Fuel 199 (2017) 358–371
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Full Length Article
Reaction phenomena of high-temperature water gas shift reaction in a membrane reactor Wei-Hsin Chen a,⇑, Ching-Wei Tsai a, Yu-Li Lin b, Rei-Yu Chein c, Ching-Tsung Yu d a
Department of Aeronautics and Astronautics, National Cheng Kung University, Tainan 701, Taiwan Energy and Environmental Laboratories, Industrial Technology Research Institute, Hsinchu 310, Taiwan c Department of Mechanical Engineering, National Chung Hsing University, Taichung 402, Taiwan d Chemical Analysis Division, Institute of Nuclear Energy Research, Taoyuan 325, Taiwan b
h i g h l i g h t s
g r a p h i c a l a b s t r a c t
Reaction phenomena of high-
CO conversion 100
Feed gas X CO (%)
80
Thermodynamic breakthrough
60
40
membrane
S/C=1 S/C=1.5 S/C=2 S/C=2.5 S/C=3
20
0
400
500
600
700
T (oC)
S/C ratio Temperature of feed gas
H2 recovery
80
75
catalyst
70
HR (%)
temperature WGSR in a membrane reactor are studied. WGSR is kinetically dominated at low temperatures. WGSR is thermodynamically governed at high temperatures. CO conversion can be improved up to 83% when membrane is in the reactor. CO conversion enhanced by membrane is higher than thermodynamic equilibrium up to 61%.
Sweep gas
65
60
55
50
400
500
600
700
o
T ( C)
a r t i c l e
i n f o
Article history: Received 1 October 2016 Received in revised form 1 February 2017 Accepted 2 March 2017
Keywords: Water gas shift reaction (WGSR) Palladium (Pd) membrane Thermodynamic breakthrough Fe-Cr catalyst CO2 capture H2 production
a b s t r a c t The membrane reactor is a promising device to produce pure hydrogen and enrich CO2 from syngas. To figure out the detailed reaction phenomena of high-temperature water gas shift reaction (WGSR) in a Pdbased membrane reactor, a computational fluid dynamics (CFD) model is developed to simulate the chemical reaction where the feed gas temperature and steam-to-CO molar ratio (S/C ratio) are in the ranges of 400–700 °C and 1–3, respectively. The predictions suggest that the WGSR proceeds from kinetically controlled reaction to thermodynamically governed one when the feed gas temperature increases. The CO conversion at high temperatures can be improved up to 83% when the membrane is in the reactor compared to that without the membrane. This is mainly attributed to the intensification of the membrane’s permeance with increasing temperature, even though high temperatures disadvantage CO conversion. The analysis also reveals that the breakthrough in the thermodynamic limit of CO conversion can be achieved in the membrane reactor when the feed gas temperature is higher than 500 °C. The CO conversion in the membrane reactor can be higher than the thermodynamic equilibrium up to 61%. Ó 2017 Elsevier Ltd. All rights reserved.
1. Introduction
⇑ Corresponding author. E-mail addresses:
[email protected],
[email protected] (W.-H. Chen). http://dx.doi.org/10.1016/j.fuel.2017.03.002 0016-2361/Ó 2017 Elsevier Ltd. All rights reserved.
Currently, a number of thermochemical methods such as steam reforming (SR) [1], dry reforming (DR) [2], partial oxidation (POX) [3], autothermal reforming (ATR) [4], and gasification [5] can be adopted to produce synthesis gas or syngas (i.e. H2 + CO) from
W.-H. Chen et al. / Fuel 199 (2017) 358–371
359
Nomenclature Ci Cp D Dcat Dij Ea F press HWGS K k k K eq K br K0 l Mi p Pe qc R r Re ri S T V wi Xi
molar concentration of species i (mol m3) specific heat (J mol1 K1) mass diffusivity (m2 s1) catalyst diameter (m) binary diffusion coefficient (cm2 s1) activation energy (J mol1) pressure scale-up factor (–) heat of reaction of WGSR (J mol1) permeance (mol m2 s1 Pan) thermal conductivity (W m1 K1) Boltzmann constant (=1.380648 1023 J K1) equilibrium constant (–) catalyst bed permeability (m2) permeance pre-exponential factor (mol m2 s1 Pan) membrane thickness (m) molar mass of species i (g mol1) pressure (atm) permeability pre-exponential factor (mol m1 s1 Pan) energy source due to the chemical reaction (J m3) universal gas constant (=8.314 m3 Pa K1 mol1) radial coordinate (m) Reynolds number (–) reaction rate of species i (mol m3 s1) source term in momentum equation (N m3) temperature (K) velocity (m s1) mass fraction of species i (–) conversion of species i (%)
natural gas, alcohol, coal, or biomass. The water gas shift reaction (WGSR), which has been widely applied in industry, can subsequently be used to enrich hydrogen from the syngas [6]. The WGSR is expressed as r forward
1
CO þ H2 O $ H2 þ CO2 DH0298 ¼ 41:1 kJ mol rbackward
ð1Þ
This reaction is not only capable of enhancing hydrogen production but also can be applied in pre-combustion process for CO2 capture. In the envisioned pre-combustion CO2 capture equipment, CO and steam are converted into CO2 and H2 via a WGSR. Thereafter, CO2 is removed from the system by physical or chemical absorption process followed by CO2 compression [7]. According to the reaction temperature, the WGSR can be categorized into two types. The first type is the high-temperature WGSR where the reaction temperature is between 310 °C and 500 °C and the oxides of iron and chromium are the most commonly used catalysts [8,9]. In addition to the Fe-Cr-based catalysts, Pt-based catalysts have also been prepared to trigger WGSR at high temperatures [10,11]. The second type is the low-temperature WGSR which is induced at temperature range of 200–260 °C, and low-temperature shift catalysts are normally composed of copper, zinc oxide, and alumina [8,12]. Co-Mo/Al2O3 and Ni-Ce catalysts have also been developed to elicit the low-temperature shift reaction [13,14]. In addition to catalyst preparation, a number of experiments and numerical simulations have been carried out to figure out the WGSR characteristics. Chen et al. [15] reported that the residence time of reactants in a catalyst bed was an important parameter in determining the cost and performance of WGSR, and suggested that the residence time should be at least 0.09 s for completing the reaction. Chen and Syu [16] designed and constructed a rotating packed bed to trigger low-temperature WGSR in a high gravity environment, and found that the CO conversion in the
