Combining Electrochemical CO2 Capture with Catalytic Dry Methane

Research Article pubs.acs.org/journal/ascecg

Combining Electrochemical CO2 Capture with Catalytic Dry Methane Reforming in a Single Reactor for Low-Cost Syngas Production Peng Zhang,‡ Jingjing Tong,‡ and Kevin Huang*,‡ ‡

Department of Mechanical Engineering, University of South Carolina, 541 Main Street, Columbia, South Carolina 29201, United States ABSTRACT: We here report a potentially low-cost catalytic dry methane reforming process to make syngas with CO2 electrochemically captured from a CO2 source via a mixed conducting membrane in a single reactor. The mixed conducting electrochemical membrane is a composite comprising an O2−-conductor and molten carbonate phase, where the catalytic bed contains a Ni-MgO-1 wt % Pt (NMP) or LaNi0.6Fe0.4O3‑δ (LNF) catalyst. The reactor with the NMP catalyst generally outperforms the LNF counterpart in CH4 conversion rate and syngas production yield. At 850 °C and over the NMP catalyst, the membrane reactor yields a CO2 permeation flux of 2.25 mL min−1 cm−2, a H2 and CO production rate of 3.75 and 3.24 mL min−1 cm−2, respectively, and a CH4 conversion of 93.9%. The LNF catalyst shows a long activation period due to the slow Ni ex-solution process but does offer a better coking and coarsening resistance. Long-term stability tests show no apparent sign of degradation within 200 h. With 3% H2O added into methane, the reactor can produce a syngas with higher H2/CO ratio preferable for liquid fuels synthesis. Overall, this work demonstrates the technical feasibility of a combined capture and conversion “all-in-one” CO2 reactor for dry reforming of CH4 KEYWORDS: Membrane reactor, Mixed conductor, Flux, Conversion rate, Catalyst

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except osmium,21,22 among which noble metals such as Rh, Ru, Pd, and Pt are the favored choice for catalysts due to their superior coking resistance, higher catalytic activity, and stability.23,24 However, the high cost and rare availability of these noble metal catalysts limit their practical applications. Instead, non-noble metals such as Ni and Co have been widely studied as a commercial catalyst for DMR. However, these catalysts suffer from severe deactivation due to carbon formation,24,25 making their practical applications also seriously limited. On the other hand, early studies have also indicated that the catalyst support plays an important role in DMR. A variety of oxides including Al2O3, MgO, TiO2, ZrO2, SiO2, CeO2, and La2O3 have been studied in the past for DMR.26,27 Key characteristics for a good catalyst support material are (1) low concentration of Lewis sites in the support (ZrO2, MgO, and La2O3);28−30 (2) strong metal−support interaction (SMSI) (CeO2, ZrO2), and (3) high oxygen storage capability (OSC).31−36 These properties render a facile intermediates reaction at the metal/support interface,37,38 oxygen spillover,39,40 CO2 adsorption,41,42 CH4 dissociation,41,42 and high

INTRODUCTION Conversion of CO2 captured from fossil fuelled power plants into useful fuels/chemicals is a strategically important technology toward a sustainable energy future. Many CO2 conversion routes have been explored in the past, including direct chemical reactions1−4 and photocatalytic5−7 and electrocatalytic reduction,8−14 just to name a few. Among these conversion technologies, direct conversion of CO2 into useful products via hydrogenation received the most attention because of its technical maturity and suitability for mass production. However, this approach is not deemed carbon neutral since H2 and electricity as produced today are largely derived from fossil fuels, which themselves produces large amounts of CO2.15,16 A better alternative method is to utilize the CO2 captured for upgrading CH4, an abundant and affordable fuel of today, into syngas as precursor for liquid fuels17−19 and methanol.3,20 This method is commonly known as dry methane reforming (DMR). Compared to its rival steam methane reforming (SMR) and partial oxidation of methane (POM), DMR is an environmentally friendly catalytic process because it combines two harmful greenhouse gases to make value-added products. The DMR reaction CH4 + CO2 = 2CO + 2H2 (ΔH0298 K = +247 kJ mol−1) is a highly endothermic and volume-expanding chemical process. Therefore, higher temperatures and lower pressures favor the reaction. From a catalysis perspective, DMR can be catalyzed virtually by almost entire VIII transition metals © XXXX American Chemical Society

