Mini CHP based on the electrochemical generator and impeded fluidized bed reactor for methane steam reforming

Abstract

The paper presents a configuration of mini CHP with the methane reformer and planar solid oxide fuel cell (SOFC) stacks. This mini CHP may produce electricity and superheated steam as well as preheat air and methane for the reformer along with cathode air used in the SOFC stack as an oxidant. Moreover, the mathematical model for this power plant has been created. The thermochemical reactor with impeded fluidized bed for autothermal steam reforming of methane (reformer) considered as the basis for the synthesis gas (syngas) production to fuel SOFC stacks has been studied experimentally as well. A fraction of conversion products has been oxidized by the air fed to the upper region of the impeded fluidized bed in order to carry out the endothermic methane steam reforming in a 1:3 ratio as well as to preheat products of these reactions. Studies have shown that syngas containing 55% of hydrogen could be produced by this reactor. Basic dimensions of the reactor as well as flow rates of air, water and methane for the conversion of methane have been adjusted through mathematical modelling.

The paper provides heat balances for the reformer, SOFC stack and waste heat boiler (WHB) intended for generating superheated water steam along with preheating air and methane for the reformer as well as the preheated cathode air. The balances have formed the basis for calculating the following values: the useful product fraction in the reformer; fraction of hydrogen oxidized at SOFC anode; gross electric efficiency; anode temperature; exothermic effect of syngas hydrogen oxidation by air oxygen; excess entropy along with the Gibbs free energy change at standard conditions; electromotive force (EMF) of the fuel cell; specific flow rate of the equivalent fuel for producing electric and heat energy. Calculations have shown that the temperature of hydrogen oxidation products at SOFC anode is 850 °C; gross electric efficiency is 61.0%; EMF of one fuel cell is 0.985 V; fraction of hydrogen oxidized at SOFC anode is 64.6%; specific flow rate of the equivalent fuel for producing electric energy is 0.16 kg of eq.f./(kW·h) while that for heat generation amounts to 44.7 kg of eq.f./(GJ). All specific parameters are in agreement with the results of other studies.

Introduction

One of the trends in modern electric power industry which allows increasing the efficiency of the electricity production from hydrocarbon fuel as well as reducing carbon emissions may be the development of power plants based on electrochemical generators (fuel cells) [1], [2]. The reactor for producing hydrogen along with the fuel cell stack is the key elements of such plants. Solid oxide fuel cells (SOFC) may be the most advanced due to the air used as oxidizer and syngas (which could contain CO along with hydrogen) used as fuel in them [3]. Syngas is usually produced by conversion of methane [4], [5] or coal [6], [7] in reactors through a catalyst in the form of a granular bed or structured surfaces [8], [9], [10]. The syngas and air in the anode and cathode channels of planar composite SOFC stacks flow transversely [11], [12].

The paper considers a configuration of the fuel cell power plant with impeded fluidized bed reactor for autothermal steam reforming (reformer) of methane by a dispersed heat carrier of the original design which could be used to create reactors producing hydrogen for high-capacity power plants. Specific fuel flow rates for producing electric and heat energy have been determined according to the proposed configuration as well as the comparison of this data with the performance of modern steam power plants has been carried out. The temperature of the methane reformer along with the fraction of hydrogen oxidized at the anode as well as EMF and electric efficiency of the planar SOFC stack has been determined to analyze the efficiency of the configuration components.

Chemical reactions proceeding in the configuration compounds have been considered as well as heat balance equations have been set up while calculating parameters of the power plant configuration [13], [14], [15]. The following equations may be considered as basic equation models for calculating the proposed cogeneration plant:

• Heat balance equation for the reformer with partial oxidation of conversion products with a view to ensuring autothermal operation (determines a fraction of combusted syngas required to reach the defined temperature in the reactor);

• Heat balance equation for the fuel cell to determine a fraction of hydrogen oxidized at the anode;

• Power balance equation for WHB to determine external heat loss with outgoing gases;

• Equation for calculating the specific fuel rate to produce electric and heat energy.

