Carbon dioxide reforming of methane in a SrCe0.7Zr0.2Eu0.1O3−δ proton conducting membrane reactor

i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 7 ( 2 0 1 2 ) 1 9 1 2 5 e1 9 1 3 2

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Carbon dioxide reforming of methane in a SrCe0.7Zr0.2Eu0.1O3Ld proton conducting membrane reactor Jianlin Li a,1, Heesung Yoon b, Eric D. Wachsman b,* a b

Florida Institute for Sustainable Energy, University of Florida, Gainesville, FL 32611, United States University of Maryland Energy Research Center, University of Maryland, College Park, MD 20742, United States

article info

abstract

Article history:

Utilizing CO2 for fuel production holds the promise for reduced carbon energy cycles. In

Received 24 July 2012

this paper we demonstrate a membrane reactor, integrating catalytic CO2 reforming of

Received in revised form

methane with in-situ H2 separation, that results in increased CO2 and CH4 conversion and

21 September 2012

H2 production compared to a Ni catalyst alone. The tubular proton-conducting

Accepted 22 September 2012

SrCe0.7Zr0.2Eu0.1O3
d membrane reactor demonstrates that the addition of the membrane

Available online 23 October 2012

improves CO2 conversion, due to in-situ H2 removal, by 10% and 30% at 900  C for CH4/

Keywords:

production at 900

Carbon dioxide reforming of methane

respectively. Further, the H2/CO in the reactor side effluent can be adjusted for subsequent

CO2 ¼ 1/1 and CH4/CO2/H2O ¼ 2/1/1 feed ratios, respectively. It also improves total H2 

C by 15% and 18% for CH4/CO2 ¼ 1/1 and CH4/CO2/H2O ¼ 2/1/1,

Steam reforming

desired Fischer-Tropsch products by combining CO2 reforming and steam reforming of

Proton conducting membrane

methane.

Hydrogen permeation

Copyright ª 2012, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights reserved.

Syngas SrCe0.7Zr0.2Eu0.1O3
d

1.

Introduction

Carbon geosequestration stores CO2 deep under the earth surface or in the ocean [1] as one option for treating CO2 from large-scale power plants. However, this technology is currently expensive [1] and alternative technologies are being developed including enhanced oil recovery [2] and methane production with simultaneous storage of CO2 [3] to offset cost, as well as high temperature steam/CO2 electrolysis [4]. H2 production is also a major source of CO2. The majority of H2 is produced through steam reforming of methane (SRM) H2 O þ CH4 4CO þ 3H2

DH0 ¼ 206 kJ mol


1

(1)

with the subsequent CO oxidation resulting in significant CO2 emissions. The global warming potential (GWP) of H2

production via the SRM process is estimated to be 13.7 kg CO2 (equiv.) per kg of net H2 produced (CO2 accounts for 77.6% of the system’s GWP) [5]. 0.3e0.4 million m3 of CO2 is produced when one million m3 of H2 is produced through a typical SRM H2 plant. The amount of CO2 emission would double if H2 is produced by coal gasification [6]. If the SRM CO2 was captured and sequestered it would cost about 25e30% of the total cost of H2 produced [7]. CH4 is also a greenhouse gas with a GWP 25 times that of CO2 and is released into the atmosphere from numerous sources such as landfills. This biogas is in fact a mixture of primarily CH4, CO2, and H2O, that if utilized could both reduce greenhouse gas (GHG) emissions and provide a source of energy. Carbon dioxide reforming of methane (CDRM) (equation (2)) can supplement other sequestration techniques with potential benefits to both the economy and the environment.