xi z
mole fraction of species i (–) axial coordinate (m)
Greek letters bF Forchheimer coefficient (kg m4) ei Lennard-Jones parameters (J) porosity (–) Pe l viscosity (Pa s) q density (kg m3) Pe Lennard-Jones parameters (Å) 1 mass transfer parameter (–) s tortuosity (–) / binding factor (–) Xd diffusion collision integral (–) Subscripts CO carbon monoxide cat catalyst eff effective eq equilibrium H2 hydrogen i species i in inlet j species j out outlet p permeate side r retentate side sim simulation
WGSR could be increased up to 70% compared to that without rotation. Chu et al. [17] investigated H2 production from the WGSR of bio-syngas with high content CO2 in a fixed-bed reactor, and pointed out that the optimal operating conditions were located at 450 °C and H2O/CO = 3 along with space velocity lower than 2500 h1. Ding and Chan [18] conducted WGSR experiments at different operating conditions to understand the reaction, and extracted kinetic rates from the experimental results for applying in a 2-D unsteady kinetic model. Their analysis showed that the WGSR should be operated at the gas hourly space velocity (GHSV) and temperature ranges of 106.09–212.18 h1 and 873–973 K, respectively. Chein et al. [10] studied the performance of WGSR using Pt-based catalysts where the reaction temperature, time factor, and S/C ratio were in the ranges of 750–850 °C, 10–20 g-cat h/ mol-CO, and 1–5, respectively. They addressed that the CO conversion could be enhanced by about 15% by using the bimetallic Pt–Ni catalyst supported by CeO2 and Al2O3 compared with Pt/Al2O3 catalyst. The operating conditions and adopted catalysts of the aforementioned are listed in Table 1. The WGSR is a reversible chemical reaction in nature so that the equilibrium gas species depend on reaction temperature and gaseous concentrations. When the WGSR is integrated in a membrane reactor, the reaction will shift toward higher CO conversion, resulting from continuous extraction of H2 from the WGSR. Membrane selection, membrane reactor design, and operational parameters are the most important considerations in designing a WGS membrane reactor [19]. Caravella et al. [20] studied the concentration polarization phenomenon of WGSR in a Pd-based membrane reactor, and found that the velocity field between particles and membranes enhanced the mass transfer towards the membrane surface. Gosiewski et al. [21] simulated WGSR in a membrane reactor applied for coal-derived gas processing, and obtained high CO conversion in the reactor at S/C ratios of 2.0–2.5. Hwang et al. [22] prepared a plate-type catalytic membrane reactor for WGSR,
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Table 1 A summary of literature review of WGSR. Feed gas
S/C (–)
T (°C)
Catalysts
GHSV (h1)
XCO (%)
Refs.
CO CO/CH4/H2/H2S = 9.7/30/60/0.3
750 400
Pt–Ni Co–Mo–Ce3K6
2000
75 84.91
[10] [13]
CO/H2O/He = 5/25/70 CO/H2/CO2 = 30/60/10 CO/N2 = 10/90
5 Steam/feed gas = 0.5 Steam/gas 5 5 Steam/gas = 0.6 4
Ni/5K/CeO2 Fe2O3–Al2O3–NiO–3BaO Iron–chromium-based catalyst
68,000 30,000
98.2 51 90
[14] [50] [15]
CO/N2 = 32.38/67.62 CO/CH4/CO2/H2/N2 = 20/5/20/30/20 CO
2 3 3
400 400 500 H: 300–500 L: 120–400 350 450 600
Cu–Zn Fe–Cr Pt
40,000 2500 106–212
70 72 80
[16] [17] [18]
H: high temperature WGSR. L: low temperature WGSR.
Table 2 A summary of literature review of WGSR in membrane reactor. Feed gas
S/C (–)
T (°C)
Pressure (atm)
Catalyst
Membrane and permeance pre-exponential (mol m2 s1 Pa0.5)
GHSV (h1)
HR (%)
XCO (%)
Refs.
CO/CO2/H2/H2O = 9.7/4.9/48.5/36.9
–
3.95–7.90
–
68
80
[20]
1 1–4
24.67
Fe-Cr-Cu
74,000 16,000–74,000
83.6
68.3
[21]
CO/H2/Ar = 60/36/4
2
400
10.86
Cu-Zn-Al
80
[22]
–
10 5–30 0.05
–
–
90
[23]
1 1–3
375 300–400 350 350–400
5000 5000–17,000 –
81
CO/H2/CO2/H2O = 7.97/43.48/ 10.99/31.88 H2/CO/CO2 = 70/18/12
1550 4000–5500
99.5
99
[24]
CO/H2O/Ar = 3/6/91
2 1–2
900 600–900
1
Ni
Pd75Ag25 – Pd 4.64 103 (n = 0.552) Pd–Ag – Pd – Pd 2.68–3.96 104 (permeance) SrCe0.9Eu0.1O3d –
20,000–40,000
CO/CO2/H2/N2/others = 61.6/1.7/ 30.6/4.8/1.3
400 360–400 400 180–400
–
–
82.5
[25]
Fe-Cr
and found that the CO conversion was increased by 26% when compared to the system without membrane. Adrover et al. [23] analyzed the influence of operating pressure on the membrane reactor performance, and found that an increase in pressure led to a significant improvement in CO conversion and the conversion in the membrane reactor was higher than that in conventional fixed-bed reactor. Sanz et al. [24] experimentally studied the WGSR with and without membrane at different operating conditions, and found that the CO conversion and H2 recovery at GHSV = 1550 h1 could be greater than 99% and 99.5%, respectively, when using the membrane. Li et al. [25] developed a membrane reactor using a SrCe0.9Eu0.1O3d tubular membrane where 92% CO conversion and 32% H2 separation were obtained. The former corresponded to a 46% increase in CO conversion. The operating conditions and some important outcomes such as CO conversion and H2 recovery in the aforementioned studies are shown in Table 2. It can be found that the reaction temperature is from 350 °C to 600 °C and even as high as 900 °C; the S/C ratio is between 1 and 4 and the pressure ranges from 1 to 30 atm. From the literature reviewed above, it is known that triggering WGSR in a membrane reactor is able to separate H2 and result in a higher CO conversion, even overcome the thermodynamic limit [25]. Nevertheless, some detailed reaction phenomena of WGSR in membrane reactors such as the interaction between WGSR and H2 separation and thermodynamic breakthrough are not figured out in sufficient detail. Accordingly, a numerical method is
developed and the effects of operating conditions on the reaction are explored. It is believed that the obtained results are able to provide practical insights into the application of membrane reactor for H2 production and CO2 capture.
