COMPUTER-AIDED ANALYSIS OF
CRITICAL TECHNOLOGIES FOR HYDROGEN MANUFACTURE
by
L. B. Evans and J. E. Strong, Jr.
MIT Energy Laboratory
Working Paper No. 86-013WP
June 1986
COMPUTER-AIDED ANALYSIS
OF
CRITICAL TECHNOLOGIES FOR
HYDROGEN MANUFACTURE
by
Prof. Lawrence
B. Evans
John E. Strong,
Jr.
Department of Chemical Engineering
M.I.T.
June 1, 1986
PREFACE
This informal working paper documents the results of a study to
demonstrate and assess the applicability of flowsheet simulation or a
screening tool to aid in process research planning. Very crude models were
used for three processes for hydrogen manufacture. Because of limitations in
the model the process and economic results presented herein should be used
only as an illustration of qualitative trends. We have not checked the data
on which these results are based and make no claim for the exact numerical
values.
The study demonstrated, however, that rigorous flowsheet simulation using
ASPEN PLUS with preliminary process economics is an excellent tool for
evaluation processes at the research stage. If more time and data were
available, the models could be readily upgraded to give realistic quantitative
comparisons.
2
ABSTRACT
The development
of new chemical
proce sses is ofteIn
hampered by exhaustive research into many var iations
several alternative production methods. There is a
in
real
need for a tool which can provide early infor mation to he:1 P
guide the selection of the most promising rou tes and
identify the critical areas of research to ma ke the final
process most efficient. This study uses the example of
hydrogen manufacture, considered very importa nt for the
future of synfuels technology, to demonstrate how computer
flowsheet simulation can fill this need.
Three different hydrogen production processes were
modeled with the ASPEN flowsheet simulation system. Steam
reforming and partial oxidation of methane were studied
separately
and then compared.
For each model the process
efficiency, defined in terms of product purity, yield, and
cost, was analyzed
as a function
of the operating
conditions. Trends in behavior were plotted and
methodologies for process optimization found.
On comparing
the processes, steam reforming was i dentified as the more
cost effective process.
Partial oxi dation, although
resulting in lower initial capital i nvestment for the same
size plant, has higher operating cos ts associated with the
need for a pure oxygen feed.
This p rocess is competitive
with steam reforming only if a very low cost source of
oxygen is available.
Th e third process simulated was electrolysis of
Th is demonstrated the method by which flowsheet
simulati on can be used to compare processes based on very
differen t technologies. It was found that because of the
cost of the large amount of electricity needed,
electrolysis produces hydrogen at several times the cost as
that of the steam reforming process.
In addition, the
capital expenditure for a large scale electrolysis plant is
much higher than the same size steam reforming faci lity
because of the high cost of the necessary electroly
equipment.
This suggests that electrolysis is not a viable
alternative for hydrogen manufacture on the scale n eeded
water.
sis
for future synfuels processes.
3
TABLE OF CONTENTS
Page
ABSTRACT
2
TABLE OF CONTENTS
INTRODUCTION
I.
.
.
.
.
.
.
.
.
.
.
.*
Steam Reforming of Methane
Partial Oxidation of Methane .
Electrolysis of Water
IV.
.
SIMULATION RESULTS
Comparisons
CONCLUSIONS
.
SOURCES CONSULTED
.
.
.
.
..
..
.
.
.
4
.
*17
*
*
*.
8
ASPENPLUS FLOWSHEET SIMULATION
Process
.
3
13
.
Steam Reforming of Methane
Partial Oxidation of Methane .
Electrolysis of Water . .
III.
.
PROCESS DESCRIPTIONS
Steam Reforming of Methane
Partial Oxidation of Methane .
Electrolysis of Water
II.
.
·. .
.
.
.
.
.
.
.·
.*
.
.
.
.
.
.
.
21
25
27
.
.
.
.
.
.
.
.
.
.* .
30
47
62
66
.
.
71
.
73
APPENDIX: ASPEN Input Files
Listing 1: Steam Reforming of Methane .
Listing 2: Partial Oxidation
of Methane
Listing 3: Electrolysis of Water
76
81
86
4
INTRODUCTION
Although
it is the third most abundant
element
by
atom, hydrogen does not appear naturally anywhere on earth
in its pure form.
Generating hydrogen therefore requires
the chemical decomposition of a heavier material containing
hydrogen,
followed
by the separation
of the gaseous
hydrogen from other side-products. (9)
feed stock in the chemical
has been a very important
petrochemical
is typically
money
industries
quite
and because
expensive,
used to generate
the production
and
of it
a great deal of work and
and optimization
has gone into the development
processes
Because hydrogen
hydrogen.
of the
In the future
hydrogen should become an even more important raw material,
for although
there
at this time, the
is a world oil surplus
certainty of long-term limited resources suggests the
eventual
need for large quantities
fuel or in synthetic
of hydrogen
gas production.
either
as a
(4)
The oil crisis in the mid 1970's focused
the
world's attention on the problem of limited resources and
the need to find alternative
inexpensive
energy sources
ways to insure
for the future.
reliable
(9)
and
One of the
critical areas identified was the need for efficient,
cheap, high-volume hydrogen production.
As a result, an
explosion of research ensued on many fronts.
People from
5
private industry, universities, and the United States
all began large scale research
government
looked at production methods
hydrocarbons
improving
4,
involvi ng partial
and
s team reforming
(e.g. 1, 4, 7, 11, and 13).
of
Some looked
to make the simplest method, direct electrolysis
water,
cost competi tive despite
inefficiency
of
the inherent thermodynamic
electrolytic
cel ls (e.g.
some explo red revolutionary
including
for
of ele ctricity generation and the high capital
cost of high-volume
Finally,
of
oxidation
Others looked at
11).
the now f avored process:
hydrocarbons
ways
(e.g. 1, 2,
Some
projects.
3,
and
4,
5).
production methods,
the once highly promising thermochemical cycles
involving complex multi-reaction systems in which water is
broken
into its elements
materials.
reaction
through the use of recycled
(e.g. 4)
Among the many advancements forwarded by this
volume of work has been one fact inherent
to hydrogen
there is a vast number of production
production:
and
separation method combinations and no efficient way to
compare
and choose among them.
Some studies
involved
detailed design and cost calculations for different
processes
method
(4), but the inflexibility
becomes
of this "brute-force"
clear when one faces the task of repeating
all of these
calculations
when a break-through
for example,
the physical
reliability
temperature reactors.
is made in,
of very high
Another comparative approach has
been the use of detailed
computer
models
to predict
operating conditions and ultimate process costs (11).
Here
6
again an inherent
the process
inflexibility
is modified
exists
in that every time
in some way, a new computer
program
for the new process must be generated.
In the hydrogen
industry,
then, there is a clear
need for a tool to aid in the study and selection
alternative process pathways.
of
An efficient method must be
developed by which processes which are vastly different in
concept
and stage of development
equal footing.
can be compared
on an
With this tool, early development
could be guided so as to eliminate
improbable
work
processes
and
point out directions of greatest need for viable, low-cost
production
methods.
is not exclusive
need in the whole
The desire for such a comparative
to hydrogen
chemical
production:
industry
tool
there is a general
for an efficient
methodology for process comparison.
This study focuses on one answer to this need:
computer flowsheet simulation.
With flowsheet simulation,
the user is given a list of unit operations
from which
may select.
By taking those unit operations
the chemical
process under study and connecting
he
which are in
them
appropriately with material flow streams, a flowsheet
simulation
is quickly
developed.
The user then has an
efficient means of studying the effects of any relevant
process
inputs on any results of interest.
The flexibility
of this method over "brute-force" calculations and
dedicated computer programs is clear when one considers the
ease with which the unit operations can be moved around,
inserted, or replaced in order to more fully explore
7
process alternatives.
the power
The goal of this study is to demonstrate
of flowsheet simulation in studying, comparing, and
optimizing alternative processes.
Various means of
hydrogen production, including steam reforming of methane,
partial oxidation of methane, and electrolysis of water,
because
of the
have been selected
for this demonstration
real and immediate
need for this tool in this field as
described above.
The flowsheet simulator ASPENPLUS,
devel oped at M.I .T. and now licensed
by Aspen Technology,
Inc., Camb ridge, MA, was used for this work.
simulator
simulating
costing
ha s al 1 of the unit operations
t hese hydrogen
processes
This
necessary
for
and has a comprehensive
pack age with which capital and operating expenses
can be easil y ca l culated.
It
s
that while every effort
houl d be noted
made to use real istic operating
simulations, the emphasis
methodology of flowsheet
operating
and cost numbers
has been
in these
of this work is on the
simulation
and not the specific
an d cost conclu sions found for hydrogen
production.
It is hoped
that this demonstration
will sh ow
the flexibil ity and appli cability of flowsheet simulation
to the study of alternative
will use simulation
processes,
in conjunction
and that engineers
with actual
data from
existing or experimental facilities to make intelligent
choices for the future.
8
I.
PROCESS DESCRIPTIONS
STEAM REFORMING OF METHANE
The first process to be simulated was the current
favored method for hydrogen production: steam reforming.
There
are many variations
on this process,
variety of ways of treating
the reformer
the synthesis
and of separating
side products.
including
in
gas produced
the pure hydrogen
Figure 1 shows a typical
a
from the
configuration,
chosen because it is not overly complex and yet represents
the average process
in use today.
The heart of this production
furnace
at the top of the flowsheet.
unit operation: steam and methane.
be used as the hydrogen
method
is the reformer
Two feeds enter this
Other hydrocarbons can
source, but methane
was selected
because of its common use and because of its high hydrogen
After the feeds enter the reformer,
to carbon ratio.
heat up and reaction
2 shows the reactions
The top of Figure
reformer.
occuring
inside.
in the
and water react to form hydrogen
Methane
carbon monoxide.
takes place over a catalyst
and
The carbon monoxide can then react with
water to produce more hydrogen and carbon dioxide.
overall reaction
they
is highly endothermic,
furnace must supply
The
so the reformer
heat to force the reactions
to go
_
__
k
s
s-
9
STEAM
Q out
Q out
CO 2
Q out
Qin
RECOVERED
MEA
WATER
.
MEA
F I GURE
1'
Process Flowsheet for Steam Reforming of Methane
.
.
ODUCT
.
10
*
REFORMER REACTOR:
CH4
+
H2 0 4---
3 H2
CO+ H20 -* C0 2
+
+ CO
H2
Temperature = 900 - 1250 K
*
Pressure
= 10- 50 ATM
Steam: CH 4
= 3.0 - 5.0
SHIFT REACTORS:
CO + H 2 0
*
-
CO2
-I
H2
Temperature
= 600 - 750 K (HIGH),
Pressure
= < 80 ATM
450 - 600K (LOW)
METHANATOR:
CO2
+ 4 H2 4
CH4 + 2 H 2 0
Temperature
FIGURE 2:
= 600 K
Chemical Reactions for Reactors in Steam
Reforming Process
11
toward completion.
Also indicated
reformer conditions.
the effective
The temperature range is limited by
range of the catalyst,
is in part controlled
hydrogen
on Figure 2 are broad ranges for the
compression
range
by the need for high pressure
at the end of the process
high throughout
and the pressure
the process
at the end).