Received: August 15, 2016 Revised: September 18, 2016

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DOI: 10.1021/acssuschemeng.6b01960 ACS Sustainable Chem. Eng. XXXX, XXX, XXX−XXX

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ACS Sustainable Chemistry & Engineering surface concentration of oxygen vacancies,43−45 all of which can facilitate a faster catalytic DMR with minimal coking. Another obstacle to a widespread adoption of DMR technology is the cost of CO2. As is well-known that industrial CO2 is captured from a rather expensive process such as “imine based chemical washing”, not to mention that compression and transportation of the captured CO2 to a DMR site add extra cost to the final product. Therefore, developing cost-effective DMR technology is of great interest. Membrane-based CO2 capture is a continuous, flow-through process. In recent years, two new types of membranes based on electrochemical transport of CO32− have been demonstrated with high flux density and selectivity: (1) mixed electron and carbonate-ion conductor (MECC) membrane consisting of a metal and carbonate phase; (2) mixed oxide-ion and carbonateion conductor (MOCC) membrane consisting of an O2−conducting ceramic phase and a carbonate phase.15,16,41,42,46−60 One of the advantages for these membrane reactors is that the continuous removal of the permeated CO2 (and O2) products can substantially promote the transport of CO2 (and O2) through the membrane, thus achieving a sustainable high flux. Since the operating temperature of a MOCC membrane matches well with the reaction temperature of DMR, a logical idea is to utilize the captured CO2 to instantly react with CH4 in the presence of catalyst in a single reactor. Figure 1

realize stable high-flux CO2 capture and high-rate production of syngas in a single reactor. The performance of the same MOCC membrane reactor with the addition of H2O into CH4 has also been evaluated.

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RESULTS AND DISCUSSION Microstructure and Permeation Rate of the Prepared MOCC. A typical SEM image showing the cross-sectional microstructure of the as-prepared GDC-MC MOCC membrane is given in Figure 2, where the dark area is the MC phase

Figure 2. Microstructure of a GDC-MC dual-phase MOCC membrane.

and the light area is the GDC phase. Evidently, the MOCC membrane has a dense microstructure with the MC completely filing the GDC matrix. This microstructure ensures that there is no physical leakage of CO2 through the membrane and that all the captured CO2 in the sweep side results from the transport of CO32− and O2− within the membrane. Figure 3 shows the CO2 permeation rate of as-prepared MOCC membrane in the temperature range from 650 to 850

Figure 1. Schematic illustration of a MOCC membrane reactor with a catalyst bed for electrochemical CO2 capture and catalytic DMR. TPB: triple phase boundary.

schematically illustrates the concept. The key CO2 capture reaction is CO2 + O2− = CO32−. If successful, such combined CO2 capture and conversion reactor can eliminate the cost for compression and transportation, thus lowering the overall cost for upgrading CH4 into syngas via DMR. In 2013, Lin and coworkers 6 1 first demonstrated the concept with a La0.6Sr0.4Co0.8Fe0.2O3‑δ (LSCF)-molten carbonate (MC) dualphase membrane. The work clearly proved the concept, but the interaction between LSCF and MC led to a poor performance: with a 10% Ni/γ-Al2O3 as the catalyst, the CH4 conversion rate and syngas production rate were only 8.12% and 0.3 mL min−1 cm−2 at 850 °C, respectively. This level of performance is too low to be practically meaningful. Furthermore, no stability was reported by the study due to the interaction. In the present work, we report that a much improved MOCC membrane reactor comprising Gd-doped CeO2 and MC and loaded with a robust Ni-MgO-1 wt % Pt (NMP) catalyst or a ceramic-based LaNi0.6Fe0.4O3‑δ (LNF) catalyst can

Figure 3. CO2 flux density as a function of temperature.

