COMPARATATIVE ANALYSIS ON VARIOUS REFORMERS SUPPLIED
WITH DIFFERENT FUELS AND INTEGRATED WITH HIGH
TEMPERATURE PEM FUEL CELLS
Harikishan R. Ellamla*, Piotr Bujlo, Cordellia Sita, Sivakumar Pasupathi
Hydrogen South Africa (HySA): Systems and Validation Centre, University of the Western
Cape, Robert Sobukwe Road, Bellville 7535, Cape Town, South Africa
* Corresponding author: Tel.: +27 21 959 9310; fax: +27 21 959 1353;E-mail address:
hariehkr@gmail.com (Harikishan R. Ellamla)
DOI: http://dx.doi.org/10.1016/j.ces.2016.06.065
Chemical Engineering Science
Received date: 19 February 2016
Revised date: 22 June 2016
Accepted date: 27 June 2016
Abstract
High-temperature proton exchange membrane fuel cells (HT-PEMFCs) have received
substantial attention in stationary sector applications, due to their high carbon monoxide (CO)
tolerance, high-quality waste heat and simplified water management system. Hydrogen rich
gas produced in a fuel reforming process can be used and can be directly supplied to the HTPEMFC stack anode omitting complex hydrogen purification process. It allows a wide range
of fuel flexibility for the reforming process. The present study is an analysis of HT-PEMFC
stack performance operating with an integrated steam reformer, operated with various fuels
like ethanol, glycerol, methanol, methane and other fuels. The HT-PEMFC stack is modelled
with a concept of varying local current density in the cathode catalyst layer.
Key words: fuel cell; steam reformer; HT-PEMFC; current density; fuel flexibility; ethanol
1. Introduction
A fuel cell is an electrochemical device which converts chemical energy of a fuel into electrical
energy with heat and water as by-products. The primary fuel is typically an alcohol or a
hydrocarbon or a substance derived from it, e.g., hydrogen, which can be supplied continuously
to the fuel cell. The conventional low temperature polymer electrolyte membrane fuel cells
(LT-PEMFCs) operate at a working temperature below 353 K and HT-PEMFCs are operated
at temperature up to 473 K (Zhang et al., 2006). The main advantages of moving to higher
operating temperature include: simplified and cheaper fuel processor construction as the HTPEMFC stack can tolerate up to 5 volume % of carbon monoxide (CO) concentration in the
fuel feed (Korsgaard et al., 2006; Zhang et al., 2007). The adsorption of CO on the Pt surface
reduces as the fuel cell operating temperature increases (Bellows et al., 1996). HT-PEMFC
allows more CO concentration in the anode fuel as comparative to the LT-PEMFC. More CO
concentration in the reformed fuel allows simplified CO purification process; the operation of
the HT-PEMFC system is simpler than that of the conventional LT-PEMFC. HT-PEMFC also
allows direct fuel supply from reformer into the HT-PEMFC. Due to high operating
temperature of the HT-PEMFC, high-quality heat is produced which can be used in thermal
integration of fuel reforming with the HT-PEMFC stack. Design of the HT-PEMFC system
depends on the type of fuel used in a fuel processor and it directly influences the overall system
efficiency and operating costs of the system. In the case of stationary fuel cell application, the
cost of the fuel processor is estimated about 80% of the total balance of plant cost (Staffel and
Green, 2013). Today worldwide, nearly 48% of hydrogen is produced from steam reforming
of methane, other reforming process (like naphtha/oil) contributes to 30%, and coal gasification
is contributing about 18% and electrolysis process contributing around 4% (Silva et al., 2015).
Steam reforming (SR) process and auto-thermal reforming (ATR) process are attractive
hydrogen production methods. The SR process gives a high yield of the hydrogen and autothermal process has a fast response to load or parameter change as well as it requires low
external energy compare to other reforming processes (Authayanun et al., 2015).
The
schematic diagram of fuel processing for an integrated fuel cell system is shown in Figure 1.
SR process is used to produce hydrogen from hydrocarbon fuels, such as ethanol, CH4,
glycerol, methanol, ethane, butane, propane or even gasoline. SR is a highly endothermic
reaction and external heat must be supplied to run the process. In general, external heat for the
endothermic reaction is supplied by catalytic burning of methane in a combustor.
Figure 1: An overview of fuel processing for fuel cell systems (Ellamla et al., 2015).