* Corresponding author. Tel.: þ1 301 405 8193; fax: þ1 301 314 8514. E-mail address: [email protected] (E.D. Wachsman). 1 Current address: Material Science and Technology Division, Oak Ridge National Laboratory, Oak Ridge, TN 37831, United States. 0360-3199/$ e see front matter Copyright ª 2012, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights reserved. http://dx.doi.org/10.1016/j.ijhydene.2012.09.134

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CH4 þ CO2 42CO þ 2H2

DH0 ¼ 247 kJ mol


1

(2)

The as-produced syngas can be converted further into synthetic liquid fuels through the Fischer-Tropsch (F-T) process providing a “drop-in” transportation fuel without changing the existing infrastructure and internal combustion engine technology [4]. Thus, the combination of CDRM and F-T offers a way to simultaneously recycle CO2 and generate fuel. In addition, various H2/CO ratios are required based on the desired F-T products, which can be achieved by combining the steam reforming, partial oxidation, and/or dry reforming reactions [8]. However, CDRM is highly energy intensive. Several technologies have been investigated to increase CO2 and CH4 conversion including catalytic conversion [9e14], plasma conversion [15e17], combinations of catalysis and plasma [18,19], and photocatalysis [20]. Among CDRM catalysts, nickel is a typical component due to its wide availability, low cost and high catalytic activity [21e24]. Integration of membrane separation and CDRM has also drawn the attention of some researchers [25,26] due to potential increase in CO2 (and CH4) conversion compared to catalysis alone via in-situ removal of H2. This process intensification can mitigate some of the energy consumption issues of CDRM, providing both operating and capital cost economic advantages. However, these membrane reactors have been based on expensive Pd based membranes with limited high temperature capability. Ceramic membrane reactors provide another option for CDRM due to their superior stability at high temperature and potential lower cost. SrCeO3
d and BaCeO3
d based perovskite oxides with multivalent cation dopants are promising for hydrogen separation membranes and have been reported by several groups [27e36]. Hydrogen permeates through these oxides by the ambipolar diffusion of protons and electrons. SrCeO3
d exhibits the highest proton transference number and undergoes no high temperature structural phase transitions up to 1000  C [37]. Thus, it was chosen in this work. Protonic defects in SrCeO3
d are formed by water dissociation into a hydroxide ion and a proton with the hydroxide ion filling an oxide ion vacancy and the proton forming a covalent bond with lattice oxygen [38]. This reaction is given in equa  tion (3), where VO ; O O and OHO are an oxygen vacancy, an oxygen-ion and a protonic defect in an oxygen lattice site, respectively (Kroger-Vink notation). 

H2 O þ VO þ O O 52OHO 

(3)

Electronic conductivity is necessary to balance the charge of protonic transport (ambipolar diffusion); however, it is insufficient in SrCeO3
d and can be increased by substituting Ce4þ with multi-valent cations, such as Eu [39,40]. The substitution reaction, expressed in Kroger-Vink notation, is equation (4), 0 where M, Ce, Ce Ce , MCe and h represent a trivalent metal, a cerium metal, a cerium-ion and a trivalent-ion in a cerium ion lattice site, and an electron hole, respectively. 

0 M þ Ce Ce 5MCe þ Ce þ h



(4)

We have reported that the electronic conductivity of SrCeO3
d and its chemical stability under CO2 atmosphere

were improved using Eu and Zr dopants, respectively [40e43]. Therefore, SrCe0.7Zr0.2Eu0.1O3
d was chosen for the membrane in the present work, and the effect of in-situ membrane separation on CDRM performance was investigated in terms of CO2 and CH4 conversion, H2 production, CO and H2 selectivity and H2/CO ratio in the syngas product. The H2/CO in syngas produced from CDRM is close to 1 while the ideal H2/ CO for F-T is w2 to produce liquid fuels. Therefore, the H2/CO in the syngas product was adjusted by combining CDRM and SRM through CH4/CO2/H2O feed mixtures. The performance of the membrane reactor was compared with and without in-situ hydrogen removal.

2.