°C with pure Ar as the sweep gas. As shown in Figure 3, JCO2 increases with temperature, suggesting a thermally activated process of the CO2 permeation through the MOCC membrane. The CO2 flux density is increased by nearly 4× from 0.16 to 0.62 mL min−1 cm−2 as the temperature is raised from 650 to 850 °C. This level of CO2 flux density is ∼4× that of the LSCFB


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the range of 1.41−2.14 mL min−1 cm−2. The DMR performance of this study is much better than the reported LSCF-based membrane reactor,61 where the CH4 conversion was only 8.1% with both H2 and CO production rates at 0.3 mL min−1 cm−2 at 850 °C.61 The possible reasons for the demonstrated better performance include (1) the CO2 permeation flux in this study is ∼4× higher than LSCF-MC membrane; (2) the catalyst in this study is NMP that has been proved better than the NiγAl2O3 catalyst used in Anderson’s61 as well as other DMR studies;63 and (3) higher inlet CH4 concentration in Anderson’s study.61 Determination of the CO2 flux density needs special caution. The amount of unconsumed CO2 can be easily measured by GC, whereas the amount of consumed CO2 is unknown due to the possible side reactions of RWGS and CH4 cracking reaction CH4 = C + 2H2 when Ni-based catalyst was used.64 To overcome this problem, we used the mass balance in H and O species. First, the amount of produced H2O in the RWGS reaction can be calculated by H balance based on the consumed CH4 and produced H2 obtained by GC. Then, the consumed CO2 can be calculated from the produced O (H2O and CO) amount. The total CO2 flux density that permeated through the MOCC membrane can, therefore, be calculated by

based MOCC membrane at the same temperature.61 The Arrhenius plot of the CO2 flux density is shown in Figure 3, and the slope of the straight-line yields an activation energy of Ea = 59.56 kJ mol−1, which is very close to that of a Gd-doped CeO2 electrolyte conductivity,62 suggesting that O2− transport is the rate-limiting step during the permeation. This is a reasonable observation considering that O2− conductivity of the GDCphase is at least 1 order of magnitude lower than that of molten carbonate.52 Our early study has also proved this hypothesis by correlating O2− conductivity with CO2 flux density.52 During the entire test, the leakage of N2 concentration varies in the range of 0.007%−0.013%. The CO2 selectivity (calculated as the concentration ratio of CO2 to N2) slightly decreases from 110 to 105 as the temperature is increased from 650 to 850 °C. DMR Conversion Performance. Temperature Effect. The DMR conversion performance using the MOCC membrane and NMP and LNF catalysts in the temperature range of 750 to 850 °C is shown in Figure 4. Both the CH4 and CO2

JCO (total) = JCO (unconsumed) + 2

2

− JH + JCO ) 2

1 (2J (consumed) 2 CH4 (1)

Thus, determined CO2 flux densities are shown in Figure 4, which increases from 0.91 mL min−1 cm−2 at 750 °C to 1.40 mL min−1 cm−2 at 850 °C. This is to compare with 0.30 and 0.62 mL min−1 cm−2 shown in Figure 3 at 750 and 850 °C, respectively, with pure Ar as sweep gas. The significant enhanced CO2 flux density by the presence of CH4 in the sweep gas is likely due to the increased ambipolar conductivity of GDC in reducing atmosphere. Furthermore, the obtained CH4 conversion rate is consistent with what was reported in the literature regarding catalytic DMR conversion.65,66 The DMR performance with LNF perovskite catalyst is shown in Figure 4b. It is clear that the CH4 and CO2 conversion rates as well as H2/CO production rates are lower than those with the NMP catalyst. For example, at 850 °C, the H2 and CO production rates are 0.45 and 1.12 mL min−1 cm−2 for LFN catalyst vs 1.51 and 2.14 mL min−1 cm−2 for the NMP catalyst, respectively. In addition, the CO production rate is almost 1.5 × the H2 production rate, suggesting that there was more H2 reacted with CO2 via RWGS to produce H2O and CO over the LFN catalyst, which was also confirmed by other study.67 CH4 Concentration Effect. The dependence of DMR performance on the CH4 inlet concentration are shown in Figure 5 for both NMP and LNF catalysts. Figure 5a of the NMP catalyst shows that increasing the CH4 concentration from 1.67 to 2.40% increases the CO production rate from 2.41 to 3.24 mL min−1 cm−2, followed by a slower increase at >2.40% CH4. Correspondingly, the H2 production rate increases more pronouncedly from 1.39 to 3.75 mL min−1 cm−2. The CH4 and CO2 conversion rates did not show significant change with the CH4 concentration. This could be explained by the concurrent increase in CO2 flux by the increased CH4 concentration. In contrast, Figure 5b of the LNF catalyst shows that the CH4 conversion rate decreases with CH4 concentration, signaling lesser catalytic activity of LNF than

