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SimultaneoussyngasproductionwithdifferentH2/COratioinamulti-tubularmethanesteamanddryreformerbyutilizingofCLCARTICLEinJOURNALOFENERGYCHEMISTRYJANUARY2015ImpactFactor:2.35DOI:10.1016/S2095-4956(15)60284-4

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Journal of Energy Chemistry 24(2015)5464

Simultaneous syngas production with different H2/CO ratio in amulti-tubular methane steam and dry reformer by utilizing of CLC

Mohsen Abbasia, Mehdi Farniaeib, Mohammad Reza Rahimpourc, Alireza Shariatica. Department of Chemical Engineering, School of Chemical and Petroleum Engineering, Persian Gulf University, Bushehr 75169, Iran;

b. Department of Chemical Engineering, Shiraz University of Technology, Shiraz 71555-313, Iran;c. Department of Chemical Engineering, School of Chemical and Petroleum Engineering, Shiraz University, Shiraz 71345, Iran

[ Manuscript received July 7, 2014; revised September 9, 2014 ]

AbstractFor syngas production, the combustion of fossil fuels produces large amounts of CO2 as a greenhouse gas annually which intensifies globalwarming. In this study, chemical looping combustion (CLC) has been utilized for the elimination of CO2 emission to atmosphere duringsimultaneous syngas production with different H2/CO ratio in steam reforming of methane (SR) and dry reforming of methane (DR) in aCLC-SR-DR configuration. In CLC-SR-DR with 184 reformer tubes (similar to an industrial scale steam reformer in Zagros PetrochemicalCompany, Assaluyeh, Iran), DR reaction occurs over Rh-based catalysts in 31 tubes. Also, SR reaction is happened over Ni-based catalystsin 153 tubes. CLC via employment of Mn-based oxygen carriers supplies heat for DR and SR reactions and produces CO2 and H2O as rawmaterials simultaneously. A steady state heterogeneous catalytic reaction model is applied to analyze the performance and applicability of theproposed CLC-SR-DR configuration. Simulation results show that combustion efficiency reached 1 at the outlet of fuel reactor (FR). Therefore,pure CO2 and H2O can be recycled to DR and SR sides, respectively. Also, CH4 conversion reached 0.2803 and 0.7275 at the outlet of SR andDR sides, respectively. Simulation results indicate that, 3223 kmolh1 syngas with a H2/CO ratio equal to 9.826 was produced in SR side ofCLC-SR-DR. After that, 1844 kmolh1 syngas with a H2/CO ratio equal to 0.986 was achieved in DR side of CLC-SR-DR. Results illustratethat by increasing the number of DR tubes to 50 tubes and considering 184 fixed total tubes in CLC-SR-DR, CH4 conversions in SR and DRsides decreased 2.69% and 3.31%, respectively. However, this subject caused total syngas production in SR and DR sides (in all of 184 tubes)enhance to 5427 kmolh1. Finally, thermal and molar behaviors of the proposed configuration demonstrate that CLC-SR-DR is applicable forsimultaneous syngas production with high and low H2/CO ratios in an environmental friendly process.

Key wordschemical looping combustion (CLC); dry reforming of methane (DR); steam reforming of methane; carbon dioxide capturing; syngas produc-tion

1. Introduction

Nowadays, the production of syngas (a mixture of CO andH2) has been increased significantly for the synthesis of keycomponents such as methanol, dimethyl ether, aldehydes, hy-drogen and ammonia in petrochemical companies [1]. Hugeamount of syngas in petrochemical companies is produced bysteam reforming of methane (SR) process. SR reaction is en-dothermic and supplementary heat for this reaction over Ni-based catalysts is provided by direct combustion of fossil fuelsin furnaces [2,3].

The H2/CO ratio in syngas is a key parameter for produc-tion of different components.

