International Journal of Engineering Research in Africa
ISSN: 1663-4144, Vol. 30, pp 144-153
doi:10.4028/www.scientific.net/JERA.30.144
© 2017 Trans Tech Publications, Switzerland
Submitted: 2016-10-03
Accepted: 2017-03-02
Online: 2017-05-25
Process Development for Hydrogen Production via Water-Gas Shift
Reaction Using Aspen HYSYS
Idowu Iyabo OLATEJU1,a, Crowei GIBSON-DICK2,b,
Steve Chidinma Oluwatomi EGEDE3,c and Abdulwahab GIWA4,d*
1-4
Chemical and Petroleum Engineering Department, College of Engineering,
Afe Babalola University, Afe Babalola Way, Ado-Ekiti, Ekiti State, Nigeria
a
ajalaid21@yahoo.com, bnyeroakporehe@gmail.com, clordsteve95@gmail.com,
d
agiwa@abuad.edu.ng
Keywords: Hydrogen, water-gas shift reaction, Aspen HYSYS, process development
Abstract. The development of a process for the production of hydrogen through water-gas shift
reaction has been developed and simulated in this work using Aspen HYSYS. This was achieved by
picking the pieces of process equipment of the plant from the appropriate section of the Aspen
HYSYS environment and connecting them together through appropriate streams. In addition, the
components involved in the process were selected from the Aspen HYSYS databank. PengRobinson Stryjek-Vera (PRSV) was used as the fluid package of the developed process for property
estimation during the simulation. The reaction of the process was modelled as an equilibrium type,
the equilibrium constant of which was estimated using Gibbs Free Energy. From the results
obtained, it has been established that pure hydrogen can be obtained from a plant comprising of a
mixer, a reactor (with approximately 80.07% conversion of the reactants), a separator and two heat
exchangers based on the fact that the mole fraction, the mass fraction and the volume fraction of
hydrogen obtained from the simulation carried out when carbon monoxide and steam were passed
into the process plant at room temperature (25 oC) and boiling temperature of water (100 oC),
respectively under atmospheric pressure was approximately 1.
Nomenclature
CO Stream
Final Product
Keq
PRSV
Qheat 1
Qheat 2
R
Reactor Bottom
Reactor In
Reactor Top
Separator Bottom
Separator In
Separator Top
Steam Stream
T
ΔG
Carbon monoxide feed stream entering the mixer
Liquid hydrogen obtained as the final product from heat exchanger 2
Equilibrium constant
Peng-Robinson Stryjek-Vera
Heat duty of heat exchanger 1
Heat duty of heat exchanger 2
Universal gas constant (kJ/kgmole-K)
Bottom stream leaving the reactor
Mixed reactants leaving the mixer and entering the reactor
Top stream leaving the reactor and entering heat exchanger 1
Bottom stream leaving the separator
Material stream leaving heat exchanger 1 and entering the separator
Top stream leaving the separator and entering heat exchanger 2
Steam feed stream entering the mixer
Temperature (oC, K)
Gibbs Free Energy change
Introduction
The hydro-treatment of finished products of refineries has become critical due to changes
occurring in fuel regulations. These changes are related to the specification of more clean fuels with
special focus on reduction of sulphur content. To achieve these goals, more hydro-treatment is
needed, and hydrogen is the main raw material for hydro-treatment units [1].
All rights reserved. No part of contents of this paper may be reproduced or transmitted in any form or by any means without the written permission of Trans
Tech Publications, www.ttp.net. (#90115484-26/05/17,23:02:45)
International Journal of Engineering Research in Africa Vol. 30
145
In refineries, a hydrogen generation unit is very necessary to supply the needed hydrogen to
all requiring processes. The most widely adopted technology for this purpose is known as steam
reforming, which has the highest energetic efficiency and the best cost-benefit ratio on large scale
[2].
The hydrogen generation of this process is normally carried out in two stages in series. The
first one, which is the steam reforming, converts steam and a hydrocarbon into synthesis gas.
Furthermore, a reactor known as water gas shift reactor (WGSR) converts the carbon monoxide
present in the synthesis gas into carbon dioxide and more hydrogen is, thereby, generated.
Consequently, the WGSR, though an intermediate step of hydrogen generation process, plays a key
role in any petrochemical plant due to increase in hydrogen demand [1].