Figure 4. Effect of temperature on the DMR performance of a GDCMC membrane reactor with an (a) NMP catalyst and an (b) LNF catalyst. The feed gas is 50 mL min−1 CO2−50 mL min−1 N2, while the sweep gas is 0.94 mL min−1 CH4−49 mL min−1 Ar. GHSV = 5800 h−1.

conversion rates exceed 85% with NMP catalyst as shown in Figure 4a. According to the DMR reaction, the theoretical ratio between H2 and CO should be 1. However, Figure 4a indicates a higher CO/H2 > 1 at T ≥ 770 °C as a result of relatively flat H2 production rate and positive CO production rate (vs T), suggesting that the reverse water gas shift (RWGS) reaction H2 + CO2 = H2O + CO is concurrent with the DMR reaction. Overall, excellent H2 and CO production rates are observed in C


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Figure 5. Effect of CH4 concentration on the DMR performance of a GDC-MC membrane reactor with an (a) NMP catalyst and an (b) LNF catalyst. The sweep gas CH4−Ar mixture has a flow rate of 60 mL min−1. GHSV = 5800−6500 h−1.

Figure 6. Stability of the DMR performance of the MOCC membrane reactor tested at 850 °C with an (a) NMP catalyst and an (b) LNF catalyst. GHSV = 5800 h−1.

NMP. The CO2 permeation and conversion rates instead increase with CH4 concentration, changing from 0.92 to 1.32 mL min−1 cm−2 for the permeation flux and 60% to 75% for the conversion rate as the CH4 concentration is increased from 1.17% to 3.57%. It is also found that the CO production rate is much higher than H2, implying that the RWGS reaction is a more active side reaction with LNF catalyst during DMR. The observed opposite trending of CH4 conversion rate with CH4 concentration for the two catalysts reflects the difference in catalytic activity of the two catalysts in catalyzing DMR and RWGS reactions. Stability Evaluation. Stability is a very important parameter to evaluate the technical feasibility of MOCC membrane reactors for commercial DMR. Figure 6 shows the stability of DMR performance tested at 850 °C with NMP and LNF catalysts. For Figure 6a of the NMP catalyst, the stability test was started after evaluating the temperature-dependence of DMR performance from 750 to 850 °C. It is evident that the DMR performance is stabilized after ∼5 h, and then the CH4 conversion rate, CO production rate, and CO2 permeation rate slightly increase during the remaining 162 h. Meanwhile, the H2 production rate decreases slightly, indicating an increasing involvement of RWGS with time; the improved CO 2 permeation rate could be a reason for this. Figure 6b of LNF catalyzed DMR shows a better stability over a 200-h period than the NMP rivalry even though its performance is lower. A noticeable step change at the 80-h

marker suggests a slow activation process of the catalyst, which could be well correlated with the ex-solution process of Ni nanoparticles to be shown later. Overall, the DMR performance with the LNF catalyst is impressive, reaching a 72.7% CH4 conversion rate, 1.18 mL min−1 cm−2 H2 production rate, 2.18 mL min−1 cm−2 CO production rate, 1.57 mL min−1 cm−2 CO2 permeation rate, and 79.1% CO2 conversion after 100 h. After the 200 h stability testing shown in Figure 6, the cell with the LNF catalyst was revaluated as a function of temperature and CH4 concentration; the results are shown in Figure 7. Compared to Figure 4b and Figure 5b of the cell with the LNF catalyst measured at the early stage of the test, Figure 7 shows a much improved performance after a 200 h stability test, during which the LNF catalyst has been considerably activated by the ex-solved Ni-NPs. For example, CH4 and CO2 conversion rates at 850 °C reached 72.8% and 79.2%, respectively, vs 30% and 60% at the same temperature shown in Figure 4b at the beginning of the testing. CO/H2 production rates and CO2 permeation flux also exhibited a dramatic increase. The trend of CH4 concentration effect shown in Figure 7b is in general similar to that in Figure 5b, but the magnitude of the performance of the activated cell has improved significantly. Bireforming of Methane (BRM). The reforming of methane with a combination of CO2 and H2O is termed bireforming of methane (BRM) in this study. One of the shortcomings of DMR is the lower H2/CO ratio than the required for F-T liquid fuels and methanol synthesis. At D