Utilizing of SR process, a syngas with a high H2/CO ratiois achieved. Many downstream processes such as methanol

production and Fischer-Tropsch synthesis process need syn-gas with a H2/CO ratio near to 1 [4]. CO2 reforming ofmethane or dry reforming of methane (DR) process is a suit-able process for syngas production with a H2/CO ratio near to1 and consumption CO2 as a greenhouse gas [5,6].

Unfortunately, there are a few industrial applications ofDR for syngas production [7]. This is due to that the amountof heat to be supplied for DR is greater than that required forSR. Another problem is catalyst development especially de-activation based on carbon formation [8,9]. However, DRhas many environmental incentives and publications aboutthis subject have been increased in literature significantly[1012]. Of course, it must be noted that similar to SR reac-tion, DR reaction is endothermic and the combustion of fossilfuels produces large amounts of CO2 as a greenhouse gas an-nually which intensifies global warming [1315].

Corresponding author. Tel: +98-917-3239077; Fax: +98-771-4551838; E-mail: [email protected]

Copyright2015, Science Press and Dalian Institute of Chemical Physics. All rights reserved.doi: 10.1016/S2095-4956(15)60284-4

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Journal of Energy Chemistry Vol. 24 No. 1 2015 55

A relatively new combustion technology which integratesair separation into the combustion process for CO2 capture ischemical looping combustion (CLC).

CLC involves the use of circulating oxygen carriers intwo interconnected reactors. The oxygen carriers are reducedwith fuel in a bubbling fluidized bed fuel reactor (FR) andthen transported to a high velocity fast fluidized bed riser airreactor (AR) for oxidation with air. Almost pure CO2 canbe captured after H2O condensation of the exit gas from FR[1620].

Feasibility of CLC instead of fired-furnace for H2 pro-duction by steam reforming of CH4 has been illustrated byRyden and Lyngfelt [21]. Results indicated that CLC can beemployed instead of fired-furnace with a good performance inSR.

Rahimpour et al. [22] have analyzed the performance ofthermal coupling of CLC and SR via the employment of Fe-based oxygen carriers. They found that CLC-SR has a betterperformance in comparison with conventional steam reformer(CSR). In addition, Abbasi et al. [23] investigated the fea-sibility of CLC application via Ni-based oxygen carriers forheat supplying in a novel stem reformer assisted by H2 perm-selective membranes. They indicated that the application ofCLC is feasible with enhancement of SR performance basedon CH4 conversion and H2 production.

It must be noted that, by application of CLC instead offired-furnace in CSR, huge amount of CO2 is only capturedbut not consumed. A good way for consumption of the cap-tured CO2 in CLC-SR is DR process for syngas production.Of course, there is not any research in literature about CO2capturing with CLC and consumption in DR simultaneously.

Application of CLC as a heat source for SR and DR reac-tions can be defined as a thermally coupled reactor [24].

In this paper, utilizing of CLC for the elimination of CO2emission to atmosphere and heat supplying during simultane-ous syngas production with different H2/CO ratio in SR andDR process has been investigated. In the other hand, CO2 iscaptured at the outlet of FR by the condensation of H2O and

then recycled to DR tubes for syngas production in CLC-SR-DR configuration consequently.

Mn40/Mg-ZrO2 particles have been employed as Mn-based oxygen carriers in CLC-SR-DR.

ZrO2 is the support of this Mn-based oxygen carrier. Thisoxygen carrier has a lot of benefits to be employed in largescales CLC from the economical view and toxicity. Addition-ally, kinetics of reduction and oxidation of this oxygen carrierwith CH4 and air has been determined in literature [25].

A mathematical model of one-dimensional heterogeneouscatalytic reaction is applied for feasibility study of CLC-SR-DR application based on the thermal and molar behaviors un-der steady state condition.

