Digester Gas Upgrading to Synthetic Natural Gas in Solid Oxide

Article pubs.acs.org/EF

Digester Gas Upgrading to Synthetic Natural Gas in Solid Oxide Electrolysis Cells Guido Lorenzi,* Andrea Lanzini, and Massimo Santarelli Department of Energy (DENERG), Politecnico di Torino, Corso Duca degli Abruzzi 24, 10129 Torino, Italy S Supporting Information *

ABSTRACT: This work focuses on the process design and performance of an innovative plant for digester gas upgrading to synthetic natural gas (SNG). The differences and advantages over traditional upgrading processes are discussed. The main strength of digester gas upgrading via high-temperature electrolysis concerns its higher synthetic natural gas productivity for a given raw digester gas feed. Electrolysis is performed through a solid oxide electrolysis cell (SOEC) system, which is fed with demineralized water and purified digester gas (made up of methane and carbon dioxide). Surplus electricity from intermittent renewable energy sources is used to supply the energy required for the SOEC stacks. The resulting methane-rich syngas is reacted in a series of methanators to yield a high CH4 content output stream. The steam reforming reaction is promoted by means of a nickel catalyst in the cathode (fuel) electrode, which reduces the methane fraction: hence, sulfur, which is present in several types of digester gas (e.g., from sewage or landfills) in the form of hydrogen sulfide, has been identified as a possible inhibitor for this reaction. However, it is also well-known that sulfur is responsible for the deterioration of the electrochemical performance of a stack. Therefore, its effect on the system has been modeled for different thermodynamic conditions. This study analyses the electrochemical and energy performance of the integrated process through which all the carbon contained in digester gas is converted/upgraded to methane-rich gas. The electrochemical dissociation of the CO2 contained in the digester gas to CH4 (with the addition of external demineralized water) is one way of cleverly exploiting the carbon content in digester gas when poor quality or limited biological substrates are available for anaerobic digestion. Finally, a comparison with other commercial digester gas upgrading techniques has been made.

1. INTRODUCTION Over the past few years, digester gas production through anaerobic digestion (AD) has become a popular way of transforming waste biomass into energy products.1 This process is part of a broader technical context that has focused on the efficient exploitation of biomass. Biological substrate sources include agricultural residues, manure, sewage sludge, the organic fraction of municipal solid waste, and energy crops. Biomass can basically be divided into wet and dry types, according to the initial water content. Dry biomass (water content 90% vol.). The possibility of coupling power-to-gas with a digester gas upgrade process can be considered in this context. With this configuration, the gas can be upgraded directly, via electrolysis, to yield a syngas rich in H2 and CO, using the feeding CO2 as the carbon source. During the digester gas to synthetic natural gas upgrade process through an SOEC, the steam reforming of methane that occurs within the stack, due to the presence of a nickelbased catalyst in the fuel electrodes, is critical. An Ni catalyst, in fact, promotes not only electrochemical reactions but also the catalytic decomposition of methane through steam reforming (CH4 + H2O → 3H2 + CO). Since this reaction is endothermic, it helps avoid methane consumption inside the SOEC in order to prevent additional heat from being supplied to the stack. This heat is generally released through the irreversible operation of the SOEC (i.e., polarization effects). Consequently, the electricity supplied to the SOEC is partially wasted as irreversible heat, which contributes to the maintenance of a constant operating temperature. Some possible solutions involve the adoption of electrodes that do not include nickel21 or the exploitation of sulfur chemisorption on nickel crystals. When digester gas obtained from proteinaceous feedstocks (such as the digester gas obtained through the anaerobic digestion of manure and sewage sludge) is considered, sulfide compounds are mostly found as contaminants in the form of hydrogen sulfide (H2S) and range from 0.005 to 2% vol.7 It is well-known that sulfur poisons nickel catalysts. However, it reduces the catalytic activity of the steam reforming reaction to a greater extent than the catalytic activity that promotes the electrochemical reactions.22,23 Therefore, the concentration of H2S should be a trade-off between the effects of the two competing phenomena (electrochemical and chemical reactions). In addition, the system should contain a sulfur removing device to eliminate all the traces of H2S from the SOEC outlet stream, in order to avoid poisoning of the methanator catalysts. The purpose this work has been to perform a detailed system analysis of the digester gas upgrading process proposed and patented by Hansen19 in one of its possible configurations, by assessing the impact of the sulfur content on the energy performance of the system. A kinetic analysis of the steam reforming reaction has also been carried out in order to take into account the presence of sulfur on the Ni-based cathode. Finally, the innovative solution has been compared with other digester gas upgrading techniques, in terms of energy consumption and overall synthetic natural gas (SNG) production relative to the amount of the initial available digester gas. An upgrade plant should ideally be located in a wastewater treatment plant or landfill, where digester gas from urban sewage and industrial waste and municipal solid wastes can be easily obtained.

Figure 1. SOEC scheme and reactions.

few years, because of its higher water electrolysis efficiency than PEM cells.24 This has led the authors to investigate its use as an initial step for carbon conversion into synthetic natural gas, using surplus electricity. The use of digester gas as the feed of the SOEC raises the problem of methane steam reforming inside the stack cathode electrodes. This phenomenon does not occur when SOECs are fed with water (simple electrolysis) or water and carbon dioxide (coelectrolysis), since the reactants are CH4-free. Kinetically, Ni favors the steam reforming reaction under SOEC stack operating conditions. Hence, the methane content in the inlet gas would drop increasingly for lower stack operating pressures. In order to avoid methane depletion in the digester gas, one possible solution is to cofeed a controlled amount of sulfur (in the form of H2S) to the inlet stream.19 In fact, it is known that sulfur covers the active sites of a catalyst and reduces the rate of the reactions that occur on the surface.25 In the present case, interest has been concentrated on the inhibition of methane steam reforming. Therefore, a kinetic study on the latter reaction was conducted (the results are provided in the Supporting Information). Three models were considered for steam reforming,26−28 and it was found that they produced indistinguishable results at the level of accuracy required for the comparative study of the upgrading processes made in Section 9. The passivation effect of sulfur on the reaction rate can be taken into account using a correction coefficient (1 − θs)3, as suggested by Rostrup-Nielsen in ref 25, which needs to be multiplied by the reaction rates of both the steam reforming and the water gas shift reactions. In ref 22, it was stated that the catalyst sites of heterogeneous catalysis (which favor the steam reforming reaction) are separate from those which promote electrocatalysis. Therefore, the presence of sulfur inhibits the decomposition of methane more than the electrochemical conversion of water and carbon dioxide. It seems that sulfur prefers to stick to the first type. Nevertheless, the electrochemical process is, to some extent, also penalized by the reduction in the number of the active sites, whose main effect is an increase in the ASR of the cells, as expressed in (1) (see the Supporting Information for further details):

