comparison
with CFB
Pressure (kPa): 2837.1,H2O/CH4: 3.561,
H2/CH4: 0.2432,
CO2/CH4: 0.1209,
N2/CH4: 0.0204,
Composition (mol %):CH4: 20.22,H2O: 72.00,
H2: 4.92,
CO2: 2.44,
N2: 0.42
Reformer
Furnace: dimensions (m): 21.834 × 35.49 ×13.72
,Furnace type: Top-fired. Reformer Tubes:Heated
length: 13.72 m, Inside diameter: 0.0978 m, Tubes
Number: 897,Catalyst shape : Rashig Rings.
Catalyst bulk density: 1362 (kg/m3).
Inlet conditions per tube : Process gas (kgmol/hr):
3.953,Temperature (K): 760,
Air for Catalyst
Regeneration
Gas+ Solid Catalyst
N2
Gas/Solid
Separator
Feed
Hydrocarbon/Steam
Hydrogen
Permselective
Membranes
Fast Circulating
Fluidized bed
Oxygen
Permselective
membranes
Sweep Gas
Catalyst circulation
H2
Combustion
Gas (mainly
CO2 +CO)
Solid
Catalyst
Regenerator
Dry reforming
Catalyst
Circulation
Air
External
Source of CO2
Fig.1. Schematic presentation of the classical industrial reforming flow sheet based on FGR
Syngas
Reactor
Hydrocarbon
Source
(e.g.; Natural
Gas)
Fig. 2. Schematic Diagram for Novel CFB Membrane Reformer. MHS not included. MHS=Membrane
Hydrogen Separator (Before Catalyst Regeneration)
8
Off-gases
Gas-Solid
Separator
H2
+
Sweep gas
N
air
2
Reformer
Regenerator
Combustion gases
Sweep gas
Gas solid
separator
air
H2 membranes
Hydrocarbon Feed
+
Water
O2 membranes
Fig.3 Schematic Representation of the CFB
Table 1: Comparison Between (FFMSR) and Industrial Fixed Bed Reformer
FFMSR
Fixed Bed
Case I
Case II
Case III
Exit Methane Conversion
Exit Steam Conversion
Total Hydrogen Yield
(per mole of methane introduced)
Methane Feed Rate (mol/hour)
Process gas exit temperature (K)
Pressure (kPa)
Length (m)
Total Reactor Volume (m3)
Membrane Diameter (mm)
Membrane Surface Area (m2)
0.8527
0.3405
2.812
0.8675
0.3226
2.884
0.913
0.3524
3.081
0.9375
0.3721
3.200
3953
1130
2200
13.72
0.1031
-
3953
1130.57
2200
0.2
0.0018
9.78
0.123
3953
1130.79
2200
2
0.018
9.78
1.229
3953
1130
2200
2
0.018
9.78
1.229
Hydrogen Yield per m3 of reactor
27.27
1602.2
171.2
177.8
9
IGEC-1
Proceedings of the International Green Energy Conference
12-16 June 2005, Waterloo, Ontario, Canada
Paper No. IGEC-1-Keynote-Elnashaie
Table 1 shows a comparison between the Industrial
Fixed Bed Reformer (IFBR) and the Fast
Fluidization Steam Reformer (FFMSR) without
oxygen membrane and with similar amount of
external heat fed as the industrial unit. It is clear
from the comparison that the FFMSR for the same
methane conversion is much smaller than the IFBR.
The volume of the FFMSR case 1 is about 1.7 %
the volume of the IFBR. A good parameter to
measure the efficiency of the unit is the hydrogen
yield per m3 of reformer, for the IFBR it is 27.7
and for FFMSR, it is 1602.2, representing a radical
5,775.30% improvement. However the hydrogen
yield per mole of methane fed is not much
improved, it is 2.812(70.3 % of the theoretical value
Table 1 shows a comparison between the Industrial
Fixed Bed Reformer (IFBR) and the Fast
Fluidization Steam Reformer (FFMSR) without
oxygen membrane and with similar amount of
external heat fed as the industrial unit. It is clear
from the comparison that the FFMSR for the same
methane conversion is much smaller than the IFBR.
The volume of the FFMSR case 1 is about 1.7 %
the volume of the IFBR. A good parameter to
measure the efficiency of the unit is the hydrogen
yield per m3 of reformer, for the IFBR it is 27.7
and for FFMSR, it is 1602.2, representing a radical
5,775.30% improvement. However the hydrogen
yield per mole of methane fed is not much
improved, it is 2.812(70.3 % of the theoretical value
of 4) for the IFBR and 2.884(72.1% of the
theoretical value of 4) for the FFMSR, which is an
improvement of only about 2.5%. Improving the
hydrogen yield depends upon increasing the volume
of the FFMSR and thus decreasing the hydrogen
yield per m3 . A typical case is case III of the
FFSMR where the fractional methane conversion
increases to 0.9375 (10 % higher than the IFBR,)
and a hydrogen yield (per mole of methane fed) of
3.2 which is 14% higher than the IFBR (and 80% of
the theoretical 4), however the hydrogen yield per
m3 drops to 177.8 which is 552% higher than IFBR.
The choice between FFMSR cases I – III depends
upon an economic optimization as shown later.
The performance of the FFMSR can be further
improved by the use of an adsorbent, e.g. CaO or
dolomite (CaO/MgO) in the form of solid powder
mixed with the catalyst to assist the hydrogen
membranes in breaking the thermodynamic
equilibrium barrier by removing another product(
CO2). The results in table 2 show clearly that it is
possible to increase the methane conversion and
hydrogen yield effectively using this technique(BrunTsekhovoi, et.al, 1988; Prasad and Elnashaie,
2004), however it is also clear that the CaO
conversion is quite low due to the fact that the CO2
sequestration is much slower than the reforming
of 4) for the IFBR and 2.884(72.1% of the
theoretical value of 4) for the FFMSR, which is an
improvement of only about 2.5%. Improving the
hydrogen yield depends upon increasing the volume
of the FFMSR and thus decreasing the hydrogen
yield per m3 . A typical case is case III of the
FFSMR where the fractional methane conversion
increases to 0.9375 (10 % higher than the IFBR,)
and a hydrogen yield (per mole of methane fed) of
3.2 which is 14% higher than the IFBR (and 80% of
the theoretical 4), however the hydrogen yield per
m3 drops to 177.8 which is 552% higher than IFBR.
The choice between FFMSR cases I – III depends
upon an economic optimization as shown later.
reactions. One of the possible techniques to
improve the performance of the CaO ( or any other
CO2 adsorbent) is to exploit the slip velocity
associated with larger particles in the FFMSR. This
can be achieved through using larger particle sizes
of CaO leading to CaO residence time, which is
higher than that of the catalyst and the gas and thus
achieving higher utilization of the CaO. The results
of this technique is shown in table 3 (Prasad and
Elnashaie, 2005), the improvement in CaO
conversion using this technique is between 50-98 %
and the improvement in methane conversion and
hydrogen yield is between 5-10%.
Autothermal
Operation
using
Formation-Combustion (CFC)
Carbon
This section presents the basic characteristics of
autothermal reforming using the CFC principle,
which is very similar to that of the Fluid Catalytic
Cracking (FCC) process (Elnashaie and Elshishini,
1996), for two feedstocks.
Feedstock Methane: The basic elementary thermal
analysis for the autothermal operation of this case
has shown that the maximum hydrogen yield per
mole of methane fed is ~ 3.3 mol H2/mol CH4,
representing 82.5 % of the theoretical yield.