SR: CH4 + H2O → CO + 3H2 ΔH = +206 kJ mol−1
PO: CH4 + 0.5O2 → CO + 2H2 ΔH = −36 kJ mol−1
ATR: Combination of PO and SR reactions
In the SR process, sulphur content is removed first followed by a water gas shift reaction and
CO absorption step. In the water gas shift reactor (WGSR), CO and steam react together
producing carbon dioxide (CO2) and hydrogen. The WGSR operates in between 473 K and 755
K with a variety of catalysts. The WGSR rich hydrogen production proceeds in the direction
of lower reaction temperature. Low temperature WGSR operates between 473 K and 513 K in
the presence of copper-zinc-aluminium catalyst (it is extremely sensitive to sulphur and
chloride poisoning) and it gives CO content lower than 1%. High temperature WGSR operates
between 560 K and 755 K in the presence of chromium or copper promoted iron-based
catalysts. Further, CO is removed from adsorption process, CO and other impurities are
removed from the outlet gas stream by leaving essentially pure hydrogen. In PO process, the
hydrocarbons react with a limited amount of oxygen (typically from the air and it is not enough
to completely oxidize the hydrocarbons to carbon dioxide and water), and the products contain
hydrogen, CO, CO2 and other compounds. The advantage of PO is an exothermic character of
the process and the fact that it is much faster than SR and finally it requires a smaller reactor.
PO process produces less hydrogen per unit of the input fuel than it is obtained by SR with the
same amount of fuel. During hydrogen production process, preferential oxidation (PO),
pressure swing adsorption or membrane separation is not necessary for the HT-PEMFC system
and it needed only a water-gas shift reactor (WGSR). In the case of an LT-PEMFC system, the
fuel purification process adds additional cost compared to the HT-PEMFC.
Jaggi and Jayanti (2013) developed a simulation model for a stand-alone power unit, which is
thermally integrated with LT- and HT-PEMFC and with ATR fuel processor. The overall
system efficiency is claimed to be about 41% and thermal efficiency is about 80%. Gardemann
et al. (2014) developed a compact ethanol ATR processor integrated with an HT-PEMFC stack
and demonstrated advantages in terms of compactness, reliability and fast start-up of the
system. The most time-consuming step is the heating of the shift reactor. George and Suresh
(2015) developed a one dimensional model of an integrated steam reformer system HTPEMFC model based on thermodynamics, transport, and kinetic equations. They concluded, in
the presence of WGSR in the fuel processor, the S/C ratio and reformer temperature are not
affecting the output cell voltage. The composition of hydrogen, CO and reformer energy
utilisation is affecting the overall system efficiency. Authayanun et al., (2013) investigated a
glycerol SR and they revealed that the water produced form electrochemical reaction in the
HT-PEMFC can be able to supply sufficient water to the SR. Park and Min (2014) developed
a quasi-three-dimensional dynamic model of a HT-PEMFC integrated with fuel processor. The
effects of the reformer operating conditions on the methane ATR process and efficiency of the
system were investigated. The process efficiency increases as increase in the ATR inlet
temperature. All these models are do not consider the effect of local current density and
temperature in the catalyst layer. In this paper we modelled an integrated HT-PEMFC system
using a combination of Aspen Plus® and ANSYS-Fluent® software. HT-PEMFC stack was
modelled in detail with consideration of local current density variation in the catalyst layer.
2. Modelling methodology of the integrated system
The integrated system is modelled using a combination of Aspen Plus® and ANSYS-Fluent®
software. The reformer is modelled as equilibrium reactor assuming it as Gibbs reactor,
combustor is modelled as stoichiometric reactor, and WGS is modelled as equilibrium reactor
and the fuel cell is modelled as stoichiometric reactor. Further HT-PEMFC stack is modelled
in-detail in ANSYS Fluent® by taking into account of CO concentration effects on the
performance and efficiency of the stack.
2.1 Model description of the SR
The schematic diagram of the overall integrated SR with HT-PEMFC stack is shown in Figure
2. In the SR, the Gibbs reactor model calculates the equilibrium compositions of the reformate
gases (H2, CO2, CO, H2O, CH4) from the SR fuel by direct minimizing of Gibb’s free energy
at the specified reactor condition like temperature and pressure. The WGS, equilibrium reactor
model requires the specification of the reaction stoichiometry, temperature, pressure and phase
of the reactor. It calculates the equilibrium constant for a given reaction first and then estimates
the equilibrium composition of the leaving stream from the reactor. In the combustion reactor,
the stoichiometric reactor model is used with unknown reaction kinetics and the purpose of this
reactor is to produce more exothermic heat, which is later used to preheat the reformer fuel.