Generally, the hydrogen generation unit that is based on steam reforming process is
responsible for approximately 95% of hydrogen generated [2]. In this process, sulphur is first
removed from the hydrocarbon stream (usually natural gas) so as to prevent catalyst poisoning and
deactivation through the use of a guard bed. Thereafter, steam is mixed with the main stream in a
fixed steam to carbon molar ratio. The steam reformer is normally a multitubular catalyst filled
furnace reactor where the hydrocarbon plus the steam are converted into synthesis gas at high
temperatures (between 700 – 850 ºC) according to the following reactions [1]:
CH4 + H2O ⇔ CO + 3H2
CO + H2O ⇔ CO2 + H2
ΔH0 298= 205.9
kJ/mol
(1)
ΔH0 298= -41.1 kJ /mol
(2)
Both reactions are reversible, but the reaction in Equation (1) is endothermic and while that
of Equation (2) is moderately exothermic. The reaction (1), due to its endothermic nature, occurs in
a lower extension in the steam reforming reactor. The synthesis gas stream composition is CO and
H2, which are later converted to CO2 and H2 that are also present in gaseous state. The main purpose
of the water gas shift reactor (WGSR) is to carry out the reaction (2) by reducing the CO fraction
and increasing the hydrogen yield. After the reaction has been accomplished, the product of the
WGSR stream is conducted to a purification section, where hydrogen purity is increased
accordingly [1].
Actually, some researches are ongoing about producing hydrogen via reactive distillation
process, which is a process that combines reaction and distillation in a single piece of equipment [3
- 23], others have already been carried out on this subject matter. For instance, Giunta et al. [24]
simulated a low temperature water-gas shift process using a heterogeneous model. Brunetti et al.
[25] studied the simulation of water-gas shift process by employing a membrane reactor. Chen et al.
[26] studied modelling and simulation of hydrogen generation from water-gas shift reactions at high
and low temperatures in an attempt to aid in recognizing the detailed phenomena of the process by
taking the effects of both the reaction temperature and the steam/CO ratio into account. Boutikos
and Nikolakis [27] studied the impacts of operating and design parameters on the performance of a
water-gas shift membrane reactor by considering factors such as the effects that relative values of
membrane permselectivity, permeation flux and reaction rate had on the performance of a water-gas
shift membrane reactor through simulation Smith et al. [28] used Computational Fluid Dynamics to
carry out the analysis of a water-gas shift membrane reactor. Falco et al. [29] studied the simulation
of water-gas shift membrane reactors using a two-dimensional model and investigated the effect of
hydrogen removal on membrane water-gas shift reactors analysing the performance of the reactor
type. Giwa and Giwa [30] simulated, analysed and optimized a hydrogen production process
involving steam reforming and water-gas shift reactions. Giwa et al. [31] employed Aspen Plus to
investigate the effects of some operating parameters, which were carbon monoxide feed molar flow
rate, steam feed molar flow rate, reaction temperature and reactor pressure, of water-gas shift
process that was used for hydrogen production.
Based on the literature review carried out, it has been discovered, to the best of our
knowledge, that no work has addressed the issue of developing a simple process route for hydrogen
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International Journal of Engineering Research in Africa Vol. 30
production via water-gas shift reaction using Aspen HYSYS. As a result of that, in order to bridge
that gas, it is aimed in this work to use Aspen HYSYS for developing a simple process for the
production of hydrogen using water-gas shift reaction.
Methodology and Process Description
The model of the plant used for the production of hydrogen via water-gas shift process in
this work was developed with the aid of Aspen HYSYS [32]. The plant had two feed streams – one
containing carbon monoxide and the other having steam.
The pieces of process equipment (a mixer, an equilibrium reactor, two heat exchangers and a
separator) of the plant were picked from the appropriate section of the Aspen HYSYS environment
and connected together with appropriate streams to form the plant as shown in Figure 1. Also, the
components involved were selected from the databank of HYSYS and attached to PRSV (PengRobinson Stryjek-Vera) fluid package.
Figure 1. Aspen HYSYS plant of water-gas shift process
The stream containing the carbon monoxide was passed into the mixer at a temperature and
a pressure of 25 oC and 101.3 kPa, respectively and at a flowrate of 55 kgmole/hr. 70 kgmole/hr of
the other stream, which was mainly steam, was passed into the mixer at a temperature of 100 oC and
at the same pressure like that of the carbon monoxide stream and the two streams were mixed
together before being passed into the reactor. The reactor was an equilibrium type whose liquid
level was set to 50%. After the reaction, which is given in Equation (3) and the equilibrium constant
of which was estimated using Gibbs Free Energy by applying the relationship given in Equation (4),
had occurred in the reactor, the top product was passed into a heat exchanger so that the given
vapour could be liquefied and maintained at a pressure of 101.3 kPa before entering the separator.