(i.e. keep the pressure
to avoid the costs of
The steam to methane
the reformer is generally about 4:1.
helps drive both reactions
toward
ratio into
The excess steam
the right, favoring
greater hydrogen production, and also helps prevent carbon
formation
on the catalyst
inside the reformer.
The next step in the process,
is a series of cooling
reformer
as shown in Figure
and shift reaction
steps.
1,
The
exit gases are cooled and sent to a high
temperature shift reactor.
converted
shift
(1)
to hydrogen
reaction.
over a catalyst
The center
reaction involved.
Here some carbon monoxide is
via the water gas
section of Figure 2 shows the
The exit from this reactor is cooled
again and sent to a low temperature
shift reactor
where
most of the remaining carbon monoxide is converted to
carbon dioxide
and hydrogen.
ranges in operating
shift
reactors.
conditions
Also shown in Figure 2 are
for the two temperature
The temperatures
are again limited
by the
catalyst constraints.
This two step shift process
is favored
because of
the reaction equilibrium and cost of catalyst versus the
process temperature.
Low temperatures favor near complete
12
conversion of carbon monoxide, but the catalysts active at
low temperature
tend to be more expensive
to use. (11)
The
high temperature catalysts are cheaper, making it more cost
effective
to shift some of the carbon monoxide
catalysts
and use a low temperature
catalyst
over these
only as a
final, clean-up step.
The stream leaving the low temperature shift
reactor should contain a large fraction of hydrogen, some
carbon monoxide,
monoxide
some methane,
on a dry basis.
The next step, as shown in Figure
1, is to remove the impurity
product
stream.
and a small amount of carbon
Many methods
carbon
dioxide from the
are in use, but for this
simulation absorption with monoethanolamine (MEA) has been
used.
Here the gaseous
multi-stage
product stream
basis with the amine.
Essentially
carbon dioxide is removed in this way.
heavy with carbon
ultimate recycle.
dioxide,
is contacted
on a
all of the
The exit amine,
is sent on for recovery
and
Usually, recovery is handled easily by
simply heating the MEA and driving off the carbon dioxide.
(4)
The product stream, now stripped of carbon dioxide,
contains hydrogen, methane, and carbon monoxide on a dry
basis.
This stream is heated up so that the small amount
of carbon monoxide which was not shifted in the shift
reactors
can be reacted
final product.
taking place
in a methanator
The bottom of Figure
in the catalytic
to produce
the
2 shows the reaction
methanator.
A small amount
of hydrogen is sacrificed to react with the carbon monoxide
13
to give methane
the water
As shown in Figure
and water.
1, all of
in the product stream from the methanator
is
removed, and the final product, consisting only of hydrogen
and methane, emerges.
PARTIAL OXIDATION OF METHANE
Another method for hydrogen production, which has
is partial
by steam reforming,
essentially
been replaced
oxidation.
This process was selected
as the second
to be
simulated, for it is similar to steam reforming, providing
an easy basis of comparison,
and it was hoped that the
simulation would demonstrate why steam reforming is
favored.
3 shows the partial oxidation
Figure
The only major
used.
the steam reforming
at the top.
difference
this flowsheet
between
one shown in Figure
Because both processes
flowsheet
and
1 is the oxidizer
a synthesis
generate
gas after their primary reactor, the exact same recovery
process was used downstream.
The feed to the oxidizer
oxygen,
used,
of methane,
Here too, other hydrocarbons
and steam.
including
consists
coal,
as the hydrogen
source.
can be
Methane was
selected simply to provide for easy comparison with the
steam
reforming
results.
desired reactions
oxygen
The top of Figure
4 shows the
taking place inside the oxidizer.
is used to oxidize
and carbon monoxide.
the methane
to produce
The
hydrogen
The steam reacts with the carbon
14
Ik
row
! I Ai
I
STEAM
Qout
Qout
CO 2
Q out
Qin
RECOVERED
MEA
WATER
MLA
·
FIGURE3: Process Flowsheet for Partial Oxidation of Methane
·
O')IMT
r
_
_
_
15
*
OXIDATION REACTOR:
CH4 +(1/2)
02*
CO
4
H20
2 H2
Co 2 +
+
CO
H2
1050 - 1300 K
Temperature
=
Pressure
Steam: CH4
= 1 - 30 ATM
= 0.5- 1.5
0.5-
02: CH4
1.5
* SHIFT REACTORS:
CO + H2 0
-
CO2 + H2
Terrperature = 600 - 750 K
Pressure
*
= < 80 ATM
METHANATOR:
CO2
+ 4H2 4 CH 4
Temperature
+ 2 H2 0
= 600 K
FIGURE4: Chemical Reactions for Reactors in Partial
Oxidation Process
16
monoxide to produce more hydrogen and carbon dioxide.
Figure 4 also shows some ranges
conditions.
in operating
Depending on these process conditions, the
overall reaction may be either exothermic or endothermic.
This points up the real advantage
to the partial oxidation
process: unlike steam reforming, little heat input is
required in this primary reactor, providing for savings on
capital
to build the reactor
The disadvantage
operating.
and fuel costs while
is the need for relatively
pure oxygen in large quantities.
This usually requires
building an on-site oxygen plant as well, making the oxygen
an expensive
feed stock.
(11)
The steam to methane
ratio for the oxidizer
less than that in the reformer
in the primary
reaction.
because water
Generally
is far
is not needed
a ratio of around
1.0
provides the necessary water to drive the second reaction
toward
the right.
generally
first
The oxygen
near 1.0.
reaction
to methane
It must be high enough to allow the
to take place but cannot
allow full oxidation
ratio is also
be so high as to
to take place, burning off the desired
hydrogen.
The remainder
as indicated
of the partial oxidation
in Figures
flowsheet,
3 and 4, is the exact same as that
for the steam reforming process.
17
ELECTROLYSIS OF WATER
The final hydrogen generation process simulated
with this study was direct electrolysis
well known
that electrical
in an electrolysis
hydrogen
when applied
It is
to water
cell, causes water to split into gaseous
and oxygen.
way to generate
current,
of water.
This process is by far the simplest
hydrogen
and has the advantage
of producing
almost 100% pure product with a minimum of separation
equipment.
associated
the large
It suffers,
however,
from the high cost
with the large amount of electricity
capital costs required
for electrolysis
necessary for high volume production.
selected
for study
used and
so as to compare
cells
Electrolysis was
to the catalytic
methods of reforming and oxidation with the completely
different
approach
of electrolysis.
simulation would demonstrate
It was hoped that the
the advantages and
disadvantages between these processes, allowing for an
intelligent selection of the appropriate process for future
development.
Figure 5 shows the simple flowsheet involved with
electrolysis.
The water is fed to a cell where electricity
is used to split
streams,
oxygen
oxygen created
it into its elements.
and hydrogen,
The the product
are then taken off.
can be a bonus, however
The pure
the high cost of
liquifaction needed to economically ship oxygen any great
distance negates this advantage unless the oxygen can be
used on site.
(4)
18
WATER
ELECTRICITY
OXYGEN
HYDROGEN
FIGURE 5:
Process Flowsheet for Electrolysis of Water
19
0
ELECTROLYSIS CELL:
H2 0
+ electricity
F I GURE 6:
-*
H2 + (1/2)02
Reaction in Electolysis
Process
20
Figure 6 shows what is happening chemically inside
the electrolysis
cell.
Again,
process makes it attractive.
the simple nature of the
The major complexity is
involved with the electrolyte which is used inside the
cell.
Some systems use liquid electrolytes
solutions of potassium hydroxide.
such as
In order to improve cell
efficiency (theoretical power required divided by actual
power used), some work was been done to develop
electrolytes
simulation,
is converted
such as solid polymers.
it has been assumed
to hydrogen
cell efficiency
is 85%.
(4)
newer
For this
that all of the feed water
and oxygen
and that the overall
21
II.
ASPENPLUS FLOWSHEET SIMULATION
STEAM REFORMING OF METHANE
All of the simulations
in this study were done with
the ASPENPLUS flowsheet simulation system.
For each
process, the flowsheet was divided into its individual unit
operations.
For each unit operation,
corresponding model or "block."
a given flowsheet
flow streams.
operations
are selected
The operating
are entered
The simulation
has a
Those blocks appearing in
and connected
conditions
by material
for the unit
and the feed streams
can then be executed,
results accessed.
ASPENPLUS
are specified.
and the desired
In addition, ASPENPLUS has a
comprehensive costing section which allows the individual
process
units to be costed,
total capital investment.
leading to a calculation
The costing
of the
section also
calculates utility and raw material usage, labor, overhead,
and depreciation
so that a total operating
cost can be
reported.
Figure 7 shows the ASPEN block diagram
steam reforming
process.
This shows all of the ASPEN
blocks used in this simulation
represent.
for the
and the types of models
Listing 1 in the appendix
ASPEN input language for this process.
shows the actual
Comparison of
they
22
FEED
REFORMER
_~~~0
(RGI BBS)
I
STEAM
I _
.
II
I
I~~~~~~~~~~~~
Qout
4
I
,\,#
COOL 1
(HEATER)
II
'I
_
HTSHI FT
(RGIBBS)
Fg
-
Qout
1*
J
COOL2
(HEATER)
-
p
/V,
~~~~~~
I
X
!
LTSHIFT
(RGIBBS)
Qout
ii
C02EXIT
COOL3
(HEATER)
I
-
METHATOR
...
I
MEA- REC
(RSTOIC)
WI
--
.4
(SEP)
Immiii,.i.
HEAT 1
I
I
(HEATER)
.
T
I
SCRUB
(SEP)
·
. lIII
..
I
,
]
v
., 4~~~~~~~~~
H20-SEP
(SEP)
RMEA
C
EXIT
H2EXIT
LMEA
FIGURE
7:
.
Ir
PRODUCT
ASPENblock diagram for the Steam Reforming Process
23
Figure 7 with Figure 1 shows that the translation
of
flowsheet to ASPEN block diagram consisted of matching each
step in the process with the appropriate ASPEN model.
The main reformer reactor is modeled with the ASPEN
reactor block
This is a model which uses Gibbs
"RGIBBS."
free energy
to find the final equilibrium
the product
stream.
of
composition
The components which are present -
carbon monoxide, and carbon
methane, water, hydrogen
dioxide - are entered in this block as are the operating
conditions
in the reform er.
Optionally,
the temperature
approach to equilibrium can be entered to simulate
incomplete reaction.
to equilibrium
approach
-15
Fo r this simulation,
for the methane
the temperature
reaction
was set at
C. (8)
All of the heaters and coolers
"HEATER" blocks.
are modeled
with
These take as input the desired exit
temperature and pressure and calculate the necessary heat
duty based on the inlet stream conditions.
is then accessed
exchanger
by the costing
and utility
This heat duty
section to determine
heat
needs for the heating or cooling
step.
and low temperature
Both the high temperature
reactors are also modeled with "RGIBBS" blocks.
Again the
species present and operating conditions are entered.
both blocks
a temperature
was used for the shift
approach
reaction.
to equilibrium
For
of +10 C
(8)
It was assumed that 100% of the carbon dioxide
removed by the monoethanolamine
shift
(MEA) in the scrubber.
is
For
24
this reason,
this step was modeled
as a "SEP"
simply
block.