The successive solution of the equations described may allow calculating the temperature and flow rate of gases in configuration components of the power plant.

The power plant configuration is shown in Fig. 1.

The basis for the syngas production is thermochemical reactor for autothermal steam reforming of methane 1 consisting of the following elements: non-sifting distributor plate 2; thermal insulation 3; industrial aluminium-nickel catalyst box KSN-2 from cylinders with a hole (18 × 18 × 4 mm) 4; retort made of high-alloy steel 5; fluidized bed of electrocorundum (0.32 mm) 6; chamber with perforations for air injection required for oxidation of a fraction of syngas 7; central pipe designed to supply syngas to a consumer 8; baffle board for breaking gas bubbles which prevent mixing of air with syngas [16] 9 fed to the combustion chamber for oxidation; cyclone dust separators for the products of conversion 10 and combustion 11; pressure equality regulator in the conversion and combustion chambers 12; valves 13; feed pump 14; syngas cooler 15; mechanical filter 16; economizer 17; SOFC stack 18; waste heat boiler 19; circulation pump 20; desulphurization unit 21; water heater 22; stack flues 23; blower 24.

The combustion chamber, where the fraction of syngas passing between retort 5 and central pipe 8 is combusted in the air flow supplied through perforations 7, has been designed to maintain the temperature of 1,100 °C in the reaction volume of the reformer. The heat generated in the fluidized bed is transferred to the catalytic volume by the circulation of the dispersed particles of electrocorundum and spent on the endothermic reaction and preheating of the reaction products up to 1,100 °C. Feed water is heated in cooler 15 and economizer 17. System water is preheated in water heater 22.

The experiments have been conducted in the reactor with retort diameter of 0.18 m and inner syngas discharge tube diameter of 0.127 m. The KSN-2 catalyst box is 0.7 m high with the bed total height of 1 m consisted of electrocorundum particles of 0.32 mm in size.

Natural gas with methane content CH₄ = 98.67% along with water steam has been used for reforming at 250 °C (mole ratio H₂O:CH₄ $\cong$1:1 has been maintained). The portion of the product (syngas) discharged from the central pipe has amounted to 0.5. The inlet natural gas and air temperature has been 20 °C. The temperature in the fluidized bed has amounted to 1,100 °C while that under the catalyst box – 800 °C. Parameters of heat carriers used for methane reforming are given in Table 1.

The combustion heat of the syngas produced is $Q_{\text{н}}^{\text{≿}}=\mathrm{17,872}$kJ/kg. Volumetric contents of dry syngas at the central tube outlet and combustion products at the combustion chamber outlet are given in Table 2.

Combustion products may contain a slight amount of carbon monoxide due to the incomplete combustion.

The following stoichiometric reaction of methane with water steam has proceeded in the reaction volume of the reactor catalyst in configuration with SOFC at mole ratio of 1:3 to prevent sooting:

$$\mathrm{C}{\mathrm{H}}_4+3{\mathrm{H}}_2\mathrm{O}=0.71\mathrm{CO}+0.29\mathrm{C}{\mathrm{O}}_2+3.29{\mathrm{H}}_2+1.71{\mathrm{H}}_2\mathrm{O}$$

with endothermic effect q_x2 = 2,777 kJ/kg of syngas. The isobaric specific heat capacity of syngas c_sg = 3.18 kJ (K·kg of syngas) at 1,100 °C has been determined from Eq. (1).

The following stoichiometric oxidation reaction of syngas has proceeded in the combustion chamber with the excess air factor of 1.2:

$$0.71\mathrm{CO}+0.29\mathrm{C}{\mathrm{O}}_2+3.29{\mathrm{H}}_2+1.71{\mathrm{H}}_2\mathrm{O}+1.2(2{\mathrm{O}}_2+7.52{\mathrm{N}}_2)=\mathrm{C}{\mathrm{O}}_2+5{\mathrm{H}}_2\mathrm{O}+0.4{\mathrm{O}}_2+9.024{\mathrm{N}}_2$$

with endothermic effect q_x1 = 14,247.5 kJ/kg of syngas.