Figure 2: Schematic diagram of material and thermal integration of HT-PEMFC stack with
fuel reformer.
2.2 Model description of the HT-PEMFC stack
The HT-PEMFC stack consists of a number of individual cells stacked together in series. Every
fuel cell has individual layers varying in thickness from micro meters to milli meters. The HT-
PEMFC components with each layer dimension are shown in Figure 3. The fuel cell consist of
two catalyst layers, two gas diffusion layer (GDL), one membrane, one flow field for oxygen
supply and another flow filed for hydrogen supply. The thickness of the catalyst layer is about
13 µm to 30 µm; anode catalyst layer in which splitting of hydrogen takes place and in the
cathode catalyst layer oxygen reduction reaction takes place. Catalyst layers are loaded with Pt
on top of carbon supports and thermal conductivity of the catalyst layer is even lower and in
the range of 0.3 Wm-1 K-1. The membrane is located in between of two catalyst layers and the
Figure 3: The HT-PEMFC cell components with dimensions.
electrolyte thickness is about 30 µm to 40 µm. Polybenzimidazole (PBI) membrane doped with
phosphoric acid is used for HT-PEMFC operation at an operating temperature of 393 K to 473
K. The main stack construction components are bipolar and monopolar plates, in which the
flow of reactants and flow of electron takes place. Bipolar plates are made-up of high thermal
conductive graphite material and plates have a thickness around 2 mm to 4 mm with flow field
configurations on both sides. Monopolar plates have the flow field configuration only on one
side. The other cell component is the gas diffusion layer (GDL), in which flow of reactants
takes place through the porous medium which is about 20 to 40 µm thick and the thermal
conductivity of the GDL is very low, about 1.7 Wm-1 K-1 as compared to the thermal
conductivity of the bipolar plate which is about 20 W-1m K-1.
2.3 Geometry description of HT-PEMFC
By taking an advantage of symmetry boundary condition, a small portion of the computational
geometry is taken and it is assumed to be repeated units consisting the whole fuel cell stack.
Symmetry boundary conditions are applied on the four sides of the computational domain,
natural convection boundary condition is applied on the free faces of the inlet and outlet side
of the stack. Schematic diagram of the computational domain for HT-PEMFC stack is shown
in Figure 4. Flow domain (Figure 4) of a length of 125 mm (in the z-direction), width of 3.466
mm (in the y-direction) and thickness of 2 mm (in the x-direction) is considered. Each bipolar
plate having 62 parallel channels (in the x-direction)of a length of 125 mm and a cross-section
of 1 mm width x 1 mm depth. The fine details of the flow distribution through the parallel
channels in the flow field are not considered and the flow rate is assumed to be uniformly
distributed through all the channels on each bipolar plate. Species transportation through the
GDL and catalyst layer is assumed as zero. Computational fluid dynamic (CFD) simulations
of flow and heat transfer have been studied along with coupled electrochemistry modelling
equations (these are introduced through the user defined function in the ANSYS Fluent®). The
details of electro chemical equations used in this modelling study are discussed in the below
section. Grid independence tested for the calculations to ensure the accuracy of the results.
Figure 4: Schematic diagram of the computational domain for HT-PEMFC stack.
2.4 Estimation of the HT-PEMFC performance
The cell voltage, Ecell, of the fuel can be calculated by
𝑟𝑒𝑣
𝐸𝐶𝑒𝑙𝑙 = 𝐸𝑇,𝑃
− |𝜂𝑎𝑐𝑡 | − |𝜂𝑜ℎ𝑚𝑖𝑐 | − |𝜂𝑐𝑜𝑛 |
(1)
rev
where ηcon is the mass transfer losses, the reversible cell voltage, 𝐸𝑇,𝑃
, can be calculated from
the equation below
0.5
𝛥𝐻𝑇 𝑇𝑐𝑒𝑙𝑙 𝛥𝑆𝑇
𝑅𝑇𝑐𝑒𝑙𝑙
𝐶𝐻2 𝐶𝑂2
(𝑅𝑇𝑐𝑒𝑙𝑙 )1.5
𝑟𝑒𝑣
𝐸𝑇,𝑃
= −(
−
)+
𝑙𝑛 [
]
𝑛𝐹
𝑛𝐹
𝑛𝐹
𝛼𝐻2𝑂
(2)
where R is the gas constant, Tcell is a cell temperature, CH2 is concentration of hydrogen, CO2 is
concentration of oxygen, αH2O is water activity, F is Faraday’s constant and n is the number of
electrons transferred. The entropy change for the reaction can be written in terms of the cell
operating temperature as (Scott and Mamlouk, 2009)
ST  9967.35 ln Tcell   12414.83
(3)
The enthalpy of water formation in gaseous phase, ΔHg,T , can be evaluated as
H g ,T  H l f  H Tvap
(4)
where the enthalpy of water formation in liquid phase, H l f , is 285830 J and the heat of
vaporization of water, HTvap , is evaluated from (Scott and Mamlouk, 2009) the following
equation
3
2
HTvap  3.6985x104 Tcell
 0.4834Tcell
 152.4258Tcell  68260.5789
(5)
where HTvap is expressed in Joule and Tcell in K.