2. Process description

The novel CLC-SR-DR configuration proposed has beenpresented in Figure 1. In this arrangement, AR and FR are op-erated in fast fluidization and bubbling fluidization regimes,respectively. The reforming reactions occurs in DR and SRsides by fixed-bed catalysts. For transferring of generated heatin AR, the reforming tubes have been inserted in AR. Also,AR has been covered by FR because AR supplies heat forendothermic reactions in FR, SR and DR sides. SR reactionis occurred in 153 reformer tubes by employing the Ni-basedcatalysts. Also, Rh/Al2O3 catalysts have been loaded in 31 re-former tubes for occurring DR reaction. This catalyst has beenselected due to advantages such as high reaction rate and nocarbon formation on the catalyst surface during H2 productionwith a CO2/CH4 ratio near 1 in feed [26].

Operating parameters and characterization of SR and DRsides in CLC-SR-DR are similar to an industrial scale CSRin Zagros Petrochemical Company, Assaloyeh, Iran, that hasbeen presented in previous work [23]. Also operating param-eters of CLC have been designed based on sufficient energyproductions for occurring SR and DR reactions in 184 reform-ing tubes.

Figure 1. CLC-SR-DR configuration

56 Mohsen Abbasi et al./ Journal of Energy Chemistry Vol. 24 No. 1 2015

Based on the experimental results of employing Mn-based oxygen carrier in CLC systems, dimensions of AR andFR, solid inventory in AR and FR and oxygen carrier circula-tion flow rate have been designed as shown in Table 1 [25,27].In CLC-SR-DR configuration, the FR feed is CH4 and ARfeed is air with 20% excess. Also, pre-reformed gas plus re-cycled H2O are inserted in SR side and the mixture of recycledCO2 plus CH4 are inserted to DR side.

Table 1. Design characteristics and input data of CLC-SR-DR

Design characteristics ValuesFR

Inlet temperature (K) 818Inlet pressure (bar) 10

Total feed gas flow (mols1) 151Heated length (m) 12

Feed velocity (ms1) 0.6Inside diameter (m) 5.1

Dense bed length (m) 6Solid inventory (kg) 90000

ARInlet temperature (K) 818Inlet pressure (bar) 20

Excess air 20%Inside diameter (m) 3.95

Feed velocity (ms1) 7Solid inventory (kg) 40000

DR sideInlet temperature (K) 818Inlet pressure (bar) 20

Total feed gas flow (kmolh1) 1087.2Number of tubes 31

Inside diameter (mm) 125Heated length (m) 12

Feed composition (mole%)CO2 50CH4 50

SR sideInlet temperature (K) 818Inlet pressure (bar) 40

Total feed gas flow (kmolh1) 7591.5Number of tubes 153

Inside diameter (mm) 125Heated length (m) 12

Feed composition (mole%)CO2 1.72CO 0.02H2 5.89

CH4 32.59N2 1.52

H2O 58.26

3. Reaction scheme and kinetics

3.1. Steam reforming of methane (SR)

During the reforming of methane with steam over Ni-based catalyst, SR and water gas shift (WGS) reactions occur

as follows:

CH4 + H2O CO+ 3H2 H0298 = 206.3 kJ mol1(1)

CH4 + 2H2O CO2 + 4H2 H0298 = 164.9 kJ mol1(2)

CO+ H2O CO2 + H2 H0298 = 41.1 kJ mol1(3)

The kinetic model of Xu and Froment is employed inmodeling as follows [28]:

r1 =k1

P 2.5H2

(PCH4PH2O

P 3.5H2 PCO

KI

)

12

(4)

r2 =k2

P 3.5H2

(PCH4P

2.5H2O

P 4H2PCO2KII

)

12

(5)

r3 =k3PH2

(PCOPH2O

PH2PCO2KIII

)

12

(6)

= 1 +KCOPCO +KH2PH2 +KCH4PCH4 +KH2OPH2OPH2 (7)

The reaction equilibrium constants, Arrhenius kinetic pa-rameters and Vant Hoff parameters for species adsorption ofgases have been represented in literatures [28,29].