Autothermal operation in this case is achieved in a
Circulating Fluidized Bed (CFB) where the riser is
the reformer and the downer is the regenerator
where the carbon is burned, catalyst is regenerated
and heat is generated and fed to the reformer
through the circulated regenerated catalyst. The
process is described by a set of two-point value
differential equations giving rise to bifurcation
behavior. The heat generated in the regenerator is
sufficient for methane reforming and cracking and
also to vaporize water to the steam needed for the
process, therefore the feed is liquid water. A sample
of the results is shown in Figure 3 (Prasad and
Elnashaie, 2005) with Steam to Methane
Ratio(SMR) as the bifurcation parameter. For SMR
lower than about 0.1 we get a temperature runaway
, this is due to that the carbon formation is too high
in the reformer and thus the heat production in the
regenerator is too high causing this temperature
runaway . On the other hand for SMR higher than
about 0.24 the system quenches to very low
temperature and almost zero hydrogen production
due to that the carbon formation in the reformer is
too small and thus the heat produced in the
regenerator is not enough to sustain autothermicity
at a high temperature and therefore the system
quenches. In the region between these two
bifurcation values the system has multiple steady
states one with high temperature and relatively high
hydrogen yield, one is a quenched state and the
third is the middle unstable saddle type state with
intermediate temperature and hydrogen yield. The
maximum hydrogen yield is 2 moles of hydrogen
per mole of methane fed which is 50% of the
theoretical maximum of 4 moles of hydrogen per
mole of methane fed( 61% of the autothermic
theoretical maximum of 3.3 described earlier).The
process optimization will make this figure reaches
close to the hydrogen yield of 3.3(82.5%) discussed
earlier (Prasad and Elnashaie, 2005).
covers the range of SCR from about 1.45 to 2.25.
For this configuration, the feed temperature is a
state variable determined through iterative solution
of the two-point boundary value problem and the
exit temperature from the riser reformer will always
be lower than the feed temperature( except when
the exothermic sequestration of CO2 is used in the
reformer).The feed temperature is a result of the
high temperature of the regenerated recycled
catalyst after it got into direct contact with the liquid
feed of heptane and water and vaporized both of
them. The feed temperature vs. SCR figure shows that
for SCR<1.45 there is a unique high temperature
steady state which is smoothly increasing with the
decrease in SCR in the range of 800- 850 0K, while for
SCR<1.25 the temperature starts to rise sharply
towards 10000K. For SCR> 2.25 the system has a
unique steady state at the low temperature of about
6750K and as SCR increases further this unique
temperature does not change much. In the region
1.45<SCR< 2.25 there are multiple steady states one
branch is high temperature statically stable steady
states with temperatures varying between about 840 0K
and 7200K, and a middle unstable saddle type branch
with temperatures almost constant at around 700 0K
and a statically stable low temperature branch with
temperatures almost constant around 680 0K. For the
hydrogen yield graph it is interesting to notice that in
the multiplicity region, unlike the methane autothermic
reforming case , the high temperature branch
corresponds to low hydrogen yield while the low
temperature branch corresponds to high hydrogen
yield. So for the this higher hydrocarbon case we can
obtain high hydrogen yield at low temperature around
6800K( 4070C = 7650F). The highest hydrogen yield is
obtained very near to the bifurcation point with SCR of
about 1.45 and this yield in this sample case is about
15.6 per mole of heptane fed. This is about 84 % of
the theoretical autothermal yield for heptane. Process
optimization with regard to operating conditions,
number of membranes, etc carried out by Chen and
Elnashaie (Chen and Elnashaie, 2005b) was able to
raise this 15.6 figure to 17 moles hydrogen per mole of
heptane fed, raising the above percentages to 91.4 %
of the theoretical yield for autothermicity of 18.6, and
77.3% of the stoichiometric yield of 22 moles per mole
of heptane.
Feedstock heptane: Heptane is used as model
component for higher hydrocarbons because most
of its kinetics and physico-chemical parameters are
available (Tottrup, 1982, Xu and Froment, 1989,
Siminski et al., 1972, Jin et al., 2000, RostrupNielsen, 1974, Snoeck et al., 1997, Tottrup, 1976,
Chen et al., 2000). The real higher hydrocarbons of
interest are gasoline, diesel and bio-oil, which
requires detailed kinetic investigation to obtain
reliable kinetic data. Notice that although the main
steam reforming reaction is irreversible the process
is still thermodynamically limited because of the
reversibility of the methanation reaction. Therefore
the performance will be strongly favored by the use
of hydrogen selective membranes and/or CO2
sequestration. It is clear from the sample results in
Fig.4 that for the reactor-regenerator CFB reformer
multiplicity of the steady state covers a critical range
of steam to carbon ratio(SCR) used in this case as
the bifurcation parameter. The multiplicity region
Table 2: Effect of CO2 Sequestration Using CaO. (*: Increase over the case with no CaO or membranes)
(= (mass of CaO/mass of catalyst) in feed
11
&
nt = no. of hydrogen membranes)
No.
Tf
(K)
P
(atm)

nt
XCH4
%
increase YH2
in XCH4*
%
increase
in YH2*
XCaO
FoCaO
(kmol/hr)
1
2
3
4
5
6
7
8
9
900
900
900
850
850
900
900
900
900
5
10
10
5
10
5
10
20
20
1
1
0.1
0.1
0.1
0.1
0.1
0.1
0.04
0
0
0
5
5
5
5
5
5
0.7637
0.7221
0.5412
0.5895
0.6347
0.7473
0.8186
0.9921
0.9127
31.27
60.36
20.19
32.89
86.90
28.45
81.79
190.6
167.34
39.42
68.45
24.27
35.79
91.24
32.90
89.14
201.76
176.03
0.0039
0.0078
0.0275
0.0117
0.0356
0.0126
0.0425
0.1315
0.2237
713.06
356.53
56.19
106.13
53.07
112.38
56.19
28.09
11.69
3.0319
2.8804
2.1249
2.3114
2.5185
2.8902
3.2341
3.9636
3.6256
Table 3: Exploitation of CaO Particles Slip Velocity
= slip factor)
Condition
s
Tf = 900K
P = 5 atm
 = 0.1
Tf = 900K
P = 10atm
 = 0.1
Tf = 900K
P = 20atm
 = 0.04
dp
(m)

20
100
500
1000
2000
20
100
500
1000
2000
20
100
500
1000
2000
1.2471
1.5492
2.2732
2.6247
2.9515
1.3768
1.6712
2.3049
2.5884
2.8554
1.6362
1.9206
2.4646
2.6932
2.9135
XCH4
%
increase
in XCH4*
0.7559
0.7662
0.7879
0.7971
0.8048
0.8457
0.8628
0.8910
0.9006
0.9085
0.9633
0.9763
0.9912
0.9945
0.9965
1.15
2.53
5.43
6.67
7.69
3.31
5.40
8.84
10.02
10.98
5.54
6.97
8.60
8.96
9.18
YH2
%
increase
in YH2*
XCaO
2.9345
2.9857
3.0905
3.1333
3.1690
3.3523
3.4263
3.5460
3.5868
3.6198
3.8395
3.8950
3.9591
3.9736
3.9830
1.53
3.30
6.93
8.41
9.65
3.65
5.94
9.64
10.91
11.93
5.90
7.43
9.20
9.60
9.86
0.0148
0.0171
0.0216
0.0234
0.0250
0.0497
0.0541
0.0613
0.0640
0.0662
0.2838
0.3011
0.3253
0.3329
0.3392
%
increase
in XCaO*
17.46
35.71
71.43
85.71
98.41
16.94
27.29
44.24
50.59
55.76
26.87
34.60
45.42
48.82
51.63
FoCaO
(kmol/hr)
110.6265
108.5606
103.9097
101.7924
99.8997
54.8651
53.8740
51.8576
51.0034
50.2244
11.4917
11.4068
11.2478
11.1823
11.1199
1600
1
(a)
1400
Methane Conversion
1200
TR (K)
(b)
0.9
1000
800
600
400
0.8
0.7
0.6
0.5
0.4
0.3
0.2
0.1
200
0
0
0.1
0.2
0.3
Steam to Methane Ratio
0
0.1
0.2
Steam to Methane Ratio
12
0.3
2.5
(c)
4.0
Carbon at riser outlet (kmol/hr)
Hydrogen Yield
2
1.5
1
0.5
0
(d)
3.5
3.0
2.5
2.0
1.5
1.0
0.5
0.0
0
0.1
0.2
0.3
0
0.1
Steam to Methane Ratio
0.2
0.3
Steam to Methane Ratio
Fig.3: Bifurcation diagram for the autothermal process with SMR as the bifurcation parameter(P = 5 atm, T f =
298.15
1000
Reformer feed temperature To (K)
Lower steady-state
950
Middle steady-state (saddle type, unstable)
900
Upper steady-state
Bifurcation points
850
800
750
700
650
1.00
1.25
1.50
1.75
2.00
2.25
2.50
2.25
2.50
Steam to carbon feed ratio (mol/mol)
Energy-based hydrogen yield
16.0
15.5
15.0
Lower steady-state
14.5
14.0
Middle steady-state
(saddle type, unstable)
Upper steady-state
Bifurcation points
13.5
1.00
1.25
1.50
1.75
2.00
Steam to carbon feed ratio (mol/mol)
Fig. 4. Typical Bifurcation Behavior of the Reactor-Regenerator CFB reformer with Heptane
13
Pure Nitrogen
Pure Hydrogen &
Sweep Gas
Effluent Gases
Detailed Cross-Section
Sweep Gas
Air Feed
Typical
Hydrogen
Membrane
Sweep Gas
Typical Oxygen Membrane
Feed & Recycle from Regenerator
Fig. 5 Preliminary Design of theast Fluidized Bed Membrane
:F
Reformer
Year
1993
1994
1995
1996
1997
1998
Table 4. Annual Chemical Engineering Plant Cost Index
C&E Index
Year
C&E Index
359.2
1999
390.6
368.1
2000
394.1
381.1
2001
394.3
381.7
2002
3956.6
386.5
2003
401.5
389.5
Table 5. General Economical Analysis Parameters
Parameter
Value
LHV of Hydrogen
0.121 GJ/kg-H2
HHV of Hydrogen
0.143 GJ/kg-H2
Heating values
Natural gas (methane)
1126.13 kJ/ft3
Natural gas (methane)
$3.34/1000ft3
Energy costs
Liquid hydrocarbon
$197.63/ton *
Currency
U.S. Dollar Equivalent
Currency
conversions
German Deutsche mark (DM)
0.606
*: The detailed information about the cost of liquid hydrocarbon is shown later.