This block takes the incoming
product
an d MEA and directs
flow of the gaseous
the individual
species present
to the out:let streams with given separation
this simul ation
For
100% of the MEA exits with the stream going
to the MEA recovery
dioxide.
fractions.
unit as does 100% of the carbon
This stream goes to another "SEP" block,
"MEA-REC" which removes the carbon dioxide and returns the
recovered MEA.
a ll
scrubber
The remaining gaseous materials in the
exit to the heater and on to the methanator.
Th ie methanator
modeled
wi th "RGIBBS."
carbon monioxide reacts
ggoes
reaction
is the only reactor which is not
It was assumed that all remaining
to form water and methane
to compl etio n).
For this; reason,
(i.e. the
the ASPEN
reactor mcidel "RSTOIC" was used
This model takes as input
the stoichiometry
occurring
conversion
of the reaction
rate of that reaction.
and the
Here, a conversion
rate
of 1.0 was used.
The last unit in the flowsheet
produce
in which
a pure, dry product
stream.
all of the water is directed
stream and all of the hydrogen
removes
any water to
This is a "SEP" block
to the water exit
and methane
is taken off as
product.
The costing
section of the simulation
accesses the
important variables from around the flowsheet to calculate
capital and operating
size units
costs.
The heat duties
such as the reformer
and the heating
furnace,
are used to
the cooling steps,
step, and the utility rates,
in terms of
25
steam and fuel oil are calculated.
such as methane
discussed
purchase
in the results
the costing
Raw material
cost, is also found.
section,
rate and operating
production
cost per million
As will be
one of the key uses for
section of the simulation
production
usage,
was to take the
cost and determine
BTU of hydrogen
the
for various
operating conditions.
PARTIAL OXIDATION OF METHANE
Because of the similarity between the processes,
the ASPEN
simulation
for partial oxidation
was built
directly on that already developed for steam reforming.
Figure 8 shows the ASPEN block diagram for the partial
oxidation
process.
Listing 2 in the Appendix
corresponding ASPEN input language.
is the
The only real
difference
is in the first block - as mentioned
downstream
processing
For this process
reactor.
of the synthesis
the oxidizer
above, the
gas is identical.
is modeled
as a "RGIBBS"
Instead of just methane and steam, though, an
additional
presence
feed of pure oxygen
of oxygen
the steam
simulation
reformer
is used.
In the block,
is added to the list of species
reactor.
that the reactions
It was assumed
taking
the
found in
for this
place go to
equilibrium before exiting.
The remaining ASPEN blocks, input language, and
operating conditions are the same as those discussed above
26
CH4
-------- No
OXIDIZER
----------
02
(RGIBBS)
STEAM
Qout
pc
iiHEATER
(HEATER)
_~~~~,
1*
i_
HTSHI FT
(RGI BBS)
_il
_
Qout
7v
COOL2
(HEATER)
-
_
1
1
.
LTSHI FT
(RGIBBS)
-- -
Qout
,v'
Ii
d ,i
-
C02EXIT
1
COOL
(HEATER)
.
Qin
METHATOR
MEA-REC
I
(RSTOIC)
w
·
F
(SEP)
.4
II
I p_
-
(HEATER)
--
imL
I
j
SCRUB
S
(SEP)
II
RMEA
"
C"
HEAT 1
1
.
H20-SEP
imml
I
(SEP)
__
-
I1
LMEA
H20EXIT
r
PRODUCT
FIGURE 8: ASPENblock diagram for the Partial Oxidation Process
27
for steam reforming.
ELECTROLYSIS OF WATER
Electrolysis, being a fairly simple process, yields
block diagram,
a simple ASPEN
3 in the Appendix
as seen in Figure
specified.
water
of water
reaction
going to hydrogen
the conversion
and oxygen
fraction
is
for this
is set at 1.0.
Because only one exit stream
block,
"RSTOIC"
easily
Here the
that all of the feed
Since it has been assumed
is separated,
cell
the electrolysis
which is modeled with an "RSTOIC" reactor.
stoichiometry
Listing
for this
shows the ASPEN input language
The feed, water, enters
simulation.
9.
separated
the block
the fact that hydrogen
for an
and oxygen
are
at the opp osite electrodes is modeled with
"CELLSEP."
of the hydrogen
is allowed
This is a "SEP" block which takes all
and puts it in one exit stream while taking
all of the oxygen and putti ng it in the other exit stream.
Because occasionally liquid water can be entrained with the
exit gases, knock-out drums are put in the flowsheet
so
that the product streams wi 11 be pure and dry.
The costing section of the simulation
primarily
of the capital
ex penditure for the electrolysis
cell and the cost of the el ectricity
literature
data,
consists
used.
the instal led capital
on the basis of $300 per kw hydrogen
Based on some
cost was calculated
heating value.
(4)
28
FEED (H20)
PRODUCT(H2)
02EXIT
FIGURE9: ASPENblock diagram for the Electrolysis
Process
29
For the operating
produced
due to the fact that additional
be necessary
assumed
cost, no credit was given for the oxygen
to liquify
the oxygen
investment
for shipping
would
and it was
that there would be no use for oxygen on site.
30
III.
SIMULATION RESULTS
STEAM REFORMING OF METHANE
The simulations developed for these hydrogen
production
processes
representative
have been executed
results and to demonstrate
flowsheet simulation.
the performance
rate.
of
First, there is the purity of
Generally, this stream will consist of
and methane,
the better.
the potential
There are several ways to measure
of a system.
the product stream.
hydrogen
to show some
and the closer
it is to 100% hydrogen
Second, there is the hydrogen production
If the feed is held constant
and the operating
conditions
varied, not only will the purity change, but the
production
rate of hydrogen
possible,
for example,
will also change.
to run the process
to get 95% pure hydrogen
It might be
in such a way as
product while producing
lower rate than the potential.
at a much
This would occur if a great
deal of carbon dioxide is produced, which is subsequently
removed.
measures
So both the purity and production
of how efficiently
into hydrogen.
rate are
the feed is being converted
Also of great importance in measuring the
performance of the system is the operating cost per unit of
hydrogen
produced.
effective.
This will tell if the plant is cost
All of these results will be looked
at for the
31
simulations developed.
There are many parameters in the steam reforming
can be varied
process which
the reformer
Because
the system.
of
the performance
to change
reactor
is the heart of
the system, however, only the operating conditions
associated with this unit have been studied.
These include
the reformer temperature, reformer pressure, and feed steam
to methane
ratio.
Figure
10 shows the fraction
of hydrogen
in the
process product stream (i.e. the purity - the remaining
is all methane)
fraction
Three curves
temperature.
different
as a function
steam to methane
of the reformer
are shown, each representing
feed ratio.
a
It is clear that as
the temperature increases the reformer reactions are driven
toward the right, and the purity of the product stream
increases.
approaches
1200K.
It is interesting
to note that the purity
100% under all conditions
This suggests
that operation
at temperatures
above
above 1200K does not
gain much and research into higher temperature catalysts
and materials
is not of great importance.
Figure 10 also shows that the higher the steam to
the greater
the final purity.
Again this is
methane
ratio,
because
the excess steam tends to push the methane
reaction
toward completion.
Figure 11 shows the product stream hydrogen purity
versus the reformer
pressure.
Here the purity drops as the
pressure increases, suggesting that higher pressure
inhibits the reformer reactions.
This is in agreement with
32
R
A
C
.95
H
Y
.9
D
R
O
.85
G
E
N
.8
N
.75
P
.7
R
0
D
U
.65
C
T .6
900
1000
1100
1200
REFORMER TEMPERATURE
FIGURE
10:
1300
(K)
Fraction of hydrogen in the steam reforming product
stream versus reformer temperature for several steam
to methane mole ratios
with pressure = 20 ATM
33
F
.
R
1
A
99
C
H
Y
.98
.g7
D
R
.96
G
.95
0
E
N
.9 4
I
.93
N
.9 2
R
.91
0
9
D
U
.8 9
C
T
10
15
20
25
REFORMER PRESSURE (ATM)
FIGURE 1: Fraction of hydrogen in the steam reforming product
stream versus reformer pressure for several steam
to methane mole ratios
with temperature = 1100K
30
34
literature
published
feed ratios show
steam to methane
representing
different
that excess
steam helps
lines
Again the different
data. (1)
to make the product stream more
highly concentrated in hydrogen.
Figure
fixed methane
same trend
12 shows the hydrogen
production
feed versus the reformer
in Figure
as observed
rate from a
The
temperature.
10 is seen here.
Higher
temperatures favor hydrogen production as do higher steam
flow rates.
Figure
13 shows a similar plot for hydrogen
as a function
production
of reformer
comes into direct conflict
that high pressure
process should be run at low pressure.
often
done, however
hydrogen
with high
This suggests
rates from a given feed.
production
It is clear
pressure.
the desirability
that the
Indeed, this is
of high pressure
becomes a problem.
at the end of the process
Rather than investing in the compressor necessary to
produce this high pressure product, usually the pressure
throughout
the process
production
rate.
is kept high at the sacrifice
The suitability
of one method
of
or the
other depends on the required condition of the product
stream.
For these simulations
pressurized
hydrogen
this pressure
is desired
is achieved
it has been assumed
that
(15 ATM or so) and that
in the process
and not with an
additional compressor.
Of primary concern
in a process
such as this is
determining the operating conditions which will produce
hydrogen
in the most cost effective
manner.
To do this,
35
H
2
400
P
R
0
350
D
U
C
T
300
R
A 250
T
E
200
m
s
c
f
150
/
d
a
Y
100
900
1000
1100
REFORMER TEMPERATURE
1200
1300
(K)
FIGURE 12: Daily hydrogen production rate from the steam reforming
process versus the reformer temperature for several
steam to methane mole ratios
with pressure = 20 ATM
36
H
2
P
R
0
D
U
C
T
370
360
350
340
330
320
R
A 310
T
E
300
290
m
S
280
C
f
270
d
260
a
y
250
10
FIGURE 13:
15
20
REFORMER PRESSURE (ATM)
25
Daily hydrogen production rate from the steam reforming
process versus the reformer pressure for several steam
to methane mole ratios
with temperature = 1100K
30
37
the cost of hydrogen
the operating
reforming
reformer
production
variables.
operating
is plotted versus each of
Figure
14 shows the steam
cost per million
temperature
the steam to methane
BTU versus
with the pressure
ratio at 4.0.
the
set at 20 ATM and
The cost drops as
temperature increases but reaches an asymptotic limit at
about 1100K.
With the pressure
and steam rate set
accordingly, then, the optimum reformer temperature would
be about 1100K.
In this way the temperature
would be the
lowest possible while still producing the least expensive
hydrogen possible.
Figure 15 shows the hydrogen
function
of the pressure
and the steam to methane
production
with the temperature
ratio set at 4.0.
cost as a
set at 1100K
Because of the
decrease in production associated with high pressures, the
operating
cost rises with pressure.
There is no way to
look at this curve and select an optimum
pressure.
As
discussed above, the pressure must be selected from a
trade-off between cost and necessary exit pressure.
the purposes
reformer
of this simulation,
For
a value of 20 ATM in the
has been assumed to be best - with pressure
drops
in the process this yields about 14-15 ATM product.
