BIOHYTHANE AS AN ENERGY FEEDSTOCK FOR SOLID OXIDE FUEL CELLS

G.K. Veluswamya,c , C.J. Laycockb , K.Shahc, A. S. Balla , A. J. Guwyb, R. M. Dinsdaleb

aCentre for Environmental Sustainability and Remediation (EnSuRe), School of Science, RMIT University, Victoria, Melbourne, Australia.

bUniversity of South Wales, Sustainable Environment Research Centre (SERC), Llantwit Road, Treforest, Pontypridd, Rhondda Cynon Taff, Wales, United Kingdom. CF37 1DL.

cAdvance Energy Technologies and Waste Utilisation Research Group, School of Engineering, RMIT University, Victoria, Melbourne, Australia.

ABSTRACT:

Biogas (60%-CH4, 40%- CO2) is a potential source of renewable energy

when used as energy feedstock for solid oxide fuel cells (SOFC), but releases

biogenic CO2 emissions. Hybrid SOFC performance can be affected by fuel

composition and reformer performance. Biohythane (58%-CH4, 35%-CO2 and

7% H2 ) can be a better alternative providing balance between energy and

biogenic emissions. Biohythane performance is studied for a 120 kW SOFC

stack using ASPEN process model and compared with other feed stocks. This

work is the first to study and report on the application of biohythane in SOFC

systems. Biohythane was found to produce less biogenic CO2 emissions and

6% less CO at the reformer than biogas. Comparisons show that biohythane

provides better efficiencies in hybrid SOFC systems. Sensitivity studies

recommends operation of stack with biohythane at Steam to Carbon Ratio

(STCR) = 2.0, i = 200 mA cm-2 and UF = 0.85 respectively.

Highlights: Biohythane can be a better alternative to biogas as a renewable fuel

for SOFC

Biohythane provides better performance than biohydrogen and

syngas

Biohythane produces less biogenic CO2 emissions than biogas

At the reformer unit, biohythane produces 6% less CO compared with

biogas

Hydrogen in biohythane have very little effect on water-gas shift

reactions

Keywords: Biohythane, Biogas, Anaerobic Digestion, Solid Oxide Fuel Cell,

Hydrogen and Biogenic CO2.

1. INTRODUCTIONThe continuous increase in energy demand due to rapidly expanding global

rural population and industrialization has created a demand for alternatives to

fossil fuels [1]. Along with the energy demand, the rise in carbon emissions has

advanced major efforts on improving alternative renewable technologies with

minimized carbon footprint (CFP) [2]. There have been many studies to improve

the carbon neutrality of renewable energy, as even renewable energy resources

make some contribution to CFP [3,4]. The CFP of renewable energy resources

can be minimised by cogeneration of energy, which is the simultaneous

production of thermal energy with either mechanical or electrical energy [5].

Fuel cell technology has gained more interest among researchers recently as

an alternative to conventional cogeneration systems, owing to its size and better

conversion efficiency [6].

Solid oxide fuel cell (SOFC) technology is a very promising technology for

stationary standalone power generation, combined heat and power, and hybrid

systems where the fuel cell is coupled with turbines to improve overall system

efficiency. Fuel cells can beneficially produce more electrical energy per unit of

input energy than conventional electrical generation systems [7,8]. Fuel cells

have a modular design that can be configured in a number of different

applications depending on the required capacity, allowing for more flexibility in

the scaling of these systems. Due to the high operating temperature of SOFCs,

they can utilise a wide range of fuels including syngas, natural gas and liquid

fuels such as methanol and kerosene [9]. The working principle of SOFC

technology is discussed elsewhere [10-14].

Combining SOFC technology with gas turbines (GT) has many advantages

including increased efficiencies and low carbon emissions, and can alleviate

current and future energy and environmental demands [15]. The fuel

composition can potentially have a considerable effect on SOFC-GT

performance. CO in the fuel significantly reduces both SOFC and hybrid module

efficiency [16, 17]. This is attributed to the lower heating value and change in

Gibbs free energy of CO, resulting in slightly lower cell temperatures and

therefore requiring more external energy to maintain the desired anode

temperature. Recently, Lv et al., 2019 [18], studied the effect of fuel composition

on SOFC-GT systems and their results indicated that H2 has a slightly positive

effect on performance, while CH4 and CO have a negative influence . Hence, it

is important to find a good energy feedstock with less environmental impact and

also provide improved SOFC-GT performance.

Anaerobic Digestion (AD) in a waste water treatment plant (WWTP) can

provide a renewable energy-rich gas with flexibility in fuel composition.