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Figure 7. DMR performance of MOCC with LNF catalyst revaluated at 850 °C after long-term testing. (a) The effect of temperature; (b) the effect of CH4 concentration. GHSV = 5800−6500 h−1.

Figure 8. (a) Performance of bireforming of methane as a function of temperature; (b) temperature-dependent CH4 ratio consumed for DMR and SRM. Catalyst: NMP; CH4 concentration in sweep gas: 4.7%; H2O concentration: 3%. GHSV = 5800 h−1.

industry, the syngas produced by DMR is usually mixed with that produced by SMR with H2/CO = 3 to raise the final H2/ CO to a level suitable for liquid fuels synthesis. This would inevitably increase the production cost. To achieve a higher H2/ CO ratio, we performed testing with an addition of 3% H2O into CH4 as the sweep gas. Such BRM was tested after 161-h DMR operation of the MOCC membrane reactor with the NMP catalyst shown in Figure 6a. The testing condition was the same as DMR, except 3% H2O was added into the CH4−Ar gas-mixture as the sweep gas. Therefore, in this experiment, it is no longer appropriate to use the H and O balance to calculate the CO2 permeation rate as described before since H2O also reacts with CH4 to produce syngas. According to our thermodynamic calculation and analysis of the post-test catalyst (see the next section), it is reasonable to assume that there is no carbon formation during BRM. Thus, the consumed CO2 permeation rate can be calculated by the C balance. The following equation was used to determine the total CO2 permeation rate (JCO2(total)) of MOCC during BRM.

both CO2 and CH4 conversion rates increase from 90.6 to 94.3% and from 77.9% to 81.9, respectively, as the temperature is increased from 750 to 850 °C. Correspondingly, CO2 permeation rate (CO2 flux density), H2, and CO production rates are also increased. For example, the CO2 permeation rate increases from 1.5 to 3.3 mL min−1 cm−2. Meanwhile, the H2/ CO ratio in the syngas decreases slightly from 1.7 to 1.3 from 750 to 850 °C, suggesting more DMR than SMR at higher temperatures. This trending change is further shown in Figure 8b, where the percentage of CH4 involved in DMR and SRM reactions are separated. The CH4 ratio consumed for the DMR reaction is increased from 35 to 75% as the temperature is increased from 750 to 850 °C, whereas it is decreased from 65 to 25% for SMR. The results suggest that the CH4 involved in either DMR or SRM reactions in a MOCC membrane reactor is mainly determined by the working temperature with the NMP catalyst. CH4 Concentration Effect. Figure 9 shows the effect of CH4concentration on performance of BRM at 850 °C. From Figure 9a, the CH4 conversion rate slightly decreases from 93.5 to 90.7% as CH4 concentration is increased from 1.7 to 4.8%. However, the CO2 conversion rate, H2 and CO production rates, and CO2 permeation rate increase significantly with CH4 concentration. For example, the CO2 conversion rate improves from 43.6 to 94.5%, H2 and CO production rates increase from 2.0 and 1.6 to 9.4 and 6.9 mL min−1 cm−2, respectively. The decreased CH4 conversion rate and increased CO2 conversion

JCO (total) = JCO (unconsumed) + JCO(produced) 2

2

− JCH (consumed) 4

(2)

where JCO2(unconsumed), JCO(produced), and JCH4(consumed) are the rates of CO2 consumed, CO produced, and CH4 consumed, respectively. Temperature Effect. Figure 8 shows the BRM performance as a function of temperature. It is evident from Figure 8a that E


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Figure 10. Stability test of the BRM performance of the MOCC membrane at 850 °C with an NMP catalyst. GHSV = 5800 h−1.