Figure 16 shows the hydrogen
the steam to methane
1100K and pressure
ratio increases,
cost as a function
ratio with the temperature
set at 20 ATM.
the cost drops.
set at
At first, as the steam
This is because,
as
observed before, excess steam helps push the methane
reaction toward completion.
of
However, above a ratio of
38
H
lb
Y
D
14
R
G
13
E
N
12
C
11
0
io
T
9
/
6
m
b
t
U
900
950
1000
1050
1100
1150
REFORMER TEMPERATURE
1200
1250
(K)
FIGURE14: Cost of hydrogen produced by the steam reforming
process versus the reformer temperature with the
reformer pressure = 20 ATM, Steam:CH4 = 4.0
1300
39
9
I
I
I
f
I
I
15
20
H
Y
D
R
8.5
0
G
E
N
8
C
0
S
T
7.5
7
/
6.5
t
u
6
10
25
REFORMER PRESSURE (ATM)
FIGURE 15: Cost of hydrogen produced by the steam reforming
process versus the reformer pressure with the
reformer
temperature
= 1 OOK, Steam:CH4 = 4.0
30
4.0
8
H
Y
D
7.9
R
0
G
E
7.8
N
C
7.7
0
S
T
7.6
$
/
7.5
m
b 7.4
t
U
7.3
2
2.5
3
3.5
4
4.5
REFORMER STEAM:CH4 RATIO
FIGURE 16: Cost of hydrogen produced by the steam reforming
process versus the reformer steam:CH4 feed ratio
with
reformer
temp = 1 1OOK, pressure
= 20 ATM
5
41
about 4.0 the cost begins to rise.
the increased
requirements
purchase
due to
cost of the steam and the heating
for bringing
temperatures.
This is probably
the incoming
steam up to reaction
The gain associated with excess steam above
a ratio of 4.0 is outweighed
by the direct operating
costs.
and pressure
With the temperature
then, the optimum
steam to methane
set as indicated,
ratio for the reformer
feed is 4.0.
It should be emphasized
that the optimizations
discussed above are reasonably simplified.
was optimized
value.
individually
In a real
process,
Each variable
with the other two fixed at some
all
three
would
have
to
be
varied simultaneously in order to obtain the overall
process optimum.
With the reformer
informative
operating
conditions
to look at a "cross-section"
see what is happening.
of the process
to
Figure 17 - obtained with data from
a run with the temperature
at 1100K, the pressure
ATM, and the steam to methane
cross-section.
set, it is
at 20
ratio at 4.0 - shows such a
The dry mole percent of each species
is
plotted versus the position in the process.
As indicated,
reformer,
the methane
the feed is pure methane.
percent
has dropped
hydrogen is the dominant species at 75%.
and carbon
dioxide
concentrations
are present
at 10%.
After the
to about 5% and
Carbon monoxide
in about equal
In the shift reactors
the carbon
monoxide is converted to hydrogen and carbon dioxide.
Figure 17 shows this, for the carbon monoxide
curve drops
42
100
P
E
90
R
C
80
E
N
T
70
60
N
50
S
T 40
R
E
A
M
,0
30
20
10
0
FEED
REFORM
FIGURE 17:
HT SHIFT
LT SHIFT
SCRUB
PRODUCT
Species profile of steam reforming process. The lines
show the mole percent of each species in the streams
exiting the indicated unit operations
43
while the carbon dioxide and hydrogen curves rise.
stays
the same, for it is considered
Methane
an inert in these
reactors.
After the scrubber, the carbon dioxide percent
for it has been removed.
drops to zero,
dioxide was acting as a diluent,
percentages
its removal
of all of the other species
of 95% hydrogen
perform economically
steam reforming
carbon
process,
it is informative
cost as a function
to look at the
of the plant capacity.
Figure
Because of the wide range of plant
(10 mscf/day
areas emerge on this plot.
economy of scale
size plants will
As a final look at the
differently.
18 show such a plot.
to 300 mscf/day)
At low capacity,
is observed.
units such as reformer
furnaces,
and operating
three distinct
a sharp
There are minimum
below the limits of these units,
rates).
The
to rise.
and 5% methane.
It is well known that different
(high capital
causes the
The resulting product consists
monoxide by the methanator.
sizes covered
carbon
shows the removal of all remaining
product stream
hydrogen
Because
sizes for
and if the process
is run
it will be very expensive
cost spread over low production
Thus the cost drops as the size increases
for low
production rates.
In the center of the figure,
region
is observed.
Here the process
a relatively
flat
is being run at rates
easily handled by the equipment available.
Increasing
production simple increases the capital and operating costs
proportionately,
and the resultant
hydrogen
cost stays
44
H
y
D
7.55
R
0 7.5
G
E
7 .45
N
7.4
C
0 7.35
S
T
7.3
7.25
m 7.2
b
7.15
t
U
7.1
0
50
100
150
200
250
PRODUCTION RATE (mscf/day)
300
FIGURE 18: Cost of hydrogen produced by the steam reforming
process versus the plant capacity with the reformer
temp = 1 1OOK, pressure
= 20 ATM, steam:CH4
= 4.0
350
45
about the same.
At high production
rates,
however,
larger pieces of equipment are necessary.
size increases,
larger
and
As equipment
the costs tend to rise more quickly.
rates the costs are increasing
at high production
than the hydrogen
produced,
Thus
faster
and the unit cost begins
to
rise.
As a basis of comparison
the actual cost figures
summarized.
medium
with the other processes,
for a selected
plant size have been
19 shows this summary.
Figure
size facility
of 300 mscf/day
As indicated,
was used.
reformer operating conditions are indicated.
operating
a
The
The total
cost was found to be $7.43 per million
BTU of
hydrogen which is in agreement with some representative
literature
data.
(9)
As the pie chart indicates,
pieces of the operating
methane
cost are raw materials
and feed steam costs) and utilities
The remainder
and cooling water).
the major
(e.g.
(e.g. fuel oil
of the cost is split
fairly evenly between labor, overhead, and capital
depreciation.
Also shown
required
in Figure
19 is the capital
for a steam reforming
investment
plant this size.
The total
bill would be about $300 million, split between process
unit purchase
costs,
setting
labor, other directs
such as
land, contingency, working capital, and start-up costs.
46
PLANT CAPACITY: 300 million scf per day
OPERATING PARAMETERS:
Reformer
Temperature
=
1 OOK
= 20 ATM
Reformer Pressure
Steam:CH4
=4.0
OPERATINGCOST= $ 7.43 per million btu hydrogen
OPERATING COST BREAKDOWN:
Depreciation
-
4.5%
Overhead 5.3%
\ Labor 4.2%
CAPITAL COST SUMMARY:
Process Units
Other Directs
Setting Labor
Contingency
Working Capital
Start up Cost
TOTAL:
$ 69,901,000
34,885,000
50,953,000
45,538,000
75,223,000
55,816,000
$ 332,316,000
FIGURE 19: Operating and capital cost breakdowns for a 300 mscf
per day hydrogen plant using steam reforming
47
PARTIAL OXIDATION OF METHANE
A similar analys is of the partia 1 oxidation process
has been made.
In this wa y, it is possi ble to compare
steam reforming
with par ti al oxidation.
Again
many operating parameterma in this proces s can be changed
affect the system performa nce.
As with the reformer,
to
this
study concentrates only on those operati ng conditions
associated with the primary reactor: the oxidizer.
These
operating parameters include the oxidize r temperature,
pressure, oxygen to methane feed ratio, and steam to
methane
feed ratio.
to methane
To li mit the scope slightly, the steam
ratio was set a t 1.0 througho ut to eliminate
this variable.
as a function
The perfor mance of the s vs tem was studied
of the remai ning three operating co nditions.
Figure 20 shows th e hydrogen
function
of the oxidizer
product
t emperature.
purl ty as a
The three curves
represent different oxygen to methane feed ratios
the reformer,
rises.
as the temperature
increases
the purity
They all begin to reach an asymptotic
temperature
slightly
over 1100K.
that the higher the oxygen
limit at a
This graph also indicates
to methane
purity of the product stream.
As with
ratio, the higher
the
This is because excess
oxygen drives the methane reaction to the right,
eliminating the methane from subsequent process streams.
Figure
21 shows
the hydrogen
function of the oxidizer pressure.
product
purity
as
As with the reformer,
higher pressure tends to inhibit the reactions, decreasing
48
F
R
A
C
H
.9
Y
D
R
0
G
.8
E
N
.7
N.
P
R
0
.6
D
U
C
T
.5
900
1000
1100
OXIDIZER TEMPERATURE
1200
(K)
FIGURE 20: Fraction of hydrogen in the partial oxidation product
stream versus the oxidizer temperature for several
oxygen to methane mole ratios
with pressure = 20 ATM
1300
49
F
R
A
C
.95
H
Y
D
.9
R
0
G .85
E
N
I
.8
,e
N
p .75
R
0
D
U
C
T
.7
~r,
10
15
20
25
30
OXIDIZER PRESSURE (ATM)
FIGURE21: Fraction of hydrogen in the partal oxidation product
stream versus the oxidizer pressure for several
oxygen to methane mole ratios
with temperature = 1100K
50
the concentration
of hydrogen
in the product.
This graph
suggests that higher oxygen to methane ratios not only
produce more pure product,
but also tend to limit the
adverse effects of the higher pressures (compare the slope
of the 02:CH4=1.5
Figure
line with the 02:CH4=0.5
22 shows the total hydrogen
versus the oxidizer temperature.
increases
line).
then begins
production
rate
The production rate
to level off at some point.
shows that similar to the reformer
process
This
there is no
great advantage to using temperatures above about 1200K for
partial
oxidation.
It is important
figure
shows that low oxygen
higher
production
to note that this
to methane
feed ratios yield a
rate from a set methane
one of those cases
with the production
in which the product
rate (compare
feed.
This is
purity conflicts
Figures 20 and 22).
Excess oxygen forces more methane to react, increasing the
purity, but this same excess oxygen causes full oxidation
to take place, burning up some of the desired
Figure
23 shows a similar
plot for hydrogen
production rate versus oxidizer pressure.
pressure
drops the production
hydrogen.
rate.
Increasing the
This highlights
again
the conflict between production rate and desired product
pressure.
Also shown
is a confirmation
of the fact that
higher oxygen to methane feed ratios give lower production
rates.
The effect of the operating
conditions
hydrogen
cost will now be examined.
Figure
hydrogen
cost per million
the oxygen
BTU versus
on the
24 shows the
to methane
51
H
2 250
P
R
0
D
U
200
C
T
R
A
T
150
E
m
s
100
C
f
/
d
a
y
50
900
1000
1100
OXIDIZER TEMPERATURE
1200
1300
(K)
FIGURE 22: Daily hydrogen production rate from the partial oxidation
process versus the oxidizer temperature for several
oxygen to methane mole ratios
with pressure = 20 ATM
52
H
2
160
P
R
150
0
D
140
U
C
T
R
130
120
A
T
E
110
100
m
S
C
90
f
60
d
a
y
70
1C
s15
20
25
30
OXIDIZER PRESSURE (ATM)
FIGURE 23: Daily hydrogen production rate from the partial oxidation
process versus the oxidizer pressure for several oxygen
to methane mole ratios
with temperature = 1100K
53
10.5
H
Y
D
R
0
G
E
N
10
9.5
9
C
0
8.5
T
8
/ 7.5
m
b
t
7
u
6.5
.5
.6
.7 .