2. Mathematical formulation 2.1. Membrane reactor The present study focuses on WGSR and H2 permeation phenomena in a Pd membrane reactor. To approach the chemical reaction in the reactor and H2 permeation across a membrane, fluid dynamics as well as heat and mass transfer at the both sides of the membrane are fully simulated using the concept of conjugate hydrogen permeation [26]. The membrane reactor is a concentric tube, consisting of an internal tube and a shell. A Pd-based membrane is on the tube and catalyst pellets are packed in the shell (Fig. 1a). The feed gas is sent into the shell from the top, while the sweep gas (i.e., steam) flows into the internal tube from the bottom (Fig. 1b). Steam is used as the sweep gas because pure H2 can be obtained through condensation. To make the physical problem more tractable, the following assumptions are adopted: (1) the flow fields are laminar and axis-symmetric and the momentum and mass transfer processes are in steady state; (2) the gas mixtures are well-mixed at the inlet and abide by the ideal gas law
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Fig. 1. Schematics of (a) geometric size, (b) flow pattern, and (c) grid system of the membrane reactor.
in the reactor; (3) the permselectivity of the membrane to H2 is infinite and is zero to other gases; (4) hydrogen permeation across the membrane obeys the Sieverts’ law; (5) the catalyst pellets are spherical (diameter = Dcat) and the packed bed formed by the pellets is treated as a homogeneous porous medium with porosity h and permeability K br [27]; (6) the temperature at the interface between the catalyst and the fluid are equivalent. The governing equations in non-porous and porous zones are given in Table 3. The source term S in the porous momentum equations is expressed as [27]
S¼
l K br
!
! !
V þbF j V j V
ð2Þ
where l is the gas mixture viscosity, K br is permeability, and bF is the Forchheimer coefficient for a packed bed with spherical particles. They are written as 2
K br ¼
dcat h3 150ð1 hÞ2
1:75q ; bF ¼ pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi 150K br h3
ð3Þ
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Table 3 Governing equations and boundary conditions. Gas phase
Catalyst phase
Continuity equation !
!
r ðq V Þ ¼ 0 Momentum equations !
!
r ðhq V Þ ¼ 0
! T
!
q V r V ¼ rp þ r l r V þðr V Þ
!
23 lðr V Þ
1 h2
!
!
q V r V ¼ rp þ r
l h
!
! T
r V þðr V Þ
! 23hl ðr V Þ þ S
Energy equation !
!
qC p V rT ¼ r ðkrTÞ
qC p V rT ¼ r ðkeff rTÞ þ qc
Species equations i !
P h r qwi j Dij rxj þ ðxj wj Þ rpp DTik rTT þ qwi V ¼ 0
r qwi
P
j Dij;eff
h
i
rxj þ ðxj wj Þ rpp DTik
Boundary
Continuity and momentum equation
Energy equation
Species equation
(1)
!
V ¼ V in
T ¼ T in
wi ¼ wi;in
(2)
r V ¼ 0; p ¼ pout
!
!
(3)
!
(4)
rV ¼ 0
V ¼0
rwi ¼ 0 rwCO2 ¼ rwCO ¼ rwH2 O ¼ 0
¼ ri
0:5 Dr rC H2 ¼ Dp rC H2 ¼ K p0:5 r;H2 pp;H2
rT ¼ 0
rwi ¼ 0
kprT ¼ kr rT
rwi ¼ 0
(6)
!
rT ¼ 0
rwi ¼ 0
V ¼0
!
þ qwi V
rT ¼ 0
(5)
V ¼0
T
kprT ¼ kr rT
!
!