2. METHANE CONVERSION IN SOECS Solid oxide electrolysis cells are devices that perform hightemperature electrolysis through a dense ion-conductive electrolyte

ASR pd = 1642

Pdrop × VTN jsf

(1) DOI: 10.1021/ef5023779 Energy Fuels 2015, 29, 1641−1652

Article

Energy & Fuels

in the Supporting Information), and the methanation process involves three adiabatic methanators with a commercially mixed Ni non-Ni catalyst, thanks to which it is possible to reach low temperatures at the inlets of the methanators. The modeled plant is fed with liquid water and with a gaseous mixture of methane and carbon dioxide, which represents digester gas. The latter is assumed to have an average molar composition of 60% CH4 and 40% CO2 on a vol. basis, with a controlled H2S partial pressure. All the reactants are assumed to be available at atmospheric pressure and 20 °C. The feed streams have to be conditioned to be suitable for SOEC operation. The SOEC inlet temperature is set to 800 °C, while the electrolytic reactions are considered to take place at 850 °C. The temperature increase from 800 to 850 °C is accomplished inside the stack by exploiting additional heat generation due to polarization effects. The configuration in which the surplus heat is used to heat up the reactants is named the “extended thermoneutral” configuration. In other words, the irreversible phenomena that occur within the SOEC supply all the heat required for the electrolytic reactions (as would happen in the case of the classical thermoneutral condition) and for the final heating of the reactants. The electrolysis section has been modeled as a 3-fold process with a stoichiometric reactor placed between two plug flow reactors, as shown in Figure 3.29 The SOEC is modeled as isothermal, with fixed reactant utilization (RU). The stack is run in an extended thermoneutral condition with a fixed current. The mass flow of reactants can be determined using the direct correlation to the Faradaic input current (2).

Figure 2. ASR for the case of sulfur coverage for different amounts of sulfur in the SOEC feed stream under thermoneutral conditions (VTN = 1.35 V; T = 850 °C). The red markers are for θs = 0.95.

The trend of the ASR, for the case of a performance drop due to an increasing sulfur content in the SOEC inlet mixture, is shown in Figure 2.

3. PROCESS LAYOUT The digester gas upgrading system comprises three main sections: pressurization and heating of the reactants, electrolysis, and methanation (Figure 3). The assumptions on the mechanical and electrical performance coefficients of the pressure changers are provided in the Supporting Information. The electrolysis process was assumed to take place in a commercial cathode supported SOEC (the details of the cell are described

Ifarad = n in ̇ × RR × RU × nel × F

(2)

Figure 3. Scheme of the plant with indication of the three main sections. 1643

DOI: 10.1021/ef5023779 Energy Fuels 2015, 29, 1641−1652

Article

Energy & Fuels

stages. The feed stream of the first methanator is subjected to a composition constraint. The optimal composition to promote the methanation process is ([H2] − [CO2])/([CO] + [CO2]) = 3 (expressed in terms of partial pressures).20 The pressure at the first reactor inlet was set at 33 bar. The compression is twostaged and intercooled. The pressure increase at each stage is set equal to the square root of the total increase, in order to obtain the maximum efficiency.32 The choice of interrefrigerated compression was made to prevent the temperature from rising too much, as this could compromise the integrity of the machine. When the molar amount of CH4 is sufficiently high, water is separated by means of condensation. The water can be used cleverly by injecting it upstream from the SOEC stack, before pumping, in order to minimize the consumption of the demineralized water. Finally, a gas rich in methane is obtained, whose quality can be adjusted in order to meet the grid specifications. The complete plant has been modeled in ASPEN Plus,33 a process engineering simulator that uses a lumped 0-D model. The Peng−Robinson34 property model was used in order to guarantee good accuracy of the hydrocarbon chemical reactions.

The knowledge of Δh̅ and TΔs ̅ across the SOEC, in extended thermoneutral conditions, allows the operating voltage (3) and the current density to be obtained, once the ASR value has been established (4). Vop =

φ + Δh ̅ −Φ + ΔḢ −Φ + ΔḢ = ̅ = Ifarad nreact,in nel × F × RU × nel × F ̇ (3)

ΔĠ ⎞ 1 ⎛ ΔḢ T ΔS ̇ − ⎜ ⎟= ASR ⎝ Ifarad Ifarad ⎠ ASR × Ifarad ̇ T ΔS TΔs ̅ = = ASR × nṘ × nel × F ASR × nel × F

jTN =

(4)

Therefore, the size of the plant (and thus the syngas production) depends on the total Faradaic current absorbed by the SOEC module, which has here been set to 10 000 A. Consequently, the main output of the design process is the value of the active area of the cells. This represents the size of the installation and is closely related to the investment costs of the whole plant. As can be seen in Figure 3, part of the cathodic outlet flow is recirculated in order to ensure a hydrogen molar fraction equal to 10% and thus to avoid catalyst degradation issues due to Ni reoxidation. The present SOEC model does not include any external sweep gas to feed the air electrode (anode). Instead, an internal partial hot recirculation of the produced oxygen stream is assumed to act like a sweep gas in the air electrode. Once the flow has been electrolyzed, it is passed through a desulfurization unit in order to remove all the traces of sulfur, thus preventing the catalyst in the downstream methanation section from being poisoned. The only effect of the desulfurization process on the thermodynamic conditions of the stream is a pressure loss. Generally, the technological solutions for these kinds of components make use of zinc oxide or zinc-based sorbent beds. These are commonly used for the removal of H2S, due to their high equilibrium constant for sulfidation and to the thermal stability of the sorbents. The following reaction occurs: ZnO + H 2S → ZnS + H 2O