Activation over potential, ηact, is a sum of anode (𝜂𝑎 ) and cathode over potential (𝜂𝑐 )
𝜂𝑎𝑐𝑡 = 𝜂𝑎 + 𝜂𝑐
𝜂𝑎 =
−𝑅-𝑇𝑐𝑒𝑙𝑙
𝑖
𝑠𝑖𝑛ℎ−1 [
]
2𝛼𝑎 𝐹
2𝑖𝑜,𝑎 (1 − 𝜃𝐶𝑂 )2
𝜂𝑐 =
−𝑅-𝑇𝑐𝑒𝑙𝑙
𝑖
𝑠𝑖𝑛ℎ−1 [
]
2𝛼𝑐 𝐹
2𝑖𝑜,𝑐
(6)
(7)
(8)
where θco is CO coverage area and it is the function of hydrogen and CO and cell
temperature.
𝜃𝐶𝑂 = 𝐴 𝑙𝑛
[𝐶𝑂]
[𝐶𝑂]
+𝐵 𝑙𝑛(𝑖) 𝑙𝑛
+𝐶
[𝐻2 ]
[𝐻2 ]
A=-0.00012784×Tcell×Tcell+0.11717499×Tcell -26.62908873
B=0.0001416×Tcell×Tcell-0ath.12813608*Tcell+28.852463626
C=-0.00034886×Tcell×Tcell+0.31596903×Tcell-70.11693333
The exchange current density of a anode, i0,a, and cathode, i0,c, can be calculated as
(9)
(10)
𝛾
𝑖𝑜,𝑎 =
𝑟𝑒𝑓
𝑖0,𝑎 𝑎𝑐,𝑎 𝐿𝑐,𝑎
𝐶𝐻2
−𝐸𝑐,𝑎
𝑇𝑐𝑒𝑙𝑙
[
] 𝑒𝑥𝑝 [
(1 −
)]
𝐶𝑟𝑒𝑓,𝑎
𝑅𝑇𝑐𝑒𝑙𝑙
𝑇𝑟𝑒𝑓,𝑎
𝑖𝑜,𝑐 =
𝑟𝑒𝑓
𝑖0,𝑐 𝑎𝑐,𝑐 𝐿𝑐,𝑐
𝐶𝐻2
−𝐸𝑐,𝑐
𝑇𝑐𝑒𝑙𝑙
[
] 𝑒𝑥𝑝 [
(1 −
)]
𝐶𝑟𝑒𝑓,𝑐
𝑅𝑇𝑐𝑒𝑙𝑙
𝑇𝑟𝑒𝑓,𝑐
𝛾
ηohmic= -0.0001667 × Tcell + 0.2289
(11)
(12)
(13)
where Cref,a is the reference concentration of anode (mole m-3), Cref,c is reference concentration
of cathode (J mole-1K-1), Ec,a is the activation energy of anode reaction (J mole-1 K-1), Ec,c is the
activation energy of cathode reaction, ac,a is the anode catalyst surface area (m2 g-1), ac,c is the
cathode catalyst surface area (m2 g-1), Lc,a is the anode catalyst loading (mg m-2), Lc,c is the
cathode catalyst loading (mg m-2). Value of the parameters used to compute the activation
losses are shown in the Table 1 (Authayanun et al., 2013; Mamlouk et al., 2011).
Table 1: Value of the parameters used to compute the activation losses.