Conventional AD processes can produce biogas (CH4-60%, CO2-40%) and

biohydrogen (H2-50%, CO2-50%), depending on the bacterial conversion

pathways. Detailed research has been conducted which reports on the

feasibility, economics and life cycle analysis of AD-derived biogas as a fuel for

SOFCs [19-24]. However, use of conventional AD-derived biogas has been

shown to be problematic towards SOFC anode operation when directly fed due

to carbon deposition [25]. Studies have been previously conducted into the use

of methane-free hydrogen [50% H2 /50% CO2] from dark fermentation

processes as a means to reduce carbon emissions from SOFCs and also to

reduce carbon deposition in SOFCs [26-29]. However, use of higher H2

contents in the SOFC fuel leads to thermal shock issues because more heat is

released due to the reduced cooling effect of the water-gas shift reaction (WGS)

[30], with a 50 °C increase in cell temperature reported at 60% H2 content in the

fuel cell. This indicates the conventional AD derived biogas and biohydrogen

can be a better alternative to renewable energy feedstocks for SOFCs.

In the last decade, researchers have successfully used food waste and bio

waste to produce biohythane in two-stage AD pilot scale systems with a

composition of CH4 (52-59%), CO2 (35-40%) and H2 (6-7%) [31-35]. This

involves co-generation of biogas and biohydrogen in parallel before blending in

the required proportions according to the energy and carbon emission demands

for SOFCs. To the authors knowledge there is no experimental or modelling

research studies reported for using biohythane as an energy feedstock for

SOFCs. This current research work will be the first of its kind to study and report

on the technical feasibility of using biohythane as fuel for SOFCs. An Aspen

plus V10 process model is used to study the performance of biohythane in a

120 kW SOFC stack. A sensitivity analysis has been conducted to determine

the ideal operating conditions of the system. Its performance against other

energy feedstocks such as syngas [36-38], natural gas, biohydrogen and biogas

was carried out for benchmarking.

1.1 ASPEN PROCESS MODEL DESCRIPTION

The ASPEN process model simulator is a widely applied research tool for

studying the thermo-physical behaviour of SOFCs under different conditions.

The comprehensive built-in process models of the ASPEN process simulation

software with extensive thermodynamic and physical property data bases, make

it a very convenient and efficient way to study different chemical process

systems. Recently, researches have used ASPEN to study SOFCs with planar

designs in reversible modes and as hybrid systems coupled with gas turbines

[39-42]. In this work, ASPEN is used to understand the technical feasibility of

biohythane utilisation as an energy feedstock. Accordingly, a zero-dimensional

SOFC model was built in to ASPEN plusV10 using available blocks for steady-

state simulation.

Design blocks and calculator blocks were used for determining the

required fuel, air and steam recycle ratios. The incoming air and fuels are

compressed to the desired pressure ratios in the compressor blocks. The high

operating temperature requirements of the SOFC are met by combustion of the

anode flue gases in a separate burner section and the same heat is utilized for

heating the inlet air to avoid thermal shocks. To avoid complexity, this work

assumes that a constant heat source is applied to the air in the heater block to

match the anode temperature. An ambient inlet air temperature of 25 is

assumed in this work. Heated air is split into nitrogen and oxygen at the cathode

section which was modelled using a separator block. Compressed fuel is

heated and reformed by the steam present in the recycled anode flue gas

coming from the anode outlet in the reformer block. In the reformer block, all the

methane reforming reactions (Equation 1, 2 and 3) are expected to occur.

Equation 1 is the methane steam reforming (MSR) reaction, which is highly

endothermic. The carbon monoxide yielded can be further reacted with steam to

produce more hydrogen and carbon dioxide via the water-gas shift reaction

(WGS), which is exothermic (Equation 2). Apart from steam present in the

anode flue gas, the carbon dioxide present in the biohythane fuel is also

considered to be a good methane oxidising agent [43]. This dry reforming

reaction (Equation 3) is highly endothermic and requires high operating

temperatures to achieve high methane conversion [44]. The net heat duty of

the reformer is set to zero; hence, the endothermic nature of the reforming

equations and exothermic nature of the water-gas shift reaction are properly

captured in the system. The hot gas from the reformer is mixed with the oxygen

coming from the cathode and passed to the anode which is set at a desired

operating temperature. High operating temperatures at the anode can lead to

direct oxidation of CO and hydrocarbons in the incoming fuel, but this is less

favourable compared with the water-gas shift of CO to H2 and reforming of

hydrocarbons to H2 [45]. Due to the complexity of the reaction system,

thermodynamic equilibrium analysis is determined using the non-stoichiometric

approach. In this approach, the equilibrium composition of the system is found

by the direct minimization of the Gibbs free energy for a given set of species

without any specification of the possible reactions that might take place in

system. The anode in this system is defined as an equilibrium R-Gibbs reactor

(Equation 1, 2, 3 and 4).