4. PLANT DYNAMIC BEHAVIOR The proposed concept is more suitable for continuous operation, due to the limited flexibility of the catalytic methanation unit and to the possible accelerated degradation of the SOEC from thermo-mechanical stress generated by a supply that can vary to a great extent. Although renewable energy sources are the ideal power supply, they are intermittent and difficult to predict. A possible alternative would be to use nuclear power to cover the demand. A second solution could foresee the presence of a syngas storage unit. According to this scenario, the SOEC operates in “supply following” mode and the generated syngas is compressed and temporarily stored in tanks in order to allow continuous operation of the downstream methanation synthesis section. The proposed integrated SOEC methanation digester gas upgrade plant concept would thus be able to run under one of the following configurations: (1) intermittent operation of the overall plant (electrolyzer and methanation synthesis unit) to act as a flexible electricity consumer coupled with the intermittent production of electricity from renewable energy sources (wind and solar); (2) continuous operation of the overall plant connected to base-load nuclear power plant production; (3) intermittent operation of only the electrolyzer, with gas storage available for the syngas from the SOEC to allow continuous operation of the methanation section. The most challenging option is certainly the first one, as dynamic operation of the methanation reactors would be required. The use of gas storage would obviate this problem but would require additional energy and investment costs to pressurize the gas and store it. The dynamic capability of the SOEC has already been demonstrated at the single-cell and stack levels35,36 but not at a system level. However, modeling studies have shown that dynamic operation and temperature control of the stack is feasible in an SOEC through the regulation of the sweep gas flow rate.37,38 The sweep gas flow rate in the oxygen electrode thus becomes the control variable that has to be used to smooth the temperature gradients across the solid stack structure. In the case of the temporary absence of surplus renewable electricity, the SOEC could even be maintained at open circuit

(5)

where ZnO and ZnS are the zinc oxide sorbent and the zinc sulfide, respectively. Additionally, ZnO is a nonpyrophoric material that can provide a high level of sulfur removal. The operating temperature is generally in the range of between 200 and 500 °C and is always below 700 °C to prevent the catalyst from reducing to elemental zinc, which in turn, would lead to a subsequent volatilization and loss of sorbent.30 The latter condition is satisfied in the present configuration because the desulfurator is positioned downstream from a cooler that decreases the temperature to 220 °C. The methanation process consists of three steps, each of which involves an adiabatic Gibbs type reactor and a cooler. CO and H2 are converted into methane through methanation and water gas shift reactions, respectively. The reaction heat produces a considerable increase in temperature, and recycling is therefore used to control the temperature rise in the first methanation reactor. The inlet reactor temperature is set at 220 °C, which is a very low value, compared to the usual methanation temperatures (generally around 300 °C), but it is achievable if a combination of Ni and non-Ni catalysts is used, as reported in ref 31. The spilled quantity is calibrated in order to limit the outlet temperature of the first methanator to 600 °C. The gas then enters the subsequent methanation 1644


Article

Energy & Fuels Table 1. Composition Table: Configurations A and Ba 1 bar θs = 0% (Case A)

H2 [%] H2O [%] CO2 [%] CH4 [%] CO [%] temperature [°C] pressure [bar] molar flow [mol/s] mass flow [kg/h] a

pre SOEC (1)

after SOEC (2)

0 72.4 11.0 16.6 0 800 2.7 0.086 6.32

65.3 13.9 3.3 0.1 17.4 220 1 0.114 4.15

θs = 95% (Case B)

after methanation after water removal (3) (4) 1.1 63.7 0.3 35.0 0 35 30 0.067 4.15

3.0 0.2 0.7 96.1 0 35 30 0.024 1.39

pre SOEC (1)

after SOEC (2)

0 64.8 14.1 21.1 0 800 3.7 0.088 6.77

50.4 16.0 3.5 18.0 12.2 220 1 0.093 3.90


2.0 0.2 0.5 97.3 0 35 30 0.032 1.82

The stream numbers refer to Figure 3.

Table 2. Composition Table (Molar Base): Configurations C and Da 30 bar θs = 0% (Case C)

H2 [%] H2O [%] CO2 [%] CH4 [%] CO [%] temperature [°C] pressure [bar] molar flow [mol/s] mass flow [kg/h] a

pre SOEC (1)

after SOEC (2)

0 65.1 14.0 20.9 0 800 32.7 0.088 6.75

51.1 15.9 3.5 17.2 12.4 220 30 0.093 3.91

θs = 95% (Case D)


2.0 0.2 0.5 97.3 0 35 30 0.032 1.80

pre SOEC (1)

after SOEC (2)

0 63.7 14.5 21.8 0 800 33.7 0.089 6.82

47.7 16.3 3.5 21.3 11.3 220 30 0.089 3.85


1.8 0.2 0.5 97.5 0 35 30 0.033 1.88

The stream numbers refer to Figure 3.

Table 3. Operating Parameters for All the Casesa A Vid [V] Vtn [V] ASRsf [Ω cm2] ASRpd [Ω cm2] jTN [A/cm2] performance drop [%] power density [W/cm2] active area [cm2] a

0.60 1.30 0.50 1.4 1.82 7142

B

C

0.89 1.34 0.50 0.18 0.66 (0.9) 18.91 0.88 15088 (11084)

0.88 1.34 0.50 0.93 1.24 10803

D 0.93 1.35 0.50 0.27 0.53 (0.83) 18.91 0.72 18768 (12111)

The values in brackets refer to the parameters that neglect the performance drop on the electrocatalytic activity of the cells.

voltage (OCV), i.e., in hot stand-by conditions, for several hours, without any significant thermal losses, as the stack units are surrounded by a thick insulation layer.39

reforming to H2 and CO, as well as a positive effect on the overall efficiency. Since the inlet pressure in the first methanator was always set to 33 bar, a syngas compressor was required downstream from the SOEC reactor to feed the methanator section. It has been stated25 that sulfur coverage values below 0.7 do not lead to any considerable differences from sulfur free conditions. This result has been confirmed in the present simulations. A value of 0.95 was used to explore the behavior of the system in rather extreme conditions. The reactant utilization (RU) of the SOEC section was set to 70% in order to be distinguished from the conditions in which diffusion limitations occur. The molar compositions and flows along the gas line are reported in Tables 1 and 2. Some operating parameters of the plant are provided in Tables 3 and 4, while the energy flows and the plant efficiency (6) are shown in Table 5.