3.2. Dry reforming of methane (DR)

During the reforming of methane with CO2 over Rh-based catalyst DR and reverse water gas shift (RWGS) reac-tions occur as follows [4]:

CH4 + CO2 2CO+ 2H2 H0298 = 247.3 kJ mol1(8)

CO2 + H2 CO+ H2O H0298 = 41.1 kJ mol1 (9)In the modeling, Richardson and Paripatyadar kinetic

model is applied, as follows [26]:

rDR = kDR

[(KCO2KCH4PCO2PCH4

(1 +KCO2PCO2 +KCH4PCH4)2

)(

1 (PCOPH2)2

KDRPCO2PCH4

)] (10)

rRWGS =kRWGSKCO2KH2PCO2PH2

(1 +KCO2PCO2 +KH2PH2)2[

1(PCOPH2O)

2

KRWGSPCO2PH2

] (11)

The Arrhenius kinetic parameters, constants of reactionequilibrium and adsorption equilibrium constants have beensummarized in previous work [23].


3.3. Oxidation and reduction of Mn-based oxygen carrier inCLC

The oxidation and reduction reactions of Mn-based oxy-gen carrier can be found below [25]:

O2 + 6 MnO 2Mn3O4 H0298 = 464.3 kJ mol1(12)

CH4 + 4Mn3O4 12MnO+ CO2 + 2H2O H0298 = 126.3 kJ mol1

(13)

To determine the oxidation and reduction rates of Mn-based oxygen carrier, the following equations are used:

rs =dXsdt =

1i

For oxidation:Xs =Xox =

mmredmoxmred

(14)

For reduction:

Xs =Xred = 1Xox

i =1

kiCng(15)

ki = koie

(EaiRT

)(16)

Specific properties and kinetic parameters of Mn40/Mg-ZrO2 oxygen carrier have been presented in literature [25].

4. Mathematical modeling

A one-dimensional heterogeneous catalytic reactionmodel for the simulation of CLC-SR-DR is applied. Thedifferential equations describing mole and energy balances inCLC-SR-DR, pressure drop and boundary conditions, hydro-dynamic concept of fluidization and the auxiliary correlationshave been presented in previous work [23].

A mathematical model of one-dimensional heterogeneouscatalytic reaction for CLC-DRM is based on the following as-sumptions:

Steady-state condition has been considered for model-ing.

Gas behaviors are ideal. In each side of reactors, plug flow mode is considered. Heat and mass axial diffusions are negligible.Bed porosities in axial and radial directions are constant. Surrounding heat loss is negligible (isolation of FR). Bubble and emulsion phases are considered for AR and

FR. The operating condition is assumed to be isothermal,

i.e. bubble and emulsion phases, due to large turbulency andmixing.

Plug flow pattern is considered to follow the bubbles. The rising velocity of the bubble is constant. Spherical bubbles with constant size are considered.

The extent of reaction in bubble phase is much less thanthat in emulsion phase.

In bubbling and fast fluidized beds, because the goodsolid mixing, the temperature in radial direction is usuallyuniform or constant, and therefore one-dimensional heteroge-neous reaction model is reasonable.

Backward finite difference method is used to solve the setof ODEs. One hundred separate segments are considered forthe reactor length and the Gauss-Newton method is used tosolve the non-linear algebraic equations in each segment forthree sides simultaneously.

For validation of the model, the simulated results are com-pared with the industrial data of CSR in Zagros PetrochemicalCompany, in Assaluyeh, Iran and a good agreement betweenthe model prediction and the plant data is observed as shownin Table 2. In fact, Table 2 presents CH4 conversion and com-position of stream at the outlet of CSR.

For evaluation of CLC-SR-DR performance, CH4 conver-sion, H2 yield and combustion efficiency (c) are determinedas follows:

CH4 conversion =FCH4,inFCH4,out

FCH4,in(17)

H2 yield =FH2,outFH2,in

FCH4,in(18)

c =(2yCO2 + yH2O)outFout

4(yCH4)inFin(19)

where, Fin and Fout are the molar flows at the inlet and outletof the reactors, respectively.