rT
Table 4 Constants in calculating viscosity, thermal conductivity, and specific heat of gas species [28]. Species
Constants in calculating viscosity A1 6
1.1127 10 1.7096 108 1.797 107 2.148 106
CO H2O H2 CO2
li ¼ 1þC
A1 T B1 1 þD1 T 2
1T
B1
C1
D1
0.5338 1.1146 0.685 0.46
94.7 0 -0.59 290
0 0 140 0
B2
C2
D2
0.6863 1.3973 0.7452 -0.3838
57.13 0 12 964
501.92 0 0 1.86 106
Constants in calculating thermal conductivity ki ¼ 1þC A2 4
5.1489 10 5.3345 106 2.2811 103 3.1728
CO H2O H2 CO2
A2 T B2 1 þD2 T 2
2T
Constants in calculating specific heat C pi ¼ A3 þ B3 T þ C 3 T 2 þ D3 T 2
CO H2O H2 CO2
A3
B3
C3
D3
6.60 8.22 6.62 10.34
1.20 103 0.15 103 0.81 103 2.74 103
0 1.34 106 0 0
0 0 0 1.955 105
In the porous energy equation, keff is the effective thermal conductivity of the catalyst bed and defined as
keff ¼ hk þ ð1 hÞkcat
ð4Þ
where kcat is the thermal conductivity of the catalyst particle and k is the gas mixture thermal conductivity. To account for the size, shape and nature of the pores inside the catalyst pellet, the effective binary gas diffusivity can be approximated as
Dij;eff ¼ Dij h=s
ð5Þ
where Dij and s are the binary gas diffusivity in free space and tortuosity of the catalyst pellet, respectively. Because the gas mixtures are assumed to obey the ideal gas law, the equation of state is adopted and expressed as
p ¼ qRT
N X i
1 xi M i
ð6Þ
The boundary conditions in the entire computational domain are made up of: (1) the upstream inflows; (2) the downstream outflows; (3) the membrane surface; (4) the centerline; (5) the tube walls; and (6) the shell wall. Details of the boundary conditions are sketched in Fig. 1b and shown in Table 3. 2.2. Chemical reaction The empirical reaction rate of high-temperature WGSR in a high-pressure environment, conducted by Adams and Barton [28] using an Fe-Cr oxide catalyst, is utilized and written as
rWGS ¼ r forward r rev erse ¼ qcat ð1 hÞF press 1:69 107 mol gh 0:31 0:156 0:05 xH2 ð1 bÞ exp 88;000RTJ=mol x0:9 CO xH2 O xCO2
where
ð7Þ
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b¼
xCO2 xH2 xCO xH2 O K eq
ð8Þ
In Eq. (7), qcat is the catalyst density and Keq is the equilibrium constant expressed as
K eq ¼ exp
4577:8 4:33 T
ð9Þ
Fpress is a pressure scale-up factor to account for the high pressure effect on the reaction rate, and expressed as [28]
(a)
550
rCO ¼ rH2 O ¼ r WGS ; r H2 ¼ r CO2 ¼ rWGS and qc ¼ r WGS DHWGS
ð11Þ
where DHWGS is the temperature-dependent heat of reaction of WGSR and expressed as [29],
840
1 2 3
546
ð10Þ
where Pr is the operating pressure at the retentate side of the membrane reactor. In the WGSR, the production rate of each species and energy source in the reaction are
(b)
Grid system
548
P
R 0:5250
F press ¼ Pr
830
544
820
T (K)
T (K)
542 540 538
810
800
536 790
534 532 530
780 0
0.001
0.002
0.003
0.004
0.005
0.01
r (m)
(c)
2.4
(d)
2.2 2
1.6
Velocity (m s-1)
-1
0.02
0.025
0.02
0.025
0.025
0.02
1.8
Velocity (m s )
0.015
r (m)
1.4 1.2 1 0.8 0.6
0.015
0.01
0.005
0.4 0.2 0
0
0.001
0.002
r (m)
0.003
0.004
0 0.005
0.01
0.015
r (m)
(e)
Fig. 2. Temperature distributions at the exits of (a) tube and (b) shell and velocity distributions at the exits of (c) tube and (d) shell (Rer = 80, Rrp = 100, S/C = 3, Tr,in = 500 °C, and Tp,in = 150 °C), and (e) comparison of temperature distribution along the membrane surface.
W.-H. Chen et al. / Fuel 199 (2017) 358–371
Table 5 Operating conditions and properties of membrane and catalyst. Condition or property
Retentate side
Permeate side
Temperature (°C) Pressure (atm) Composition (%) [37]
400–700 10 H2: 34.69% CO: 47.96% CO2: 17.35% 1.3 104 98.11 (at 400 °C)
150 1 Steam: 100%
Mass flow rate (kg s1) Re (–) S/C ratio (–) Permeability pre-exponential factor (mol m1 s1 Pa0.5) [35] Membrane thickness (lm) [35] Activation energy (kJ mol1) [35] Selectivity to H2 [35] Selectivity to CO, CO2 and H2O Porosity (–) [36] Catalyst density (kg m3) [28] Catalyst diameter (m) [36]
100
(a)
80
1.02 105 100
60
X CO (%)
364
40
1–4 8.45 108
S/C=1 S/C=2 S/C=3 S/C=1(Equilibrium) S/C=2 (Equilibrium) S/C=3 (Equilibrium)
20
5 12.2 Infinite 0 0.4 2476 2 103
0 350
400
450
500
550
600
650
700
750
o
T ( C) 1
(b)
DHWGS ¼ 4:912 104 þ C pCO2 þ C pH2 C pH2 O C pCO ðT 298Þ
0.8
The temperature-dependent specific heat cp of the four gas species are given in Table 4. 2.3. Properties of gas mixtures
X CO,sim / X CO,eq
ð12Þ 0.6
0.4
In the membrane reactor, four gaseous species H2, CO, CO2, and H2O at the retentate side are considered to account for hydrogen production and permeation processes. At the permeate side, a binary gas mixture of H2/steam is taken into account. The viscosity, thermal conductivity, and specific heat of a gas species i are calculated by [28].
1 þ C1 T
1
þ D1 T
2
; ki ¼
A2 T B 2 1 þ C 2 T 1 þ D2 T 2
C pi ¼ A3 þ B3 T þ C 3 T 2 þ D3 T 2
and
ð14Þ
N X xi C pi
ð15Þ
Dij ¼ 1:881 10
T 1:5
1 Mi
þ M1j
pr2ij Xd
550
600
650
700
750
Catalyst bed 0.25
0.2
400 450 500 550 600 650 700
0.15
0.1
0.05
0
0.05
o
C C C o C o C o C o C o o
0.1
0.15
0.2
Z (m)
i¼1
Meanwhile, the binary diffusion coefficients in the gas mixture are evaluated by the Chapman–Enskog equation [31,32]:
500
0.3
j¼1
3
450
T (oC)
where / is the binding factor and x is the mole fraction. The specific heat of the gas mixtures is given by
Cp ¼
400
ð13Þ
N N X X xi li xi k i and k ¼ n n X X i¼1 i¼1 xj /ij xj /ij j¼1
350
(c)
The constants in the preceding equations are tabulated in Table 4. Furthermore, the viscosity and thermal conductivity of a gas mixture are calculated by the Wilke semi-empirical correlations [30,31]:
l¼
0.2
CO mole fraction
A1 T B1
li ¼
S/C=1 S/C=2 S/C=3
Fig. 3. Distributions of (a) CO conversion, (b) XCO,sim/XCO,eq, and (c) CO mole fraction along Z axis without membrane. (S/C = 2 and r = 14.75 mm.)