The assumption of the R-Gibbs reactor to define the water-gas shift

reaction was validated with experimental data on the SOFC [46]. The

experiments were carried out for different H2 and CO2 mixtures. The R-Gibbs

assumption is able to correctly predict the formation of CO and CO2 from the

WGS reaction as shown in Table 1 below. A fraction of the flue gases from the

anode outlet which are rich in steam is recycled back to the reformer unit while

the remainder is assumed to leave for the burner unit. The recycled steam to

carbon ratio (STCR) is calculated using a design specification block according

to incoming hydrocarbon percentile. The ASPEN base case model is

developed and validated for T-SOFC of 120 kW stack, which has been widely

used in earlier literature [47,48], where a similar anode gas recycle [AGR] has

been used for reforming. Voltage and fuel requirement calculations are carried

out in separate calculator blocks and are discussed in the following sections.

C H 4+H2O↔3H 2+CO (1)

CO+H 2O↔CO2+H2 (2)

C H 4+CO2↔2H2+2CO (3) H 2+0.5O2→H2O (4)

2. MATERIALS AND METHODS

2.1 Voltage and Current density calculation

The cell voltage is first calculated based on the Nernst Equation. The fuel

cell achieves maximum reversible voltage when no current from an external

load is applied (Equation 5) [37]. The same voltage is less than the reversible

voltage when current inform an external load is applied due to irreversible

ohmic, activation and concentration losses (Equation 6). In Eq. (5), g is the

molar Gibbs free energy of formation (J/mol) at standard pressure (1 bar), 2

represents the number of electrons produced per mole of H2 fuel reacted, F is

the Faraday constant (96 485 C/mol), Tavg is the average temperature between

the SOFC inlet and outlet streams (K), Rg is the molar gas constant of 8.314 J

mol-1 K and Pi is the partial pressure (in bar) of the gaseous component i. The

partial pressures were taken as average values of the anode and cathode inlet

and outlet streams.

The ohmic losses due to resistance of electron flow through the fuel cell

is calculated using Equations 7-10 [49]. The subscript terms A, C, E and int

refer to anode, cathode, electrolyte and interconnection respectively. Ap

(radians) and Bp (radians) are the angles related to the extent of electrical

contact and interconnection respectively. ρa, ρc, ρe and ρint (Ω m) are the

resistivity terms and are calculated based on the temperature-dependent

relationships shown in Table 2 [50]. The cell diameter is given by Dm (m), while

the thickness of the components is given by tA, tC, tE and tInt (m) respectively.

The interconnection width is defined by Wint (m). The current density is defined

as i (A/m2) (Equation (10)) and is calculated based on the imposed current I (A)

and total active area of the fuel cell S (m2). The corresponding ohmic losses are

calculated as the summation of the aforementioned individual losses (Equation

12). The activation losses due to the electrochemical reactions is defined by

Equations (13-14) as outlined in [51]. Ract,a, Ract,c (Ω m2) represent the specific

resistance at the anode and cathode respectively, while Ka, Kc are the pre-

exponential factors. The activation losses are calculated as the product of the

resistance terms and current flux density (i) as shown in Equation 15. The

losses due to mass transfer limitations are calculated as voltage concentration

losses. A simplified method based on Fick’s Law is used (Equation 16-17), [37].

The terms iLa, ILc, are the limiting current density and their values are given

Table 2. The cell concentration over potential losses can be obtained using

Equation 18.

The calculations described above are carried out using a design

specification and calculator blocks, where the input fuel flow is varied until the

SOFC stack power equals a specified value as a function of imposed current

(Equation 19-23). Uf and Ua are fuel and air utilization factors.

V R=−∆ g f2. f

±Rg .T avg

2. flnPH 2 .P

0.5O 2

PH2O(5)

V=V R−V ohm−V act−V con (6)

V ohm ,a=i . ρa .(A .π . Dm)

2

8.t a(7)

V ohm ,c=i . ρ c . (A .π . Dc)

2

8.t c. A .¿ (8)

V ohm ,e=i . ρa .t e (9)

V ohm ,∫ ¿=i . ρ∫ ¿ .(π .Dm )

t∫¿

w∫¿¿

¿¿¿ (10)

i= IS (11)

V R=V ohm , a+V ohm, c+V ohm ,e+V ohm,∫¿ ¿ (12)

1Ract , a

=Ka .2.FRg.T ( PH 2

P0 )M 1

exp ( −Ea

Rg . T ) (13)

1Ract , c

=K c .4. FRg .T (PO 2

P0 )M 2

exp ( −Ec

Rg . T ) (14)

V act=(Ract , a+Ract , c ). i (15)