Table 2. Agreement between model prediction and plant dataCH4 Composition (mole%)Parameters

conversion (%) CO2 CO H2 CH4 N2 H2Oplant data 26.5 5.71 3.15 31.39 20.41 1.29 38.05

CSR 26.0 5.72 3.19 31.53 20.33 1.30 37.94

5. Results and discussion

5.1. Thermal behavior of CLC-SR-DR

Variation of temperature along the reactor axis in AR, FRSR and DR sides of CLC-SR-DR have been presented in Fig-ure 2. For better explanation of thermal behavior and patterns,the knowledge of heat generation and consumption plus trans-ferred heat between sides are useful. Therefore, heat genera-tion and consumption plus transferred heat between sides areillustrated in Figures 3 and 4. As shown in Figures 24, it canbe said that AR temperature increased at the entrance of thereactor rapidly due to the occurrence of high exothermic oxi-dation reaction of oxygen carriers. After that, AR temperaturedecreased at the rest due to the complete oxidation of oxygencarriers that did not cause heat generation in the second halfof AR (see Figure 3a). In fact, sharp increasing of FR temper-ature in the first half of reactor is due to a large heat transferfrom AR to FR (see Figure 4a).


Figure 2. Variation of temperature along reactor axes in all the sides of CLC-SR-DR

Figure 3. (a) Variation of heat generation in AR and heat consumption inFR, SR and DR sides along reactor axes in CLC-SR-DR, (b) comparison be-tween heat consumption in FR and sum of heat consumption in SR and DRsides in CLC-SR-DR

Temperature reduction of SR and DR sides at the entranceof CLC-SR-DR is due to that heat generation in AR is lessthan the sum of heat consumption in FR, SR and DR sides asshown in Figure 3(b) and Figure 4(a). Also, the reduction ofSR and DR temperature sides at the second half of the reactoris due to the reduction of heat transfer from AR to SR and DRsides which have been clearly presented in Figure 4.

Figure 4. (a) Variation of heat transfer from AR to FR, DR and SR sidesalong reactor axes in CLC-SR-DR, (b) comparison between heat transfer fromAR to DR and SR sides in 1 tube of CLC-SR-DR

It must be noted that, in the second half of CLC-SR-DR,sensible heat of very hot gases in AR supplies the transferredheat from AR to SR and DR sides. Larger heat transfer fromAR to DR in comparison with SR per one tube is due to lowertemperate of DR side at the second half of CLC-SR-DR.

5.2. Molar Behavior of CLC-SR-DR

Variations of component mole fractions along the reactoraxis in AR, FR, SR and DR sides of CLC-SR-DR have beenshown in Figure 5.

It is obvious from Figure 5(a) that oxygen was consumedrapidly in the first half of AR and constant amount of oxygenmole fraction at the rest indicated complete oxidation of oxy-gen carrier in AR. Also, a complete consumption of CH4 inFR was obvious in Figure 5(b).

Figure 5(c) and 5(d) show that syngas was produced witha H2/CO ratio equal to 9.826 and 0.986 in SR and DR sides ofCLC-SR-DR, respectively.

Figure 6(a) illustrates that the combustion efficiencyreached to 1 at the outlet of FR and therefore pure CO2 canbe captured by the condensation of H2O.

Figure 6(b) indicates that CO2 molar flow ratedecreased from 633.5 to 167.3 kmolh1 in DR side and thus


Figure 5. Variation of components mole fractions along reactor axes in AR (a), FR (b), SR side (c) and DR side (d) of CLC-SR-DR

Figure 6. (a) Variation of combustion efficiency along reactor axes in FR of CLC-SR-DR, (b) variation of CO2 molar flow rate along reactor axes in DR sideof CLC-SR-DR, (c) variation of H2O molar flow rate along reactor axes in SR side of CLC-SR-DR


541 kmolh1 CO2 was produced in FR. Therefore, all of theproduced CO2 in FR can be recycled and consumed in DR sideof CLC-SR-DR. In addition, Figure 6(c) demonstrates thatH2O molar flow rate decreased from 633.5 to 167.3 kmolh1in SR side and 1082 kmolh1 H2O was produced in FR.Therefore, all of the produced H2O in FR can be recycled andconsumed in SR side of CLC-SR-DR.