0:5 ð16Þ
where T, p, rij , and Xd are the temperature, pressure, interaction value for binary mixture, and diffusion collision integral, respectively, and Xd is correlated by
A
C E G Xd ¼ B þ þ þ and 1 expðD1Þ expðF 1Þ expðH1Þ The parameters equations
!
1¼
kT
eij
ð17Þ
rij and eij are obtained by the following
W.-H. Chen et al. / Fuel 199 (2017) 358–371
rij ¼
ri þ rj 2
and
eij ¼
pffiffiffiffiffiffiffi
ei ej
ð18Þ
The values of ri, ei and constants A-H can be found elsewhere [33]. 2.4. Numerical method The commercial software COMSOL Multiphysics 4.0a was utilized to solve the physical problem in which SPOOLES solver was
365
used. The physical configuration was constructed using an orthogonal grid system to reduce the numerical truncation error from the calculation. The entire computational domain is represented by 5 length scales and their numbers of grids are denoted by (a, b, c, d, e), as shown in Fig. 1c. Three different grid systems of (a, b, c, d, e) = (14, 65, 300, 110, 30), (28, 130, 600, 220, 60) and (56, 260, 1200, 440, 120) were tested and compared with each other. The inlet Reynolds numbers and temperatures at the retentate side and the permeate side were 80 and 773 K as well as 100 and
Fig. 4. Distributions of (a) isothermal, (b) reaction rate, and (c) (1-b) contours in the catalyst bed. (S/C = 2).
366
W.-H. Chen et al. / Fuel 199 (2017) 358–371
423 K, respectively, while the S/C ratio was 3. The distributions of temperature and velocity along the r-direction at the exits of the tube and the shell shown in Fig. 2a–d indicate that the difference between the second and the third grid systems is almost imperceptible, whereas there is a small difference between the first and the
(a)
100 CO conversion (w membrane) CO conversion (w/o membrane) CO conversion (Equilibrium)
80
X CO (%)
60
second grid systems. In view of satisfying the requirement of grid independence of the second grid system, it was adopted for numerical simulations. Meanwhile, the predicted temperature distributions along the membrane surfaces in this study were compared with the those of Chein et al. [27] to validate the numerical method, using the same operating conditions, and the temperature distributions along the membrane surface at the retentate and permeate sides are shown in Fig. 2e. It depicts that the predicted distributions are in good agreement with those of Chein et al. [27], achieving the numerical validation. The concentrations of all chemical species in thermodynamic equilibrium were calculated using the commercial software HSC Chemistry 7.0. The validation of the calculations had been carried out previously for the WGSR [4] and ethanol autothermal reforming [34]. 2.5. Membrane properties and operating conditions
40
A Pd-based membrane was used to separate H2 from gas mixtures, and its permeance is a function of permeability, thickness, and temperature. An Arrhenius-type equation correlating these parameters is used and expressed as [35].
20
0 350
400
450
500
550
600
650
700
750
K¼
Pe Ea Ea exp ¼ K 0 exp l RT RT
ð19Þ
o
T ( C)
(b)
where Pe, l, Ea, R, T, and K0 are given in nomenclature. The operating conditions and properties of the membrane [35] and catalyst [28,36] are tabulated in Table 5. The feed gas at the retentate side included a dry stream and a wet stream. The dry stream consisted of H2 (34.69 vol%), CO (47.96 vol%), and CO2 (17.35 vol%), which was obtained from the Wabash River coal gasification [37]. The wet stream was steam and controlled according to the steam-tocarbon monoxide molar ratio (S/C ratio). The mass flow rate of the feed gas was fixed at 1.3 104 kg s1 where its Reynolds number is 200 at 25 °C and S/C = 1. Steam at the permeate side was adopted as the sweep gas with Rep = 100. The definitions of Reynolds numbers at the retentate side (Rer) and the permeate side (Rep) can be found in a previous study [33]. The flows at the permeate side and the retentate side were in a counter-current pattern, resulting from yielding a higher H2 separation efficiency when compared to that in a co-current pattern [38]. The pressures at the permeate side was 1 atm, while it was 10 atm at retentate side.
100
80
X CO (%)
60
40
CO conversion (w membrane) CO conversion (w/o membrane) CO conversion (Equilibrium)
20
0 350
400
450
500
550
600
650
700
750
T (oC)
3. Results and discussion 3.1. WGSR without membrane
(c)
100
80
X CO (%)
60
40
20
0 350
CO conversion (w membrane) CO conversion (w/o membrane) CO conversion (Equilibrium) 400
450
500
550
600
650
700
o
T ( C) Fig. 5. Profiles of CO conversion at S/C = (a) 1, (b) 2, and (c) 3.
750
The profiles of CO conversion in the reactor without membrane at the feed gas temperatures of 400–700 °C and S/C ratios of 1–3 are shown in Fig. 3a. The profiles increase with increasing temperature, followed by decrease after the temperature is higher than 500 °C, regardless of the S/C ratio. The distributions of equilibrium CO conversion from thermodynamic analysis are also shown in Fig. 3a. Thermodynamically, a higher CO conversion is favored at lower temperatures due to the exothermicity of the WGSR. The predicted CO conversion in Fig. 3a decreases when the temperature increases from 400 °C to 500 °C, revealing that the chemical reaction is dominated kinetically in this temperature range. An inverse trend in CO conversion is observed at temperatures of 500–700 °C, and this trend is consistent with the thermodynamic analysis [11,13]. It follows that thermodynamics reigns over the WGSR at temperatures higher than 500 °C. Based on the results in Fig. 3a, the profiles of the ratio of predicted CO conversion to equilibrium one (i.e., Xco,sim/Xco,eq) is shown in Fig. 3b. The ratio is around 0.22 at the feed gas temperature of 400 °C, suggesting that the temperature is too low to
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W.-H. Chen et al. / Fuel 199 (2017) 358–371 Table 6 A summary of literature review of CO conversion increment in membrane reactor. Feed gas
S/C
Temperature (°C)
Pressure (atm)
Catalyst
Membrane and permeance pre-exponential (mol m2 s1 Pa0.5)
GHSV (h1)
HR (%)
XCO (%)
Increment (%)
Refs.