The liquid bottom product of the separator was expected not to have any hydrogen while the top
vapour leaving the separator was condensed by passing it through a heat exchanger and setting its
vapour fraction to zero (0).
CO  H 2 O
ln K eq 

G reaction
RT
CO2  H 2
(3)
(4)
International Journal of Engineering Research in Africa Vol. 30
147
Results and Discussion
The results obtained from the simulation of the developed Aspen HYSYS model as the heat
of reaction of the process was  4.110 4 kJ kgmol , and this revealed that the developed model of
the process was a valid one because the value was found to correspond to the one obtained from the
literature and given in Equation (2).
Furthermore, the balances across each of the pieces of equipment involved in the plant were
taken for the inputs and the outputs, and they are outlined thus.
Balances across the Mixer
Given in Table 1 are the conditions of the input materials entering and those of the outputs
coming out of the mixer of the plant. As can be seen from the table, the carbon monoxide entered at
room temperature (25 oC) while the steam was at the boiling point of water (100 oC), and the
pressures of the two input streams were both 101.32 kPa. Also looking at the table, the mixed
stream coming out of the mixer was found to be at the same pressure as those of the inputs, but its
temperature was found to be between those of the two input streams (83.95 oC). This result has
demonstrated one of the principles of mixing that heat lost by one material is gained by the other.
All other conditions of the streams around the mixer are given in Table 1.
Table 1. Conditions of the inputs and the outputs of the mixer
Name
CO Stream
Steam Stream
Vapour
1.00
1.00
Temperature [C]
25.00
100.00
Pressure [kPa]
101.32
101.32
Molar Flow [kgmole/h]
55.00
70.00
Mass Flow [kg/h]
1540.60
1261.06
Std Ideal Liq Vol Flow [m3/h]
1.93
1.26
Molar Enthalpy [kJ/kgmole]
-110597.80
-239347.67
Molar Entropy [kJ/kgmole-C]
159.09
181.15
Heat Flow [kJ/h]
-6.08E+06
-1.68E+07
Reactor In
0.99
83.95
101.32
125.00
2801.66
3.19
-182697.73
177.36
-2.28E+07
In Table 2, the molar compositions of the inputs and the outputs around the mixer are given.
According to the results presented in the table, mixing of the input materials was achieved because
the mole fraction of the output of the mixer was different from those of the input streams. This was
found to be an indication that the mixer used to simulate the process was functioning properly.
Table 2. Molar compositions of the inputs and the outputs of the mixer
Component
CO Stream
Steam Stream
Reactor In
CO
1.00
0.00
0.44
H2O
0.00
1.00
0.56
Hydrogen
0.00
0.00
0.00
CO2
0.00
0.00
0.00
Component
CO
H2O
Hydrogen
CO2
Table 3. Mass compositions of the inputs and the outputs of the mixer
CO Stream
Steam Stream
Reactor In
1.00
0.00
0.55
0.00
1.00
0.45
0.00
0.00
0.00
0.00
0.00
0.00
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Table 4. Volume compositions of the inputs and the outputs of the mixer
Component
CO Stream
Steam Stream
Reactor In
CO
1.00
0.00
0.60
H2O
0.00
1.00
0.40
Hydrogen
0.00
0.00
0.00
CO2
0.00
0.00
0.00
Tables 3 and 4 showing the mass compositions and the volume compositions of the
materials around the mixer also supported the fact that proper mixing was achieved because each of
both of them was different from the values of the input streams.
Balances across the Reactor
The conditions of the material streams in and out of the reactor used in this process were as
given in Table 5. As can be observed from the table, the main output stream of the reactor was the
top one because the molar flow, the mass flow and the volume flow of the bottom stream of the
reactor was zero (0). Also noticed from the results given in Table 5 was that, even though the
pressure was still atmospheric, the temperature of the output (top stream) of the reactor was far
higher than that of the input stream. This was as a result of the exothermic nature of the reaction
occurring in the reactor as was observed from the negative sign of the reaction heat.
Table 5. Conditions of the inputs and the outputs of the reactor
Name
Reactor In
Reactor Bottom
Vapour
0.99
0.00
Temperature [C]
83.95
469.65
Pressure [kPa]
101.32
101.32
Molar Flow [kgmole/h]
125.00
0.00
Mass Flow [kg/h]
2801.66
0.00
Std Ideal Liq Vol Flow [m3/h]
3.19
0.00
Molar Enthalpy [kJ/kgmole]
-182697.73
-182698.59
Molar Entropy [kJ/kgmole-C]
177.36
197.11
Heat Flow [kJ/h]
-2.28E+07
0.00E+00
Reactor Top
1.00
469.65
101.32
125.00
2801.64
4.47
-182698.59
197.11
-2.28E+07
From the reaction that occurred, the equilibrium constant and the conversion of which were
estimated to be 6.81 and 80.07% respectively, the molar compositions, the mass compositions and
the volume compositions of the streams around the reactor were obtained, and they were as given in
Tables 6 – 8, respectively. From the table, it was discovered that the desired reaction has, actually,
occurred because the products of the process, which were hydrogen and carbon dioxide, are now
seen to have values on the list of the compositions.