Experimental section

Polycrystalline SrCe0.8Zr0.2O3
d and SrCe0.7Zr0.2Eu0.1O3
d powders were prepared by conventional solid-state reaction. NiOeSrCe0.8Zr0.2O3
d was chosen as a tubular support to maintain mechanical integrity and fabricated by tape-casting and rolling techniques. NiO was used to create porosity in the membrane support as well as be the catalyst when it was reduced to Ni by exposure to H2. Eu was not used in the support tube composition since electronic conduction is not functionally necessary for the support. The tubular support was w1 mm thick with a diameter of w6 mm. NiOeSrCe0.8Zr0.2O3
d tubular supports were sealed at one end and pre-sintered at 1300  C in air. SrCe0.7Zr0.2Eu0.1O3
d was coated on the inner side of the pre-sintered support by colloidal coating. The tubular membranes were finally sintered at 1520  C in air. Detailed powder synthesis and membrane fabrication have been discussed in our previous works [30,35]. The dense SrCe0.7Zr0.2Eu0.1O3
d membrane used in this experiment was w33 mm thick (Fig. 1(a)) with an active area of 12 cm2 [29,30]. The experimental setup (Fig. 1(b)) is the same as that in reference [30]. The outer side of the membrane (feed side) was exposed to CH4 and CO2 and steam. The influence of temperature, feed flow rate, CH4/CO2, and CH4/CO2/H2O were evaluated in terms of CO2 and CH4 conversion, H2 production and H2/CO ratio. Steam was achieved by vaporizing desired amount of water generated by a syringe pump fed heater. The inner side (sweep side) of the membrane was flushed with He at 20 cm3/min, in co-current flow with the feed gas. The flow rates of CH4, CO2, and He were controlled by mass flow controllers. The CH4/CO2/H2O feed ratios were adjusted based on a constant flow rate of methane. The reactants were flowed into the quartz chamber and exposed to the Ni catalyst on the outside of the membrane. The produced and/or unreacted steam in the reactor side effluent was condensed, by a cold trap filled with ice, prior to analysis of the composition of the reactor side effluent by gas chromatography (GC) (Varian CP 4900). The concentrations of the permeated H2 in the sweep gas (He) were analyzed by a mass spectrometer (Dycor QuadLink IPS Quadrupole Gas Analyzer). Before the experiment, the NiO in the support was reduced to create porosity by exposing the outer side of the membrane to 5 cm3/min H2 and 15 cm3/min Ar mixed with 3% H2O and the inner side of the membrane was exposed to 20 cm3/min He at 900  C overnight. Using the mass spectrometer, a plateau in

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Fig. 1 e Membrane morphology, SEM cross section of the membrane after experiment (a); and experimental setup (b).

H2 permeation flux was observed indicating the NiO was completely reduced to Ni. Moreover, the Ar was used as tracer to confirm there was no trans-membrane leakage (within mass spectrometer detection limits). The conversion of CH4 and CO2, and the ratio of H2/CO in the reactor side effluent are defined as:

XCH4 ¼

XCO2 ¼

out Fin CH4
FCH4

Fin CH4 out Fin CO2
FCO2

H2 =CO ¼

Fin CO2

 100%

 100%

out Fout H2 =FCO

(5)

(6)

(7)

The hydrogen from reacted CH4 exists as both H2 and H2O. The H2 selectivity (relative to H2O) is defined as the percentage of H2 produced from converted CH4. SH2 ¼

Fout  H2   100% out 2  Fin CH4
FCH4

(8)

Similarly, the carbon from reacted CH4 and CO2 exists as CO and C. The CO selectivity is defined as the percentage of CO in the reacted carbon mass balance.

Fout  CO   100% SCO ¼  out out in Fin CH4
FCH4 þ FCO2
FCO2

(9)

out (i ¼ CH4, CO2, H2, and CO) are the where Xi ; Si ; Fin i and Fi conversion, selectivity, input and output flux of i, respectively.

3.