V con , a=−R .Tne . F ( i

iL, a ) (16)

V con , c=−R .Tne .F ( i

iL ,c ) (17)

V con=V con ,a+V con , c (18)

I=2. F .nH 2 ,Consumned (19)

nH 2 ,∈¿=nH2 ,gas+1(nCO gas)+4 (nCH 4gas )+… ..¿ (20)

U f=nH 2 ,consumed

nH2 ,∈¿ ¿ (21)

nO2 , consumed=0.5nH 2 ,consumed (22)

U a=nO2, consumed

nO2 ,∈¿¿(23)

2.2 Model Validation

The developed ASPEN model predictions are compared and validated

against published data from earlier work in which gas compositions and power

capacities were varied (Table 3, Table 4 and Table 5). The gross and net

efficiencies are calculated by Equations 24 and 25. The power output is

calculated based on the gas compositions at the anode for a 92% DC to AC

conversion efficiency [47, 48]. The compressor loads are also taken in to

consideration in the evaluation of net efficiency, where LHV is defined as the

lower heating value of the fuel. The model predictions are in good agreement

with earlier works. The outlet gas compositions from the anode and pre-

reformer are similar to earlier reported literature results. The voltage and current

density predictions are well within the error limits of 2-3%. These slight

deviations are possible due to the assumptions that the reformer net heat duty

is zero and that the inlet air temperature is not heated by energy from stack gas

combustion. Assumptions made in the model development do not have any

major impact on the results and match well with previous published data. The

aim of this research is to determine the feasibility of using biohythane derived

from two-stage AD in order to reduce the carbon footprint, rather than to

optimize current SOFC technology. The validated ASPEN process model is

used for further studies in the work. The optimal blended mixture proportion is

identified, and its results are discussed in the following sections.

ƞeff ,gross=PEL , AC

nFuel ,∈¿. LHV Fuel¿ (24)

ƞeff ,net=PEL, AC−PCompnFuel ,∈¿ .LHV Fuel

¿ (25)

3. RESULTS AND DISCUSSIONThe validated model was run for a biohythane composition (CH4-58%,

CO2-35% AND H2-7%) derived from a two-stage AD pilot facility utilising food

waste [35]. As there are no previous pilot scale studies on the integration, it was

assumed the incoming fuel is at 200 °C and the air is at 25 °C. To understand

the effect of various operating parameters, a sensitivity analysis has been

conducted for an industrial-sized SOFC stack of 120 kW, as reported in earlier

publications with baseline operating parameters of Uf/Ua/STCR at 0.85/0.19/2.

3.1 Effect of Current Density

The imposed current density was varied from 50 to 550 mA/cm2 to study

its effect on the fuel cell stack operating performance (see Figure 2). An

increase in current density has a positive effect on the power while it has a very

significant and negative impact on the voltage which drops by 80%. The current

density has an undesirable effect on the system efficiencies by creating higher

fuel demands. An increase in fuel flow rates affects the loading capacity of

compressors, lowering the efficiencies. Peak power performance is observed at

around 450 mA/cm2; after that the power reduces and fuels cells are usually

operated left of this range [48]. Hence, it is desirable to operate the stack at

higher power outputs but at a reduced efficiency. However, a trade-off between

operating cost and performance has to be achieved. It is therefore

recommended to operate the stack between 180-200 mA/cm2, where both

voltages and efficiencies are higher for this biohythane mixture.

3.2 Effect of Fuel Utilization Factor

The fuel utilization factor was varied from 0.60 to 0.95 to determine the

effects on overall SOFC stack performance (Figure 3). Any increase in the fuel

utilization factor leads to an increase in power density as more H2 is utilised for

power generation. However, this increase in power density also leads to higher

voltage losses across the cell, as according to Nernst Equation. This trend is

captured in the model and reflected in Figure 3. A positive effect on efficiencies

was observed, resulting in a higher power output; at higher fuel utilization, less

fuel is required and hence higher efficiencies are observed. The demand for

higher power generation and better efficiency will drive the operation at higher

Uf; however, it is important to consider that SOFCs are a better technology

when used for driving cogeneration systems irrespective of the nature of the

fuel. At higher Uf, much less fuel is available for burners downstream and hence

less heat energy will be generated. Owing to these characteristics, it is

recommended to operate the stack for this biohythane blend at a typical range

between 0.80 and 0.85, where there is still depleted fuel available for

downstream combustion.