OXIDIZER
.8
.9
1
1.1
OXYGEN:CH4 RATIO
FIGURE 24: Cost of hydrogen produced by the partial oxidation
process versus the oxidizer oxygen:CH4 feed ratio
w i th oxidi zer temp= 1 OOK, pressure=20 ATM
54
feed ratio with the temperature
pressure
set at 20 ATM.
rates, increasing
hydrogen
cost.
set at 1100K and the
Because of decreasing
the oxygen
to methane
ratio increases
is no way to select an optimum
curve.
There
the
This is also in part due to the increasing
cost of the oxygen which must be supplied
There
production
is a trade-off
between
to the oxidizer.
ratio from this
cost and desired
product purity which can only be solved when the purity is
specified.
steam
In order
reforming
about 95%.
pressure,
Figure
to compare
process,
to the results
the desired
purity was set at
Based on this, for the given temperature
the necessary
oxygen
to methane
and
ratio is 1.0 (see
20).
Figure 25 shows the hydrogen
the oxidizer
temperature
at 1.0 and the pressure
the temperature
reformer
lowest
from the
at which
optimum
temperature
pressure.
of
to steam ratio set
The cost drops as
and then levels off.
temperature
the cost is a minimum.
oxidizer
increases.
set at 20 ATM.
the optimum
Figure 26 shows
oxidizer
with the oxygen
increases
process,
cost as a function
As with the
was taken as the
From Figure 25 this
is seen to be about 1100K.
the hydrogen
cost versus
The cost rises steadily
the
as the pressure
Again, an optimum cannot be determined without
a specification
of the desired
product
pressure.
For
comparison with the reformer process, an oxidizer pressure
of 20 ATM was selected.
Figure
27 shows a cross-section
of the partial
oxidation process similar to that presented for steam
55
22
H
y
D
R
20
0
G
18
E
N
16
C
0
S
14
T
12
$
/
m
b
t
10
8
U
6
900
950
1000
1050
1100
1150
OXIDIZER TEMPERATURE
FIGURE 25:
1200
1250
(K)
Cost of hydrogen produced by the partial oxidation
process versus the oxidizer temperature with
oxidizer
pressure=20
ATM, 02:CH4 = 1.0
1300
56
10
H
Y
D
R
9.5
0
G
E
N
9
C
0
8.5
T
8
$
/
m
b 7.5
t
u
7
10
15
20
25
OXIDIZER PRESSURE (ATM)
FIGURE 26: Cost of hydrogen produced by the partial oxidation
process versus the oxidizer pressure with the
oxidizer
temp= 11 OOK, 02:CH4 = 1.0
30
57
100
P
E
90
R
C 8o
E
N
T
70
60
N
50
S
T 40
R
E
A
30
M20
%
10
FEED
OXIDIZER
HT SHIFT
LT SHIFT
SCRUB
PRODUCT
FIGURE27: Species profile of partial oxidation process. The lines
show the mole percent of each species in the streams
exiting the Indicated unit operations
58
reforming.
oxidizer
The feed is pure methane,
the methane
a dry mole basis.
concentrations
shown,
concentration
drops to less than 5% on
The carbon monoxide
and carbon
are about equal at 20%.
for it was found to completely
oxidizer.
and after the
Oxygen
is not
react in the
This leaves hydrogen at about 57%.
this to the steam reforming
dioxide
Comparing
result shows that more carbon
oxides are formed in this process, yielding less hydrogen
from the primary
reactor.
The two shift reactors
show the carbon monoxide
being almost completely converted to carbon dioxide and
hydrogen.
The scrubber removes all of the carbon dioxide,
and the methanator
monoxide.
reacts the trace amounts of carbon
The final product stream, as desired, is about
95% hydrogen
and 5% methane.
Figure 28 shows the cost of the hydrogen
function of the plant capacity.
as a
Three curves are shown,
each representing a different oxygen purchase cost.
been suggested
that the primary
reason
It has
steam reforming
is
favored over partial oxidation is because of the investment
necessary
to produce
be translated
the effect
As would
the oxygen.
into the oxygen
(11)
purchase
can
cost in order study
of this cost on the viability
be expected,
of the process.
this graph shows that cheaper oxygen
produces a cheaper hydrogen product.
compared
This investment
These curves will be
to the one from steam reforming
later in this
section.
The same general
cost trends as those observed
in
59
14
H
Y
D
02 Cost = 0.08 $/kg
13
R
0
G
1
2
E
N
C
11
0
S
T
Io
$
9
02 Cost = 0.016 $/Kg
/
m
b 8
t
02 Cost = 0.00 1 $/kg
u
-7
_
0
I
50
I
IJ
I
250
200
150
100
PRODUCTION RATE (mscf/day)
-
I
I
300
350
FIGURE28: Cost of hydrogen produced by the partial oxidation process
versus the plant capacity for several oxygen purchase costs
with oxidizer
temp=1 lOOK, pres=20 ATM, 02:CH4 = 1.0
60
steam reforming
are seen in Figure
rate there
is a sharp economy
increases,
the costs
hydrogen
partial oxidation
The operating
of scale.
rise faster
and the
breakdown
is quite different
supplied
The operating
for this process
reactor
(a raw material)
the heat is supplied
externally
cost
than that of
and
This is because
the heat in the primary
cost
- about $1.30
The major cost is raw materials
by the feed oxygen
reforming
BTU hydrogen
take up less of the cost.
oxidation
in Figure 29.
are shown and the operating
than steam reforming.
reforming.
utilities
costs for a 300 mscf/day
plant have been tabulated
conditions
more expensive
oil
Yet as production
than the capacity
and capital
is seen to be $8.74 per million
partial
At low production
cost slowly rises.
The operating
steam
28.
in
is
while in steam
by burning
fuel
(a utility).
Figure 29 also shows the capital
for a partial oxidation
investment
plant of this size.
summary
The total
investment would be about $186 million - $150 million lower
than the steam
capital
reforming
process.
The fact that the
is lower for partial oxidation
to the fact that the primary reactor
is in large part due
is far less
complicated, requiring a lower purchase , setting, and
start-up cost.
61
PLANT CAPACITY: 300 million scf per day
OPERATING PARAMETERS:
Oxidizer Temperature = 1100K
= 20 ATM
Oxidizer Pressure
Oxygen:CH4
Steam:CH4
= 1.0
= 1.0
OPERATINGCOST = $ 8.74 per million btu hydrogen
OPERATING COST BREAKDOWN:
~~-Depreciation 0.5%
~'-
Overhead 0.6%
Labor 0.5%
CAPITAL COST SUMMARY:
Process Units
Other Directs
Setting Labor
Contingency
Working Capital
Start up Cost
$ 10,937,000
4,924,000
6,120,000
6,427,000
96,225,000
62,005,000
-
TOTAL:
$ 186,638,000
FIGURE 29: Operating and capital cost breakdowns for a 300 mscf
per day hydrogen plant using partial oxidation
62
ELECTROLYSIS OF WATER
In the electrolysis
simulation
there are no real
operating conditions which can be varied.
temperature
and pressure
can have some effect on the cell
efficiency
(4), for simplicity
simulation
was assumed
parameter
rate versus
the cell efficiency
constant
of interest,
at 85%.
Figure
in this
The only operating
then, is the hydrogen
the feed flow rate.
relationship.
Although the
production
30 shows this
As would be expected, the production
increases linearly with the water feed rate.
can be used to find the necessary
This graph
water feed rate given a
desired production rate.
Figure
process.
31 shows
the cost performance
In the oxidation
the key cost factor.
critical.
process
of this
the price of oxygen was
Here the cost of electricity
is
The three curves shown in Figure 31 represent
different electricity purchase prices.
In all cases the
cost of hydrogen increases with the plant capacity because
of the accelerating
cells.
cost of high capacity
It is interesting
is observed
electrolysis
to note that no economy
at low production
rate.
of scale
This is because
unlike
reformer or oxidizer reactors, electrolysis units do not
really have a minimum
cost of electricity,
Figure
size.
As expected,
the lower the product
32 shows the operating
the lower the
hydrogen
and capital
cost.
cost
breakdowns for an electrolysis plant producing 300 mscf/day
hydrogen.
The necessary
feed rate is shown and the
63
H 300
2
P
R
0
D
U
270
240
210
C
T
180
I
0
N
150
120
m
s
90
f
60
c
/
d 30
a
y
0
0
5000
10000
15000
20000
25000
WATER FEED RATE (kmol/hr)
FIGURE 30: Daily hydrogen production rate from the electrolysis
system versus the water feed flow rate
30000
64
45
H
Y
D
R
Elec Cost = 0.06 $/kwhr
40
0
G
E
N
35
C
0
S
Elec Cost = 0.03 $/kwhr
30
T
25
/
m
b
Elec Cost = 0.01 $/kwhr
20
t
u
c
15
0
~
-------L
I
~~~
50
II
100
150
PRODUCTION RATE
FIGURE 31;
~
~I1
~
,,I,
200
~~
"I
250
(mscf/day)
Cost of hydrogen produced by electrolysis
versus the
plant capacity for several electricity purchase costs
300
65
PLANT CAPACITY: 300 million scf per day
OPERATING PARAMETERS:
Water Feed Rate = 30,000 kmol/hr
OPERATINGCOST= $ 28.95 per million
btu hydrogen
OPERATING COST BREAKDOWN:
.
Labor
14.8% /
CAPITAL
Raw Materials 0.2%
Overead
\
Deprec
18.7%S
COST SUMMARY:
Process Units
Other Directs
Setting Labor
Contingency
Working Capital
Start up Cost
TOTAL:
$ 1,179,678,000
447,250,000
369,746,000
583,823,000
289,498,000
259,989,000
$ 3,129,984,000
FIGURE32: Operating and capital cost breakdowns for a 300 mscf
per day hydrogen plant using electrolysis
66
resulting operating cost is $28.95 - much higher than
either
of the other two processes.
operating
However,
A large part of the
cost is the utility bill (i.e. electricity).
an unexpectedly
labor, overhead
large part of the cost is in
and depreciation.
These costs are all
associated
with the capital
investment
investment
leads to high depreciation
(e.g. high capital
costs),
and as seen
at the bottom of Figure 32 the capital for this plant is
huge.
For a plant this size running only on electrolysis
units, over $3 billion would be required, and over $1
billion
units
excess
of it would
be spent on enough
to handle the capacity.
large electrolysis
These costs are far in
of those from the previous
processes.
PROCESS COMPARISONS
One of the advantages
as a tool is the ability
to using flowsheet
to easily
compare
alternatives
look for viable versus impossible processes.
shows a comparison
of steam reforming
based on plant capacity.
simulation
and
Figure 33
to partial
oxidation
Two partial oxidation curves are
shown, each representing a different oxygen purchase cost.
The dashed
line represents
reforming.
Clearly,
if the
oxygen cost is high, steam reforming is much more cost
effective than partial oxidation.
oxidation
However, if cheap oxygen
is available,
the partial
process
can become
competitive.