CO/H2O/Ar = 3/6/91
2
900
1
Ni
–
–
92.5
29.5
[25]
CO
3
440
20
Fe–Cr–Cu
48,000
90
98.5
8.9
[41]
CO/H2/CO2 = 64.5/33/2.5
3
375
14.80
–
–
–
98
4.5
[42]
Ar/CO/H2O = 11/1/0.75
–
360
1
CuO/CeO2
–
–
68
14
[43]
CO
1
450
15
Fe–Cr–Al2O3
–
100
98
25
[27]
H2/CO2/CO = 70/12/18
1
400
4
Fe–Cr oxide
50,000
15
59
27
[44]
CO
2.6
450
14
Fe-Cr
2100
81.2
98.2
5.2
[45]
H2/ CO/CO2 = 34.69/47.96/17.35
1
700
10
Fe–Cr–Al2O3
SrCe0.9Eu0.1O3d – – – Pd/Ag23 7.73 107 (Permeability) Pd/Ag – Pd 1.43 103 Pd – Pd/Ag – Pd 1.69 102
1596
75
52
23.57
This study
sufficiently drive the WGSR. When the temperature is raised to 500 °C, the ratio rises to approximately 0.87–0.89, reflecting that the kinetically governed WGSR is fairly sensitive to the temperature. Once the temperature is higher than 500 °C, the ratio merely varies a bit. This verifies the thermodynamically dominating mechanism of the WGSR at temperatures higher than 500 °C. It is noted that the ratio Xco,sim/Xco,eq is slightly affected by the S/C ratio. To proceed farther into an analysis of the reaction behavior, the distribution of CO concentration (mole fraction) in the catalyst bed along z-axis at r = 14.75 mm (the center of the shell) is plotted in Fig. 3c. For the feed gas at 400 °C or 450 °C, the CO concentration linearly decreases downstream, elucidating the insufficient time for the reactants reaching the equilibrium state. Similar phenomenon has also been observed in the study of Basile et al. [39]. On the other hand, the CO concentration in the catalyst bed reaches the steady state downstream for the feed gas temperature between 500 °C and 700 °C. The higher the temperature, the more rapid drop the CO concentration. From the reactor design point of view, an increase in temperature can significantly reduce the catalyst bed size, as a consequence of intensifying chemical reaction rate or shortening the residence time. Nevertheless, the optimal temperature develops at 500 °C inasmuch as this temperature yields the highest CO conversion and thereby H2 production, as observed in Fig. 3a.
3.2. Reaction phenomena in membrane reactor The isothermal, reaction rate, and (1-b) contours in the catalyst bed of the reactor with membrane are demonstrated in Fig. 4 where three feed gas temperatures of 400 °C, 500 °C, and 600 °C along with S/C = 2 are taken into account and where b has been defined in Eq. (8). On account of the exothermic WGSR involved, corresponding to the three feed gas temperatures, the maximum temperatures in the catalyst bed are 425 °C, 568 °C, and 651 °C (Fig. 4a), respectively. The sweep gas (steam) temperature is 150 °C which has a cooling effect on the membrane. Consequently, the temperature in the catalyst bed adjacent to the membrane is lower. The reaction rate contours at 400 °C are apparently different from those at 500 °C and 600 °C (Fig. 4b). For the feed gas temperature at 400 °C, the intensity of the chemical reaction grows downstream, due to heat release from the WGSR. As for the cases of 500 °C and 600 °C, the WGSR is rapidly triggered at the entrance of the catalyst bed because of higher temperatures. Once the downstream reaction rate is down to a certain level, stemming
from more CO consumption, the isothermal contours become uniform, especially at 600 °C. The parameter (1-b) is greater than the zero when the forward reaction rate is higher than the backward one. Once the two reaction rates are equivalent, the WGSR reaches the chemical equilibrium and (1-b) is zero. Fig. 4c depicts that the contours at 400 °C are close to unity, suggesting that the forward reaction rate is by far larger than the backward one. This also implies that the WGSR is controlled by the kinetics process. When the feed gas temperature is 500 °C, the value of (1-b) progressively decreases from around 0.98 at the upper left corner to 0.16 at the bottom right corner. It follows that the chemical reaction is gradually thermodynamically dominated downstream. As for the temperature of 600 °C, the contours show that (1-b) approaches zero in the most of the zone, except for the area in the vicinity of the entrance. This reveals that 600 °C is sufficiently high to rapidly trigger the WGSR so that the chemical reaction is mainly ruled by thermodynamic equilibrium. In view of some H2 permeating through the membrane in the region adjacent to the membrane surface, the H2 concentration goes down nearby and the value of (1-b) is higher. 3.3. Influence of membrane The profiles of CO conversion at S/C = 1, 2, and 3 in the presence and absence of the membrane are shown in Fig. 5. It is known that the membrane at the retentate side of the reactor acts as the H2 sink [26]. While the WGSR proceeds, the produced H2 is separated by the membrane. The reduced H2 concentration in the catalyst bed facilitates the forward reaction of the WGSR, according to the Le Chatelier’s principle [40]. This is the reason why the CO conversion in the presence of the membrane is higher than that in the absence of the membrane. In particular, when the temperature is greater than 500 °C, the CO conversion in the presence of the membrane is higher than the thermodynamic CO conversion, rendering the breakthrough of the thermodynamic limit. For example, Fig. 5a reveals that the CO conversion at S/C = 1 and 700 °C (=52.07%) is higher than the equilibrium CO conversion (=32.44%) by 19.63%. Past studies [25,27,41–45] regarding the thermodynamic breakthrough are given in Table 6. It can be seen that the increment in CO conversion due to membranes are between 5.2% and 29.5%. Accordingly, the predicted results in this study are in a reasonable range. As far as the CO conversion is concerned, there are two major forces at play when the temperature increases. The first factor in play is the retardation of forward WGSR which abates the CO
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W.-H. Chen et al. / Fuel 199 (2017) 358–371
(a)
(b)
Tfeed gas
400
500
600
xCO
Tfeed gas
400
500
600
xCO
400
400
500
500
600
600
xH 2 ,p
xH 2 ,r
xH 2 ,p
xH 2 ,r
Fig. 6. Distributions of CO and H2 concentration contours in the reactor (a) without and (b) with membrane. (S/C = 1.)