Results and discussion

3.1. Membrane reactor effect on CO2 and CH4 conversion and H2 production The effect of the membrane was investigated by comparing CO2 and CH4 conversion and H2 production under two reactor configurations [34]: (1) with just Ni catalyst and (2) with Ni catalyst and in-situ H2 removal. In configuration (1), the permeate side was blocked, no sweep gas was used or H2 removed. Therefore, the membrane reactor (support tube) functioned only as a catalyst. In contrast, in configuration (2) the permeate side was connected to a mass spectrometer with sweep gas flowing, so that the permeated H2 could be removed in-situ and its concentration analyzed. Fig. 2(a) and (b) compare CO2 and CH4 conversions with and without in-situ H2 removal in 3 for feed gas compositions of Fin CH4 ¼ FCO2 ¼ 10 cm =min and in in in 3 FCH4 ¼ 2FCO2 ¼ 2FH2 O ¼ 20 cm =min, respectively.

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CO2 and CH4 conversion increased with increasing temperature, under both reactor configurations, as CDRM and SRM are endothermic. Moreover, the CO2 and CH4 conversion increased significantly with in-situ H2 removal as shown in Fig. 2. The CO2 conversion was 78% at 900  C and CH4/CO2 ¼ 1/ 1 with in-situ H2 removal compared to 70% without, a 10% improvement. Similarly, the CO2 conversion was 70% with insitu H2 removal at 900  C and CH4/CO2/H2O ¼ 2/1/1 compared to 54% conversion without, a 30% improvement. The improvement in conversion can be ascribed to the lower H2 partial pressure in the product side, from in-situ H2 removal, which drives the reactions (equations (1) and (2)) forward. H2 production was also significantly improved with in-situ H2 removal due to the higher CO2 and CH4 conversions (Fig. 3).

a

90 CH /CO =1/1 4

2

80

X

CO2

CH4

X

70

CO2

60

X

CH4

X

CH4

&X

CO2

(%)

X

The improvement was 15% and 18% at 900  C for CH4/CO2 ¼ 1/ 1 and CH4/CO2/H2O ¼ 2/1/1, respectively, and as expected there was much greater H2 production with H2O in the feed stream since it is an H2 source.

3.2. Effect of CH4/CO2 feed ratio with membrane reactors The effect of CH4/CO2 on CO2 reforming of CH4 was investi3 gated using Fin CH4 ¼ 10 cm =min with additional desired amount of CO2. Fig. 4 shows CO2 and CH4 conversion as a function of temperature and CH4/CO2 feed ratio. Both CO2 and CH4 conversions increased with increasing temperature since CDRM is endothermic. The CO2 conversion also increased with increasing CH4/CO2 feed ratios and was 52%, 60% and 78% at 900  C for CH4/CO2 feed ratios of 1/2, 1/1.5 and 1/1, respectively. The corresponding CH4 conversion was 93%, 89% and 87% under the same respective conditions. The conversion of CO2 and CH4 depends significantly on the catalyst and it’s supports [44] and our results are consistent with the literature [9,10,23,45e47]. Ideally the CO2 and CH4 conversions should be equal to each other with CH4/CO2 feed ratio of 1/1. However, the measured CH4 conversion is higher in this work implying side reactions, most likely carbon deposition through CH4 decomposition and/or the Boudouard reaction: CH4 4C þ 2H2

DH0 ¼ 75 kJ=mol DG0 ¼ 21960
26:45T J=mol

(10)

50

2CO4CO2 þ C DH0 ¼
172 kJ=mol DG0 ¼
39810 þ 40:87T J=mol 40 650

700

750

800

850

900

950

o

b

Temperature ( C) 90

Carbon deposition is further indicated by the gas phase carbon balance, sum of CH4, CO and CO2, in the reactor output. It is 2e5% less than the input amount for CH4/CO2 feed ratio of

CH /CO /H O =2/1/1 4

2

(11)

45

2

2/1/1

CH /CO /H O

80

4

2

2

40

70

H production (cm /min)

CH4

X

2/1/1 35

3

CO2

60

X

CH4

X

CO2

50

2

X

CH4

&X

CO2

(%)

X

40

30

25

20 CH /CO 4

1/1

2

15

30 650

700

750

800

850

900

950

o

Temperature ( C) Fig. 2 e Conversion of CO2 and CH4 under membrane reactors with and without in-situ H2 removal (solid symboldwith H2 removal; hollow symboldwithout H2 removal) (a) CH4/CO2 [ 1/1 and (b) CH4/CO2/H2O [ 2/1/1.