This figure shows that if the oxygen
cost was
-57
9
I
I
1
-
I
H
Y
D
R
8.75
G
E
N
8.5
0
" Partial Oxid: 02 Cost = 0.016 $/kg
8.25
C
0
S
8
T
7.75
7.5
m
b
t
7.25
\ Steam Reforming
U
7
_
0
I1
50
·I
100
I·
150
1I
200
i1
250
I
300
PRODUCTION RATE ('mscf/day)
FIGURE 33: Cost of hydrogenversus production capacity for steam
reforming and partial oxidation
350
68
about $0.001
reforming
per kilogram,
would yield
partial oxidation
and steam
about the same cost hydrogen.
So a
decision of which process is better to pursue can be made
based on the expected
Figure
34
oxygen cost for the site.
shows
similar compiarison of steam
a
and electrolysi s.
reforming
inexpensive
Even
withi relatively
e lectricity, electrol ysis
steam reforming at these
cannot compare to
producti on raites.
The high
co mbined with the electrical
capital cost of electrolysis
utility cost make electrolysis on this scale prohibitive.
This
su ggests
that in terms of hydroge.n production
scale, electrolysis
Finally,
is not
Figure
operati ng and capital
plant
f or the three
terms
of
a vi able
35 prov ides
ternative.
a comparison
of the
costs for a 300 mscf/day hydrogen
processes
s tudied.
As
indicated,
in
operating costs steam reformi ng is favored
somewha t over partial oxidation
favored over electrolysis.
partial
al
on this
and both are greatly
As for capital investment,
the
oxidation plant, using a less complex primary
reactor, is least expensive.
The steam reforming plant is
somewhat more expensive, and the electrolysis
extremely expensive compared to the others.
plant is
59
30
I
,
I
I
I
I
H
Y
27.5
D
R
0
Electrolysis:
Elec Cost = 0.03 $/kwhr
Electrolysis:
Elec Cost = 0.01 $/kwhr
25
G
E
22.5
N
20
C
0
S
T
17.5
15
/
m 10
b
t 7.5
Steam Reforming
.
I
I
,
II
U
5
-
0
-
I
---
50
100
IiI
i
200
250
150
PRODUCTION RATE (mscf/day)
i
I
300
FIGURE 34: Cost of hydrogen versus production capacity for
steam reforming and electrolysis
-I
350
70
·
OPERATING COST COMPARISON FOR 300 mscf/day
Process
Steam Reforming
Partial Oxidation
Electrolysis
·
PLANT:
Hydrogen Cost ($/mbtu)
$ 7.43
8.74
28.95
CAPITAL INVESTMENT COMPARISON FOR 300 mscf/day
Process
Steam Reforming
Partial Oxidation
Electrolysis
PLANT:
Total Capital Investment
$ 332,316,000
186,638,000
3, 129,984,000
FIGURE35: Operating and capital cost comparison for the three
hydrogen processes simulated
71
IV.
CONCLUSIONS
Three different hydrogen production process have
been simulated
and analyzed.
Two of them, steam reforming
and partial oxidation, were similar enough for direct
comparison
results.
of many of the operating
The third, electrolysis,
parameters
and
showed a completely
different type of technology, yet through the simulations,
it was possible
to compare
the ultimate
cost performance
of
this system with the other two.
With the steam reforming
processes,
it was demonstrated
and partial oxidation
how the simulations
used to set the primary reactor operating
to achieve
the most cost effective
could be
conditions
product.
so as
This
optimization not only helps to determine proper operating
conditions
given
the process,
but also the optimum
cost
results could be used during the design process to select
the best process
results
research
above
pathway to follow.
from the simulations
work.
For example,
1200K yielded
In addition,
the
could be used to help guide
in both cases,
no advantage,
temperatures
so work need not be done
on catalysts for higher temperatures.
The final comparison
of steam reforming
to partial
oxidation showed that steam reforming is favored, primarily
because
pure oxygen
is not used in the process.
As had
72
been expected,
unless
the oxygen
very cheap,
needed
in partial oxidation,
imposes a high operating
cost on the
process.
The electrolysis simulation showed that steam
reforming is much more cost effective than electrolysis.
This occurs
even at low electricity
the capital cost associated
of 300 mscf/day
results
cost, suggesting
with electrolysis
is the prohibiting
can guide research.
factor.
that
on the scale
Here again the
The key need is not for
cheaper electricity or more efficient electrical
production.
Instead, research should be concentrated on
manufacturing
a high capacity
same or better efficiency
electrolysis
at a fraction
cell with the
of the capital
investment.
This study has demonstrated
simulation
as a tool
in technological
the power of flowsheet
development.
By
combining process and cost performance, several process
alternatives, even based on different technologies, can be
readily compared.
This comparison aids in eliminating
improbable processes early in development and deciding on
which process is the most viable to pursue in the future.
With a process selected, flowsheet simulation further helps
in driving the research necessary to make the desired
process possible
and cost effective
in the long run.
73
SOURCES CONSULTED
1.
W. Balthasar and D.J. Hambleton, "Industrial Scale
Production of Hydrogen from Natural Gas, Naphtha and
Coal," Int. J. Hydrogen Energy, Vol. 5, pp. 21-33
(1980).
2.
D.F. Balz, H.F. Gettert, and K.H. Gruendler,
"Production of S ynthesis Gas by Partial Oxidation and
High-Pressure Sh ift Conversion," Plant/Operations
Vol. 2 , No. 1, pp. 47-49 (1983).
Progress,
3.
R. Bidard and U. LaRoche,
"Modern Technology
Electrolysis for Power Appication - I. Fundamentals of
Efficient Thermolysis of Water," Int. J. Hydrogen Energy
, Vol. 4, pp. 123-131
4.
(1979).
K.E. Cox and K.D. Williamson,
eds., Hydrogen:
Its
Technology and Implications, Volume I: Hydrogen
Production Technology, CRC Press, Inc., Cleveland, Ohio
(1977).
5.
W.J.D. Escher and T.D. Donakowski, "Competitively
Priced Hydrogen Via Hi gh-Efficiency Nuclear
Electrolysis,"
Int. J. Hydrogen Energy, Vol. 1, p p.
389-399 (1977).
6.
7.
ed., Selected Papers From the Energy
Workshop: Industry Perspectives on Pioneer Process
Pla-nts, Rand, Santa Monica, CA (1981).
R.E. Horvath,
"Hydrogen,
8.
Steam Reforming," Hydrocarbon Processing,
T
April
1982, p. 162 (1982).
"Hydrogen," Kirk-Othmer Encyclopedia of Chemical
Technology, Vol. 12, pp. 938-982, John Wiley & Sons, NY
(1980).
9.
Hydrogen as a Fuel, National
Washington, D.C. (1979).
10. E.W. Merrow,
S.W. Chapel,
Academy of Sciences,
and C. Worthing,
A Review of
Cost Estimation in New Technologies - Implications for
Energy Process Plants, Rand, Santa Monica, CA (1979).
74
11. R.B. Moore, "Economic Feasibility of Advanced
Technology for Hydrogen Production from Fossil Fuels,"
Int. J. Hydrogen
(1983).
Energy,
Vol. 8, No. 11/12,
pp. 905-911
12. R. Shinnar, G. Fortuna, and D. Shapira, "Use of Nuclear
Energy in Production of Synthetic Natural Gas and
Hydrogen from Coal," I&EC Process Design & Development,
Vol. 23, No. 2, pp. 183-194
13. W.F. van Weenen
(1984).
and J. Tielrooy,
"How to Optimize
Hydrogen Plant Designs," Int. J. Hydrogen Energy, Vol.
8,
No.
9,
pp.
689-700
(1983).
14. A.M. Watson, "Use Pressure Swing Adsorption for Lowest
Cost Hydrogen," Hydrocarbon Processing, March 1983, pp.
91-95
(1983).
75
APPENDIX:
ASPEN
Input Files
76
LISTING
:itDROGE
1
N MANUFACTURE --
:
BY JOHN
STEAM REFORMING
E. STRONG. JR.
TITLE 'HYLDROGEN MANUFACTURE BY STEAM REFORMING OF METHANE GAS:
DESCRIPTION
iN-UNIT5
OuT-U'IIT
'"TIS FLOWSHEET MODE-S THE MANUFACTURE OF HYDROGEN
i YMEANS OF STEAM REFORMING OF METHANE GAS."
MET
MET
SYS-OFTIONS INTERFRET=
7
COMPONENTS
COMPONENTSH2 H2/C02 C02/C0 CO/H20 h20/CH4
PROPERTIES
CH4/DEA C4HlIN02-1
SYSOF3
FLOWSHiEET DATA
9
FLOWSHEET
BLOCK REFORMER
BLOCK COOL I
BLOCK HTSHIFT
BLOCK COOL2
BLOCK LTSHIFT
BLOCK COOL3
IN=FEED
BLOCK HEAT1
BLOCK METHATOR
BLOCK M20-DEP
ItI
OUT=HTSFTOUT
OUT=LTSFTIN
OUT=LTSFTOUT
iN=LTS-FTIN
IN=LTSFTi UT
BLOCK SCRB-MIX
BLOCK SCRB-SEP
BLOCK MEA-HEAT
BLOCK MEA-REC
uUT=REFOUT
OUT=HTSFTIN
STEAM
IN=REFOUT
IN=mTSFTIN
IN=HTSFTOUT
IN=SCRUBIN LMEA
I N=SCRUBCOM
OuT=SCRUBIN
OUT=SCRUBCOM
OUT=iMr-A
SCRUBOUT
N=HMEA OUT=MOTMEA
IN=MiTMEA
OUT=RMEA C02EXIT
I N=SCRUBOLT
OUT=MOTSCRUB
IN=HOTSCRuB
OUT=WET-PROD
I N=wET-FROD
OUT=H2EX IT PRODUCT
INPuT STREAM SPECIFICATIONS
STREAM FEED TEMF=500 PRES=20 NPHASE=1 PHASE=V MOLE-FLOW=10000.0
MOLE-FRAC CH4 1.0
STREAM STEAM TEMP=550 FRES=20 NPHASE=1 PHASE=V MOLE-FLOW=40000.O
MOLE-FRAC H20 1.0
STREAM LMEA TEMP=320
PRES=20
MOLE-FRAC DEA 0.0b84/
I
NPHASE=1
PHASE=L
MOLE-FLOW=50000.0
H20 O.9316
BLOCK DATA
FORTRAN SET-INIT
DEFINE XFM TREAM-VAR STREAM=FEED VARIABLE=MOLE-FLOW
DEFINE XC1 CBLOCK-VAR CBLOCK=CHTS SENTENCE=SIZING-DATA &VAR I ABLE=CAP
DEFINE XC2 CBLOCK-VAR CBLOCK=CLTS SENTENCE=SIZING-DATA &
VARIABLE=CAP
DEFINE XC3 CBLOCK-vAR CBLOCK=CMETH SENTENCE=SIZING-DATA &
VAR 1ABLE=CAP
F
XCI=xFM
F
F
XC2=XFM
XC3=XFM
EXECUTE FIRST
BLOCK REFORMER RGiBBS
R
NATC,
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NATOM=3 NREAC=2
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BLOCK CL3
HEATER
FARAM TEMP=320 PRES=O
BLOCK SCB.E-MIX
CO 1
t1H15
.