Parameters
Reference exchange current density (A m-2)
Catalyst loading (mg cm-2)
Reference cell temperature (K)
Catalyst surface area (m2 g-1)
Reference concentration (mol cm-3)
Activation energy (J mol-1 K-1)
Transfer coefficient
Cathode
0.004
0.4
373.15
64
0.004
72400
0.75
Anode
1440
0.2
433.15
32.25
0.002
16900
0.5
The current produced by the stack, Ist, is given by
Ist = i Acell
(14)
where i is the current density and Acell is the total active area of a single cell, the stack voltage,
Vst, is given by
Vst = Vcell Ncell
(15)
where Vcell is the cell operating voltage and Ncell is the number of cells connected in series,
which is determined from the requirement of the power of the stack, which is given by,
Pst = Ist Vst
(16)
3. Results and discussion
3.1 Fuel reformer
The proportion of reactants needed to feed through the reformer depends on the steam to carbon
(S/C) ratio and also on the reformer operating temperature. Higher operating temperature with
excess steam yields more hydrogen as a product. The S/C ratio defined as number moles of
water per carbon atom in a feed stream. The variation of CO concentration in the reformate
gas, produced with the aid of various SR with respect to the change in S/C and reformer
operating temperature, is shown in Figure 5. A reduction of the S/C ratio will result in energy
savings in the form of reduced heat input to the primary reformer and in savings in process
steam. In addition, it will lower the pressure drops in the reactor, resulting in energy savings in
the process. As the S/C ratio increases from 1 to 6, the percentage of CO in reformate gas is
decreasing irrespective of the type of fuel used. As the reformer operating temperature
increases, the percentage of CO in the reformate gas also increases. S/C ratio should be greater
than one, to avoid choking. The various chemical reactions occurring in the SR are presented
in Table 2. The reaction pathway proceeds in different directions based on the S/C ratios,
operating temperature and pressure. Due to the restriction in the CO concentration in reformate
gas; CO needs to be converted into CO2 and hydrogen. WGS reactor is employed in the
conversion of CO into hydrogen rich fuel. Ethanol SR is performed at a temperature range of
800–1000 K and at an atmospheric pressure. The ethanol SR gives a high yield of hydrogen
along with by-products like water, CO, methane and CO2. Figure 5a shows highest CO% for
% CO (dry basis)
30
25
800K
20
900K
15
1000K
10
5
0
1
2
3
4
5
6
S/C ratio
Figure 5a: CO concentration in reformate gas from the ethanol SR calculated for various S/C
ratios and temperatures.
25
800K
% CO (dry basis)
20
900K
15
1000K
10
5
0
1
2
3
4
5
6
S/C ratio
Figure 5b: CO concentration in reformate gas from the methanol SR calculated for various S/C
ratios and temperatures.
25
800K
% CO (dry basis)
20
900K
15
1000K
10
5
0
1
2
3
4
5
6
S/C ratio
Figure 5c: CO concentration in reformate gas from the glycerol SR calculated for various S/C
ratios and temperatures.
25
800K
% CO (dry basis)
20
900K
15
1000K
10
5
0
1
2
3
4
5
6
S/C ratio
Figure 5d: CO concentration in reformate gas from the methane SR calculated for various S/C
ratios and temperatures.
Table 2: Reaction pathways of SR.
Fuel/set of
substances
Ethanol
Methane
Methanol
Glycerol
Butane
Propane
Iso-octane
Ethane
CO
Reaction
conditions
Sufficient steam
supply
Reaction formula
Remarks
C2H5OH + 3H2O 2CO2 +
6H2
Insufficient steam
supply
C2H5OH + H2O 2CO + 4H2
Autothermal steam
reforming
C2H5OH + 2.2H2O 1.47
CO2 + 4.66H2 + 0.26CO +
0.27 CH4
CH4 + H2O CO +3 H2
Ideal pathway, the
highest hydrogen
production
Undesirable products,
lower hydrogen
production
Hydrogen selectivity is
around 70% at a
temperature of 773 K
Steam-methane
reforming reaction
Steam-methanol
reforming reaction
Steam-Glycerol
reforming reaction
Butane reforming
reaction
Propane reforming
reaction
Iso-octane
reforming reaction
Ethane reforming
reaction
Water gas shift
reaction
CH3OH+ H2O CO2 + 3H2
CH3OH CO + 2H2
Reforming reaction
occurs at relatively low
temperature 513-533 K
C3H8O8 + 3H2O 3CO2 +
7H2
C4H10 + 4H20 4CO + 9H2
C3H8 + 3H20 3CO + 7H2
C8H18 + 8H20 8CO + 17H2
C2H6 + 2H20 2CO + 5H2
CO + H2O CO2 + H2
Reduces the coke
formation, enhance
hydrogen production
an S/C=1 and it is gradually decreasing as increase in S/C value. S/C= 1 means 3 molecules of