14
Exit Gases (mainly CO2) for
Dry Reforming in Downstream
Sweep Gas
External
Hydrogen
Separator
Hydrogen Poor
Gas and Catalyst
H2
Air
Catalyst
Regenerator
N2
Hydrogen
Permselective
membranes
Gas/Solid
Separator
Oxygen
Permselective
membranes
Circulating
Fluidized
Bed Membrane
Reformer
Sweep
Gas
Syngas
Regenerator
For Dry
Reforming
Dry
Reforming
Reactor
Dry Reforming
Catalyst Circulation
Feed
Water/Hydrocarbon
Air
External Sources
of CO2
Hot Catalyst Circulation
Solid Flow
Control Valve
Hydrocarbon
for Dry
Reforming
(e.g., Natural Gas)
Figure 6. Novel autothermal reforming process for efficient pure hydrogen production
15
This novel process is very promising, however
extensive research is needed, which includes
modeling, optimization pilot plant verification and
investigation of the complex hydrodynamics of the
gas solid flow/mixing/separation. Also since
bifurcation behavior is dominating the efficient range
of operation, process dynamics and control
investigation should be carried out. A schematic
presentation of the reformer part of this novel
design is shown in Fig.5, the mechanical design of
this part of the pilot plant is mostly carried out in cooperation with Professor Faisal Abdelhady of
Auburn University.
AUTOTHERMAL HYDROGEN
COST ESTIMATION
configuration is the autothermal operation shown
in Figure 6. The optimal design and operating
parameters for the maximum net hydrogen yield
are summarized in Table 6, which are used for
the hydrogen production cost estimation. For the
hydrogen production cost estimation, we also
need some information about the process
performance
under
the
above
optimal
autothermal conditions (Table 6). Accordingly,
Table 7 presents the process operation
information for the above optimal performance in
this novel autothermal reformer-regenerator
system ( Chen and Elnashaie, 2005 a,b )
2). Service lives of properties : Plant building and
land life: ~20 years; Reformer and other main
equipments: ~12 years; Lives of nickel reforming
catalyst and desulfurization catalyst: ~5years(Peters
and Timmerhaus,1991).
3). Operating time: 330 days/year, 24 hours /day.
4). Feedstocks: Heptane is used as a model
component for liquid hydrocarbons.
5). Heat exchangers operating mode: countercurrent operation.
PRODUCTION
Economic Analysis: Methodology and Basis
Hydrogen production costs are very diverse,
because the hydrogen production cost reported in
the literature each uses its own basis, e.g.: some
reports uses a high price of $~6/GJ for natural gas
(Scholz,1993;Braun, 2003), while others uses a
much lower price of $~3/GJ (Blok, et.al.,1997;Padro
and Putsche, 1999). It is important to unify the basis
of cost estimation in order to be able to compare
different technologies and to compare hydrogen as
a fuel with other fuels. Therefore, the following
economical analysis basis for converting the
economical data to the same basis is used (Chen,
2004).
Process Technical Calculations:
Details of these calculations are given elsewhere
(Chen, 2004), the condensed summary is given
here. The following is given for the technical
calculations final results of a pilot plant with a
hydrogen production capacity of 100 kg-H2/day
1. Storages for heptane and water : volume of
heptane and water storage tanks are 4.1and 2.0 m3
respectively.
2. Feed pumps for heptane and water: flow rates
for heptane and water are: 0.0803 and 0.0917 GPM
respectively.
3. De-sulfurization unit for hydrocarbon feed:
volume of catalytic de-sulfurization unit is: 0.0674 L
4. Membrane reformer: Membrane reformer ( riser
of the CFB) internal diameter is 0.07 m, the
numbers for hydrogen and oxygen permselective
membrane tubes are 9 and 37, respectively. The
diameter of membrane tubes is the same as used in
the previous works, which is 0.00489 The design
length of the reformer and the membrane tubes are
the same as the optimal reformer length, which is,
rounded to1.75 m.
Annual Chemical Engineering Plant Cost Index:
All the costs/prices are scaled to the same basis,
i.e., the year of 2003 using the appropriate Annual
Chemical Engineering Plant Cost Index (C&E
index), which is listed in Table 4 (Chemical
Engineering, 1983-2004]) and table 5 gives the
heating values, energy costs and currencies
conversions.
Economical Analysis Parameters: For the
expression of hydrogen production cost/price,
different units are used in the literatures, for
example, $/GJ, $/Nm3 or $/kg-H2. Unless otherwise
specified, the general economical analysis
parameters listed in Table 5 are used for all the
calculations.
5. External hydrogen separator: from the
earlier investigation(Chen and Elnashaie, 2005
a,b) there is some hydrogen remaining (<6% of
hydrogen produced in the reformer) in the exit
gases of the riser reformer after most of the
hydrogen has been removed through the
hydrogen permselective membranes, in order to
produce hydrogen efficiently, an external
hydrogen separator is used for the further
hydrogen removal, SP1 in table 8.
Main Process Data/Parameters for Hydrogen
Production Cost Estimation: In this section, the
cost estimation of hydrogen production by steam
reforming of liquid hydrocarbons heptane using our
novel autothermal reformer-regenerator process(
Chen and Elnashaie, 2005 a,b ) is performed. The
main data and parameters are summarized for
the cost estimation:
1) From the earlier optimization results (Chen
and Elnashaie, 2005 a,b ) the best process
6. Catalyst regenerator: for the efficient
hydrogen production and the maximum use of
heat of reactions generated in the catalyst
16
Hydrogen Production Cost
Capacity of 100 kg-H2/day
regenerator, it should has almost 100% catalyst
regeneration efficiency. Thus it should be well
designed or over-designed. In this cost
estimation an over-designed catalyst regenerator
is used with a long length of 3.5 m.
7. Gas solid separator: diameter of the gassolid separator is calculated to be 0.163 m. The
design height is about 4.5-4.8 times the diameter
of gas-solid separator. Therefore the height of the
gas-solid separator is about 0.734-0.782 m. we
finalize it to 0.8 m.
8. Air compressor: volumetric air supply rate at
30 atm is calculated to be: 0.0594 Nm3/min
9. Sweep gas steam in hydrogen membrane
tubes: total sweep gas steam flow rate is
calculated to be 99.00 kg/h.
10. Downstream hydrogen separation from
sweep gas: four heat exchangers are used in
this section to separate the hydrogen from sweep
gases and to maximize energy efficiency. The
specifications of the 4 exchangers are given in
table 8.