NPHASE=I
PFASE='v
HEATER
PARAM TEMP=320 PRES=O
BLOCK SCSB-SEP
SEP
FRAC STEEAM=MEACOMPS=DEAC02 h20 FACS=1.0
FLASH-SPECS
SCRuBkuT NPF'ASE=1 PHASE=V
BLOCK MEA-HEAT
EATER
PRES=O
PARAM TEMP=400
BLOCK MEA-REC
NPHASE=1
PHABE=V
SEP
FRiACSTRE=RMEA COMPS=DEA FRAhS=1.0
1.0 0.5
Fk;H:
TMP-=58.v: F
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F'HAE=v
78
BLOCK.MiETHATr:STOIC
F'RA:
STOiC
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TEM'=o.':
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PRES=-2
CO -1/
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NHASE=
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SI Z i NG-"ATA LA=`:,
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REF-CU=15
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CELOCK CSCETu,
u,-T
ER:
SIZ iNG-DATA DIM5FT
CNT
COSTIN
G-T TE
ES=5 TEM=3
=
NEUIU=iF=
T
USER-CORR REF-COST=5 0Ci
REF-C=50O)
EXPON-'0.&5
100 MiX-CAF = 1, C.,E¢
M I -CAP=
TRA-TOWER^
D i kM=5FT
> NTR,
CBLOCi. CSCRBU
S I Z I N-ATA
= 1,
ES=5
TMF'=350
IhiEQiOIP=1
ON
COSTINC-LATA T YE=rue:' I TF'T
CBLOCK CNOC.k V-VESMSEL
SIZING-DATA
COST IN-DATA
DIAM=5-.FT
NE2CuP =1
CELOCti.
CMEA H'ETX
REFERENCE SELL
REFERENCE TUBE
Si Z ING-D
T
CBLOCK CCGOL
TT-LENGTH=1TH
10FT'.
TEMP=50(.. FRES=5
LTILIiTY=STM1
LCCk=MEA-HEAT
U=35t. ETU./HR-SCFT-R
rHEATX
REFERENCE SHELL TlLITY=COOL-H20
REFERENCE TBEE BLOCKi=COL1
SIZ ING-DA
U=50)..EiTu/H
R-SFT-R.R>
CELOCK CCOOL2 HEATX
REFERENCE SHELL UTILITY=COOL-H20
REFERENCE TUBE BLOCK=COOL2
SI ZING-DATA
U=50<BTU/HR-SQFT-R>
CBLOCK CCOOL3 HEATX
REFERENCE SHELL UTILITY=COOL-H20
REFERENCE TUBE BLOCK=COOL3
SI ZING-DATA
U=5'i<BTU/HR-SQFT-R>
CELOCK CHEAT1
EATX
REFERENCE SELL UTILITY=STM1
REFERENCE TUiE BLOCk.=HEAT1
SI Z ING-DATA
=50<.BTU/HR-SQFT-R.>
FORTRAN REF-DUT
DEFINE XD1
LOCK-VAR BLOCK=REFORMERSENTENCE=F'RAm &
~Li'E±
i
jEt
fL
__r_.-vi.
EC
L L i-I .- 1-
-
S~=
I I
-LiT
&I
79
V RiLE
i- =DO
GF
i
X=X
C8~C~:
RE:T
hiNACE
uTA=I,0i
SiZi -- DTA
C:-STiNG-LAT
TYPE=REF0FRmERF
FRS=E
UTiiTY GAS=FUEL-hAS
STi
UL-TY
STEAM
OiuCE
F ':uCimASED
F OFERTiES
E YSO-P12
FARMi COMPONENT=i:H20 TIN=70)C
TOUJ=700
FRES=15
COST FRICE= .CE-3$/.LB>
UTILITY FUEL-GAS GAS
SOURCE FiURCmASEO
COST
i</k:hR.>
FRICE=Ol!1
:.
UTILITY COOL-m20 wWTER
SOURCE PURCHHSED
PROPERTIES SYSOF12
FRAM
COMF'ONENT=H2
COST FFRICE=1.2E-4$/L >
RAW-MATERiAL Cm4- I
REFERENCE STREAMr=FEEL.
COST FRICE=
.Ci$5(/LB,
RAW-MATERiALSM-iN
REFERENCE STREAM=STEAM
COST F'RICE=2 ..E-3..$,/LB>
UNIT REFORM
CBLOCKS CREF CrTS
CLTS CMETH
CSCRB
NOCK CMEA CCOOL1
CCO L2 CCOOL3 CMEATi
OPERATI NG-COST
OPERATiNG-LADR
AVAILAiILITY
NOPER=2 RATE=10 .00
FACTOR=0.5
PROJECT-DATES
START APRIL
19%=
LABOR-COSTS
WAGES RATE=20.0
JOB-CONDITION
LOCATION=MiT
COST I NG-OPT I ONS OFER-COST
SENSITIVITY RESULTS
DESCRIPTION 'STEAM REFORMING RESuJTS'
DEFINE S1 MOLE-FLOW STREAM=FRODUCT COMPONENT=H2
DEFINE S2 MOLE-FLOW STREAM=PRODUCT COMPONENT=C4
DEFINE S3 OFER-COST-VAR SENTENCE=RESULTS &
VAR iA LE=NET-OF-COST
DEFINE S4 GF'rER-ClST-VAR SENTENCE=AVAILABILTY
VARIABLE=FACTOR
9;
F
:
F
t
F URIT:
ULCULATE
5 = ' i S1+
j
CALCaULATEVOLUME FER DAY FRODUCED (SCF):
6= 110
CALCULATE
F
.0*24. 0*S4*0 .791/ Eo
COST PER
7=S3i S i iu
MILLION
STU:
.0*24.0*3o5.0*54*27 1 .3/1Eb)
&
&
TAiBULH2E i 5
CO-L-uhktBEL= '
iN
'PFRODCTi
'STREAM'
TBUL'A'E
2 SO COL-LbEL='
FDUCT'
'RATE
'OF H2'
'
LNiT-LA~E;"=
MsCs riA
TSdJLATE 3 57 C_-LAs.EL ='PD3'J
' ST
UNiT-LABEL='
VARY STREAM-vAi
STREAM=STEAM VARIABLE=MOLE-FLOw
RAN6E LIST=2C!C
$
2300.j0300v0 35 ii)
4000't) 45000 50000
80
/MET
81
LISTING
2
--A F
RiHL AILTN
JC]OiHr
E. ~S.TAO'if:, JR.
-iLLOGEN MANUFLkCTE
:
SB
TiTLE
hT'LOcGEN
MANUFACTUREBY FATiAiL OXIDATlON OF HYDROCABONS'
ltSCF-TI- IO
"T iS FLOWSHEET MODELS THE MiAN'FkCTuRE
OF HYDROGEN
E, MEANS OF FARTI IAL OXIDATIOiN COF HTfDOCARi3ONS:
r-EL:
LCAN iNCLDEL METHANE = CH4
iNA~'HTHA= C
.,"
COAL
= CH . i
IN-UNITS
ET
ET
i NTER:FET=Co
OUT-UNITS
S'YS-iT -TION
RUN-CONTROLMA-T IME=500
COMPONENTS
COMPONENTSH2
DEA Cm11
PROFERT IS
2/C02 LC02/C C20
H2/CH4
CH4/02 02/
NO2-1i
rS
fSOF3
FLlW.SmiEET DATA
II
t LUWr:E
BLOCK OXiDIZE R
BUOCK LL3.2
BLOCK
BLUICK
BLOCK
BLOCK
BLOCK
BLOCK
TSHIFT
COiOL
S5CR-M i X
SCRB-SE
MEA-mEATH
MEA-REC
BLOCK HEAT1
BLOCK METHATOR
BLOCK h20u-SEF
iN=FEED ST EHM OFEED OUT=GXOUT
iN=OXOUT
OUT=HTSFTIN
iN=HTSFTiN
IN=HTSFTO'T
iN=LTSFTIN
IN=LTSFTOuT
IN=SCRUBIN LEA
IN=SCRuBCOM
OUT=HT FTOUT
OUiT=LTFTIN
OuT=LTSFTOUT
OuT=SCRuBIN
OUT=SCRUBCOM
OUT=IHMEA SCRUBOUT
iN=HfEA
OUT=OTMEA
iN=HOTMEA
OUT=RMEA CO2EXIT
iN=SCRUBOUT
OiT=HOTSCRUB
iN=HOTSCRUB
OUT=wET-FROD
IN=WET-FROD
OUT=H2OEXIT PRODUCT
INFUT STREAM SPECIFICATIONS
STREAM FEED TEMP=500 PRES=20 NPHASE=I PHASE=V MOLE-FLOW=10000.0
MOLE-FRAC CH4 1.0
STREAM STEAM TEMP=550 PRES=20 NPHASE= PHASE=V MOLE-FLOW=10000
MOLE-FRAC h20 .0
STREAM 02FEED TEMP=500 PRES=20 NFHASE=1 PHASE=V MOLE-FLOW=1000)
MOLE-FRAC 02 1.0
STREAM LMEA TEMP=320 PRES=20 NPHASE=
MOLE-FRAOC EA 0.0684/ H0
:
PFHASE=L MOLE-FLOW=50000.0
0.9316
BLOCK DATA
BLOCK OXIDIZER RGIBBS
PARAM PRES=20 TEMPF=10C,)NPHASE=1 VAPFOR==i&
NATOM=3
PROD H2/ C/
ATOMS
ATOM H2
1
C02/ CH4/ H20/ 02
C
(/
H
2
O
-0 - . ' ;
Cb4 i
F;
i . .'
i i
-,'
4i.
ii
( /i
i/
2
_2
1 0t. /
±
Ii5
82
BLOCK COO L
hEA-ER
FARAv;
i
TEiF'70
PRES= ) NFiHASE= 1 Pi4SE=
/'
BL-OC: HTSHIFT
RGIBBS
F'ARAM TEP=670
FP'RES=-2 NPrHASE=l
NM.TOM'!=3NREAC=3
H20/M4I/
ii/
F;RDDH2' 20/ C2/
ATOM!S
C
/
4
/
(
;
i
0
U
CM4 1
2!Z,2
;AFOF:=1 &
/
/
_1 /
1
TAPFF-SP·E
STOIC 2 C4
EXTENT-F'EC
;
1 1
-,/ h.'
2 .0
STOIC 3 r2 --
r,
E T E'-.
.-
-i/
20 2
-.. 1
BLOCK COOL2 HEATER
PARAM TEMPF=470 PFRESI=
NPHASE=l
BLOCK LTSHIFT RG6iBS
PARAM TEMP=470 PRES=-2
NMTOM=3 NREAC=3
PROD H2/ C0/
02/ H20/
;
ATOMS
ATOM
/
-. O 1 H2 3
NPFHASE=1 VAPOR:
H
C
1
/ 2 /
4 i
I
C2
1
/
02z 1
i /
H20
i
O/i
02
1'
0 /
STOIC i CO -/i H20 -1/
TAFP-SPECl
1 14
&
CM4/ 02
C
H2
CH4
PAISE=V
0
/
O
0 /
H2 ii
0 /
0 /
1/
2 /
/1
2
C02 1
STOIC
CH4 -ii
H20 -1/ CO 1/ M2 3
EXTENT-SPEC 2 0.0
STOIC 3 2 -2/ 02 -1/
EXTENT-SFEC 3 0.0
BLOCK COL3
HEATER
PARAM TEMP=320
PRES=O
H20 2
NPHASE=
PHASE=V
BLOCK SCRB-MIX HEATER
PARAM TEMP=320 PRES=O
BLOCK SCRB-SEP SEP
FRAC STREAM=HMEACOMPS=DEA C02 H20 02 FRACS=1.0
FLASH-SPECS SCRUBOUT NPHASE=1 PHASE=V
BLOCK MEA-HEAT HEATER
PARAM TEMP=400 PRES=O NPHASE=1 PHASE=V
BLOCK MEA-REC SEP
FRAC STREAM=RMEA COMPS=DEA FRACS=1
BLOCK EAT 1 HEATER
PARAM TEMP=580 PES=O
.