3.3 Effect of Steam to Carbon Ratio (STCR)

The overall performance of the SOFC is also determined by the

reforming unit characteristics. Reforming ensures maximum hydrogen is

available through methane reforming reactions and reduces the risk of coke

poisoning of the anode. Apart from the H2O available in the anode flue gas, CO2

present in the incoming fuel also act as a reforming agent. Earlier literature

studies by Tjaden et al., 2014 and Cozzolino et al., 2017 [24, 52] have shown

that for a typical biogas composition and reformer operation at 675-700°C, a low

STCR of 1 is sufficient to avoid coke formation. However, coke formation is not

a direct correlation for system efficiency, irrespective of whether biogas or

biohythane compositions are used. Hence a detailed sensitivity analysis (Figure

4 (a-d)) needs to carried out to understand the effect of STCR on system overall

performance. Figure 4a, depicts the effect of the STCR on system overall

performance for various anode operating temperatures. It is observed that there

is an overall improvement of 5% in system efficiency when the anode

temperature is increased from 800°C to 1000ºC. The increase in temperature

leads to an increase of reforming flue gas temperature and it assists with the

endothermic methane reforming reaction, as shown in Figure 4b. This

observation is in good agreement with Cozzolino et al., 2017 [52]. Also, it is

observed that there is an improvement of 3% when STCR is increased from 0.5

to 3. This result matches similar work conducted using biogas (CH4-55%, CO2-

45%) reformed with recycled anode flue gas [24]. Even though SOFC systems

are designed to operate at higher temperatures, the high temperatures can

make the material more susceptible to faster degradation [53].

Also, as discussed earlier by La Licata et al 2011 [30], the presence of

additional H2 in the incoming fuel can increase the heat stress on the reformer

side due to the WGS reaction, which is exothermic. Therefore, further studies

were carried out at 900 ºC operating temperature in order to avoid heat stress

issues.

Figure 4c depicts the performance of the reformer unit at 900°C as a

function of STCR. An increase in the recycled steam flow rate to the reformer

increases the reformer inlet temperature leading to a positive effect on the

reforming reaction and therefore the methane conversion increases as

observed in Figure 4c. Methane conversion reaches almost 100% as maximum

steam is recycled.; The reformer outlet temperature is lower than the inlet

temperature because the reforming reactions are endothermic. Figure 4d shows

the reformer unit outlet gas composition as a function of STCR. As discussed

above, an increase in steam flow rate will result in an increase of reforming

reactions and an increase in H2O in the reformer flue gas. There is a very

marginal decrease in the CO2 mole fraction by 1% as the recycled steam flow

rate increases. This indicates that along with methane steam reforming,

methane dry reforming is also occurring. Reforming reactions are highly

endothermic and the required heat is provided by the increasing steam flow

rate. This is observed by the increase in CO and H2 mole fractions. A good

reformer operating characteristic is defined by high H2 and high CO output as

this reduces coking and demonstrates adiabatic operation [54]. Hence, it is

recommended to operate the SOFC with high STCR between 2 and 2.5, where

a maximum efficiency of 49.5% is observed, beyond which the efficiency is

asymptotic in nature. The extra steam can potentially be used for other heating

purposes in downstream.

3.4 Performance Comparison

Based on the above studies, the recommended operating stack

characteristics will be STCR -2.0, i-200 and UF-0.85 for biohythane mixtures

(CH4-58%/ CO2-35% / H2-7%). The performance of biohythane is first compared

with biogas for reformer performance to understand the influence of the

presence of H2 in the biohythane. The composition of biogas mixtures varies

according to the source even though all are derived from AD process in WWTP.

Methane percentiles range from 60-70 % and CO2 varies between 30-40%, with

the remaining components being mostly nitrogen, oxygen and sulphides [55-

58]. Therefore, a standard biogas composition of 60% CH4 and 40% CO2 was

used for this comparison with a recommended STCR of 2, as reported earlier in

[24]. The comparison study was conducted for a fuel and an air inlet

temperature of 200 ºC and 25 ºC respectively, with fuel utilization at 85% for an

anode working temperature of 900 ºC. It is observed that both renewable AD

gases offer similar efficiencies as shown in Table 6. In the reformer unit, it is

observed that the methane conversion is increased by 2% for biogas when

compared to biohythane. Most importantly, the CO content in the biohythane

reformed gas is lower by 6% when compared with biogas. The similar efficiency

is due to the similar LHV energy content of biogas (CH4-60%, CO2-40% - 1

Kmol = 481 MJ) and biohythane (CH4-58%, CO2-35% , H2-7%- 1 Kmol = 482

MJ). This indicates biohythane can provide similar efficiencies as biogas and

can provide better performance of hybrid SOFC systems due to low CO levels

entering the anode. This observation agrees with Lv et al., 2019 [18].