CO2 permeation rate (CO2 flux density) is seen, which leads to declining H2 and CO production rates. The H2 production rate remains relatively flatter than that of CO during the 105 h test. The decrease of CO2 permeation rate could be associated with the loss of MC in the MOCC membrane starting after 220 h, which will be further discussed in the next section. The H2/CO ratio increases gradually from 230 to 303 h, as a result of the decreased CO production rate. The CH4 and CO2 conversion rate started to fluctuate after 270 h, indicating loss of catalytic activity due to accelerated coarsening of Ni particles in the presence of steam;68 the latter is further proved by SEM imaging of the catalyst morphology before and after the test to be shown in the next section. Post-test Analysis of MOCC Membrane and Catalyst. The microstructures of the MOCC membrane and catalysts before and after the stability test for DMR and BRM were characterized by SEM and EDS. Figure 11a, b, and c show the microstructures of the MOCC membrane after the long-term test at locations of sweep side, center, and feed side, respectively. The surface and subsurface regions of the sweep side located at the bottom of the membrane, as shown in Figure 11a and d, are covered by MC, implying a downward movement of carbonate during the test. The surface and subsurface regions of the feed side located at the top of the membrane are shown in the Figure 11c and e with partially filled MC and porosity, respectively, indirectly supporting the downward movement of the carbonate. However, plenty of porosities are observed in the midsection of the membrane in the Figure 11b. These images indicate that MC at the midsection moved faster than the top layer toward the sweep side during operation, which may arise from the effect of the surface tension of the MC. The microstructures of NMP and LNF catalysts before and after the stability test are shown in Figures 12 and 13, respectively. Figure 12a and b shows the SEM images of the NMP catalyst before and after the performance test, respectively. The particle sizes, which are approximately 50− 100 nm of the before-test catalyst in Figure 12a, have been increased to 50−800 nm after the stability test in Figure 12b. The sintering behavior of Ni-based catalysts after hightemperature exposure is well-known to the catalysis community.26 However, there is no form of carbon observed on the surface of catalyst after DMR and BRM tests. The EDS analysis at sites 1 and 2, as marked in Figure 12b, is shown in the inset,

Figure 9. (a) Performance of bireforming of methane; (b) CH4concentration dependence of CH4 ratio involved in DMR and SRM at 850 °C. Catalyst: NMP. GHSV = 5800−6500 h−1.

rate with the CH4 concentration are mainly due to a decrease in CO2/CH4 ratio. The CO2 permeation rate increases from 1.5 to 2.9 mL min−1 cm−2 with CH4 concentration and finally flattens at 2.9 mL min−1 cm−2 at 4.78% CH4, suggesting that the CO2 permeation rate has reached the limit. Comparison of Figure 5a of DMR and Figure. 9a of BRM indicates a sharp increase in H2/CO ratio by introducing H2O into the sweep gas. For example, the H2/CO ratio is increased from 0.81 and 1.08 to 1.24 and 1.22 at CH4 concentrations of 2.7% and 3.4%, respectively, by adding 3% H2O into CH4. Figure 9b shows that ∼70% CH4 was used for the DMR reaction at 1.7% CH4 and that it was increased to 81.7% as the CH4 concentration is increased to 3.4%, which corresponds to the sharp increase of the CO2 conversion rate shown in Figure 9a. However, it decreases slightly with CH4 concentration from 3.4% to 4.8%. This may be attributed to a decrease of CO2/ CH4 ratio at higher CH4 concentration. Overall, the observed H2/CO enhancement arises clearly from the involvement of SMR. Stability Evaluation. The performance stability of the BRM was tested at 850 °C with a 4.8% CH4−3% H2O sweep gas and 50 mL min−1 N2/50 mL min−1 CO2 as the feed gas. Since this test was a continuation of the previous long-term test, the timeaxis starts at 198 h in Figure 10. From Figure 10, it appears that BRM performance is stable for the first 20 h, then a decrease in F


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Figure 11. SEM images of the MOCC membrane after the DMR and BRM test: (a) at the sweep side; (b) at center; (c) at the feed side; (d) sweep-side surface; and (e) at the feed-side surface.

Figure 13. SEM images of the LNF catalyst before-test (a) and aftertest (b), and (c) the elemental distributions of the yellow box.