11. Cooling the nitrogen-rich air exiting from
the oxygen membrane tubes: heat exchanger
HE5 is used to cool the exiting nitrogen rich air
and preheat the feed air, the area of heat transfer
is estimated as 4.12 m2
12. Cooling the off-gases from the gas-solid
separator: In addition to HE1-HE5 two other
HE6 and HE7 are used to maximize the
utilization of the heat with the exit gases from the
gas-solid separator.
for
a
Plant
Based on the total cost for the main
units/equipments summarized in Table 8, we can
estimate the fixed-capital investment and then
add the working capital to obtain the total capital
investment. For the estimation of capital
investment, the method of process plant
component cost factors presented by Peters and
Timmerhaus, 1991 is used. The detailed cost
factors and the cost estimation for the capital
investment are summarized in Table 10.
Therefore, for steam reforming of liquid
hydrocarbons using this novel autothermal
reformer-regenerator process, the estimated cost
for purchasing the main units/equipments is
$73,249.72. Including other costs such as for
installation, land, fittings, piping, control, services,
instrument and other fees, the total fixed-capital
investment is $228,905.38. Assumed the working
capital is 15% of total capital investment, the
total capital investment is estimated at
$269,300.44. For the hydrogen production cost
estimation, we need the operating cost, which
can be estimated from the unit consumptions of
the raw materials, utilities and operating labor.
The following prices are used for the operating
cost estimation, which are already converted to
2003 prices using Annual Chemical Engineering
Plant Cost Index (Table 4):
Liquid hydrocarbon: Notice, in this investigation
heptane is just used as a model component for
higher/liquid
hydrocarbons,
examples
are
naphtha, gas-oil, gasoline, diesel or bio-oils. The
price for liquid hydrocarbon or heptane can be
estimated using the available current prices of
naphtha, diesel, gasoline or bio-oils. The price of
naphtha in the international market is about $204250/ton (website, 2002; website, 2003a) or $148188/ton (Haus, 2003). The price of diesel is $217256/ton (Haus, 2003). The price of fuel oil using
Peters and Timmerhaus’s data is $132-186/ton.
(Peters and Timmerhaus, 1991 ) . Therefore,
taking the average price for those petroleum
products, we find that the average price for
liquid
hydrocarbons
is
about:
$197.63/ton,Electricity: $0.0466/kWh (website,
2003b) , Process water (city water):
$0.275/ton(Peters
and
Timmerhaus,
1991),Cooling water (tower): $0.0475/ton(Peters
and Timmerhaus, 1991 ), Steam (100 psig):
$5.82/ton(Peters and Timmerhaus, 1991 ). Note,
although the prices reported by Peters and
Timmerhaus are based on the year of 1990, the
prices listed above are already converted to the
year 2003 prices as we did earlier using Annual
Chemical Engineering Plant Cost Index listed in
Table 4. Labor: $30,600/year/person (for a
Specifications
and
Prices
for
Main
Units/Equipments: Based on the detailed process
technical calculations, the final specifications for the
main units/equipments of the novel autothermal
circulating fluidized bed membrane reformerregenerator process can be determined. Then the
prices for these units/equipments can be obtained
from the available market data or estimated from
the reported price using statistical corrections
(Peters and Timmerhaus, 1991;Coulson and
Richardson,1983;ColeParmer,2003/2004;Fisher,2004/
2005;
http://www.4tanks.com/stainless_tanks1.htm).
Because the reported price data can be different in
different years, in order to eliminate this kind of
error, all the prices are standardized to the year
2003 using Annual Chemical Engineering Plant
Cost Index(Chemical Engineering, 1983-2004),
which is presented earlier in Table 4. Table 8
summarizes the final specifications and prices for
those main process units/equipments. Most of the
final specifications are larger than the calculated
values
for
the
sake
of
operation
flexibility/uncertainty. Techniques for calculating the
cost of each item are given elsewhere (Chen,
2004).
17
chemical operator, the reported US national total
compensation is $30,517/year/person(website
2003c)). Usually, for small pilot plants, “we can
assume that the reformers would operate
unattended except for necessary maintenance
and emergency repairs. The control computer
would on a routine basis electronically
transfer operating data to a central monitoring
different type of biomass and waste materials. This
can be achieved through the novel process (Flow
sheet shown in Fig.13) with the Transport Membrane
Reactor Hydrogen Producer (TMRHP) as its
heart(Fig.14) ,being developed through co-operation
between Tourtellotte & Associates in Birmingham,
Alabama and Elnashaie groups. A brief presentation
of the process is given here. The basic principles of
the process are:
1) improving performance by optimizing gasifier
configuration and fundamental processes.
2) use of hydrogen permeable membranes and/or
CO2 sequestration to “push” the thermodynamic
equilibrium
towards
maximum
hydrogen
production, and minimum CO production.
3) use of small amount of Nickel based catalyst to
promote reforming and water-gas-shift reactions
for maximization of hydrogen and minimization of
CO production.
The purpose is to achieve important integrated
advances in the gasification of biomass process in
order to produce pure hydrogen, which is suitable for
fuel cells, at cost approaching the present strategic
aim of $2.5 per Kg H2( according to DOE recent
solicitations). The suggested novel TMRHP for
biomass gasification to hydrogen does not use air
nor oxygen and
utilizes catalyst, hydrogen
permselective membranes and CO2 adsorbent to
shift thermodynamics of the process towards more
H2 and less CO, and to accelerate its kinetics and to
produce pure hydrogen suitable for fuel cells directly
from the membrane side. This novel process will
bring down the cost of hydrogen production from
biomass towards this strategic aim through its
characteristic novel features.
station that has responsibility for multiple
reformers”(Myers,et.al.,2002).
Therefore
we
assume 1 person is needed to work in the central
monitoring station which is responsibility for 3 pilot
plants/reformers. One person for one shift, three
shifts a day. Based on the total capital investment
and unit consumptions of raw materials, utilities and
labor, the hydrogen production cost by steam
reforming of liquid hydrocarbon using this novel
autothermal
reformer-regenerator
process
is
estimated. As shown in Table 10, for the pilot plant
with a capacity of 100 kg-H2/day, the total hydrogen
production cost is about $2.224/kg-H2. For the same
hydrogen production capacity, the reported
hydrogen production cost by steam reforming of
methane in traditional fixed-bed steam reformer is
about $9.094/kg-H2(Raissi,et.al., 2002)
The
hydrogen production cost reduction is about 75.54%
for such small pilot plant. Therefore the hydrogen
production cost can be reduced significantly using
this novel autothermal circulating fluidized bed
membrane reformer-catalyst regenerator process
utilizing higher hydrocarbons.
Hydrogen Production Cost for Industrial Plant
Scale with a Capacity of 1,000,000 kg-H2/day: The
typical industrial plant for hydrogen production by
steam reforming of hydrocarbons is in the order of
214,286 kg-H2/day) (Scholz,1993; Zittler and
Wurster, 1996). Thus in this section the hydrogen
production cost estimation is performed for a large
industrial scale plant with a capacity of 1,000,000 kgH2/day. Table 11 shows that the total hydrogen
production cost for it is $0.639/kg-H2. This novel
technology uses liquid/higher hydrocarbons, while
methane/natural gas is used in industrial steam
methane
reforming.
Usually,
higher/liquid
hydrocarbons are not suitable for classical fixed-bed
steam reformers because of the excessive carbon
formation and the typical industrial process for
hydrogen production is the partial oxidation for which
Scholz, 1993; reported the cost to be about
$1.534/kg-H2 Thus the hydrogen cost reduction is
about 58.34% for the same higher/liquid
hydrocarbon feed. This novel process when using
bio-oil will represent an important integral part of biorefineries
Main Features of the Suggested Novel
Technology Targeting and Overcoming the Main
Barriers to the Efficient and Economical
Production of Pure Hydrogen from Biomass for
Fuel Cells: As explained earlier, hydrogen is the
perfect fuel; fuel cells utilize it to efficiently produce
electricity (>50%) with zero emissions. The main
objective of this research is to produce hydrogen
from renewable biomass economically, by reducing
the cost of its production towards 2.5 per kg of H 2.