water per molecule ethanol. The ethanol reformer operating at S/C >2 and at reactor
temperature lower than 800 K gives better conversion producing reformate with CO
concentrations lower than 5%. Methanol SR reaction is recognized as a very attractive and
promising process of hydrogen production. From Figure 5b, the methanol reformer operating
at S/C >2 and temperature in between 800-900 K gives lower CO (<5%) concentration in the
reformate gas. If the methanol reformer operating at S/C >2 and temperature above 800 K gives
CO% less than 10. Further, CO concentration in the reformate stream can be reduced with help
of using WGSR. Chein et al. (2012) performed a methanol SR at a temperature from 513-533K
and at an atmospheric pressure. The reforming reaction occurring at a relatively low
temperatures. Methanol SR is widely studied and the most common catalysts are based on
copper, such as Cu/ZnO/Al2O3 (Iulianelli et al., 2014). Similar type of study are reported in
the literature, Authayanun et al., (2014) investigated the effect of WGSR effect on performance
of HT-PEMFC system and reported without WGSR, the system performance achieved more
than 50% with compare to presence of WGSR. The WGS reactor is not necessary for the
methanol SR process, when it integrated with the HT-PEMFC stack. Figure 5c, the glycerol
reformer operating below 800 K gives lower than 5% of CO concentration in the reformate
gas. Authayanun et al., (2013) reported that the glycerol SR process is independent of S/C ratio
and the reformer temperature, when operated at S/C > 2 and T > 900 K. Glycerol SR shows a
higher hydrogen production and lower carbon formation than methane SR. However, an
external heat requirement was needed to maintain the SR to achieve high system efficiency at
high fuel utilization. Adhikari et al (2007) studied a thermodynamic analysis of hydrogen
produced from the glycerine steam reforming process. Their study revealed that the best
conditions for producing hydrogen are at a temperature of > 900 K, a pressure of 1 atm and a
molar ratio of water/glycerine of 9:1. Under these conditions, methane production is minimized
during the reforming process and carbon formation is thermodynamically inhibited.
Natural gas contains CH4 that can be used to produce hydrogen with thermal processes, such
as steam-methane reformation and PO. In the industrial CH4 SR process, CH4 reacts with a
steam under 1–25 atmospheric pressure in the presence of a catalyst to produce hydrogen, CO,
and a relatively small amount of CO2. Figure 5d, the methane SR operating at S/C >3 and
operating temperature below 1000 K gives less than 10% CO conversion in the reformate gas.
Other hydrocarbon fuels like ethane, propane, butane and iso-octane can be used for a hydro
production via steam reforming process. Propane, n-butane is a major constituent of liquid
petroleum gas. Commercial SR processes use Ni-based catalysts because of their acceptably
high activity and significantly lower cost in comparison with the alternative precious metal
based catalysts. Schadel et al. (2009) studied a steam reforming of methane, ethane, propane,
butane, and natural gas over a rhodium-based catalyst. Ethane, propane, and butane are
converted at much lower temperatures than methane, also in natural gas mixtures. Trimm et al.
(2004) developed an auto-thermal reforming model of gasoline mixtures to produce hydrogen
for fuel cell applications. They predicted a 70% fuel conversion at S/C ratio between 2 and 3.4,
with selectivity to hydrogen of between 65 and 70%. In our study, SR of butane, ethane,
propane, and isooctane show almost identical conversion as a function of temperature and S/C.
However, a high S/C ratio is disadvantageous for fuel cell performance due to the lower
hydrogen content caused by the dilution of water. The SR operations of glycerol, ethanol and
CH4 strongly affect the HT-PEMFC performance, especially at a high current density operation
(Authayanun et al., 2014). The inclusion of the WGS reactor in the reforming processes of
glycerol, ethanol and CH4 can improve the efficiency of the HT-PEMFC system, whereas the
WGS reactor is not necessary for methanol SR integrated with HT-PEMFC. The design value
for S/C ratio is decided based on % CO and operating condition of the reformer temperature.