The isobaric specific heat capacity of the products of the syngas complete combustion c_c = 8.44 kJ/(K·kg of syngas) has been determined from Eq. (2). The calculated volumetric contents of wet syngas as well as of the products of its complete combustion are given in Table 3.

As seen in Table 3, hydrogen concentration in syngas has decreased to 54.84% while the concentration of carbon dioxide and water steam has increased with the decrease in the methane to steam ratio to 1:3.

Heat from the combustion of a portion of syngas (1−x)q_x1(1−q₃−q₅) brought in with superheated steam q₆, hot air q₇ and methane q₈ may be spent for the endothermic reaction of syngas production x·q_x2 as well as for preheating this reaction products x·c_sg·t and combustion products (1−x)·c_c·t up to 1,100 °C.

$$(1-x)q_{\text{х}1}(1-q_3-q_5)+\sum _6^8{q}_i=x·q_{\text{х}2}+x·c_{\text{sg}}·t+(1-x)·c_{\text{≿}}·t\text{,}$$

where (1−q₃−q₅) = 0.8; $\sum _6^8{q}_i=\mathrm{2,950}$kJ/kg of syngas.

Substitution of t = 1100°С in Eq. (3) yields the following value for x content:

$$x=\frac{q_{\text{х}1}(1-q_3-q_5)+\sum _6^8{q}_i-c_{\text{≿}}·t}{q_{\text{х}1}(1-q_3-q_5)+q_{\text{х}2}+c_{\text{sg}}·t-c_{\text{≿}}·t}=\mathrm{0.6.}$$

The following syngas flow rate may be required for the power plant with the electric power of Q_E = 5.44 kW, efficiency η = 0,628 [2] and reformer dimensions described in the experimental section:

$$B_{\mathrm{sg}}=Q_{\mathrm{E}}/(q_{\mathrm{x}1}·\eta )=5.44/\left(\mathrm{14,247.5}·0.628\right)=60.97·{10}^{-5}kg/s·\left(2.195kg/h\right)\text{.}$$

Here, input flow rates at x = 0.6 would amount to: methane G_M = 0.836; water G_H2O = 2.82; air G_a = 6.89 kg/h; combustion products at the outlet G_cp = 8.351 kg/h. The chemical efficiency of the methane reformer is

$${\eta }_x=\frac{B_{\text{sg}}{q}_{{\text{х}}_1}}{G_{\text{M}}{Q}_{\text{н}}^{\text{p}}}=0.6097·{10}^{-3}·\frac{14\mathrm{,}247.5}{\frac{0.836}{\mathrm{3,600}}\mathrm{49,090}}=\mathrm{0.76.}$$

According to [17], the fraction of EMF from CO oxidation at SOFC anode is less than 1% of the hydrogen oxidation fraction, so it has been neglected in calculations. Only hydrogen has been oxidized at SOFC anode. The water-gas shift reaction is not a hydrogen donor since it proceeds only at iron-chromium catalyst [18] that has not been in the anode channel.

At SOFC anode, hydrogen may be oxidized from syngas by oxygen coming from electrolyte by the following stoichiometric reaction:

$$3.29{\mathrm{H}}_2+0.71\mathrm{CO}+0.29\mathrm{C}{\mathrm{O}}_2+1.71{\mathrm{H}}_2\mathrm{O}+1.645{\mathrm{O}}_2=5{\mathrm{H}}_2\mathrm{O}+0.71\mathrm{CO}+0.29\mathrm{C}{\mathrm{O}}_2$$

with endothermic effect $\Delta {\mathbf{H}}_1^0=-\mathrm{11,375}$kJ/kg of syngas.