conversion, stemming from thermodynamic disfavor. The other factor is the intensification of membrane permeance, as expressed in Eq. (19), which is conducive to the CO conversion. Obviously, the second factor is insufficient to counteract the first one at temperatures higher than 500 °C. As a result, the CO conversion declines with increasing temperature. Nevertheless, the decreasing tendency of the CO conversion in the presence of the membrane is not as pronounced as that in the absence of the membrane so that the difference of CO conversion in the presence and absence of the membrane tends to increase with increasing temperature, as shown in Fig. 5. The CO and H2 concentration contours in the reactor at S/C = 1 and 400 °C, 500 °C, and 600 °C with and without membrane are
shown in Fig. 6. For the reactor without the membrane (Fig. 6a), the CO and H2 concentration contours are characterized by onedimensional behavior. That is, the concentrations only alter along the z-direction, and are almost independent of the r-direction. For the feed gas temperature of 400 °C, the CO concentration declines downward whereas the H2 concentration is lifted gradually, owing to the kinetically dominated mechanism. Alternatively, the chemical reaction at 600 °C is violent and behaves as a scavenging front to consume CO. This leads to a rapid drop in CO concentration and a dramatic growth in H2 concentration. Unlike the behavior without membrane, the CO and H2 concentration contours in the reactor with membrane become two-dimensional (Fig. 6b). At a given z position, the CO concentration at the membrane surface
W.-H. Chen et al. / Fuel 199 (2017) 358–371
(a)
369
100
80
X CO (%)
60
40 S/C=1 S/C=1.5 S/C=2 S/C=2.5 S/C=3
20
0
400
500
600
700
600
700
o
T ( C)
(b)
80
75
HR (%)
70
65
60
55
50
400
500
T (oC)
(c)
Fig. 8. Three-dimensional profile of (a) XCO,w./XCO,w.o. and (b) XCO,sim/XCO,eq.
70
between H2 production from WGSR and H2 permeation across the membrane, the highest H2 concentrations at 400 °C and 600 °C are located near the upper right corner of the catalyst bed. By virtue of disadvantaging the forward reaction of the WGSR at higher temperatures, the high H2 concentration zone in the catalyst bed at 500 °C is larger than that at 600 °C. It is noteworthy that a layer of attenuated H2 concentration along the membrane surface at the retentate side clearly develops. This layer will cause the concentration polarization, thereby reducing the H2 permeation rate [46].
60
Pure H 2 index (%)
50 40 30 20
3.4. Influence of S/C ratio
10 0
400
500
600
700
T (oC) Fig. 7. Profiles of (a) CO conversion, (b) H2 recovery, and (c) pure H2 index.
is higher whereas the H2 concentration is lower, as a result of H2 permeating through the membrane. For this reason, the H2 concentration at the permeate side is raised markedly upward. This is especially obvious at 500 °C and 600 °C. Under the interaction
The profiles of CO conversion and H2 recovery (HR) in the membrane reactor at various S/C ratios are displayed in Fig. 7a and b, respectively. It is not surprising that an increase in S/C ratio enlarges CO conversion. This can be explained by pushing the equilibrium state toward the right side of the WGSR when more steam is added into the reaction system [47]. The maximum values of CO conversion at S/C = 2–3 are located at 550 °C which are somewhat higher than that (i.e., 500 °C) in the reactor without membrane (Fig. 3a). The profile is promoted markedly when the S/C ratio increases from 1 to 1.5, whereas the increment in the conversion for the ratio increasing from 2.5 to 3 is not pronounced. More
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W.-H. Chen et al. / Fuel 199 (2017) 358–371
steam addition in the WGSR will cause higher energy consumption [48]. Accordingly, from the compromise viewpoint among CO conversion, CO2 capture and energy consumption, Fig. 7a suggests that the WGSR in the membrane reactor operated at 550 °C along with S/C = 2 may be appropriate. HR is a crucial indictor to identify the performance of a membrane reactor for H2 separation. Fig. 7b indicates that the HR descends with increasing S/C ratio. This is attributed to the reduced H2 partial pressure at the retentate side of the reactor at a higher S/C ratio, resulting in a lower H2 permeation rate [49]. HR in the reactor is influenced by both the membrane permeance and CO conversion. An increase in temperature intensifies the former but lowers the latter. Despite the decrease in CO conversion with increasing temperature, the increase in HR reveals that the influence of increasing temperature on the permeance is over on the CO conversion. However, it is noted that the HR decreases when the temperature increases from 400 °C to 450 °C. Fig. 5 has suggested that the CO conversion increases sharply in this temperature range. This corresponds to a substantial increase in H2 production so that relatively less H2 is separated. The multiplication of CO conversion and HR can be considered as an index to account for relative amount of pure H2 obtained in the membrane reactor. The index is termed pure H2 index (PHI) and given by
PHIð%Þ ¼
CO conversion HR 100
ð20Þ