10 650

1/1

700

750

800

850

900

950

o

Temperature ( C) Fig. 3 e H2 production under membrane reactors with and without in-situ H2 removal (solid symboldwith H2 removal; hollow symboldwithout H2 removal).

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solid symbol-X

X

CH4

&X

CO2

(%)

90

80

1/2 CH4

hollow symbol-X

CO2

1/1

1/1.5

1/1

70

60

50

40 650

1/2

1/1.5

700

750

800

850

900

950

o

Temperature ( C) Fig. 4 e CO2 and CH4 conversion as a function of temperature and CH4/CO2 feed ratio.

1/1. In contrast, the total gas phase carbon remained essentially constant for CH4/CO2 feed ratios of 1/1.5 and 1/2 indicating that carbon deposition was negligible under these conditions, in agreement with Wang [44]. Carbon deposition is thermodynamically possible for CH4/CO2 feed ratio of 1/1 at temperature up to 870  C at 1 atm. Lower CH4/CO2 feed ratio can suppress carbon deposition, down to 760  C and 710  C for CH4/CO2 feed ratios of 1/1.5 and 1/2, respectively. Higher CO2 than CH4 conversion has been reported [23,45,48] and ascribed to the reverse water gas shift reaction (RWGS): H2 þ CO2 4CO þ H2 O DH0 ¼ 40:6 kJ=mol

(12)

The higher CH4 than CO2 conversion in this work may be due to different catalysts and higher operating temperature. Equations (10) and (11) are highly dependent on catalysts. Sacco et al. [49] reported that the primary source of surface carbon on Ni catalyst was CH4 indicating that most of the carbon deposition was through equation (10). Although the Boudouard reactions should not occur spontaneously over 700  C [44], CH4 decomposition is favorable at high temperature. In addition, H2 permeates through the membrane reactor lowers the H2 partial pressure limiting the RWGS, which would lower CO2 conversion. Thus, the CH4 conversion was higher than CO2 conversion due to carbon deposition with CH4/CO2 feed ratio of 1/1. In contrast, low CH4/CO2 feed ratio not only suppresses carbon deposition but also enhances the RWGS (equation (12)), resulting in higher CO2 conversion. For example, the CO2 conversion was 52.4% with CH4/CO2 feed ratio of 1/2 at 900  C, higher than half of the CH4 conversion (93.2%). For CDRM with a CH4/CO2 feed ratio of 1/2, CO2 conversion should be only half that of CH4. Thus, some of the CO2 was most likely consumed by H2 through RWGS reaction. The H2 and CO selectivity as a function of temperature and CH4/CO2 feed ratio are shown in Fig. 5. It is noted that total H2