The flow rate of gases at the anode channel outlet is B^∗_sg = 3.84 kg/h (1.068·10⁻³kg/s), C^∗∗_sg = 2.04 kJ/(kg·K). Excess entropy of the reaction (4) is ΔS⁰ = 1.955 kJ/(K·kg of syngas). Gibbs energy change of the reaction (4) at standard conditions is ΔG⁰ = 13,570 kJ/kg of syngas. According to [19], the temperature at SOFC anode is:

$$T_3^{\ast }=\frac{\Delta H_1^0-\Delta G^0}{\Delta S^0}=\frac{-\mathrm{11,375}+\mathrm{13,570}}{1.955}=\mathrm{1,22}\text{К}\left(850°\text{C}\right)\text{.}$$

The temperature level is 783 °C being in accordance with [20]. At $T_3^{\ast }=1\mathrm{,}122\text{К}$, EMF of one planar element is:

$$E_E=-\frac{\Delta G_{\mathrm{1,122}\text{К}}}{2F}=\frac{186·{10}^3}{2·9.648·{10}^4}=0.985\text{V,}$$

where ΔG_{1,122 К} is Gibbs energy change at 1,122 K; K is 186·10³ kJ/kmol of water.

Conventionally, fuel cells may operate under isobaric-isothermal conditions when the electric power is equal to the Gibbs free energy decrease ΔG kJ/kg of syngas at 1,122 K multiplied by the syngas flow rate B_sg, kg/s:

$$Q_E=\Delta G_{1,122\mathrm{k}}·B_{sg}=\mathrm{8,742}·60.97·{10}^{-5}=5.33\text{kW}\text{.}$$

On the other hand, the electric power can be expressed [20] by the following formula:

$$Q_E=B_{sg}·q_{\times 1}\eta ,\mathrm{kW}\text{.}$$

Gross electric efficiency of the electrochemical generator (ECG) η = 0.61 (61%) may be determined by equating these formulas.

In the equation, the sum of power generated at SOFC anode during hydrogen oxidation from syngas as well as brought in by syngas flowing from cooler 15 (Fig. 1) and the air entering the cathode channel from WHB $G_{\text{a}}^{\text{*}}·c_{\text{a}}^{\ast }·t_2$is equal to the electric power taken away from ECG B_sg·q_x1·η as well as of syngas exiting the anode channel $B_{\text{sg}}^{\ast }·c_{\text{sg}}^{\ast \ast }·t_3^{\ast }$ and the air exiting the cathode channel $G_{\text{a}}^{\ast \ast }·c_{\text{a}}^{\ast \ast }·t_2^{\ast }$:

$$B_{\text{sg}}·{\phi }_{\text{S}}·\Delta H_1^0+B_{\text{sg}}·c_{\text{sg}}^{\ast }·t_3+G_{\text{a}}^{\text{*}}·c_{\text{a}}^{\ast }·t_2=B_{\text{sg}}·q_{\text{х}1}·\eta +B_{\text{sg}}^{\ast }·c_{\text{sg}}^{\ast \ast }·t_3^{\ast }+G_{\text{a}}^{\ast \ast }·c_{\text{a}}^{\ast \ast }·t_2^{\ast }\text{,}$$

where φ_s is a fraction of oxidized hydrogen at SOFC anode; c^∗_sg is true specific heat of syngas at constant pressure equal to 3.0 kJ/(K·kg of syngas) at 750 °C (determined from Eq. (1); t₃ is the temperature of syngas at the anode channel inlet from coolant 15 which has been assumed 350 °C lower than the temperature of conversion products at the coolant 15 inlet, i.e. t₃ = 1100–350 = 750 °C. The fraction of hydrogen oxidized at the anode may be determined from the heat balance equation as follows:

$${\phi }_{\text{S}}=\frac{B_{\text{sg}}·q_{\text{х}1}·\eta +B_{\text{sg}}^{\ast }·c_{\text{sg}}^{\ast \ast }·t_3^{\ast }+G_{\text{a}}^{\ast \ast }·c_{\text{a}}^{\ast \ast }·t_2^{\ast }-B_{\text{sg}}·c_{\text{sg}}^{\ast }·t_3-G_{\text{a}}^{\text{*}}·c_{\text{a}}^{\ast }·t_2}{B_{\text{sg}}·\Delta H_1^0}.$$

Hydrogen unoxidized at the anode and carbon monoxide coming from the anode channel with the air depleted by oxygen coming from the cathode channel has burned down in WHB.