3.2 HT-PEMFC performance
Cell voltage is calculated from Eq.(1) based on Eq. (2-13) at various cell temperatures, current
densities and CO concentrations. The variation of cell voltage with respect to current density,
% CO and at a constant operating temperature of 453 K is shown in the Figure 6. The
concentration of CO has no effect on the fuel cell performance when operated at low current
density. At high current densities, the fuel cell performance decreases with an increase in the
reformer temperature and decrease in the S/C ratio, ie., at higher CO concentrations. In present
study, we have taken high current density value to predict the cell performance. From above
data, problem is formulated as a constant voltage of 0.43 V; Eq. (17) is expressed in terms of
current density as the dependent variable and temperature as the independent variable at a 5%
CO concentration.
Figure 6: The variation of cell voltage with respect to current density, % CO and at a constant
operating temperature of 453.
A variation of the local current density is calculated by using Eq. (17).
icell = -6.803×10-05×TC×TC+0.06374×TC-14.05
(17)
where icell is the local current density. The total heat released from each cell in the stack can be
written as, Qcell,
𝑄𝑐𝑒𝑙𝑙 = [−
∆𝐻𝑔,𝑇
−𝑉𝑐𝑒𝑙𝑙 ] 𝐴𝑐𝑒𝑙𝑙 𝑖𝑐𝑒𝑙𝑙
𝑛𝐹
(18)
The total amount of heat released from the stack can be written as, Qst,
Qst= Qcell Ncell
(19)
The spatial variation of the current density on the surface of the catalyst layer is predicted under
the condition of the stoichiometric ratio of hydrogen at three and varying inlet temperature of
the cathode feed. The spatial variation of the local current density in the cathode catalyst layer
is shown in Figure 7. It can be seen from Figure 7 that there is nearly ± 18% variation in the
current density over as compared to the mean value. As a result, the heat source also varies
with respect to local current density. The predicted temperatures in the cathode catalyst layer
is shown in Table 3. The temperatures in the cathode catalyst layers varies between 406 K and
452 K with a relatively cold spot forming on the catalyst layer closest to the inlet of the cathode
air. The temperature differences in the cathode catalyst layer are decreasing as the increase in
the air inlet temperatures.
Table 3: Temperature variations in the cathode catalyst layer at a constant flow rate of air.
Air inlet
temperature (K)
400
410
420
Temperature in the catalyst layer (K)
Maximum Minimum Volume average
452
406
439
452
414
442
463
439
456
Current density Volume
average (A cm-2)
0.823
0.833
0.864
Average current density in the catalyst layer is increasing as the increase of cathode air inlet
temperature. The spatial variation of the temperature profile in the cell is shown in Figure 8.
The temperature at the inlet of the cathode air to the outlet of the cathode air is around 30 K to
50 K, it will vary based on the cathode air inlet temperature. Figure 8 shows the temperature
variation along the flow direction as well as in the thickness and width of the cell. Estimated
number of cells required for a 5 kWe stack is shown in Table 4. The average current density
varies with respect to inlet temperature of the cathode air in the HT-PEMFC stack. The preheated cathode air reduces the temperature variations within the catalyst layer; it will result in
the increasing average current densities within the cell. For a cell voltage of 0.43 V, the
estimated electrical efficiency of the stack is 35, and 91 cells are required to produce the desire
output of 5 kWe. Further, these simulation results are used for the calculation of thermal energy
and electrical energy of the integrated HT-PEMFC system.
(a)
(b)
Figure 7: Spatial variation of the current density profile in the catalyst layer at various cathode
inlet temperatures (a) 400 K and (b) 410 K.
Table 4: HT-PEMFC stack specifications.
Power requirement (W)
Cell voltage (V)
Stack voltage (V)
Number of cells
Stack current (A)
Current density of a cell (A cm-2)
Cell area (cm2)
CO concentration in the feed (%)
Estimated electrical efficiency (%)
Estimated thermal efficiency (%)
5000
0.43
39
91
218
0.82
156
5
35
65
Figure 8: Spatial variation of the temperature profile in the HT-PEMFC.
3.3 Integration of HT-PEMFC stack with SR
Internal SR is thermally integrated with a 5 kW HT-PEMFC stack. The specification of 5 kW
is shown in Table 5. The 91-cell stack with an active cell area of 156 cm2 operates at a cell
voltage of 0.43 V, and 5% CO, and with an average current density of 0.82 Acm-2. Nominal
current density values of the fuel cell deviates from the actual values as shown in Table 4. At
higher current density values of fuel cell operation, the amount of heat released from the stack
is much higher than the electrical value. In such case the thermal efficiency of HT-PEMFC
stack is higher than the electrical efficiency. The overall energy balance of the HT-PEMFC
stack supplied with SR fuel is shown in Table 6. The energy balance calculations are tabulated
based on the flow sheet as shown in Figure 2.