As shown in Figure 7(a), CH4 conversion reached to0.2803 and 0.7275 at the outlet of SR and DR sides, re-spectively. Also, H2 yield obtained 0.9894 and 1.441 inSR and DR sides of CLC-SR-DR, respectively, as given

in Figure 7(b). Higher H2 yield of DR side in com-parison with SR side is due to higher CH4 conversion inDR side.

It can be seen in Figure 7(c) that, 3223 and 1844 kmolh1syngases were produced in SR and DR sides of CLC-SR-DR,respectively. Of course, the lower syngas production in DRside in comparison with SR side can be attributed to that thenumber of DR tubes was 31 tubes against 153 tubes in SRside. Finally, by mixing outlet streams of SR and DR sides,5067 kmolh1 syngas with a H2/CO ratio equal to 3.13 couldbe achieved.

Figure 7. (a) Variation of CH4 conversion along reactor axes in SR and DR sides of CLC-SR-DR, (b) variation of H2 yield along reactor axes in SR and DRsides of CLC-SR-DR, (c) variation of syngas flow rate along reactor axes in SR side of CLC-SR-DR, (d) variation of syngas flow rate along reactor axes in DRside of CLC-SR-DR

5.3. Ef fect of ratio of DR tubes to total tubes on CLC-SR-DRperformance

Number of DR and SR tubes by considering fixed totaltubes equal to 184 tubes in CLC-SR-DR, is a key parame-ter that influences CLC-SR-DR performance. In this part ofpaper, number of DR tubes changed from 20 to 50 tubes thatmeans the ratio variation of DR tubes to total tubes from 0.109to 0.272.

It must be noted that, during changing of DR tubes, allthe operating parameters and dimensions of CLC-SR-DR didnot change as shown in Table 1 and the number of total tubesremained fixed equal to 184 tubes. By increasing DR tubesand reducing of SR tubes, feed flow rate in DR and SR sidesincreased and decreased, respectively. In the other hand, feed

flow rate per one tube in DR side remained constant but in-creasing number of tubes raised feed flow rate of DR side.

Figures 8 and 9 illustrate the effects of ratio of DR tubesto total tubes on temperature and CH4 conversion in DR andSR sides of CLC-SR-DR, respectively. As shown in thesefigures, by increasing DR tubes, temperature and CH4 con-version in SR side decreased due to the receiving of more heatby DR side. Results showed that by increasing the numberof DR tubes from 20 to 50 tubes, CH4 conversion in SR sidedecreased from 0.2949 to 0.2534 (see Figure 9a). Although,as shown in Figure 9(b), CH4 conversion in DR decreasedfrom 0.7454 to 0.6944 because the heat generation in AR wasfixed, by increasing DR tubes, the transferred heat from AR toone tube of DR side was reduced. This phenomenon has beenclearly shown in Figure 9(b) that by increasing the number of


DR tubes, the temperature of DR side along the reactor axiswas reduced.

According to the results of Figure 10, by increasing the

ratio of DR tubes to total tubes from 0.109 to 0.272, H2 yielddecreased from 1.035 to 0.905 in SR side and 1.48 to 1.378 inDR side.