The design and operating conditions are optimized
to achieve maximum hydrogen production combined
with the use of suitable catalyst, CO2 adsorbent and
hydrogen permselective membranes to shift the
reactions towards higher hydrogen and lower CO
production and to produce pure hydrogen from the
membrane side, which is directly suitable for fuel
cells. The CO2 rich gas stream(s) are directed to the
novel dry reformer proposed by Elnashaie (Prasad
and Elnashaie, 2002) for the production of syngas to
be used for methanol production. The main part of
the novel technology is the TMRHP, which is a
transport (entrained) circulating flow unit. It is
basically a Circulating Fluidized Bed (CFB) system
designed to convert biomass (initially southern
A NOVEL TRANSPORT REACTOR FOR THE
PRODUCTION OF HYDROGEN FROM BIOMASS
One of the main challenges facing hydrogen
economy is moving from producing hydrogen with
fossil fuels as the feedstock towards other
feedstocks, especially renewable ones such as
18
variety softwood/hardwood whole trees, as well as
other biomass feedstocks) directly to hydrogen. The
Membrane Hydrogen Separation (MHS) and the
Reactor-ReGenerator Dry Reformer (RGDR) are
also important parts of this new technology. The
integrated unique features of this novel process can
be summarized in the following:
1- No use of air or oxygen leading to gases of
higher BTU values.
2- Use of suitable catalysts to increase the
hydrogen and decrease the CO production.
3- Continuous regeneration/recycling of catalyst.
4- Use of hydrogen membranes to produces
pure hydrogen suitable for fuel cells.
5- Use of CO2 adsorbent to assist the
membranes in improving conversion, push
the thermodynamic balance of the mixture
towards more hydrogen, less CO and supply
part of the endothermic heat of the process.
6- Continuous regeneration/recycling of the CO2
adsorbent.
7- Thermally efficient process with minimum
external heat approaching autothermicity
through:
 Carbon burning for the regeneration of the
catalyst, and the burning of other
combustibles in the regenerator and
recycling of the hot catalyst to the gasifier.
 In-situ heat supply in the gasifier due to the
exothermic sequestration reaction.
8- Optimization of the catalyst deactivationregeneration
process
for
autothermal
operation.
9- Separating remaining
hydrogen in the
effluent gases using novel polymeric
Membrane Hydrogen Separator (MHS).
10- Utilization of all the CO2 produced from the
gasifier and the regeneration of the catalyst
and CO2 adsorbent, for the production of
syngas using novel RGDR.
11- Quantification
of
the
complex
reaction/diffusion/hydrodynamic
processes
taking place in the unit to develop a rigorous
computer simulation model.
12- Application of advanced digital control for online optimization and regulation.
Preliminary lab experiments have shown that
catalyst can increase H2 from 12% to 40%, while
decreasing the CO from 40% to 18% and methane
from 15% to 9%. We are aiming through our novel
optimized membrane gasifier to make the catalyst
,coupled with hydrogen membranes and CO2
sequestration, improve the H2 yield much more. The
catalyst is chosen to promote both reforming
reactions (increasing H2 on the expense of methane)
and gas-shift reaction (increasing H2 on the expense
of CO). Laboratory experiments with some of these
catalysts have shown very clearly the validity of the
concept (Wang, et.al., 1997; Demirbas, et.al., 1996;
Spiro, et.al., 1984; Demibras and Caglar, 1998;
Corte, et.al., 1985; NIRE, 1997). The flow sheet of
mechanical design of the TMRHP is shown in
Fig.14. The pilot plant receives the biomass in chip
form and it includes facilities to dry the chips to
desired moisture content (10-50 %). The dried chips
are ground to wood flour, carefully metered into a
high temperature re-circulating gas stream (together
with regenerated catalyst and CO2 adsorbent) and
indirectly heated to rapidly bring the mix to the
gasification temperature. Reaction times are in
milliseconds, eliminating the polymerization, which
occurs during slower pyrolysis reactions. The output
from the membrane side steam carrier is mixed with
hydrogen from MHS, condensed to obtain pure
hydrogen directly, and passed through a drying unit
to decrease its moisture content. The output stream
from the first Cyclone (C1) is passed through a
regenerator where all solids including carbon
deactivating the catalyst, carbonate form of the CO2
adsorbent as well as other combustible are burned
using air or oxygen. An efficient cyclone (C2) is then
used to remove the hot regenerated catalyst and
CO2 adsorbent, which will be recycled to the TMRHP
to supply most of the heat, needed for the
endothermic gasification process. The cleaned
(hydrogen free, CO2 rich) gas will be mixed with the
gas out from the MHS and will be split into two parts,
with main fraction going to novel RGDR suggested
by Elnashaie (Prasad and Elnashaie, 2002). The
rest of the process is clear from Fig.13. DOE states
in its recent solicitations that it is not interested in
biomass feedstock cost reduction, therefore the
feedstock cost of $0.6/ Kg H2 given by DOE in
recent solicitations will be considered a give
constant.
Preliminary/approximate
calculations,
show that this novel technology for a unit of 100Kg
H2 per day will have the following cost indices:

Capital investment based on unit life of 10
years is: $ 0.35 /Kg H2

Utilities (electricity, steam, cooling water, etc)
is: $ 0.05/ Kg H2
Labor is: $ 1.02/Kg H2


Raw material (as fixed by the solicitation) is: $
0.6/Kg H2
__________________________________________
Total: $ 2.02 /Kg H2.
for a unit of 1500 Kg H2 per day, the indices
become:

Capital investment based on unit life of 10
years is: $ 0.25 /Kg H2

Utilities ( electricity, steam, cooling water, etc)
is : $ 0.03 / Kg H2
Labor is: $ 0.3 /Kg H2


Raw material ( as fixed by the solicitation) is
: $ 0.6/Kg H2
__________________________________________
Total: $ 1.18 /Kg H2
Both are below the DOE target of $2.5/Kg H2.
This novel process will represent an important critical
part of bio-refineries.
19
CONCLUSIONS
Hydrogen economy has been shown to be quite
promising,
provided
well-directed
innovative
research is carried out and sufficient well managed
funding is made available. Hydrogen/ethanol from
renewable sources, such as biomass and cellulosic
waste are the most promising clean and renewable
fuels for the future. These fuels represent a bright
future for humanity providing everlasting clean fuels,
which are also free of deadly conflicts for the control
of energy sources. It is also shown that autothermic
efficient production of hydrogen using novel
reformers configurations and wide range of
feedstocks is a very promising route towards
achieving a successful hydrogen economy. A novel
autothermic process using circulating membrane
fluidized beds (Autothermal Circulating Fluidized
Bed Membrane Reformer (ACFBMR)) is presented
and is shown to be able to efficiently handle higher
hydrocarbons such as gasoline, diesel and bio- oils
to produce pure hydrogen. From an engineering
design and operation points of view, it is shown that
the process shows bifurcation behavior over a wide
range of the practical range of parameters. This
bifurcation behavior (Elnashaie and Elshishini, 1996)
can be intelligently exploited to maximize the
hydrogen productivity of the unit through rigorous
well organized, modeling, optimization and
experimental verification plan. However operating in
the bifurcation region requires the development of
rigorous/reliable design equations as well as process
dynamic models in order to design tight optimal
control systems. From an economic point of view the
hydrogen production by steam reforming of
hydrocarbons in the novel ACFBMR is evaluated.
Heptane is used as a model component for
liquid/higher hydrocarbons. Based on earlier
optimization results (Chen et.al. 2003c ; Chen and
Elnashaie, 2005b), a detailed flowchart is used for
technical calculations, from which the specifications
and costs of the main units/equipments are
determined. Using the published statistical
correlations and cost factors, the total capital
investment and operating cost are determined. The
investigated range of hydrogen production capacity
for the cost estimation varies from a pilot plant of
100 Kg-H2/day to a very large industrial plant of
10,000,000 kg-H2/day. The results show that the
hydrogen production cost generally decreases from
a cost of $2.224/kg-H2 for a small pilot plant of 100
Kg-H2/day to a much lower cost of $0.625/kg-H2 for
a very large plant of 10,000,000 Kg-H2/day. The
comparison of the economics of hydrogen
production shows that the hydrogen production cost
using this novel autothermal reformer-regenerator
process is lower than the cost reported by the most
economical steam methane reforming in industrial
fixed-bed reformers. For example, with the same
capacity of 100 Kg-H2/day, the hydrogen cost in
industrial steam methane reforming process is
around $9.10/kg-H2, while the hydrogen costs are
$2.054/Kg-H2 for methane feed and $2.224/Kg-H2
for heptane feed in this novel autothermal reformerregenerator system, the cost reductions are 77.43%
for methane and 75.56% for heptane, respectively. If
the hydrogen production capacity is a typical
industrial plant capacity of 214,286Kg-H2/day, the
reported hydrogen cost in industrial fixed-beds by
steam methane reforming is about$0.739-0.966/kgH2, while using this autothermal process, the
hydrogen costs are $0.664/kg-H2 for heptane steam
reforming and $0.501/Kg-H2 for steam methane
reforming, respectively. The cost reductions are
10.15%-31.26% for steam reforming of liquid/higher
hydrocarbons and 32.21%-48.14% for methane
steam reforming, respectively. Therefore the
comparison suggests that this autothermal
circulating fluidized bed membrane reformer can be
a more efficient and more economical pure hydrogen
producer. Another novel process for the production
of hydrogen from different renewable biomass
sources is presented and discussed. The process
combines the principles of pyrolysis and the
simultaneous use of catalyst, membranes and CO2
sequestration to produce pure hydrogen directly
from the unit. Preliminary cost estimation shows that
for a small pilot plant of 100Kg H2 / day the cost will
be about $ 2.02 /Kg H2, while for a capacity of 1500
Kg H2 / day the cost will be about $ 1.18 /Kg H2.