Table 6. Molar compositions of the inputs and the outputs of the reactor
Component
Reactor In
Reactor Bottom
Reactor Top
CO
0.44
0.09
0.09
H2O
0.56
0.21
0.21
Hydrogen
0.00
0.35
0.35
CO2
0.00
0.35
0.35
Table 7. Mass compositions of the inputs and the outputs of the reactor
Component
Reactor In
Reactor Bottom
Reactor Top
CO
0.55
0.11
0.11
H2O
0.45
0.17
0.17
Hydrogen
0.00
0.03
0.03
CO2
0.00
0.69
0.69
International Journal of Engineering Research in Africa Vol. 30
149
Table 8. Volume compositions of the inputs and the outputs of the reactor
Component
Reactor In
Reactor Bottom
Reactor Top
CO
0.60
0.09
0.09
H2O
0.40
0.10
0.10
Hydrogen
0.00
0.28
0.28
CO2
0.00
0.53
0.53
Balances across Heat Exchanger 1
The vapour obtained from the reactor as the top stream was passed into a heat exchanger
(denoted as heat exchanger 1) in order to get it condensed before entering the separator. Given in
Table 9 are the conditions of the input and the output of the heat exchanger. Based on the results
obtained and shown in the table, reasonable level of condensation was achieved from the
condensation operation carried out in the heat exchanger.
Table 9. Conditions of the inputs and the outputs of heat exchanger 1
Name
Reactor Top
Separator In
Vapour
1.00
0.20
Temperature [C]
469.65
-271.08
Pressure [kPa]
101.32
101.32
Molar Flow [kgmole/h]
125.00
125.00
Mass Flow [kg/h]
2801.64
2801.64
3
Std Ideal Liq Vol Flow [m /h]
4.47
4.47
Molar Enthalpy [kJ/kgmole]
-182698.59
-335707.62
Molar Entropy [kJ/kgmole-C]
197.11
-937.63
Heat Flow [kJ/h]
-2.28E+07
-4.20E+07
Qheat 1
1.91E+07
Table 10. Molar compositions of the inputs and the outputs of heat exchanger 1
Component
Reactor Top
Separator In
CO
0.09
0.09
H2O
0.21
0.21
Hydrogen
0.35
0.35
CO2
0.35
0.35
Table 11. Mass compositions of the inputs and the outputs of heat exchanger 1
Component
Reactor Top
Separator In
CO
0.11
0.11
H2O
0.17
0.17
Hydrogen
0.03
0.03
CO2
0.69
0.69
Table 12. Volume composition of the inputs and the outputs of heat exchanger 1
Component
Reactor Top
Separator In
CO
0.09
0.09
H2O
0.10
0.10
Hydrogen
0.28
0.28
CO2
0.53
0.53
Since there is not any material transfer across the heat exchanger, the molar compositions,
the mass compositions and the volume compositions of its input and output streams were found to
be the same, and they are given in Tables 10 – 12, respectively.
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Balances across the Separator
The material leaving the reactor, which was condensed using a heat exchanger, was passed
into a separator in order to have a very high purity of the desired product (hydrogen). From the
results of the conditions given by the balances taken across the separator, shown in Table 13, it was
observed that the stream entering the separator had been divided into two (top and bottom). Even
though the temperatures and the pressures of the two output streams of the separator were the same,
their flow rates were different as a result of the separation that has taken place.