The input fuel of the reformer is needed to preheat upto reformer operating temperature and
also fuel cell stack input air also need to heat up to stack operating temperature. The amount
of heat required for the preheating reformer fuel is equal to half the heat generated by fuel cells.
The heat rejected from the fuel cell having a lower temperature value and the amount of thermal
energy not enough to pre-heat reformer fuel.
Table 6: Heat balance of integrated SR with 5 kW HT-PEMFC stack.
Type of
fuel
S/C
ratio
Refor
mer
operat
ing
tempe
rature
(K)
Methane
4
Methanol
Ethane
Ethanol
Glycerol
Propane
Butane
Isooctane
Heat sinks (kW)
Amoun
t of
heat
require
d to
preheat
reform
er fuel
Amount of
heat
required
for an
endotherm
ic reaction
800
5.50
6
1000
4
Heat sources (kW)
Net
availabl
e heat
form
system
Amoun
t of
heat
availabl
e from
the
reforme
r gases
Amoun
t of
heat
availab
le from
the fuel
cell
stack
Amo
unt
of
heat
avail
able
from
burne
r
4.40
Amou
nt of
heat
requir
ed to
prehe
at fuel
cell
inlet
air
0.64
1.51
9.96
2.88
3.81
9.26
3.15
0.64
2.00
9.96
5.39
4.29
800
6.49
2.38
0.64
2.68
9.96
5.78
9.32
6
1000
10.32
1.39
0.64
3.50
9.96
7.51
9.32
4
800
6.52
5.49
0.64
2.79
9.96
3.02
3.12
6
1000
10.99
4.44
0.64
3.77
9.96
5.81
3.48
4
800
5.97
2.96
0.64
2.47
9.96
5.29
8.15
6
1000
9.72
2.05
0.64
3.29
9.96
7.20
8.04
4
800
7.94
2.59
0.64
3.19
9.96
6.50
8.48
6
1000
12.79
1.37
0.64
4.23
9.96
8.26
7.65
4
800
7.11
4.54
0.64
1.78
9.96
2.91
2.37
6
1000
11.87
3.00
0.64
2.41
9.96
5.86
2.72
4
800
7.37
5.77
0.64
3.01
9.96
2.94
2.14
6
1000
12.30
4.61
0.64
4.09
9.96
5.95
2.46
4
800
7.83
5.95
0.64
3.15
9.96
3.05
1.74
6
1000
13.10
4.71
0.64
4.31
9.96
6.16
1.97
The external heat source is needed to preheat the reformer fuel. The amount of heat needed for
the endothermic reforming reaction can be supplied by burning methane in the combustor. The
increase in the S/C ratio value more amount of thermal energy is required for the preheating of
the feed to reformer temperature. Around 60 % heat is required to preheat the reformer fuel as
increase in S/C ratio from 4 to 6. For an example, methanol need a 5.50 kW of thermal energy
to preheat the reformer fuel at reformer operating temperature of 800 k. Methanol is having
more excess thermal energy followed by ethanol and glycerol. At higher current densities the
presence of CO is not the only factor that affects the HT-PEMFC performance and system
efficiency. The content of hydrogen, and its purity, and the amount of energy consumed by the
system are also key factors that influence the system efficiency. The overall heat balance of the
system can be achieved by configuring heat exchangers in a multiple heating/cooling steps.
Conclusions
Integration of HT-PEMFC with reformer adds additional advantage, to improve the overall
system efficiency. The HT-PEMFC stack was modelled with a concept of varying local current
density in the cathode catalyst layer. There is nearly ± 18% variation in the current density over
as compared to the mean value. This local current density variation, which will effect in the
overall system efficiency. Around 60 % heat is required to preheat the reformer fuel as increase
in S/C ratio from 4 to 6. Methanol is having more excess thermal energy followed by ethanol
and glycerol. The recovery of the waste heat from the hot reformate gas stream, from the
cathode air and the anode off-gas aids a considerable increase in the overall efficiency of the
system. Proper location for the installation of heat exchanger in the system flow sheet diagram
needs to be configured in order to optimize all the structure of the integration process.
Acknowledgments
This work is supported by Hydrogen and Fuel Cell Technologies RDI Programme (HySA),
funded by the Department of Science and Technology in South Africa (project KP1-S03).
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