Both numbers are promising for hydrogen economy.
20
Table 6. Optimization Results for Autothermal Reformer-Regenerator System( Chen and Elnashaie, 2005 a,b )
Feed temperature to the riser reformer (K)
845.5 *
Number of hydrogen membrane tubes **
16.471
Number of oxygen membrane tubes **
71.842
Steam to carbon feed ratio (mol/mol)
1.325
Reaction pressure (kPa)
29.811
Reactor length (m)
1.731
Total feed gas flow rate (kmol/h)
2.517
Solid fraction in bed (v/v)
0.0187
Efficiency of catalyst regenerator
1.00
Hydrogen production per unit volume of reformer (kg/h/m 3)
630.40
Optimal net hydrogen yield (moles of hydrogen per mole of heptane fed)
16.732
Notes for table 6 : *Under autothermal operation reformer feed temperature is a system variable not feed parameter
because it is automatically determined by the reforming process configuration. **For practical applications the
numbers should be the closest non-fraction/integer figures. During the reformer scale up/down, we use these optimal
fraction figures for calculations only.
Table 7. Process Information for the Optimal Autothermal Reformer-Regenerator System (Chen and
Elnashaie, 2005 a, b)
Riser Reformer Construction Parameters and Nickel Catalyst Properties
0.0978
Internal diameter of the reformer tube (m) [(Elnashaie and Elshishini,1993)
Outside diameter of hydrogen/oxygen permselective membrane tubes (m) (Adris ,et.al., 1994c)
Total percentage of cross-sectional area occupied by membrane tubes
Nickel catalyst particle density (kg/m3) (Elnashaie and Elshishini, 1993)
Mean diameter of catalyst particles ( μm )(Adris ,et.al., 1994c)
0.00489
22.08%
2835
186
Riser Reformer Operation Data
Reformer temperature (K)
Reaction side
Flow
rate
(kmol/h)
Hydrogen membrane side
Oxygen membrane side
Heptane
Methane
Carbon dioxide
Carbon monoxide
Hydrogen
Water/steam
Oxygen
Carbon
Solid catalyst (kg/h)
Hydrogen
Sweep gas steam
Oxygen from air feed
Nitrogen from air feed
Inlet
845.5
0.245
0.000
0.000
0.000
0.000
2.272
0.000
0.000
329.8
0.000
8.236
0.629
2.365
Outlet
742.9
0.000
0.041
1.147
0.0524
0.0146
0.0516
0.000
0.475
329.8
4.184
8.236
0.570
2.365
Catalyst Regenerator Operation Data
Inlet
742.9
Heptane
0.000
Methane
0.041
Carbon dioxide
1.147
Carbon monoxide
0.0524
Hydrogen
0.000
Flow rate (kmol/h)
Steam
0.0516
Oxygen from air feed
1.166
Nitrogen from air feed
4.388
Carbon
0.475
Solid catalyst (kg/h)
329.8
*Process data are based on the optimal performance for autothermal reformer-regenerator process
feed rate of 0.245 kmol/h.
Regenerator temperature (K)
21
Outlet
1011.3
0.000
0.000
1.715
0.000
0.000
0.1336
0.5828
4.388
0.000
329.8
with a heptane
IGEC-1
Proceedings of the International Green Energy Conference
12-16 June 2005, Waterloo, Ontario, Canada
Paper No. IGEC-1-Keynote-Elnashaie
Unit
No.
Table 8. Summary of the Specifications and Prices of the Main Units/Equipment for the
Novel ACFBMR
Dimension
Power Weight Price
Description
Specification
Material
(mm)
(kW)
(kg)
(Dollar)
C1
Air compressor
4.7cfm, 450 psig
838X41X69
2
HE1
Heat exchanger
Evaporator (0.34 m )
HE2
Heat exchanger
Shell& Tube (6.42 m2)
2
S.S.
1.61
54.9
1,964.00
304 S.S.
2.8
2,130.00
304 S.S.
53.1
5,350.00
HE3
Heat exchanger
Shell & Tube (13.06 m )
304 S.S.
108.0
6,335.00
HE4
Heat exchanger
Shell & Tube (1.92 m2)
304 S.S.
15.9
4,685.00
HE5
Heat exchanger
Shell & Tube (4.12 m2)
304 S.S.
34.1
5,010.00
304 S.S.
1.8
1,685.00
304 S.S.
5.1
3,065.00
2
HE6
Heat exchanger
HE7
Heat exchanger
Evaporator (0.62 m2)
Liquid hydrocarbon
6.2 GPH, 800 psig
pump
Water pump
6.2 GPH, 800 psig
P1
P2
R1
R2
S1
S2
S3
S4
S5
S6
SP1
SP2
Shell & Tube (0.22 m )
330X350X203
303 S.S.
0.552
20.9
855.00
330X350X203
303 S.S.
0.552
20.9
855.00
Membrane reformer
Catalyst
regenerator
Gas hydrocarbon
storage
Liquid hydrocarbon
storage
Water storage
1000K, 30 atm
1750XID70
S.S.
22,237.55
1000K
3500XID70
S.S.
5,890.92
360 ft3 2400 psig
S.S.
855.00
1100 Gallon
S.S.
1,650.00
550 Gallon
S.S.
825.00
Desulfurization tank
Hydrogen
production storage
Nitrogen rich air
storage
External hydrogen
separator
Gas-solid separator
Flow rate 1-10 GPM
Total
327XID70
304 S.S.
3.2
158.00
600 Gallon, 60 atm
S.S.
2,750.00
600 Gallon, 60 atm
S.S.
2,750.00
3784.25
1000K, 30atm
1750XID70
S.S.
3-6 GPM
800X165
C.S.
5.4
2.714
415.00
$73,249.72
Table 9. Estimation of Capital Investment Using Plant Component Cost Factors (Peters and
Timmerhaus, 1991 )
Components
Assumed % of total
fixed-capital
Cost, $
investment
% of total fixed-capital
investment
Direct costs
Purchased equipment
32
73,249.72
32.0%
Purchased-equipment installation
8
18,312.43
8.0%
Instrumentation (installed)
6
13,734.32
6.0%
Piping (installed)
8
18,312.43
8.0%
Electrical (installed)
4
9,156.22
4.0%
Building (including services)
4
9,156.22
4.0%
Yard improvements
2
4,578.11
2.0%
Service facilities (installed)
13
29,757.70
13.0%
Land
2
4,578.11
2.0%
180,835.25
79.0%
Subtotal direct costs
Indirect costs
Engineering and supervision
5
11,445.27
5.0%
Constructive expense
9
20,601.48
9.0%
Contractor’s fee
2
4,578.11
2.0%
Contingency
5
11,445.27
5.0%
Subtotal indirect costs
48,070.13
Total fixed-capital investment
228,905.38
Working capital (15% of total capital investment)
40,395.07
Total capital investment
269,300.44
23
21.0%
100.0%
IGEC-1
Proceedings of the International Green Energy Conference
12-16 June 2005, Waterloo, Ontario, Canada
Paper No. IGEC-1-Keynote-Elnashaie
Table 10. Hydrogen Production Cost in a Pilot Plant with a Capacity of 100 kg-H2/day
Description
Capital Investment
Total Capital Investment
Operating Cost
Unit
Raw Materials
Consumption
(kg/kg H2)
Liquid hydrocarbon 2.9883
Process water
4.9909
Subtotal
Utilities
Electricity (kWh/kg) 0.6514
Steam*
-0.900
Cooling water
0.900
Subtotal
Unit
Consumption
(people/plant)
Labor
1.00
Subtotal
Total Hydrogen Production Cost:
Cost ($)
Cost ($/kg-H2)
269,300.44
0.6801
Unit
($/kg)
Cost ($/kg-H2)
Price
0.19763
0.000275
0.5906
0.0014
0.5920
0.0466
0.00582
0.00005
0.0304
-0.0052
0.0000
0.0252
Cost($/people
/year)
30600.00
Cost ($/kg-H2)
0.9273
0.9273
2.2244
* The unit consumption of steam is negative( in tables 10&11) because the autothermal process produces net steam as a
by-product.