production (sum of the permeated H2 and the H2 in the reactor side effluent) was used in calculating the H2 selectivity. The CO selectivity increased, whereas, the H2 selectivity decreased with increasing temperature. This can be ascribed to endothermic nature of the RWGS reaction consuming some of the H2 to further produce more CO with increasing temperature. At CH4/CO2 feed ratio of 1/1 the H2 and CO selectivities are equivalent, 96.2%, at w780  C. This crossover point moves toward lower temperature and higher selectivity as the CH4/ CO2 feed ratio decreases. Moreover, the desired product selectivity was 92% under all conditions investigated. H2 production as function of CH4/CO2 feed ratio and temperature is shown in Fig. 6. The H2 permeation, H2 in reactor side effluent, and total H2 production increased with increasing temperature due to the higher membrane ambipolar conductivity and higher CH4 conversion at higher temperature. A maximum H2 permeance of 2.2 cm3/min (w0.2 cm3/cm2 min) was achieved at 900  C with CH4/CO2 feed ratio of 1/1. H2 permeance decreased with decreasing CH4/CO2 feed ratios due to the increasing PO2 in the system, which causes lower electronic, and thus ambipolar, conductivity of the membrane material. A maximum total H2 production of 17.1 cm3/min (w1.5 cm3/cm2 min) was achieved at 900  C with CH4/CO2 ratio feed ratio of 1/2. The H2/CO ratio in the reactor side effluent as a function of temperature and CH4/CO2 feed ratio is shown in Fig. 7. H2/CO was less than 1 and decreased with increasing temperature due to endothermic RWGS and higher H2 permeation with increasing temperature. The decrease in H2/CO with lower CH4/CO2 feed ratio is also most likely due to RWGS.

3.3. CH4/CO2/H2O effect on conversion, H2/CO ratio and H2 production The theoretical H2/CO ratio in syngas from CDRM is 1/1, but is less than 1/1 with this hydrogen membrane reactor due to the 100 1/2

99

1/1.5 98

1/1

97

CO

2

& S (%)

4

H2

CH /CO feed ratio

S

100

96 95 1/1 94

CH /CO feed ratio 4

93 92 650

1/1.5

2

solid symbol-S

H2

hollow symbol-S

CO

1/2 700

750

800

850

900

o

Temperature ( C) Fig. 5 e CO2 and CH4 selectivities as a function of temperature and CH4/CO2 feed ratio.

950

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18

1/2 1/1

15 14 13 12 11 10

3

H production (cm /min)

17 16

1/2

1/1.5 1/1.5

1/1 solid symbol- total H production 2

hollow symbol- H in reactor side effluent 2

CH /CO feed ratio 4

2

3

2

1/1 2

1

650

1/2

H permeation 2

1/1.5 700

750

800

850

900

950

o

Temperature ( C) Fig. 6 e H2 production as a function of temperature and CH4/CO2 feed ratio. All aspects of H2 permeation increased with increasing temperature due to the higher membrane ambipolar conductivity and higher CH4 conversion.

in-situ removal of H2 and the RWGS. A low ratio is preferred for synthesis of oxygenated compounds and long-chain hydrocarbons [44,50,51]. However, the ideal H2/CO is w2 to produce liquid fuels through the F-T process and the H2/CO ratio can be high in syngas produced through SRM. Therefore, it is possible to increase the H2/CO in the reactor side effluent by combining CDRM and SRM. CH4/CO2/H2O feed ratios of 2/1/1 and 2/1/1.5 were investigated in terms of H2/CO and CH4 and CO2 3 in 3 conversion with Fin CH4 ¼ 20 cm =min; FCO2 ¼ 10 cm =min, and desired amount of steam.