5. RESULTS The presence of sulfur has a reduced impact on the energy efficiency of a plant when the SOEC is pressurized. Four different operating conditions were simulated, as reported in Tables 1 and 2. The operating temperature of the stack is set to 850 °C.29 It has already been proved40 that pressurizing the reactants of the SOEC (especially liquid water), rather than compressing the products of electrolysis, implies energy savings that increase with pressure. On the other hand, the steam reforming reaction is favored at low pressures; therefore, operating at high pressure levels has positive consequences on the inhibition of the CH4 1645


Article

Energy & Fuels

same conclusions can be drawn for the pressurized configurations (C and D), even though the difference, due to the presence of H2S, is less evident. In fact, in this case, the percentage molar difference in the carbon inlet content is slightly less than 5% (referring to case D; see Table 2), while the electric input does not vary. However, in both cases (atmospheric and pressurized), sulfur-free operation entails lower efficiencies. The system is affected for atmospheric operation and in pressurized conditions by the performance drop, with an increase in the overall ASR. Hence, a lower current density is sufficient to maintain thermoneutral operation. Nevertheless, this fact implies an increase in the active area that is necessary to produce the constant Faradaic current value (10 000 A in the present case). The result is a reduction in the power density and an increase in the number of the installed cells, which in turn leads to the same electrochemical performance. Moreover, the benefits obtained from sulfur passivation are much more evident for lower pressures, that is, when the steam reforming reaction is thermodynamically favored.

Table 4. Flows and Recirculation Rates for All the Cases RR after mixing [%] RR before mixing [%] inlet molar flow [kmol/h] condensed water [kg/h] total inlet mass flow [kg/h] recirculation rate stack [%] recirculation rate I stage methanation [%]

A

B

C

D

73 83 0.308 2.77 6.33 12 71

67 79 0.319 2.08 6.77 19 48

67 79 0.318 2.11 6.76 19 49

66 78 0.320 1.97 6.82 21 41

Table 5. Plant Energy Results total AC electric need [kW] pumping power [kW] stack power [kW] upgraded gas (SNG) [kg/h] methane inlet [kg/h] LHVSNG [MJ/kg] SNG chemical energy flow [kW] ηpl [%]

ηpl =

A

B

C

D

15.465 2.232 13.233 1.390 0.816 49.2 18.975 71

15.892 2.246 13.646 1.821 1.078 49.4 24.988 81

14.240 0.585 13.655 1.806 1.069 49.4 24.786 86

14.328 0.599 13.729 1.885 1.116 49.4 25.881 87

LHVSNG × ṁ SNGout LHVCH4ṁ CH4in + Wel,pl

6. SNG INJECTION INTO THE GAS GRID The product gas should meet all the specifications necessary for grid injection. Hence, it is worth comparing the quality of the SNG with the present standards. The threshold values were taken from the grid code which is currently in force in Italy (Wobbe index: 47.31−52.33 MJ/Sm3; gas gravity: 0.5548−0.8; higher heating value: 34.95−45.28 MJ/Sm3).41 Since the upgraded gas has been shown to fulfill the minimum requirements in terms of Wobbe index, but not in terms of specific gravity, it is necessary to condition the product gas with a higher molar weight substance. Possible solutions include blending the SNG with nitrogen or propane. (The molar weight of propane is higher than that of nitrogen; hence, a smaller amount is sufficient to reach the specific gravity threshold. In addition, nitrogen would reduce the heating value of the mixture, while propane would increase it.) In this work, a small amount of propane was mixed with the SNG in order to comply with the grid specifications (although propone might not be a feasible or economic blending solution in each and every location). In the present case, the quantity that proved sufficient to meet the grid specifications was 4.0 × 10−4 mol/s, which is equivalent to 1.76 × 10−2 g/s. A comparison between the quality of the outlet stream before and after the injection of propane is shown in Tables 6 and 7. The efficiency has not

(6)

As can be seen from Table 5, pressurizing the SOEC entails a better performance in SNG production for kinetic reasons. Moreover, switching from sulfur-free to high coverage operation implies an increase in energy efficiency, which is more accentuated as the pressure rises. The overall electric power demand appears to vary to a great extent between atmospheric and pressurized conditions. This is attributable the lower electric consumption that is obtained from the pressurization of the liquid water, before it enters the SOEC, instead of the compression of the steam entering the series of methanators. Moreover, sulfur passivation entails a slight increase in the electricity needs of the plant; nonetheless, the benefits, in terms of productivity, outbalance the extra energy costs (see Table 5). Thermal integration of the plant has been calculated for all the four different cases, and thermal self-sufficiency has been proved for each simulated condition (see Section 7). From a comparison of the two configurations at atmospheric pressure, it is possible to notice that the overall electric input is roughly the same. Nevertheless, the main difference, especially under atmospheric operation, pertains to the productivity of SNG, in terms of flow rate and, in turn, of chemical energy flow. This is due to the fact that the percent difference in carbon molar content entering the first methanator between configurations A and B is −24.1% (referring to case B; see Table 1). Another parameter that appears to vary considerably is the recirculated fraction at the first methanation stage. It is much higher in configuration A, because the methanation reaction is faster and releases more heat, due to the absence of methane in the feed stream. For this reason, a higher outlet stream fraction is necessary to restrain the outlet temperature to the fixed value of 600 °C. The composition of the methanation section outlet stream does not vary to any great extent, but the condensed water flow is much higher for cases with no sulfur passivation. This explains why the higher and the lower heating values appear approximately unchanged in the two configurations. The

Table 6. Specifications for Grid Injection and Global Efficiency before Conditioning of the Upgraded Gas conditions without propane injection Wobbe index [MJ/Sm3] gas gravity higher heating value [MJ/Sm3] global efficiency [%]

A

B

C

D

49.43 0.5456 36.51 71

49.81 0.5477 36.86 81

49.80 0.5476 36.85 86

49.86 0.5480 36.91 87

been recalculated for the case of propane addition, since this expedient is introduced to increase the density of the mixture and does not affect the performance of the upstream process. The only contribution that should be included is the pumping consumption, but it has proved to be negligible compared to the electric power supplied to the SOEC stack (the ratio is lower than 10−3). 1646


Article

Energy & Fuels Table 7. Specifications for Grid Injection and Global Efficiency after Conditioning of the Upgraded Gas conditions without propane injection Wobbe index [MJ/Sm3] gas gravity higher heating value [MJ/Sm3]

A

B

C

D

49.96 0.5613 37.43

50.42 0.5598 37.72

50.28 0.5598 37.62

50.32 0.5597 37.6

7. THERMAL INTEGRATION OF THE PLANT A thermal integration has been performed for all four cases through the pinch analysis methodology,42 and the results show that the thermal requirements of the plant are matched internally for all four configurations. The calculations performed for the atmospheric operation for the case of a sulfur coverage equal to 95% (case B) are given in this section as an example of the applied methodology. A list of the heat exchangers (see Figure 4 for their definition), divided into hot and cold bodies, is given in Table 8, while the composite curves are reported in Figure 5. In this plot, the curve of the cold fluids was shifted to the right by a quantity equal to the total external thermal need (−2153 W).