Table 11. Hydrogen Production Cost in a Large Plant with a Capacity of 1,000,000 kg-H2/day
Description
Cost ($)
Cost ($/kg-H2)
Capital Investment
Total Capital Investment
Operating Cost
Unit
Raw Materials
Consumption
(kg/kg H2)
Liquid
hydrocarbon
2.9883
Process water
4.9909
Subtotal
Utilities
Electricity
(kWh/kg)
0.6514
Steam*
-0.900
Cooling water
0.900
Subtotal
Unit consumption
(people/year/plant)
Labor
52.00
Subtotal
Total Hydrogen Production Cost:
67,645,212.12
0.0171
Unit
($/kg)
Cost ($/kg-H2)
Price
0.19763
0.000275
0.5906
0.0014
0.5920
0.0466
0.00582
0.00005
0.0304
-0.0052
0.0000
0.0252
Cost
($/people/year)
30600.00
Cost($/kg-H2)
0.0048
0.0048
0.6390
.
Hydrogen+Sweep Steam
Gasifier
TMRHP
Cyclone- C1
Membran Sweep
Steam
Ash&
Solid
Wood Flouring M/C
Wood
ReGen
Wood Crusher
Air
Wood Dryer
CO2 Recycle
MHS
Feed Heater
Syngas to
methanol
converter
Cyclone-C2
Carrier
Gas
1300-1500F
130-170F
Booster
H2
Retension
Chamber
H1
Reactor
Regenerator
Cooler
Liquid&Solids
Hydro-Cyclone
Dry Reforming Catalyst
Circulation
Hydrocabon Fuel
Source (Natural
Gas)
Pure Hydrogen
for Fuel Cell
Dryer
Condencer
Fig.13. A Preliminary Flow Chart for the TMRHP based process
To Cyclone
Hydrogen &
Sweep gas
Membrane
Sweep gas
Detailed Sectional View
Biomass Flour
& Carrier Gas
Fig.14.Preliminary Design for the Novel TMRHP
25
IGEC-1
Proceedings of the International Green Energy Conference
12-16 June 2005, Waterloo, Ontario, Canada
Paper No. IGEC-1-Keynote-Elnashaie
NOMENCLATURE
dp
diameter of catalyst particle (m)
Fi
molar flowrate of species i (kmol.hr)
ΔHj
heat of reaction for reaction j (kJ.mol-1)
nt
number of membrane tubes
T
temperature (K)
XCH4
conversion of methane
XCaO
conversion CaO
YH2
yield of hydrogen
Greek letters:

(mass of CaO/mass of catalyst) in feed

slip factor
Subscripts:
f
feed conditions
REFERENCES
Adris, A. M., C.J Lim,. and J.R Grace. 1994c. The
Fluidized Bed Membrane Reactor (FBMR) System: a Pilot
Scale Experimental Study. Chem. Eng. Sci. 49: 58335843.
Adris, A., S.S.Elnashaie, and R. Hughes. 1991 Fluidized
Bed Membrane Steam Reforming of Methane. Can. J.
Chem. Eng., 69:1061–1070.
Adris, A., J.Grace, C. Lim, and S.S. Elnashaie. 1994b.
Fluidized Bed Reaction System for Steam/ Hydrocarbon
Gas Reforming to Produce Hydrogen. US Patent no
5,326,550.
Adris, A.M., J. Grace,C. and S.S. Elnashaie. 2002.
Fluidized Bed Reaction System for Steam/ Hydrocarbon
Gas Reforming to Produce Hydrogen. Canadian Patent
no. 2,081,170.
Adris, A. M., Lim, C.J. and Grace, J.R. 1994a. The
Fluidized Bed Membrane Reactor (FBMR) System: a Pilot
Scale Experimental Study. Chem. Eng. Sci. 49:58335843.
Blok, K. R. H. Williams, R. E. Katofsky and C. A.
Hendriks. 1997. Hydrogen production from natural gas,
sequestration of recovered CO2 in depleted gas wells and
enhanced natural gas recovery. Energy. 22, Issues 2-3 :
161-168.
Braun, H. 2003. Calculating Hydrogen Production Costs,
Hydrogen News, the Official Publication of the Hydrogen
Political Action Committee. (www.h2pac.org.)
Brun-Tsekhovoi,
A.R.,
A.N.,
Zadorin,
Ya.R.
Katsobashvili, and S.S. Kourdyumov.1988. The Process
of Catalytic Steam Reforming of Hydrocarbons in the
Presence of Carbon Dioxide Acceptor. Hydrogen Energy
Progress VII, Proceedings of the World Hydrogen Energy
Conference. Pergamon Press. 885-900
Chemical Engineering, (1983-2004).
Chen, Z.
and S.S.Elnashaie. (in Press, 2005b)
Optimization of Reforming Parameter and Configuration
for Hydrogen Production. AIChE Journal
Chen, Z. 2004. A Novel Circulating Fluidized Bed
Membrane Reformer for Efficient Pure Hydrogen
Production for Fuel Cells from Higher Hydrocarbons. PhD
Thesis, Auburn University. Alabama,USA. Chapter 16,
387-428.
Chen Z. and S.S.E.H. Elnashaie. 2002 Efficient
Production of Hydrogen from Higher Hydrocarbons using
Novel Membrane Reformer, Proceeding of the 14th World
Hydrogen Energy Conference, Montreal, Canada.
Chen, Z.., and S.S.E.H. Elnashaie. (In Press, 2005c).
Steady-State Modeling and Bifurcation Behavior of
Circulating
Fluidized
Bed
Membrane
ReformerRegenerator for the Production of Hydrogen for Fuel Cells
from Heptane, Chem. Eng. Sci.
Chen, Z.., Y. Yan, and S.S.E.H. Elnashaie. 2003a. Novel
Circulating Fast Fluidized Bed Membrane Reformer for
Efficient Production of Hydrogen from Steam Reforming of
Methane. Chemical Engineering Science, 58(19):43354349.
Chen, Z.., Y. Yan, and S.S.Elnashaie. 2003b.
Nonmonotonic Behavior of Hydrogen Production from
Higher Hydrocarbon Steam Reforming in a Circulating
Fast Fluidized Bed Membrane Reformer, Industrial &
Engineering Chemistry Research, 42:6549-6558.
Chen, Z.. Y. Yan, and S. S.E.H. Elnashaie. 2004a.
Hydrogen Production and Carbon Formation during the
Steam Reforming of Heptane in a Novel Circulating
Fluidized Bed Membrane Reformer. Ind. Eng. Chem. Res.
43(6):1323-1333.
Chen, Z., Y. Yan, and S.S.E.H. Elnashaie. 2004b.
Catalyst Deactivation and Engineering Control for Steam
Reforming of Higher Hydrocarbons in a Novel Membrane
Reformer. Chem. Eng. Sci. 59(10):1965-1978.
Chen, Z. and S.S.E.H. Elnashaie. (in Press, 2005a).
Steady-State Modeling and Bifurcation Behavior of
Circulating
Fluidized
Bed
Membrane
ReformerRegenerator for the Production of Hydrogen for Fuel Cells
from Heptane, Chem. Eng. Sci.