The CO2 and CH4 conversions as a function of temperature and CH4/CO2/H2O feed gas ratio are shown in Fig. 8. Both the CO2 and CH4 conversion increased with increasing temperature, since both CDRM and SRM are endothermic. The CH4 conversion also increased with increasing H2O concentration, while CO2 conversion decreased. Higher H2O concentration favors steam reforming of CH4, enhancing CH4 conversion, but also means lower CH4 and CO2 feed concentrations since the total pressure is 1 atm. This favors the water gas shift reaction (reverse of equation (12)), thus lowering CO2 conversion. The CH4 and CO2 conversion were 85% and 70% at 900  C with CH4/ CO2/H2O feed ratio of 2/1/1, indicating more than half of the converted CH4 was through the steam reforming reaction (equation (1)). SRM is more dominant than CDRM at low temperature since it is less endothermic and the water gas shift reaction is exothermic. H2 production (reactor effluent, permeated, and total) and H2/CO ratio in the reactor effluent as function of CH4/CO2 feed ratio and temperature are shown in Fig. 9. The H2/CO in the reactor effluent decreased with temperature, similar to that in Fig. 7. However, the H2/CO ratio from 700  C to 900  C is between 1.7e1.9 and 2.0e2.5 for CH4/CO2/H2O feed ratios of 2/ 1/1 and 2/1/1.5, respectively, demonstrating the ability to adjust the H2/CO for the F-T process. All aspects of H2 production increased with increasing temperature due to both greater conversion and membrane conductivities with increasing temperature. A maximum H2 permeance of 4.7 cm3/min (w0.4 cm3/cm2 min) and a maximum total H2 production of 48.0 cm3/min (w4.0 cm3/ cm2 min) were achieved at 900  C with CH4/CO2/H2O feed ratios of 2/1/1 and 2/1/1.5, respectively. The reactor effluent and total H2 production also increased with increasing H2O concentration due to higher CH4 conversion. In contrast, H2 permeance decreased with increasing H2O due to higher PO2 (from thermodynamic equilibrium with H2O) that decreases electronic conductivity, and thus H2 permeability [30]. Moreover, the increase in total H2 production from that of reactor

1.00

90

CH /CO feed ratio 4

2

CH /CO /H O feed ratio 80

0.95

2

2/1/1.5

2

2/1/1

X

CH4

1/1

(%)

70 2/1/1

60

CO2

0.80

0.85

&X

2

H /CO

4

X

CH4

1/1.5

50 40 X

CO2

0.80

30

1/2

2/1/1.5 20

0.75 650

700

750

800

850

900

950

o

Temperature ( C) Fig. 7 e H2/CO in the reactor side effluent as a function of temperature and CH4/CO2 feed ratio. H2/CO was less than 1 due to endothermic RWGS.

10 650

700

750

800

850

900

o

Temperature ( C) Fig. 8 e CH4 and CO2 conversion as a function of temperature under various CH4/CO2/H2O ratios.

950

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50

2.6 total H production 2

2

35 2

30 8

CH /CO /H O reactor effluent 4 2 2 solid symbol--2/1/1 hollow symbol--2/1/1.5

7

H /CO 2

2.0

900

1.5 950

6

2

3

H in feed side effluent

40

H /CO

H production (cm /min)

45

5 H permeation

4

2

3 2 650

700

750

800

850 o

Temperature ( C) Fig. 9 e H2 production and H2/CO as a function of temperature. All aspects of H2 production increased with increasing temperature due to both greater conversion and membrane conductivities with increasing temperature.

effluent alone further demonstrates the membrane effect (as shown in Figs. 2 and 3), but under higher H2O concentrations that decrease this effect due to lower membrane conductivities with increased PO2 .

4.

Conclusions

Higher CO2 and CH4 conversions for CDRM and combined CDRM/SRM were demonstrated using a tubular SrCe0.7Zr0.2Eu0.1O3
d proton-conducting membrane reactor that co-produced pure H2 and syngas product streams. 10% and 30% higher CO2 conversion was achieved for CH4/CO2 ¼ 1/ 1 and CH4/CO2/H2O ¼ 2/1/1, respectively, at 900  C. The total H2 production was also increased 15% and 18% for CH4/CO2 ¼ 1/1 and CH4/CO2/H2O ¼ 2/1/1, respectively, at 900  C. The ability to tailor the H2/CO ratio in the reactor side effluent for the F-T process was also demonstrated. Thus the proton-conducting membrane reactor can, not only produce pure H2, but be combined with the F-T for simultaneous CO2 sequestration and liquid fuel production with increased yields.

Acknowledgments This work was supported by the Florida Institute for Sustainable Energy. Dr. Jianlin Li would like to thank Dr. Sean Bishop for valuable discussions.

references

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