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CONCLUSIONS In summary, the present work demonstrates that catalytic DMR can produce high-yield syngas with CO2 captured directly from a mixed conducting membrane in a single reactor. The CH4 conversion rate and H2/CO production rates in general increase with temperature and CH4 concentration. The performance of a reactor with NMP catalysts is better than that of the LNF counterpart. The latter catalyst shows a long activation period due to the slow kinetics of the Ni-NPs exsolution but possesses a better coking and coarsening resistance. Long-term stability tests indicate that such combined capture and conversion reactor show no apparent sign of degradation within 200 h. Furthermore, bireforming (CO2/ H2O) of methane produces a higher H2/CO ratio preferable for liquid fuels production. Future research needs to focus on developing high CO2-flux membranes such as thin films and better catalysts so that higher reforming capacity at higher CH4 concentrations can be achieved without invoking carbon formation. Alternatively, MECC membranes will be tested for DMR because half of O2 is permeated with CO2, which can significantly mitigate coking and raise the H2/CO ratio in the syngas.

Figure 12. SEM images of the NMP catalyst: (a) before and (b) after the test.

which further proves no obviously carbon formed on the catalyst surface. From Figure 13a, the particle size of the LNF catalyst before-test is approximately 1 μm and with a smooth surface. However, from Figure 13b of the LNF catalyst after the test, some small particles rich in Ni were found precipitated out from the La2O3/FeOx matrix. Such dispersed small size (∼200 nm) Ni-particles would not only serve as the high catalytic activity site for DMR and BRM reactions but also enable a high coking resistance due to the strong Ni-oxide interface.69

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EXPERIMENTAL SECTION

Preparation of the MOCC Membrane. A 20 mol % Gd2O3 doped CeO2 (GDC, 20GDC fuel cell material) was selected as the ceramic matrix since it is chemically inert to MC. Porous GDC matrix was made by intimately mixing GDC powers with a pore-former carbon black (99%, Alfa Aesar) in a ratio of 60:40 (vol %), followed by pressing the mixture into pellets (ϕ20 mm) with a static pressure of G