Figure 8. Effect of ratio of DR tubes to total tubes on the temperature of SR side (a) and DR side (b) along reactor axes in CLC-SR-DR

Figure 9. Effect of ratio of DR tubes to total tubes on CH4 conversion along reactor axes in SR side (a) and DR side (b) of CLC-SR-DR

Figure 10. Effect of ratio of DR tubes to total tubes on H2 yield in SR side (a) and DR side (b) along reactor axes in CLC-SR-DR

Based on the reduction of CH4 conversion in SR and DRsides by increasing DR tubes, syngas production per one tubeof SR side decreased from 22 to 19.32 kmolh1 and 60.93to 56.76 kmolh1 in DR side (see Figure 11a and 11b). Ofcourse, because of larger CH4 conversion in DR in compari-son with SR, by increasing DR tubes from 20 to 50 tubes, totalsyngas production in SR and DR sides (all of 184 tubes) wasenhanced from 4827 to 5427 kmolh1.

By increasing the number of DR tubes and reducing SRtubes, the necessary amount of CO2 in feed of DR side wasincreased. Also, the needed amount of H2O in feed of SR sidewas reduced as shown in Figure 12. Therefore, by increasingthe ratio of DR tubes to total tubes, the production of CO2 inFR was not sufficient for the feed of DR side. Excess CO2should be provided by another process such as separation ofCO2 from outlet gases of SR side by amine process. From


economical view, by increasing the number of DR tubes inCLC-SR-DR, more high cost Rh-based catalyst and CO2 areneeded and therefore the costs of CLC-SR-DR increases sig-nificantly. Considering the huge amount of CO2 emissionsto atmosphere annually, the application of CLC-SR-DR in-stead of CSR is logical because 574.8 ton/day CO2 can becaptured and converted to syngas. An investigation in rela-tion to the commercial viability and economical feasibility of

this novel configuration plus experiments in pilot plant scaleare necessary in order to consider the commercialization ofCLC-SR-DR configuration. Finally, for application of thisconfiguration, by compaction and inclusion of three reactors(including AR, SR and DR) inside FR, oxygen carrier parti-cles collision and abrasion should be well considered. There-fore, it is mandatory to the addition of makeup oxygen carrierto FR.

Figure 11. Effect of ratio of DR tubes to total tubes on the syngas production along reactor axes in the 1 tube of SR side (a), DR side (b) and DR and SR sides(c) in CLC-SR-DR

Figure 12. Effect of ratio of DR tubes to total tubes on CO2 (a) and H2O (b) molar flow rates in DR side (a) and SR side (b) along reactor axes in CLC-SR-DR

6. Conclusions

In this study, utilizing of CLC for the elimination of CO2emission to atmosphere and heat supplying during simulta-

neous syngas production with different H2/CO ratios in SRand DR processes has been investigated in a CLC-SR-DRconfiguration. In fact, CO2 and H2O at the outlet of FR bythe condensation of H2O are separated and recycled to DR and


SR tubes for syngas production in CLC-SR-DR configuration.Mn40/Mg-ZrO2 particles have been employed as Mn-basedoxygen carriers in CLC-SR-DR. Based on the thermal andmolar behaviors under steady state condition, a mathematicalmodel of one-dimensional heterogeneous catalytic reaction isapplied for feasibility study of CLC-SR-DR application.

SR reaction occurres in 153 reformer tubes by employ-ing Ni-based catalysts. Also, Rh/Al2O3 catalysts have beenloaded in 31 reformer tubes for occurring DR reaction.

Simulation results show that combustion efficiencyreaches to 1 at the outlet of FR and therefore pure CO2 canbe captured by the condensation of H2O. Syngas is producedwith a H2/CO ratio equal to 9.826 and 0.986 in SR and DRsides of CLC-SR-DR, respectively.

In addition, CH4 conversion reaches to 0.2803 and0.7275 at the outlet of SR and DR sides, respectively. Also,H2 yield obtains 0.9894 and 1.441 in SR and DR sidesof CLC-SR-DR, respectively. In CLC-SR-DR, 3223 and1844 kmolh1 syngases are produced in SR and DR sides,respectively.