Figure 5. Composite curves for cold and hot flows at 1 bar.

The pinch point temperature difference (ΔTPP) was set to 20 °C, but the entire system was placed below the pinch point, which means that there was no need to supply heat to the system but only to cool it down. In practical terms, this situation made it possible to avoid inserting additional sources of heat, while the surplus heat could be used in other low temperature applications. However, the excess heat could be released to the environment. Since the entire system is below

Figure 4. Scheme of the plant with indication of the stream names used for the pinch analysis calculations.

Table 8. List of the Flows Involved in the Thermal Integration of the Plant (Configuration B)a #

name

inlet temperature (TIN) [°C]

outlet temperature (TOUT) [°C]

heat demand [W]

G × CP [W/K]

T*IN [°C]

T*OUT [°C]

1 2 3 4 5 6 7 8 9 10

ECO (C) EVA (C) SURR (C) biogas heater (C) mix heater (H) oxygen cooler (H) cooler 1 (H) cooler 2 (H) cooler 3 (H) cooler 4 (H)

20 102 107 36 400 850 850 600 406 265

102 107 400 400 800 30 220 220 220 35

383 2429 601 514 1670 −669 −2229 −2314 −544 −1995

4.7 485.7 2.0 1.4 4.2 0.8 3.5 6.1 2.9 8.7

30 112 117 46 410 840 840 590 396 255

112 117 410 410 810 20 210 210 210 25

−2153

total external thermal need [W]

a G [kg s−1]: mass flow. CP [J kg−1 K−1]: specific heat capacity. TIN * [K]: corrected inlet temperature. For cold fluids (1−5), TIN * = TIN +ΔTPP, while for hot fluids (6−10) T*IN = TIN − ΔTPP. T*OUT [K]: corrected outlet temperature. For cold fluids (1−5), T*OUT= TOUT + ΔTPP,, while for hot fluids * = TIN − ΔTPP. (6−10) TIN

1647


Article

Energy & Fuels the pinch point, the minimum temperature difference between the two fluids is 50 °C and is between the two top edges of the composite curves (Figure 5).

8. SENSITIVITY ANALYSIS A sensitivity analysis has been performed in order to understand how different parameters affect the methane conversion throughout the stack. In particular, the analysis was carried out by varying the pressure and the temperature for different sulfur coverage values. Figure 6 shows the effects of pressure on the conversion of methane.

Figure 8. Methane consumption profile at 10 bar for different SOEC operating temperatures.

9. DIGESTER GAS UPGRADING TECHNIQUES Once the performance of the electrolysis system had been calculated, it was considered interesting to compare it with other conventional techniques currently used to improve the quality of digester gas. Some of these routes are examined in this work: water scrubbing (WS), pressure swing adsorption (PSA), and membrane separation (MS), whose energy consumption is given in ref 43. The efficiency and the specific productivity (SP) of each process can be defined as LHVSNG × ṁ SNG,out ηsep = (LHVCH4 × ṁ CH4in + Wel + Eextra) (7)

Figure 6. Methane consumption profile at 850 °C for different SOEC operating pressures.

It is possible to notice that, as long as the S-coverage is below 70%, the quantity of the methane converted through the cells does not change compared to the sulfur free operation. CH4 is almost unreacted for an almost complete sulfur coverage (θS → 1), due to the saturation of the active catalyst sites. The limited influence of the low coverage values on the steam reforming reaction rate has been found in other works.25 A pressure increase leads to significant benefits on the methane content outlet of the SOEC: as expected, the steam reforming activity diminishes for growing pressure. CH4 consumption has been analyzed for the case of temperature variations (Figures 7 and 8). It can be seen that, for each pressure, the

SP = YCH4(1 − LCH4)

(8)

where Eextra is a term that includes all the energy costs associated with the conditioning of the SNG to the grid quality (e.g., energy content of the possible admixtures of hydrocarbons and compression to grid pressure). In particular, it was assumed that PSA produces an SNG that needs to be mixed with propane in order to be suitable for dispatchment. It has been estimated that propane consumption is 0.013 [Nm3/NmRG3], while the electric energy requirement is 0.053 [MJ/NmRG3].43 The efficiency and the specific productivity of each technology are given in Table 9, assuming that the outlet SNG has a Table 9. Efficiencies and Specific Productivity Values for Three Different Digester Gas Upgrading Methodologies PSA WS MS

efficiency (ηsep) [%] specific productivity [NmSNG3/NmRG3] efficiency (ηsep) [%] specific productivity [NmSNG3/NmRG3] efficiency (ηsep) [%] specific productivity [NmSNG3/NmRG3]

92.9 0.576 94.2 0.591 86.6 0.564

methane content of 0.97 vol. (The productivity values refer to a molar composition of biogas equal to 60% CH4 and 40% CO2 vol. The lower heating value of methane is assumed to be equal to 35 MJ/Nm3 for the efficiency calculations.) The electrochemical-aided upgrade process that has been analyzed in this work is another SNG production option. This can be either a one-step process (Figure 3), i.e., the digester gas is fed directly to the SOEC, or a two-step process, i.e., first CO2 is removed from the digester gas and then a mixed H2O/CO2 stream is fed to the integrated SOEC-methanation plant

Figure 7. Methane consumption profile at 1 bar for different SOEC operating temperatures.

decrease in temperature produces a beneficial effect on the methane conservation in the stack, and this is more evident for higher pressures. 1648


Article

Energy & Fuels

Figure 9. Scheme of the upgrading process from digester gas to SNG through an initial CO2 separation and subsequent coelectrolysis and methanation.

Table 10. Summary of the Efficiencies and Specific Productivities for All the Digester Gas Upgrading Routesa process

action(s)

efficiency [%]

specific productivity [NmSNG3/NmRG3]

pressure swing adsorption (PSA) water scrubbing (WS) membrane separation (MS) PSA + H2O/CO2 CoE (two steps) WS + H2O/CO2 CoE (two steps) MS + H2O/CO2 CoE (two steps) digester gas CoE (one step), atmospheric digester gas CoE (one step), pressurized

CO2 separation CO2 separation CO2 separation CO2 separation + H2O/CO2 coelectrolysis CO2 separation + H2O/CO2 coelectrolysis CO2 separation + H2O/CO2 coelectrolysis H2O/CO2 coelectrolysis H2O/CO2 coelectrolysis

93.0 94.2 86.6 88.0 90.7 84.4 71.0−81.0 86.0−87.0

0.576 0.591 0.564 0.976 0.991 0.964 1.000 1.000

a

The efficiencies are calculated with (7) and (9). The specific productivity values for the two-step processes are obtained by adding the molar fraction of CO2, namely, 0.4, to the separation process values (since no carbon losses are considered). The range of efficiencies for the one-step biogas upgrading via electrolysis refers to Table 5, in which the different operating conditions are specified.