NPHASE=1 PHASE=V
1.0
.. 5 1.0
F'LrA. TEF'=o.C. ~ F'r:-:E=-2 NFMASE=i F·'ASE=.,
..M1' -3
STOIC 1.,C.2-i
!ESR
C it"... i 'ii:
83
CH4 1/ H20 1
,':
_ >
ti' -!.. k_'!--,i
_
_-_ _ _ _
tz
PCMPS=HO
M=
E i xIT
r
RE
S
FRACS=1 .0
-,GTINGSECTION
COS;7IN
CLOC. CT
LSEf
CoST NG-DATA
I
NEQLii=,
CELOCKC CLTS USER
T .O'tJ;At
USER-COST PURfCllr-T=52:
COSTI NG- TAHNC:U I F= 1
l
CBLOCK ClETH LiESE
USEF-COST
PJRCH-COS5T==1 iCiih,':,u
COSTI NG-ATA
I F= 1
NE.
CBLOCK CSCRUE TA.:F-TOwER
SIZINel-IT
Z, ih'=
F
NTR=10i
r;FT.
,PRtS=COSTING-EL AMT
TTF0E=SORFTiOlN
NECQ!UIP=1
C'.KNOCr V-vESSEL
DH
Li==5<FT.'
CiOSTi NG-DAT
*
E'Q.
IF=1
TEMPi=35O
CBLOCk
SIZING-iTIA
C'LOCK
CMrEiA
TT-LENT-=1'::.
FT> TEMF=500J FRES=5
.EATX
RE.EN
.E
rlEL UTliITi=ST!:
REFEENC.E TEE
GC..=MEA-EAT
S.i ZI-OA
C
_LOC..
: u=35J.
.. . u/-tFT-..
Oi
.ETa
uTI Li=CO.'.-i2
T =
BLiOC.=COOL1
REFERENCE
HELL
REFERENCE
TE
SI ZING-DATA U=50K<BT/HR- -FT-R>
CELOCK CCOOL2 HEATX
REFERENCE SHELL uTILITY=COOL-H20
REFERENCE
TUBIE
LOC..K=COOL2
SI ZING-DATA U=50<BTU/HR-S!FT-R.>
CBLOCK. CCOOL3 HEATX
REFERENCE SHELL UTILITY=COOL-H20
REFERENCE
TUBE
BLOCK=COOL3
SI Z ING-DATA U=50<BTU/HR-SQFT-R
CBLOCK CEATI HEATX
REFERENCE SHELL UTILITY=STM1
REFERENCE TUBE BLOCK=HEATI
SIZ ING-DATA U=50<BTU/HR-SQFT-R>
,FORTRAN Ox-DUTr
DEFINE XD1 BLOCK-vAR BLOCK=OXIDIZER
VAR I ABLE=CALC
DEFINE XD2 CBLOCK-VAR
VARIiABLE=DUT!
XD3=xE!
CBLOCK=COXID
SENTENCE=PARAM &
SENTENCE=SIZING-DATA
&
.... ;UJr,
- _
84
U,,iL'i'-'E
CLLOCK COXiD FRNACE
SiZiNG-LTAi
DiuT Y= (i
TYFE=F.OCH
C'STiNG-4AT
UTILITY GAS=FUEL-GAS
UT I LTT
F ES=t=
STm1 STEAM
SOROCE PHURCHASED
FRF'ERTiES SYSOF12
PARAt COrM'RONENT=H2 TIN=700
TOUT=700 F'RES=15
CE=2. E-3$/LB
COST F'
UTILITY FUE_-GAiS GAS
SOURCE F'URC
;;HASED
COST F'RI CE=i:.
UTILITY
SOURCE
OOL-20
1 ..
FW
MR>
wATER
'dRCHEASED
PROFERTIES SF
12
FARAM COMF;O,!NENT=i2-
COST
RICE=1 .ZE-4:..$L
5
RAw-MAiTERIAL Ci4-IN
REFERENCE
STREAM=EED
COST F-RI E=C.(i:5. iLb.
RAw-MATER iAL
REFERENCE
ST- I N
STREAM=STEHM
COST FRICE= 2.E-
3 .$LB.l>
RAW-MATERI AL 02- i N
REFERENCE ST7R
EAm=]r'
COST FRiCE=.t0l7"
.:$, L)'>
UNIT
REFORM
CELOCKS COXID CHTS CLTS CMETH CSCRUB CKNOCK CEA
CCOOL2
CCOOL3 CHEAT1
OPERAT ING-COST
OPERATING-LABOR NOPER=2 RATE=10.00
AVAILABILITY FACTOR=0.5
PROJECT-DATES
START
MAY
1 So
LABOR-COSTS
WAGES RATE=20.0
JO-COND ITION LOCATION=MIT
COSTING-OFT IONS OPER-COST
SENSITIVITY RESULTS
DESCRIFPTON 'PARTIAL OXiDATiON RESULTS'
DEFINE S1 MOLE-FLOW STREAM=FRODUCT COMPONENT=H2
DEFINE 52 MOLE-FLOW STREAM=PRODUCT COMPONENT=CH4
DEFINE 53 OFER-COST-VAR SENTEiNCE=RESULTS &
VARIABLE=NET-OF-COST
DEFINE 54 OFER-COST-vAR SENTENCE=AVAILABILiTY &
VARIABLE=FACTOR
F
IF(S1.EQ.O.0) 51=1.O
2=1.
IF(S2.E0).C.)
F
;
CALCULATE PURITY:
A=SS .' i ! F
CCOOL1 &
F
So=S * i
00 i*.
C,*;4*j.
71
1 Eo
TE COST PER MILLION BTU:
S7=S/(Si*i0i0.0i*24.O*3bS.0*4*271
3/IE
85
CALLL
F
..
TABULATE 1 5 COL-LABEL=i% H2' 'iN' 'PRODUCT 'STREAM'
TABdLATE 2 S CO-LHbEL='PRDLiL-TT TE' 'OF H2' &
UN7T-LABEL='MSCFiDYAY
TABULATE 3
7 COL-LABEL=PRODULCT
'COST'
UNiT-LABEL='$/METU'
STREAM-LAR STREAM=2-FEED VR I ABLE=MOLE-FLOW
9RANE
L ST=5000 10000 15(500
VAR
VARY
L5--.-VAR BLGCK=OAiDiZE
RANGE LIST=i0
15 20 25 30
R SENTENCE=PARAM
VARIASLE=FRES
86
LISTING
r YDROGEN MANUFACTUE --
i:
Bt
:
3
ELECTROLSSIS
JOHN E. STRONGl,
JR.
TITLE 'htDROGEN MANUFACTURE BY ELECTROLYSiS OF WATER'
DESCRIP iihD "H IdS FLOWSHEET MODELS THE MANUFACTURE OF HYDROGEN
BY MEANS
OF ELECTROLYSIS
OF WATER"
IN-UNiTS MET
OUT-UNITS MET
SYS-OFTION23 INTERPRET=O
:
COMPONENTS
COMPONENTS H2 H2/02 02i20
PROPERTIES
H20
S fSOP3
FLOSnMEET DATA
FLOWSiHEET
BLOCK CELL
IN=FEED
OUT=GASES
BLOCK CELLSEP IN=GASES OUT=H2UP 02uP
BLOCK M2t.NOCK
I N=m2UP
OUT=FRODUCT
BLOCK
:
2KNOC.
iN=02uP
INFUT STREAM
OUT=O2EXiT
SPECIFICATIONS
STREAM FEED TEMF=340 FPRES=1NPASE=l
MOLE-FRAC
M20
PASE=L
MOLE-FLOW=30000
1i.
DATA
;Z BLGOC.
BLOCK CELL RSTOIC
PARAM TEMP=340 PRES=l
STOIC 1 MIXED H20 -1/ H2 1/ 02 0.5
CONV 1 MIXED H20 1.0
BLOCK CELLSEP SEP
FRAC STREAM=H2UP COMPS=H2 FRACS=1.0
BLOCK H2KNOCK MIXER
BLOCK 02KNOCK MIXER
;
COSTING SECTION
FORTRAN COST-SET
DEFINE XF STREAM-VAR STREAM=FEED VARIABLE=MOLE-FLOW
DEFINE XC1 CBLOCK-VAR CBLOCK=CCELL SENTENCE=USER-COST &
VARi ABLE=PURCH-COST
DEFINE XC2 CBLOCK-VAR CBLOCK=CCELL SENTENCE=UTILITY &
VARIABLE=ELEC-RATE
F
F
XCl=20000.O*xF
XC2=93.47*XF
EXECUTE BEFORE CCELL
CBLOCK CCELL USER
COSTING-DATA NEQUIP =
'ii 1
2Y :
E t-t
E=;OCuE
t=O..'
87
UNi T EECTOi
_;'!
CELOtC:.S CCELL
UTILITY PFJwER ELECTRICITY
SOURCE PFOuChhASED
uCOT FIh E=i.03
RAw-MATERI-L
REFERENCE
20-IN
STREAM=FEED
COST FPiCE=2 .3E-4:$/LB>
OFERATI NG-COST
vAiLHBi.iT
FACTO=.5:
OFERATiSG-LABORNOFER=i ,RttE=1 .0'
FROC.ECT-DATES
START M
io
LABOR-COSTS
wAGEt Ri-TEs`T':i .00
JOB-CONDL I u!iO LOCiATiON=M IT
COST I NG-OFPTNS
OF'ER-COST
SENSITIVITY RESULTS
DESCRIPTION
ELECTROLYSiS RESULTS'
DEFINE S MOLE-FLOw STREAm=FRODUCT COMFONENT=H2
DEFINE 53 OPER-COST-VAR SENTENCE=RESULTS &
VARIABLE=NET-OP-COST
DEFINE S4 OFER-COST-VAR SENTENCE=AVAILABILITY &
VARIABLE=FACTOR
CALCULATE VOLUME PER DAY PRODUCED
F
(MSCF):
R1=Si* 100. (O0*24.0*S4*0 .791i1E6
CALCULATE COST PER MILLION BTU:
F
R2=53i/,Si*100(0.
0.*24.t0*365.0*S4*271.3/1E6)
TABULATE 1 R CL-LABEL='FRODUCT'
UNIT-LABEL='MSCF/DAY'
TABULATE 2 2 COL-LAEL='PRODUCT'
'RATE' 'OF H2' &
'COST' UNIT-LABEL='$/MBTU'
VARY UTILITY-VAR UTILITY=POWER SENTENCE=COST VARIABLE=FPRICE
RANGELIST=.C1 0.03 0.0b
VARY STREAM-vAR
STREAM=FEED VARIABLE=MOLE-FLOW
RANGE LIST=500 1
$
v00
2500 5000
10000 20000 300,0