Chen, Z., P. Prasad and S.S.E.H. Elnashaie. 2002a.The
Coupling of Catalytic Steam Reforming and Oxidative
Reforming of Methane to Produce Pure Hydrogen in a
Novel Circulating Fast Fluidized Bed Membrane
Reformer, Fuel Chem. Div. Preprints 47(1):111-113, ACS,
Orlando, FL.USA
Chen, Z., Y. Yan and S.S.E.H. Elnashaie. 2002b. Using
Coking and Decoking Model in a Circulating Fast
Fluidized Bed Membrane Reformer for Efficient
Production of Pure Hydrogen by Steam Reforming of
Higher Hydrocarbons. Proceedings of the Regional
Symposium on Chemical Engineering.. Kuala Lumpur,
Malaysia. p1239-1247.
Christensen, T. S. 1996. Adiabatic Prereforming of
Hydrocarbons. Important Step in Syngas Production,
Applied Catalysis A: General, 138:285-309.
Chen, Z., Y. Yan, and Said S.S.Elnashaie.
2003c.Modeling and Optimization of a Novel Membrane
Reformer for Higher Hydrocarbons. AIChE J., 49(5):12501265.
Cole-Parmer,
Cole-Parmer
Instrument
Company.
2003/2004. Vernon Hills, IL, www.coleparmer.com.
Corte, P., C. Lacoste, and J.P.Traverse. 1985.
Gasification and Catalytic Conversion of Biomass by
Flash Pyrolysis. J. Anal. Appl. Pyrolysis. 7:323-335.
Coulson, J.M., and J.F. Richardson. 1983. Chemical
Engineering, vol.6, Pergamon Press Ltd., Oxford,
England.
Demirbas, A. and A. Caglar. 1998. Catalytic Steam
Reforming of Biomass and Heavy Oil Residues to
Hydrogen”, Energy Educ. Sci. Technol. 1:45-52.
Demirbas, A.,S. Karshoglu and A. Ayas. Hydrogen
Resources: Conversion of Black Liquor to Hydrogen Rich
Gaseous Products. Fuel Sci. Technol. Intl., 14, pp:451463, 1996.
Elnashaie, S.S.and A. Adris. 1989. Fluidized Bed Steam
Reformer for Methane. Proceedings of the IV International
Fluidization Conference, Banff, Canada.
Elnashaie, S.S.E.H.; S.S. Elshishini. 1993. Modelling,
simulation and optimization of industrial fixed bed catalytic
reactors, Gordon and Breach Science Publishers: London,
UK.
Elnashaie, S.S.E.H. and S.S. Elshishini. 1996. Dynamic
Modeling, Bifurcation and Chaotic Behavior of Gas-Solid
Catalytic Reactors, Gordon and Breach Science
Publishers, London, UK.
Elnashaie, S.S and P. Garhyan. 2003. Conservation
Equations and Modeling of Chem. and Biochem.
Processes. Marcel Dekker, USA.
Fisher
Catalog.
2004/05.
Fisher
Scientific,
www.fishersci.com.
Haus, G.M. 2003. GTL Prospects in Bolivia, 3rd Annual
GTL World summit, London, UK.
Jin, W., X.Gu, S.Li, P.Huang, N.Xu, J.Shi. 2000.
Experimental and Simulation Study on a catalyst Packed
Tubular Dense Membrane Reactor for Partial Oxidation of
Methane to Synga. Chem.Eng.Sci., 55:2617-2625
Myers, D.B., G.D. Ariff, B. D. James, J. S. Lettow, C.E.
(Sandy) Thomas, & R. C. Kuhn. 2002. Cost and
Performance Comparison Of Stationary Hydrogen Fueling
Appliances. Task 2 Report, The Hydrogen Program
Office, Office of Power Technologies, U.S Department of
Energy, Washington, D.C. Under Grant No. DE-FG0199EE35099.
Nemati,M., S.T.L Harrison, G.S. Hansford,G.S. and C.
Webb,C. 1998. Biological Oxidation of Ferrous Sulphate
by Thiobacillus Ferrooxidans: A Review on the Kinetic
Aspects”. Biochemical Engineering Journal. 1:171-190.
NIRE Annual Report. 1997. Hydrogen Production by
Catalytic
Thermochemical
Conversion.Biomass
Division.Padro, C.E.G., and V. Putsche. 1999. NREL/TP570-27079: Survey of the Economics of Hydrogen
Technologies, NRE Laboratory.
Peters, M.S. and K.D. Timmerhaus. 1991. Plant Design
and Economics for Chemical Engineers, 4th edition,
McGraw-Hill Book Company, London.
Prasad, P. and S. S. E. H. Elnashaie. 2004. Novel
Circulating Fluidized-Bed Membrane Reformer Using
Carbon Dioxide Sequestration. Ind. Eng. Chem.
Res.43(2):494-501.
Prasad, P., Z. Chen, and S.S.E.H. Elnashaie. (in Press,
2005b). Static Bifurcation Characteristics of an
Autothermal Circulating Fluidized Bed Hydrogen
Generator for Fuel Cells. AIChE J.
Prasad, P. and S.S.E.H. Elnashaie. 2002 Novel
Circulating Fluidized-Bed Membrane Reformer for the
Efficient
production
of
Ultraclean
Fuels
from
Hydrocarbons. Ind. Eng. Chem. Res., 41(25):6518-6525
Raissi, A., L. Gu, T. Robertson and D.L. Block. 2002.
System Analysis of H2 Production and Utilization at KSC,
NASA Hydrogen Review Meeting.
Rostrup-Nielsen,J.R. 1974. Coking on Nickel Catalysts
for Steam Reforming of Hydrocarbons. J. Catal. 33:184201.
Sammels, A. F., M. Schwartz , R.A. Mackay, T.F.Barton.,
and D.R. Peterson. 2000. Catalytic Membrane Reactors
for Spontaneous Synthesis Gas Production, Catalysis
Today, 56, 325-333.
Scholz. W.H. 1993. Processes for Industrial Production of
Hydrogen and Associated Environmental Effects. Gas
Sep. Purif. 7:131-139.
Shah, M.M., and R.F. Drnevich. 2000. Integrated Ceramic
Membrane System for Hydrogen Production. Proceedings
of the 2000 Hydrogen Program Review, DOE.
Siminski, V.J., F.J. Wright, R.B. Edelman, C. Economos
and O.F. Fortune. 1972. Research on Methods of
Improving the Combustion Characteristics of Liquid
Hydrocarbon Fuels, AFAPL TR 72-74, vols. I and II, Air
Force Aeropropulsion Laboratory, W.P. Air Force Base,
OH.
Snoeck, J.W., G.F. Froment and M. Fowles, 1997. Kinetic
study of the Carbon Filament Formation by Methane
Cracking on a Nickel Catalyst. J. Catal., 169:250-262
Spiro,C.L.; Mskee, D.W. and Kosky, P.G. 1984.
Comparison of Effect of Adding Boudouard Catalysts
Before and After Coal Charring. Fuel.63:1333-1345.
Tottrup P.B. 1976. Kinetics of Decomposition of Carbon
Monoxide on a Supported Nickel Catalyst. J. Catal.42:2936.
Tottrup, P.B. 1982. Evaluation of Intrinsic Steam
Reforming Kinetic Parameters from Rate Measurments.
Applied Catalysis. 4:377-389.
Wang, D.; S.Czernik, D.Montane, M. Mann and E.
Chornet. 1997. Biomass to Hydrogen via Fast Pyrolysis
and Catalytic Steam Reforming of the Pyrolysis Oil or its
Fractions. Ind.Eng.Chem.Res. 36:1507-1518.
www.eia.doe.gov/emeu/aer/txt/ptb0806.html, 1960-2001.
Accessed on December 2003a. Table 8.6 Average
Retail Prices of Electricity.
www.rediff.com/money/2003/apr/24petro.htm.April,
2003b. Petrochem Prices Drop 20 Percent .
www.salaryexpert.com. 2003c Basic Salary Report for a
Chemical Operator in US (national).
www.thehindubusinessline.com/2002/07/10/stories/20020
71002170400.htm. 2002. HPL still Sourcing Naphtha from
IOC.
Xu, J and G.F. Froment, 1989. Methane Steam
Reforming, Methanation and Water-Gas Shift: I. Intrinsic
Kinetics. J. AIChE J.35 (1):88-96.
Zittel, W., and R. Wurste, 1996. Hydrogen in the Energy
Sector, Ludwig-Bölkow-Systemtechnik GmbH, (1996).
27