Another important observation from this comparative study is the role of

H2 in the fuel. The presence of extra H2 in the biohythane fuel composition has

very little effect on the reverse water gas shift reaction as less CO is produced

when compared to reformed biogas fuels. A similar observation is made by [59],

when they reformed biogas with different proportions of steam. Hence, it will be

safe to assume that the presence of H2 in the biohythane, reduces the

stoichiometric carbon content in the incoming fuel, rather than significantly

influencing the reverse water gas shift reaction. The extra H2 is available as fuel

for the anode reactions, which is evident from the high H2 and H2O content in

both the reformer and anode flue gases for biohythane.

On comparison of the anode flue gas compositions, it is observed that

the equivalent biogenic CO2 emissions for biohythane are 6% less when

compared with biogas. Hence, it is expected that the use of biohythane over

biogas as energy feedstock in SOFCs will definitely reduce the equivalent

biogenic CO2 emissions of the overall AD process and will improve the carbon

foot print of waste water treatment considerably.

The performance comparison is further extended for other commonly

used fuel mixtures like natural gas and biomass derived syngas used in SOFC

stack. The STCR for natural gas and syngas is taken from previous published

literature. All of the gases are compared for a 120 kW SOFC stack.

Biohydrogen (50% H2-50% CO2) derived from AD is also compared for

performance. Table 7 lists all of the gaseous fuel mixtures studied in this work.

Figure 5 compares the performance of the SOFC stack running on

different gas mixtures. Natural gas has the highest operating efficiency and the

lowest CO2 emission when compared with other hydrocarbon based fuels. A

very high CH4 percentage and minimal CO2 volume makes natural gas the ideal

fuel for an SOFC system. Although natural gas gives a better overall

performance (51%-ƞ and 25% GHG-CO2), it is a fossil fuel and hence other

alternative gases have to be considered. The other alternative is the biomass

derived syngas, but it has a low performance overall with lower efficiency (40%)

and CO2 emissions (38%). The reduced efficiency of syngas is attributed to its

low energy content and the increased fuel and air required to ensure power

requirements are met. Also, production of syngas has its own Capex and Opex

burdens due to the purification and tar removal processes involved. The other

alternative, biohydrogen derived from AD process, gives higher biogenic CO2

emissions (45%) and also works with a lower efficiency (40%). This is due to

low calorific value of the fuel and its dilute nature due to presence of more CO2.

Hence the best fuels for a long-term and sustainable renewable energy will be

AD derived biogas and biohythane due to the better efficiency and biogenic CO2

emissions. However, as discussed previously, the two-stage AD derived

biohythane is expected to provide better performance when compared to single-

stage AD derived biogas in terms of reforming and biogenic CO2 emissions for

hybrid systems.

3.5 Summary

The above results and discussions are summarised as follows:

1. The ASPEN process model sensitivity studies recommend operation of an

SOFC stack operating on biohythane at STCR = 2.0, i = 200 mA cm-2 and UF

= 0.85.

2. Two-stage AD derived biohythane provides similar efficiencies to natural gas

and biogas. Also, biohythane provides better efficiencies than syngas and

biohydrogen.

3. H2 in the biohythane has very little effect on the water-gas shift reactions and

it directly contributes to electrochemical energy generation at the anode.

4. CO in the anode fuel can negatively impact SOFC-hybrid systems

performance due to its cooling effect. Biohythane when reformed produces

6% less CO when compared with biogas. Hence, biohythane is expected to

provide better performance than biogas when used as an energy feedstock

for SOFC-Hybrid systems.

5. Biogenic CO2 emissions from biohythane are also 6% less compared with

biogas at the anode flue gas section. Hence biohythane can decrease the

overall carbon foot print of the process.

Overall, the use of biohythane as an energy feed stock for SOFCs provides

many advantages to WWTPs and the energy industry with minimal

environmental impact. Promotion of biohythane production from two-stage

AD is more energy efficient compared with biogas produced from single-

stage AD [60]. The energy industry will benefit from higher efficiencies when

SOFCs are used in a hybrid setup. Most importantly, lower biogenic

emissions from biohythane provides a better alternative to fossil fuels like

natural gas, syn gas and biogas, which gives high biogenic CO2 emissions.

6. CONCLUSIONSTwo-stage AD derived biohythane was studied as an energy feedstock for

SOFC systems and is proposed to be a better alternative to single-stage AD

derived biogas. Biohythane produces 6% less biogenic CO2 at the anode and

6% less CO in the reformer, which will benefit the anode temperature. The H2 in

the fuel affects the fuel carbon stoichiometric coefficient negatively and has very

little effect on the water gas shift reactions. Overall, biohythane would give good

performance in SOFC hybrid systems and reduce the carbon footprint of AD

processes. Biohythane has the potential to be a long-term and sustainable

renewable energy resource.