The profiles of PHI shown in Fig. 7c depict that the higher the S/ C ratio, the higher the PHI value, implying that more amount of pure H2 based on the feed gas is obtained in the reactor. 3.5. Enhancement of CO conversion by membrane The three-dimensional profile of CO conversion ratio in the reactor with membrane to that without membrane, namely, Xco,w./Xco,w.o. versus S/C ratio and feed gas temperature is shown Fig. 8a. The entire trend suggests that the WGSR associated with membrane operated at high temperatures and lower S/C ratios renders better performance. Within the investigated ranges of S/C ratio (i.e., 1–3) and feed gas temperature (i.e., 400–700 °C), the ratio is between 1.09 and 1.83, implying that the CO conversion can be improved at least 9% and up to 83% by the membrane. The profile of the ratio of the simulated CO conversion in the membrane reactor to the equilibrium CO conversion without membrane (i.e., Xco,sim/Xco,eq) shown in Fig. 8b depicts that the breakthrough of thermodynamic limit can be accomplished when the feed gas temperature is higher than 500 °C, irrespective of the S/C ratio. The maximum ratio is 1.61 which occurs at 700 °C and S/C = 1. This implies, in turn, that the CO conversion in the membrane reactor can be higher than the thermodynamic equilibrium up to 61%. These results clearly reveal that the membrane reactor in association with appropriate operation is able to intensify CO conversion, enrich CO2, and obtain more pure H2, which is conducive to H2 production and CO2 capture. 4. Conclusions A computational fluid dynamics (CFD) model accounting for the WGSR in a Pd-based membrane reactor has been successfully developed, and detailed reaction mechanism under the interaction of H2 production and separation has been investigated. The profiles of CO conversion, featured by increase followed by decrease with temperature, reveal that the WGSR is kinetically dominated at low temperatures and thermodynamically governed at high temperatures. From CO conversion point of view, the optimal
temperature is between 500 °C and 550 °C, whether the membrane is installed or not. Despite lower CO conversion at high temperatures, the effect of increasing temperature on the membrane’s permeance obviously prevails over the retardation of CO conversion. As a result, the higher the feed gas temperature, the better the improvement of CO conversion by the membrane reactor compared to that without membrane. In particular, the CO conversion in the membrane reactor is higher than that at thermodynamic equilibrium when the feed gas temperature is greater than 500 °C, achieving the breakthrough of thermodynamic limit. In the membrane reactor, an increase in S/C ratio facilitates pure H2 production from the feed gas, and the optimal temperature is between 600 °C and 650 °C. However, more energy is consumed when the S/C ratio increases. Within the investigated ranges of feed gas temperature (400–700 °C) and S/C ratio (1–3), the CO conversion in the membrane reactor can be improved up to 83% when compared to the reactor without membrane. Meanwhile, the CO conversion in the membrane reactor can be higher than the thermodynamic equilibrium up to 61%. Acknowledgments The authors acknowledge financial support from the Ministry of Science and Technology (MOST 105-3113-E-042A-001) and the Bureau of Energy, Ministry of Economic Affairs, Taiwan, R.O.C., for this research. References [1] Farshchi Tabrizi F, Mousavi SAHS, Atashi H. Thermodynamic analysis of steam reforming of methane with statistical approaches. Energy Convers Manage 2015;103:1065–77. [2] Chen WH, Lin SC. Reaction phenomena of catalytic partial oxidation of methane under the impact of carbon dioxide addition and heat recirculation. Energy 2015;82:206–17. [3] Al-Hamamre Z, Voß S, Trimis D. Hydrogen production by thermal partial oxidation of hydrocarbon fuels in porous media based reformer. Int J Hydrogen Energy 2009;34:827–32. [4] Chen WH, Lin MR, Lu JJ, Chao Y, Leu TS. Thermodynamic analysis of hydrogen production from methane via autothermal reforming and partial oxidation followed by water gas shift reaction. Int J Hydrogen Energy 2010;35:11787–97. [5] Chen WH, Hsu PC, Lin BJ. Hydrogen permeation dynamics across a palladium membrane in a varying pressure environment. Int J Hydrogen Energy 2010;35:5410–8. [6] Chen W-H, Jheng J-G. Characterization of water gas shift reaction in association with carbon dioxide sequestration. J Power Sources 2007;172:368–75. [7] van Dijk HAJ, Cohen D, Hakeem AA, Makkee M, Damen K. Validation of a water–gas shift reactor model based on a commercial FeCr catalyst for precombustion CO2 capture in an IGCC power plant. Int J Greenhouse Gas Control 2014;29:82–91. [8] Luengnaruemitchai A, Osuwan S, Gulari E. Comparative studies of lowtemperature water–gas shift reaction over Pt/CeO2, Au/CeO2, and Au/Fe2O3 catalysts. Catal Commun 2003;4:215–21. [9] LeValley TL, Richard AR, Fan M. The progress in water gas shift and steam reforming hydrogen production technologies – a review. Int J Hydrogen Energy 2014;39:16983–7000. [10] Chein RY, Lin YH, Chen YC, Chyou YP, Chung JN. Study on water–gas shift reaction performance using Pt-based catalysts at high temperatures. Int J Hydrogen Energy 2014;39:18854–62. [11] Mendes D, Mendes A, Madeira LM, Iulianelli A, Sousa JM, Basile A. The watergas shift reaction: from conventional catalytic systems to Pd-based membrane reactors—a review. Asia-Pac J Chem Eng 2010;5:111–37. [12] Gradisher L, Dutcher B, Fan M. Catalytic hydrogen production from fossil fuels via the water gas shift reaction. Appl Energy 2015;139:335–49. [13] Zhang Y, Zhang G, Zhao Y, Li X, Sun Y, Xu Y. Ce–K-promoted Co–Mo/Al2O3 catalysts for the water gas shift reaction. Int J Hydrogen Energy 2012;37:6363–71. [14] Ang ML, Oemar U, Kathiraser Y, Saw ET, Lew CHK, Du Y, et al. Hightemperature water–gas shift reaction over Ni/xK/CeO2 catalysts: suppression of methanation via formation of bridging carbonyls. J Catal 2015;329:130–43. [15] Chen WH, Hsieh TC, Jiang TL. An experimental study on carbon monoxide conversion and hydrogen generation from water gas shift reaction. Energy Convers Manage 2008;49:2801–8. [16] Chen W-H, Syu Y-J. Hydrogen production from water gas shift reaction in a high gravity (Higee) environment using a rotating packed bed. Int J Hydrogen Energy 2010;35:10179–89.
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