Oxidation of carbon monoxide has proceeded according to the following stoichiometric equation:

$$0.71\mathrm{CO}+0.29\mathrm{C}{\mathrm{O}}_2+5{\mathrm{H}}_2\mathrm{O}+(2-1.645){\mathrm{O}}_2+7.52{\mathrm{N}}_2=\mathrm{C}{\mathrm{O}}_2+5{\mathrm{H}}_2\mathrm{O}+7.52{\mathrm{N}}_2\text{.}$$

The exothermic effect of reaction (5) is $\Delta {\mathbf{H}}_2^0=-2\mathrm{,}872$kJ/kg of syngas. The isobaric specific heat capacity of combustion products outgoing from WHB is c_out = 1.55 kJ/(K·kg).

The flow rate of the air fed to the cathode channel for the complete syngas combustion and cathode cooling has been determined by the stoichiometric equation of reaction (2): ${\mathbf{G}}_a^{\ast }=10.32kg/h$(2.87·10⁻³kg/s); the air flow rate at the cathode channel outlet is ${\mathbf{G}}_a^{\ast \ast }=8.314kg/h$(2.31·10⁻³kg/s); the flow rate of gases outgoing from WHB is G_out = 12.16 kg/h (3.37·10⁻³kg/s).

The balance equation may contain: heat power coming from the anode channel $B_{\text{sg}}^{\ast }·c_{\text{sg}}^{\ast \ast }·t_3^{\ast }$; from the cathode channel $G_{\text{a}}^{\ast \ast }·c_{\text{a}}^{\ast \ast }·t_2^{\ast };$ with feed water from economizer$G_{{\text{H}}_2\text{O}}·h_{\text{ECO}}$; with air for the cathode channel $G_{\text{a}}^{\ast }·c_{\text{a}}·t_{\text{a}}$; with air for the reformer$G_{\text{a}}·c_{\text{a}}·t_{\text{a}}$; with methane $G_{\text{м}}·c_{\text{м}}·t_{\text{м}}$; from oxidation of unoxidized hydrogen at the anode $(1-{\phi }_{\text{S}})·\Delta H_1^0·B_{\text{sg}}$and carbon monoxide $\Delta H_2^0·B_{\text{sg}}$(the heat power that is spent for the production of the superheated steam for the reformer $G_{{\text{H}}_2\text{O}}·h_{\text{ss}}$); with preheated air at the cathode inlet $G_{\text{a}}^{\ast }·c_{\text{a}}^{\ast }·t_2$; with preheated air at the methane reformer inlet $G_{\text{a}}·c_{\text{a}}^{\ast \ast \ast }·t_4$; with preheated methane at the reformer inlet $G_{\text{м}}·c_{\text{м}}^{\ast }·t_{\text{м}}^{\ast }$; and with the combustion products at the WHB outlet $G_{\text{out}}·c_{\text{out}}·t_{\text{out}}$.

$$B_{\text{sg}}^{\ast }·c_{\text{sg}}^{\ast \ast }·t_3^{\ast }+G_{\text{a}}^{\ast \ast }·c_{\text{a}}^{\ast \ast }·t_2^{\ast }+G_{\text{H}2\text{O}}·h_{\text{ECO}}+G_{\text{a}}^{\ast }·c_{\text{a}}·t_{\text{a}}+G_{\text{a}}·c_{\text{a}}·t_{\text{a}}+G_{\text{м}}·c_{\text{м}}·t_{\text{м}}+(1-{\phi }_{\text{S}})·\Delta H_1·B_{\text{sg}}++\Delta H_2·B_{\text{sg}}=G_{\text{H}2\text{O}}·h_{\text{ss}}+G_{\text{a}}^{\ast }·c_{\text{a}}^{\ast }·t_2+G_{\text{a}}·c_{\text{a}}^{\ast \ast \ast }·t_4+G_{\text{м}}·c_{\text{м}}^{\ast }·t_{\text{м}}^{\ast }+G_{\text{out}}·c_{\text{out}}·t_{\text{out}}$$

Substitution of the values listed in Table 4 may give the temperature t_out = 520 °C instrumental for defining losses.