ACKNOWLEDGMENTSThe authors would like to thank the Universities UK International (UUKi- RF-

2018-74) Rutherford fellowship programme for supporting and funding GKV and

the FLEXIS research project (Grant Number: WEFO 80835).

.

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Figure 1: ASPEN PROCESS MODEL FLOW SHEET

0

200

400

600

800

0

40

80

120

160

200

0 100 200 300 400 500 600

Volta

ge, m

A

Pow

er D

ensi

ty, k

W/c

m2;

AC

Pow

er, k

W;

Gro

ss a

nd N

et e

ffeci

ency

Current Density, mA/cm2

Current Density

Power Density, kw/cm2AC,PowerGross EffeciencyNet EffeciencyCell Voltage,mV

Figure 2: The effect of Current Density on Fuel Cell Stack Performance

0

100

200

300

400

500

600

700

800

0

10

20

30

40

50

60

0.55 0.65 0.75 0.85 0.95

Volta

ge, m

V; C

urre

nt D

ensi

ty, m

A/c

m2

Fuel

Flo

w R

ate,

km

ol/h

r; G

ross

and

Net

Ef

feci

ency

Fuel Utilization Factor, Uf

Uf@ 900 C

Fuel Flow Rate, kmole/hrGross effeciencyNet EffeciencyVoltage, mVCurrent Density, mA/cm2

Figure 3: The effect of Fuel Utilization Factor on Fuel Cell Performance.

Figure 4a: The Effect of Temperature on SOFC Performance

Figure 4b: The Effect of Temperature on Methane conversion in Reformer

0

20

40

60

80

100

600

700

800

900

1000

1100

1200

0 0.5 1 1.5 2 2.5 3 3.5

Rec

ycle

d St

eam

Rat

e, k

mol

e/hr

; Met

hane

C

onve

rsio

n %

Ref

orm

er In

let a

nd O

utle

t Tem

pera

ture

, K

Steam To Carbon Ratio (STCR)

Reformer Characterstics

Refromer Inlet TReformer outlet TConversion%STEAM

Figure 4c: The Effect of STCR on Reformer performance

0

0.05

0.1

0.15

0.2

0.25

0.3

0.35

0.4

0 0.5 1 1.5 2 2.5 3 3.5

Mol

e Fr

actio

n

Steam to Carbon Ratio, (STCR)

Refromer outlet gas composition

CH4 CO2 CO H2 H2O

Figure 4d: The Effect of STCR on Reformer Outlet Gas Composition

0

10

20

30

40

50

60

Natural gas Biogas Bio hythane Biohydrogen

Syngas

Syst

em E

ffeci

ency

, CO

2 in

flue

gas

%

Gross Effeciency Net Effeciency CO2

Figure 5: Comparison of SOFC performance for different gas mixtures

Table 1: Validation of ASPEN MODEL –GIBBS Reactor Assumption

Gas Composition (50% H2, 50% CO2)

[46] ASPEN MODEL (Gibbs Reactor)

CO2 0.48 0.46CO 0.04 0.04O2 0 0H2 0 0.04H20 0.48a 0.46

a, Back calculated based on mole balance.

Table 2: Model input parameters

Cell Length/ Diameter (m) 1.5/0.022

Anode Thickness, ta (m) 0.0001

Cathode thickness, tc (m) 0.0022

Electrolyte thickness, Te (m) 0.00004

Interconnection thickness tint (m) 0.000085

Interconnection width Wint (m) 0.009

Anode resistivity ρA (Ω m) 2.98*10^-5 exp(-1392/Top)

Cathode resistivity ρc (Ω m) 8.114*10^-5 exp (600/Top)

Electrolyte resistivity ρE (Ω m) 2.94*10^-4 exp (10350/Top)

Interconnection resistivity ρint (Ω m)

0.025

A/B 0.804/0.13

Pre-exponential factor, Ka/Kc (A/m2)

2.13*108 / 1.49 *1010

Slope, m 0.25

Activation Energy, Ea/Ec (J/mol) 110000/160000

Anode limiting current density, iLa

(A/m2)29900

Cathode limiting current density, iLc (A/m2)

21600

Table 3: Natural Gas Composition-1

[47] InputsNatural Gas Composition,

(Mole fraction)

CH4-81.3%, C2H6-0.29%, C3H8-0.4%, C4H10-0.2%, N2-

14.3, CO2-0.9%

Active area, (S), m2 96.0768

Anode Operating

Temperature (Top), K

1183.15

Input Air temperature, C 630

Input Fuel temperature, C 200

DC Power , kW 120

UF/UA/STCR 0.85/0.19/1.8

DC-AC inverter efficiency 92%

Results ComparisonLiterature Model

Current Density (mA/cm2) 178 179.5

Voltage (mV) 700 685

Net AC efficiency NA 50%

Gross AC efficiency 52% 51%

Anode Inlet gas

composition (mole %)