(whose efficiency, ηCoE, is 81.4%). The H2O/CO2 integrated SOEC-methanation plant was studied in depth in ref 48 on the basis of a previous study.40 A scheme of the two-step process is depicted in Figure 9. This process is considered to cause no carbon loss. Since the only external effluent from the plant, other than SNG, is condensed water and given the sealed structure of the SOEC, no carbon slip from the plant was taken into consideration. All the carbon present as carbon dioxide is eventually transformed into methane throughout the SOEC and the methanation section. Moreover, CO is almost absent after water removal (see Tables 1 and 2). Hence, the conversion of CO2 into CO, due to the water gas shift reaction, is negligible in the methanation process. The two-step process efficiency can therefore be defined as

(9)

reductions in the energy efficiency values that can be observed. The atmospheric one-step process yields a lower energy efficiency than the two-step routes but also leads to reduced plant complexity, due to the absence of an upstream CO2 separation unit. In fact, all the current commercial techniques have shown problems in using CO2, which has to be released into the environment or stored in appropriate sites, usually of a geological nature. In the former case, this involves raising the greenhouse emissions of the plants and compromising the carbon closed cycle,44 which is one of the main reasons for a rational use of biomass. In the latter scenario, CO2 should be cooled and/or pressurized before being sent to storage, with a consequent higher cost and energy consumption. Another possibility is to install a plant in which the carbon is recovered and turned into other useful chemicals through various processes (dry reforming of methane,45 synthesis of cyclic carbonate from CO2 and epoxide,46 CO2 hydrogenation to methanol47).

It is clear, from (9), that the resulting efficiency is affected by the CO2 separation efficiency, and ηts can assume different values that depend on the separation process to which it is coupled. A comparison of the different available solutions, in terms of energy efficiencies and specific productivities, is presented in Table 10. The efficiencies are clearly higher for the conventional methodologies. However, these methodologies do not account for the CO2 conversion into extra fuel. It is worth noting that the CO2 separation routes are characterized by much lower productivity values, due to the absence of the C-recycling of the separated CO2 stream. Productivity increases considerably in the two-step processes, in spite of the slight

10. DISCUSSION A system analysis of digester gas upgrading plant configurations, despite all the uncertainties connected to a preliminary sizing of the components, has shown a marked influence of pressure on the energy performance. The effect of sulfur passivation leads to a benefit in terms of efficiency, with an increase in the overall plant energy performance of 10 percentage points, as shown in Table 5. However, the presence of hydrogen sulfide causes a decrease in power density (i.e., a higher installed stack capacity is required to upgrade the same amount of digester gas), and a trade-off should therefore be considered from an economic point of view.

ηts =

(LHVSNG × ṁ SNG,out)separation + (LHVSNG × ṁ SNG,out )electrolysis (LHVCH4 × mCH4in + Wel + K )separation +

(LHVSNG × ṁ SNG,out)electrolysis ηCoE

1649


Article

Energy & Fuels

separated CO2, which is converted directly into fuel, thus avoiding the need of an upstream dedicated separation plant which would lead to increased plant complexity. Future work should deal with a techno-economic assessment of the proposed plant configurations and compare them with conventional digester gas upgrade solutions.

Operation with a pressurized SOEC leads to a decrease in the benefits that can be derived from sulfur passivation and to a very limited efficiency improvement between operation with sulfur (D) and without (C). In this case, the inhibition of the steam reforming activity, via sulfur passivation, does not seem to be a good solution because of the considerable increase in the active area, due to the performance drop side effect. The analysis of the four different configurations (cases A to D) has shown that the pressurization of the electrolysis module, without resorting to the use of sulfur as a passivator of the heterogeneous catalysis reactions, is beneficial in terms of energy consumption (Table 5). As far as the size of the SOEC stack is concerned, case A results in the lowest active area value, but it is much less efficient in terms of energy balance. This means that the most convenient process needs be assessed through a thermo-economic analysis. It is necessary to weigh the contributions against the total costs, due to operation (connected to energy consumption), and the investment costs (related to the size of the plant). One way of limiting the increase in the active area would be to use ceramic electrodes,19 which have a very limited reforming activity. (Possible ceramic materials are LSCM, CeO2, titanates, or combinations thereof19). This solution prevents a dramatic increase in the active area. However, ceramic electrodes are still not as fully developed as Ni-YSZ cermets for SOFC/SOEC applications. The overall SOEC-based electrochemical upgrade process is completely self-sufficient, from a thermal point of view, due to an optimal thermal integration between the methanation section (heat source) and the gas preheating section (which includes the H2O vaporization phase). The produced SNG is compatible with the acceptable Wobbe index range but fails to meet the specifications in terms of gravimetric density. It is therefore suggested that the produced SNG needs to be conditioned with propane to yield a heavier synthetic gas which would comply with the grid requirements.

■

ASSOCIATED CONTENT

* Supporting Information S

Nomenclature; SOEC thermodynamics and operation in coelectrolysis; electrolysis thermodynamic efficiency; steam reforming kinetic models; effects of kinetic formulations on methane conversion (optimal ratio for methanation); sulfur effects (sulfur performance drop); pinch analysis additional information; additional information about Aspen Plus simulations (reactors’ description; catalyst quantity; pressure losses; machines efficiencies). This material is available free of charge via the Internet at http://pubs.acs.org/.

■

AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

■

ACKNOWLEDGMENTS The authors would like to express their gratitude to Mr. Valerio Novaresio for his support with the kinetic simulations in the Cantera environment.