Simulation results present that by increasing the ratio ofDR tubes to total tubes from 0.109 to 0.272, CH4 conver-sion in SR side decreases from 0.2949 to 0.2534. Also, CH4conversion in DR decreases from 0.7454 to 0.6944. In ad-dition syngas production per one tube of SR side decreasesfrom 22 to 19.32 kmolh1 and 60.93 to 56.76 kmolh1 inDR side. Although, by increasing the ratio of DR tubes tototal tubes from 0.109 to 0.272, total syngas production inSR and DR sides (all of 184 tubes) is enhanced from 4827to 5427 kmolh1.

Finally, based on the thermal and molar behaviors, it canbe concluded that CLC-SR-DR has been designed correctlyand CLC can provides heat, CO2 and H2O for the productionof syngas in SR and DR reactions simultaneously.

Nomenclaturesav Specific surface area of catalyst pellet, m2m3

Ac Cross section area, m2

Ai Inside area of inner tubes, m2

Ao Outside area of inner tubes, m2

Ct Total concentration, molm3

Cp Specific heat of the gas at constant pressure, Jmol1

dp Particle diameter, mDi Tube inside diameter, mDo Tube outside diameter, mDij Binary diffusion coefficient of component i in j,

m2s1

Dim Diffusion coefficient of component i in the mixture,m2s1

Fi Flow rate of component i, mols1

F b Molar flow in bubble phase, mols1

F e Molar flow in emulsion phase, mols1

hf Gas-solid heat transfer coefficient, Wm2K1

hi and ho Heat transfer coefficient between fluid phase andreactor wall in exothermic and endothermic sides with

convection, Wm2K1

k1 Reaction rate constant for the 1st rate equation,molkg1s1

k2 Reaction rate constant for the 2nd rate equation,molkg1s1

k3 Reaction rate constant for the 3nd rate equation,molkg1s1

ki Rate constant of reaction i, molkg1s1bar1/2

k0i Pre-exponential factor of chemical reaction rateconstant for oxidation and reduction of OCs,mol1nm3n2s1

kg,i Mass transfer coefficient for component i, ms1

K Thermal conductivity of fluid phase, Wm1K1

KCH4 CH4 Adsorption equilibrium constant, Pa1

KCO2 CO2 Adsorption equilibrium constant, Pa1

KH2 H2 Adsorption equilibrium constant, Pa1

KDR Constant of DR reaction equilibriumKRWGS Constant of RWGS reaction equilibriumKw Thermal conductivity of reactor walls, Wm1K1

L Reactor length, mMi Molecular weight of component i, gmol1

n Reaction order for oxidation and reduction of OCsP Total pressure, PaPi Partial pressure of component i, PaR Universal gas constant, Jmol1K1

rDR Rate of DR reaction, molg1cat s1

rRWGS Rate of RWGS reaction, molg1cat s1

Re Reynolds numberSci Schmidt number of component iT Temperature, Kt Time, su Superficial velocity of fluid phase, ms1

ug Linear velocity of fluid phase, ms1

U Overall heat transfer coefficient between exothermicand endothermic sides based on convection, Wm2K1

UAD Overall heat transfer coefficient between AR side andDR side, Wm2K1

UAF Overall heat transfer coefficient between AR sideand FR side, Wm2K1

UAS Overall heat transfer coefficient between AR sideand SR side, Wm2K1

XOC Conversion of oxidized OCsyi Mole fraction of component iz Axial reactor coordinate, m

Greek letters Viscosity of fluid phase, kgm1s1

Density of fluid phase, kgm3

b Density of catalytic bed, kgm3

i Time needed for full conversion, sci Critical volume of component i, cm3mol1

Hf,i Enthalpy of formation of component i, Jmol1


b Void fraction of catalytic bed Bubble phase volume as a fraction of total bed

volume Volume fraction of catalyst occupied by solid particle

in bubble Oxygen carrier circulation flow rate to fuel flow

rate ratioc Combustion efficiency

Superscriptsg In bulk gas phases At surface catalyst

Subscripts0 Inlet conditionsi Chemical speciesj Reactor sideAR Air reactorFR Fuel reactorDR Dry reformingOC Oxygen carrierSR Steam reforming

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