Table 13. Conditions of the inputs and the outputs of the separator
Name
Separator In
Separator Bottom
Separator Top
Vapour
0.20
0.00
1.00
Temperature [C]
-271.08
-271.08
-271.08
Pressure [kPa]
101.32
101.32
101.32
Molar Flow [kgmole/h]
125.00
80.96
44.04
Mass Flow [kg/h]
2801.64
2712.86
88.78
3
Std Ideal Liq Vol Flow [m /h]
4.47
3.20
1.27
Molar Enthalpy [kJ/kgmole]
-335707.62
-353313.67
-9019.50
Molar Entropy [kJ/kgmole-C]
-937.63
-600.52
-44.58
Heat Flow [kJ/h]
-4.20E+07
-2.86E+07
-3.97E+05
Table 14. Molar compositions of the inputs and the outputs of the separator
Component
Separator In
Separator Bottom
Separator Top
CO
0.09
0.14
0.00
H2O
0.21
0.32
0.00
Hydrogen
0.35
0.00
1.00
CO2
0.35
0.54
0.00
Table 15. Mass compositions of the inputs and the outputs of the separator
Component
Separator In
Separator Bottom
Separator Top
CO
0.11
0.11
0.00
H2O
0.17
0.17
0.00
Hydrogen
0.03
0.00
1.00
CO2
0.69
0.71
0.00
Table 16. Volume compositions of the inputs and the outputs of the separator
Component
Separator In
Separator Bottom
Separator Top
CO
0.09
0.12
0.00
H2O
0.10
0.15
0.00
Hydrogen
0.28
0.00
1.00
CO2
0.53
0.73
0.00
Looking at the results of the molar compositions, the mass compositions and those of the
volume compositions given in Tables 14 – 16 respectively, it was discovered that the separation
carried out was very efficient because hydrogen has been separated from the mixture that entered
the separator and given in pure form as the top stream of the separator. The separator employed in
this case was so efficient that there was not any hydrogen in the bottom stream as it can be noticed
from the zero values of the mole fraction, the mass fraction and the volume fraction of hydrogen
present in the bottom stream of the separator.
Balances across Heat Exchanger 2
In order to collect the desired product in liquid form, the pure hydrogen obtained as the top
stream of the separator was then passed into another heat exchanger, which was denoted as heat
exchanger 2. The results obtained from the heat exchange operation carried out were as given in
Table 17. According to the table, there was not any material loss as a result of the heat exchange
International Journal of Engineering Research in Africa Vol. 30
151
operation that occurred because the molar, the mass and the volume flowrates of both the input and
the output of the heat exchanger were found to be the same. The variations in the conditions of the
heat exchanger wat noticed to occur in the temperature, which, of course, showed an effect on the
vapour fraction of the final product.
Table 17. Conditions of the inputs and the outputs of heat exchanger 2
Name
Separator Top
Final Product
Qheat 2
Vapour
1.00
0.00
Temperature [C]
-271.08
-251.90
Pressure [kPa]
101.32
101.32
Molar Flow [kgmole/h]
44.04
44.04
Mass Flow [kg/h]
88.78
88.78
3
Std Ideal Liq Vol Flow [m /h]
1.27
1.27
Molar Enthalpy [kJ/kgmole]
-9019.50
-8716.12
Molar Entropy [kJ/kgmole-C]
-44.58
6.28
Heat Flow [kJ/h]
-3.97E+05
-3.84E+05
-1.34E+04
Table 18. Molar compositions of the inputs and the outputs of heat exchanger 2
Component
Separator Top
Final Product
CO
0.00
0.00
H2O
0.00
0.00
Hydrogen
1.00
1.00
CO2
0.00
0.00
Table 19. Mass compositions of the inputs and the outputs of heat exchanger 2
Component
Separator Top
Final Product
CO
0.00
0.00
H2O
0.00
0.00
Hydrogen
1.00
1.00
CO2
0.00
0.00
Table 20. Volume compositions of the inputs and the outputs of heat exchanger 2
Component
Separator Top
Final Product
CO
0.00
0.00
H2O
0.00
0.00
Hydrogen
1.00
1.00
CO2
0.00
0.00
The compositions of the input stream as well as that of the output stream of the heat
exchanger is given in Table 18 for the molar fractions, Table 19 for the mass fractions and Table 20
for the volume fractions of the components involved in the process. According to the results given
in the tables, it was obvious that the quality of the material leaving the heat exchanger 2 was the
same as that of the input material to the heat exchanger based on the fact that there was not any
change in the values of the compositions.
Conclusion
It has been shown in this work that it is possible to obtain pure hydrogen from a plant
comprising of a mixer, a reactor, a separator and two heat exchangers because hydrogen having a
mole fraction, a mass fraction and a volume fraction of 1 has been obtained from the simulation
carried out when carbon monoxide and steam were passed into the process plant at room
temperature (25 oC) and at boiling temperature of water (100 oC), respectively under atmospheric
pressure.
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International Journal of Engineering Research in Africa Vol. 30
Acknowledgement
Special thanks go to Aare Afe Babalola, LL.B, FFPA, FNIALS, FCIArb, LL.D, SAN, OFR,
CON – The Founder and President, and the Management of Afe Babalola University, Ado-Ekiti,
Ekiti State, Nigeria for providing a very conducive environment that enabled the accomplishment of
this research work.
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