■

NOMENCLATURE

List of Acronyms

AC = alternate current CoE = Coelectrolysis RG = raw gas SNG = synthetic (or substitute) natural gas

11. CONCLUSIONS An SOEC-based upgrade of digester gas has been compared with other methods for the upgrading of digester gas to synthetic natural gas. The key advantages of electrochemical-aided upgrading are (1) enhanced synthetic natural gas yields, due to the conversion of the carbon contained in the CO2 to synthetic natural gas; (2) direct upgrading of digester gas to synthetic natural gas with CO2 management included (i.e., no venting to the atmosphere). The electrochemical upgrading process can be carried out considering two different configurations: (1) one-step process with direct feeding of the digester gas to the SOEC reactor and further upgrading of the resulting syngas to SNG in a methanation unit; (2) two-step process with upstream CO2 removal from the digester gas and subsequent coelectrolysis to produce a syngas suitable for further upgrading to SNG. The role of sulfur passivation on high-temperature electrolysis has been reviewed and modeled to identify the feasible stack operating conditions for syngas production. The electrochemical upgrade process (one-step) has been shown to be penalized by about 3 percentage points, in its best plant configuration, compared to the most efficient two-step methodology (Table 10). However, both in the case of pressurized sulfur free operation and in the case of atmospheric sulfur passivated operation, the one-step process can compete with the others in terms of energy efficiency. In both cases, it appears that the new approach simplifies the management of the

List of Subscripts and Apices

id = ideal op = operating pd = performance drop sc = sulfur coverage sf = sulfur-free TN = thermoneutral List of Parameters

ΔḢ = enthalpy flow difference, kW ΔĠ = free Gibbs energy flow difference, kW ΔṠ = entropy flow difference, kJ·mol−1 ΔTPP = pinch-point temperature difference, K Δh̅ = molar enthalpy difference, kJ·kmol−1 Δs ̅ = molar entropy difference, kJ·kmol−1·K−1 Φ = heat flow (>0 inlet), kW ηCoE = water and carbon dioxide coelectrolysis efficiency ηts = two-step process efficiency ηsep = CO2 separation efficiency ηpl = plant efficiency φ̅ = molar heat flow, kJ·mol−1 θs = sulfur coverage ASR = area specific resistance, Ω·cm−2 Eextra = extra energy necessary for digester gas upgrading techniques with CO2 separation, kW F = Faraday constant, 96 485 C·mol−1 Ifarad = Faradaic current, A

1650


Article

Energy & Fuels

(15) Hahn, H.; Krautkremer, B.; Hartmann, K.; Wachendorf, M. Review of concepts for a demand-driven biogas supply for flexible power generation. Renewable Sustainable Energy Rev. 2014, 29, 383− 393. (16) Ibrahim, H.; Ilinca, A.; Perron, J. Energy storage systems Characteristics and comparisons. Renewable Sustainable Energy Rev. 2008, 12, 1221−1250. (17) Koohi-Kamali, S.; Tyagi, V. V.; Rahim, N. A.; Panwar, N. L.; Mokhlis, H. Emergence of energy storage technologies as the solution for reliable operation of smart power systems: A review. Renewable Sustainable Energy Rev. 2013, 25, 135−165. (18) Zhan, Z.; et al. Syngas Production By Coelectrolysis of CO2/ H2O: The Basis for a Renewable Energy Cycle. Energy Fuels 2009, 7, 3089−3096. (19) Hansen, J. B. Process for converting biogas to a gas rich in methane. WO 2012003849 A1, 2012; https://www.google.ch/patents/ WO2012003849A1?cl=en&dq=WO2012003849&hl=it&sa=X&ei= QiqfVImZAuGpygOo9YGICw&ved=0CBwQ6AEwAA. (20) From solid fuels to substitute natural gas (SNG) using TREMP. (21) Jeonghee, K.; et al. Fabrication and characterization of allceramic solid oxide fuel cells based on composite oxide anode. J. Power Sources 2013, 241, 440−448. (22) Rostrup-Nielsen, J. R.; Hansen, J. B.; Helveg, S.; Christiansen, N.; Jannasch, A.-K. Sites for catalysis and electrochemistry in solid oxide fuel cell (SOFC) anode. Appl. Phys. A: Mater. Sci. Process. 2006, 85, 427−430. (23) Hansen, J. B. Correlating sulfur poisoning of SOFC nickel anodes by a temkin isotherm. Electrochem. Solid-State Lett. 2008, 11, B178. (24) Ferrero, D.; Lanzini, A.; Santarelli, M.; Leone, P. A comparative assessment on hydrogen production from low- and high-temperature electrolysis. Int. J. Hydrogen Energy 2013, 38, 3523−3536. (25) Rostrup-Nielsen, J. Sulfur-passivated nickel catalysts for carbonfree steam reforming of methane. J. Catal. 1984, 85, 31−43. (26) Xu, J.; Froment, G. F. Methane steam reforming, methanation and water-gas shift: I. Intrinsic kinetics. AIChE J. 1989, 35, 88−96. (27) Rostrup-Nelsen, J.; Sehested, J.; Nørskov, J. K. Hydrogen and synthesis gas by steam- and CO2 reforming. Adv. Catal. 2002, 47, 65− 139. (28) Mogensen, D. Methane steam reforming kinetics over Ni-YSZ anode materials for solid oxide fuel cells. Ph.D. Thesis, Technical University of Denmark, 2011. (29) Sun, X.; et al. Thermodynamic analysis of synthetic hydrocarbon fuel production in pressurized solid oxide electrolysis cells. Int. J. Hydrogen Energy 2012, 37, 17101−17110. (30) Garces, H. F.; et al. Low temperature H2S dry-desulfurization with zinc oxide. Microporous Mesoporous Mater. 2010, 127, 190−197. (31) Jensen, J. H.; Poulsen, J. M.; Anderson, N. U. From coal to clean energy. Nitrogen+Syngas, March−April, 2011; http://www.topsoe. com/sites/default/files/topsoe_from_coal_to_clean_energy_ nitrogen_syngas_march_april_2011.pdf. (32) Anzalone, G.; Bassignana, P.; Brafa Musicoro, G. Fondamenti di Meccanica e Macchine (Book in Italian); Hoepli: Milano, 2009. (33) Design and Optimize Chemical Processes with Aspen Plus. https://www.aspentech.com/products/aspen-plus.aspx (accessed Jan 31, 2015). (34) Peng, D.; Robinson, D. B. A new two-constant equation of state. Ind. Eng. Chem., Fundam. 1976, 15, 59−64. (35) Petipas, F.; Fu, Q.; Brisse, A.; Bouallou, C. Transient operation of a solid oxide electrolysis cell. Int. J. Hydrogen Energy 2013, 38, 2957−2964. (36) Bucheli, O., Bertoldi, M., Pofahl, S.; Montinaro, D. Development and manufacturing of SOFC-based products at SOFCpower SpA. In 4th European PEFC & H2 Forum, Lucerne, Switzerland, July 2−5, 2013. (37) Udagawa, J.; Aguiar, P.; Brandon, N. P. Hydrogen production through steam electrolysis: Model-based dynamic behaviour of a cathode-supported intermediate temperature solid oxide electrolysis cell. J. Power Sources 2008, 180, 46−55.