The absolute loss with outgoing gases may be written as follows:

$$Q_2=G_{\text{out}}{\succsim }_{\text{out}},t_{\text{out}}=3.379·{10}^{-3}·1.55·520=2.796\text{kW,}$$

The fractional loss with outgoing gases:

$$q_2=\frac{Q_2}{G_{\text{м}}{Q}_{\text{H}}^{\text{р}}{\eta }_x}=\frac{2.796}{0.232·{10}^{-3}·\mathrm{49,090}·0.763}=\mathrm{0.321.}$$

The fractional external heat loss through the thermal insulation:

$$q_5=1-\eta -q_2=1-0.61-0.321=\mathrm{0.069.}$$

An incomplete combustion is assumed to be zero.

Power coming to the water heater 22 to preheat system water may be written as follows:

$$Q_{\text{cp}}=G_{\text{out}}·c_{\text{out}}·\left(t_{\text{out}}-t_{\text{out}}^{\ast }\right)=3.379·{10}^{-3}·1.55·\left(520-120\right)=2.09\text{kW,}$$

where $t_{\text{out}}^{\ast }$ is the temperature of outgoing gases after the water heater (120 °C is assumed).

Rather high values of the temperature and power loss with outgoing gases may be due to the fact that 35.4% of hydrogen passes through the anode channel in transit to WHB where it is oxidized along with anodic carbon monoxide.

The specific flow rate of the equivalent fuel to produce electricity may be written as follows:

$$b_{\text{E}}=\frac{1·\mathrm{3,600}}{Q_{\text{eq}\text{.f}\text{.}}·{\eta }_x}=\frac{1·\mathrm{3,600}}{\mathrm{29,330}·0.763}=0.16\text{kg}\text{eq}\text{.f}\text{.}/\left(\text{kW·h}\right)\text{.}$$

The specific flow rate of the equivalent fuel to produce heat energy is:

$$b_{\text{S}}=\frac{{10}^6}{Q_{\text{eq}\text{.f}\text{.}}·{\eta }_x}=\frac{{10}^6}{\mathrm{29,330}·0.763}=44.7\text{kg}\text{of}\text{eq}\text{.f}\text{.}/\text{GJ}\text{.}$$

The specific flow rate values calculated by standard methods closely coincide with parameters of modern electric power plants operating in heat-extraction modes [26].

Calculations of the power plant configuration have shown that the temperature level at the anode (850 °C) during hydrogen oxidation by atmospheric oxygen and EMF of one fuel cell (0.985 V) lie within the range of values given in Refs. [20], [21]. Gross electric efficiency of a fuel cell (61.0%) is consistent with data given in Refs. [22], [23], [24] while the fraction of hydrogen consumed in the anode channel ${\phi }_{\text{S}}=64.6\%$ is in agreement with data provided in Ref. [25]. Only 64.6% of hydrogen coming with syngas may be oxidized in the anode channel when the fuel cell efficiency is rather high. Low oxidation level may be related to low hydrogen content (54.84 m³/m³) in syngas produced in the methane reformer at steam to methane ratio of 3:1. Syngas with higher hydrogen concentration (e.g. 78 m³/m³) produced at water to methane flow rate ratio (FRR) of 1:1 could not be used due to possible carbon deposition along with blocking of pores of the anode channel material of the fuel cell.

Low hydrogen concentration in syngas may limit the rate of hydrogen diffusion into anode since it depends on differences in hydrogen concentration in syngas and anode material. Thus, 35.4% of hydrogen has no sufficient time for oxidation and flows to WHB. Since WHB capacity is intended for balance-of-plant needs only, the combustion of a significant quantity of unreacted hydrogen in it may result in rather high temperature values of outgoing gases.