H2-27%, CO-5.6%,CH4-

10.1%,H2O-27.9%,CO2-

23.1%,N2-6.2%

H2-26.9%, CO-5.7%,CH4-

10.6%,H2O-27.6%,CO2-

23%,N2-6.1%

Anode exhaust gas

composition (mole %)

H2-11.6%, CO-7.4%, H2O-

50.9%,CO2-24.9%,N2-

5.1%

H2-11.7%, CO-7.3%, H2O-

50.8%,CO2-25%,N2-5%

Table 4: Natural Gas Composition -2

Extrapolated data [48] InputsNatural Gas Composition,

(Mass fraction)

CH4-93.8%, N2-3.8%, CO2-2.4%

Active area, (S), m2 96.0768

Anode Operating

Temperature (Top), K

1193.15

Input Air temperature, C 20

Input Fuel temperature, C 200

DC Power , kW 127.4

UF/UA/STCR 0.85/0.2/2

DC-AC inverter efficiency 92%

Results ComparisonLiterature Model

Current Density (mA/cm2) 200.6 194

Voltage (mV) 661 672

Net AC efficiency NA 50%

Gross AC efficiency 48 51%

Anode Inlet gas

composition (mass %)

H2-3.16%, CO-

11.2%,CH4-5.81%,H2O-

27.3%,CO2-51.2%,N2-

1.24%

H2-3%, CO-9%,CH4-

7.2%,H2O-27%,CO2-

52.8%,N2-1.26%

Anode exhaust gas

composition (mass %)

H2-1.39%, CO-11.9%,

H2O-39.88%,CO2-

45.9%,N2-0.94%

H2-1%, CO-9%, H2O-

41.1%,CO2-47.8%,N2-

0.9%

Table 5: Syngas results comparison

[48] InputsSyngas Composition,

(Mole fraction)

H2-34%, CO-16%, CH4-7.4%, CO2-15.8%, N2-1%,

H2O-25.7%

Active area, (S), m2 96.0768

Anode Operating

Temperature (Top), K

910

Input Air temperature, C 24.9

Input Fuel temperature, C 200

DC Power , kW 120

UF/UA/STCR 0.85/0.167/2.5

DC-AC inverter efficiency 92%

Results ComparisonLiterature Model

Current Density (mA/cm2) 188 182

Voltage (mV) 662 676

Net AC efficiency 37% 39%

Gross AC efficiency 43% 43%

Anode Inlet gas

composition (mass %)

H2-29.4%, CO-7.2%,CH4-

3.6%,H2O-33.1%,CO2-

25.8%,N2-1%

H2-32%,CO-9%,CH4-

2%,H2O-31%,CO2-

24.7%,N2-0.9%

Anode exhaust gas

composition (mass %)

H2-6.2%, CO-4.2%, H2O-

58.7%,CO2-30%,N2-0.9%

H2-7%, CO-5%, H2O-

57.7%, CO2-29%,N2-0.9%

Table 6: Performance comparison for biohythane vs biogas

Biogas (CH4-60%,

CO2-40%)

Biohythane (CH4-58%,

CO2-35%, H2-7%)

Fuel required, kmol/hr 1.65 1.66

Reformer inlet T, K 1031 1023

Reformer outlet T, K 813 812

Recycled anode gas, Kmol/hr 6.63 5.95

Methane conversion at

Reformer

90% 88%

Reformer

outlet gas

composition

CO 0.08 0.075

CO2 0.37 0.35

H2 0.22 0.24

H2O 0.26 0.27

Anode outlet

gas

composition

CH4 0.00 0.00

CO 0.082 0.079

CO2 0.37 0.35

H2 0.08 0.086

H2O 0.46 0.48

System Gross efficiency, % 49.9 49.5

Table 7: Model Input Parameters

Natural Gas

Syngas BioHythane Biogas Biohydrogen

CH4 0.82 0.07 0.58 0.6 0C2H4 0.00 0.0 0 0 0C2H6 0.03 0.01 0 0 0C3H8 0.004 0.0 0 0 0C4H10 0.002 0.0 0 0 0N2 0.143 0.0 0.0 0.1 0CO 0.009 0.23 0 0 0CO2 0.0 0.19 0.35 0.3 0.5H2 0.0 0.26 0.07 0 0.5H2O 0.0 0.23 0 0 0Uf 0.85 0.85 0.85 0.85 0.85STCR 1.8 2.71 2.0 2.0 NAAnode Temp,°C

900 900 900 900 900

Inlet air ,°C 20 20 20 20 20Fuel ,°C 200 200 200 200 200