LCH4 = methane losses, % LHVSNG = lower heating value of the SNG, kJ·kg−1 LHVCH4 = lower heating value of methane, kJ·kg−1 Pdrop = performance drop coefficient on cell voltage, % RU = reactant utilization (ṅR/ṅreact,in) RR = reactant ratio (ṅreact,in/ṅin) SP = specific productivity of SNG per unit of raw digester gas, NmSNG3·NmRG3 −1 V = voltage, V Wel = electric power (>0 inlet), kW Wel,pl = electric power demand of one-step upgrading plant, kW YCH4 = methane molar fraction, NmSNG3·NmRG3 −1 j = current density, A·cm−2 ṁ SNGout = SNG outlet mass flow, kg·s−1 ṁ CH4in = methane inlet mass flow, kg·s−1 ṅ = molar flow, kmol·s−1 nel = number of electrons involved in the electrochemical reactions ṅin = total inlet molar flow, kmol·s−1 ṅR = effectively reacting species molar flow, kmol·s−1 ṅreact,in = reacting species inlet molar flow (ṅin·RR), kmol·s−1

■

REFERENCES

(1) Weiland, P. Biogas production: Current state and perspectives. Appl. Microbiol. Biotechnol. 2010, 85, 849−860. (2) Bridgwater, A. V.; Toft, A. J.; Brammer, J. G. A techno-economic comparison of power production by biomass fast pyrolysis with gasification and combustion. Renewable Sustainable Energy Rev. 2002, 6, 181−248. (3) Song, H.; Starfelt, F.; Daianova, L.; Yan, J. Influence of drying process on the biomass-based polygeneration system of bioethanol, power and heat. Appl. Energy 2012, 90, 32−37. (4) Holmberg, H.; Ahtila, P. Evaluation of energy efficiency in biofuel drying by means of energy and exergy analyses. Appl. Therm. Eng. 2005, 25, 3115−3128. (5) Mckendry, P. Energy production from biomass (part 2): Conversion technologies. Bioresour. Technol. 2002, 83, 47−54. (6) Basic data on biogas; Swedish Gas Technology Centre Ltd.: Sweden, 2012; http://www.energigas.se/~/media/Files/www_ energigas_se/Publikationer/Infomaterial/BasicDataonBiogas2012. ashx. (7) Ryckebosch, E.; Drouillon, M.; Vervaeren, H. Techniques for transformation of biogas to biomethane. Biomass Bioenergy 2011, 35, 1633−1645. (8) Scholz, M.; Melin, T.; Wessling, M. Transforming biogas into biomethane using membrane technology. Renewable Sustainable Energy Rev. 2013, 17, 199−212. (9) Makaruk, A.; Miltner, M.; Harasek, M. Membrane biogas upgrading processes for the production of natural gas substitute. Sep. Purif. Technol. 2010, 74, 83−92. (10) Tippayawong, N.; Thanompongchart, P. Biogas quality upgrade by simultaneous removal of CO2 and H2S in a packed column reactor. Energy 2010, 35, 4531−4535. (11) Alonso-Vicario, A.; et al. Purification and upgrading of biogas by pressure swing adsorption on synthetic and natural zeolites. Microporous Mesoporous Mater. 2010, 134, 100−107. (12) Hansen, J. B.; et al. GreenSynFuels: Economical and Technological Statement Regarding Integration and Storage of Renewable Energy in the Energy Sector by Production of Green Synthetic Fuels for Utilization in Fuel Cells; EUDP project journal number: 64010-0011; Danish Technological Institute: Taastrup, 2011. (13) Giddey, S.; Badwal, S. P. S.; Kulkarni, A. Review of electrochemical ammonia production technologies and materials. Int. J. Hydrogen Energy 2013, 38, 14576−14594. (14) Socolow, R.; et al. Direct Air Capture of CO2 with Chemicals; APS: College Park, MD, 2011. 1651


Article

Energy & Fuels (38) Cai, Q.; Adjiman, C. S.; Brandon, N. P. Optimal control strategies for hydrogen production when coupling solid oxide electrolysers with intermittent renewable energies. J. Power Sources 2014, 268, 212−224. (39) Hanasaki, M.; et al. SOFC durability against standby and shutdown cycling. J. Electrochem. Soc. 2014, 161, F850−F860. (40) Hansen, J. B.; et al. Production of sustainable fuels by means of solid oxide electrolysis. ECS Trans. 2011, 35 (1), 2941−2948 http:// ecst.ecsdl.org/content/35/1/2941.abstract. (41) SNAM RETE GAS Codice di rete snam rete gas (Italian Natural Gas Network Standard); http://www.autorita.energia.it/allegati/gas/ codicerete/codicerete_srg.pdf. (42) Kemp, I. Pinch Analysis and Process Integration; 2007; http:// sutlib2.sut.ac.th/sut_contents/H112960.pdf. (43) Pertl, A.; Mostbauer, P.; Obersteiner, G. Climate balance of biogas upgrading systems. Waste Manage. 2010, 30, 92−99. (44) Jury, C.; Benetto, E.; Koster, D.; Schmitt, B.; Welfring, J. Life cycle assessment of biogas production by monofermentation of energy crops and injection into the natural gas grid. Biomass Bioenergy 2010, 34, 54−66. (45) Er-rbib, H.; Bouallou, C.; Werkoff, F. Production of synthetic gasoline and diesel fuel from dry reforming of methane. Energy Procedia 2012, 29, 156−165. (46) North, M. Synthesis of cyclic carbonates from epoxides and carbon dioxide using bimetallic aluminium (salen) complexes. ARKIVOC 2012, 2012, 610−628. (47) Tao, X.; Wang, J.; Li, Z.; Ye, Q. Theoretical study on the reaction mechanism of CO2 hydrogenation to methanol. Comput. Theor. Chem. 2013, 1023, 59−64. (48) Giglio, E.; Lanzini, A.; Santarelli, M.; Leone, P. Synthetic natural gas via integrated high-temperature electrolysis and methanation: part I energy performance. Journal of Energy Storage 2015, accepted for publication.

1652


Digester Gas Upgrading to Synthetic Natural Gas in Solid Oxide

Recommend Documents