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ACS Sustainable Chemistry & Engineering 200 MPa and calcining in air at 600 °C for 2 h to remove the carbon, and finally sintering at 1450 °C for 5 h to achieve a mechanically strong ceramic body. The thickness of the porous GDC was approximately 0.83 mm. The molten carbonate (MC) phase used in this study was a binary eutectic mixture of Li2CO3 (99%, Alfa Aesar) and Na2CO3 (99%, Alfa Aesar) in a molar ratio of 52:48. It was first homogenized by melting at 600 °C for 2 h. The obtained carbonate was then broken up and ballmilled into submicron powders, followed by pressing into a pellet with a similar diameter to the active surface area of the GDC matrix, which will be later placed on the surface of porous GDC in a testing permeation cell for in situ making of dense MOCC membrane. Catalyst Fabrication. Metallic Ni-based Ni0.2Mg0.8O-1 wt % Pt (NMP) and ceramic-based LaNi0.6Fe0.4O3‑δ (LNF) catalysts were selected for this study. For the NMP catalyst, Mg(NO3)2·6H2O (99%, Alfa Aesar) and Ni(NO3)2·6H2O (99%, Alfa Aesar) powders were used as precursors directly dissolved into deionized water with the molar ratio of 8:2, and then 1 wt % H2Cl6Pt·6H2O (Sigma-Aldrich) was added into the aqueous solution and stirred thoroughly for 12 h. After drying on a hot plate at 80 °C, the precipitate was finally calcined in air at 800 °C for 5 h to obtain the final product. For the LNF catalyst, a conventional solid-state reaction method was used. The stoichiometric amounts of La2O3 (99.9%, Alfa Aesar, dried at 1000 °C prior to use), NiO (99%, Avantor Performance Materials), and Fe2O3 (99.9%, Alfa Aesar) powders were intimately mixed by ball milling for 24 h in an agate jar with the aid of ethanol alcohol. Pellets were then pressed from the powder under 200 MPa and calcined at 1200 °C in air for 12 h. The partially sintered pellets were then reground and calcined again at the same temperature. The sintered pellets were broken up and ball milled into ∼1 μm powders. Finally, both catalysts powders were pressed into pellets under 200 MPa and then pulverized into 20−40 mesh size powders. Prior to testing, the catalysts were reduced in pure H2 for 2 h at 850 °C. CO 2 Capture and Conversion Test. The rates of CO 2 permeation and DMR conversion using the prepared MOCC membrane in the presence of a catalyst were evaluated by a homemade permeation cell, which has been previously described in detail.52 Briefly, to assemble the cell, a porous GDC matrix was first sealed to a supporting alumina tube by a commercial silver paste as the sealant (Shanghai Research Institute of Synthetic Resins). A short alumina tube was then mounted to the top surface of the porous GDC to shield the feed gas. A 0.2 g MC pellet was then gently placed on the top surface of the porous GDC. Upon heating to above 490 °C, the MC pellet melts and flows into the underlying porous GDC, forming a dense MOCC membrane in situ. The amount of MC has been predetermined to be able to fill all the pores in the porous GDC matrix. The feed gas was a mixture of 50 mL min−1 CO2 and 50 mL min−1 N2; N2 was used as a tracer gas for leak correction if any. The sweep gas was in general a mixture of CH4 and Ar but varied slightly for different testing, which will be given in detail in the next section. The concentrations of CO2, CH4, H2, CO, and N2 in the effluent were analyzed by a gas chromatographer (Agilent 490). Commercial mass flow controllers (Smart-Trak, 50 Series) specifically calibrated for each gas under use were employed to control the gas flow rates. Appropriate amounts of LNF and NMP catalysts were added into the sweep chamber right beneath the MOCC membrane to form a fixed bed in approximately 1 cm height. The gap between the membrane sweep-side surface and the catalyst bed is ∼1 mm. The testing temperature for CO2-flux evaluation was in a range of 650 to 850 °C in an interval of 25 °C with pure Ar as the sweep gas, while for the DMR conversion, the operating temperature range was raised from 750 to 850 °C in an interval of 20 °C with CH4−Ar as the sweep gas. At each temperature, ∼ 1 h was given to allow the MOCC membrane to reach a steady state before GC sampling. The permeation rate of CO2 (JCO2) in the CO2 capture experiment with pure Ar as the sweep gas was calculated using the following equation: cCO2 F JCO = × Ar 2 (1 − cCO2) S (3)

where, CCO2 is the measured concentrations of CO2 in the effluent. FAr (ml min−1) is the flow rate of sweeping Ar; S is the effective reaction area of the MOCC membrane, 0.921 cm2. The CO2 permeation rate in the CO2 conversion experiment with CH4−Ar mixture as the sweep gas was calculated by a different method, which will be elaborated in the next section. The CH4 and CO2 conversion rates (XCH4 and XCO2) were calculated by the following equations:

XCH4 =

XCO2 =

FCH4(in) − FCH4(out) FCH4(in)

× 100 (4)

JCO (total) − JCO (unconsumed) 2

2

JCO (total)

× 100

2

(5)

where, FCH4 (in) and FCH4 (out) are the mass flow rates of methane into and out of the system, respectively; JCO2 (total) and JCO2 (unconsumed) are the calculated total CO2 permeation rate and the outlet CO2 flow rate, respectively. The H2 and CO production rates are calculated by c H2 F JH = × Ar 2 (1 − c H2 − cCH4 − cCO − cCO2) S (6) JCO =

cCO F × Ar (1 − c H2 − cCH4 − cCO − cCO2) S

(7)

where, CH2, CCH4, CCO2, and CCO are the measured concentrations of H2, CH4, CO2, and CO, respectively, in the effluent. Other Characterization. The microstructure and elemental distributions of the MOCC membrane and catalysts before and after testing were captured and analyzed by a field emission scanning electron microscope (FESEM, Zeiss Ultra) equipped with an energy dispersive X-ray spectroscopy (EDS) analyzer.

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AUTHOR INFORMATION

Corresponding Author

*Tel: +1-803-777-4185. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS Financial support from NSF (CBET-1340269 and CBET1401280) is greatly appreciated. REFERENCES

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J


Combining Electrochemical CO2 Capture with Catalytic Dry Methane

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