Shashi Kumar Reforming

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INTERNATIONAL JOURNAL OF CHEMICALR EACTOR ENGINEERING

Volume 6 2008 Article A109

The Production of Syngas by DryReforming in Membrane Reactor Using

Alumina-Supported Rh Catalyst: ASimulation Study

Shashi Kumar Mohit Agrawal

Surendra Kumar Sheeba Jilani

Indian Institute of Technology Roorkee, India, [email protected] Indian Institute of Technology Roorkee, India, [email protected] Indian Institute of Technology Roorkee, India, [email protected] Institute of Technology Roorkee, India, sheeba [email protected]

ISSN 1542-6580

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The Production of Syngas by Dry Reforming in

Membrane Reactor Using Alumina-Supported RhCatalyst: A Simulation Study

Shashi Kumar, Mohit Agrawal, Surendra Kumar, and Sheeba Jilani

Abstract

A one-dimensional, isothermal mathematical model for methane reforming

with carbon dioxide in a conventional xed bed reactor (FBR) and a porous Vy-cor glass membrane reactor (MR) has been developed. The reactors are packedwith alumina-supported Rh catalyst. A simulation study shows that conversion of methane is higher in MR than that of FBR at all temperatures. In order to analyzethe overall performance of MR, a detailed simulation study has been carried out toelucidate the effect of temperature, sweep gas ow rate, dilution ratio, feed ratioon percent conversion of methane, yield of hydrogen and carbon monoxide, andratio of hydrogen to carbon monoxide in produced syngas. Besides, a comprehen-sive investigation on the effectiveness of the catalysts for methane reforming withcarbon dioxide reaction has been made and presented.

KEYWORDS: syngas, dry reforming, membrane reactor, Vycor glass mem-brane, alumina-supported Rh catalyst

Please send correspondence to Shashi Kumar, [email protected]. Sheeba Jilani is grateful toAligarh Muslim University, Aligarh, for granting her study leave to carry out doctoral studies atthe I.I.T. Roorkee.




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1. INTRODUCTION

In last few years, natural gas, a non-renewable energy source of primary energy,is utilized as a feed stock for several industrial high value-added productions, andalso as environmentally clean and easily transportable fuel due to its abundanceand enormous surplus in remote areas and underground resources. Natural gas is amethane enriched fuel (contains more than 80% of methane) and its use causes arise in global concentration of green house gases, CH 4 and CO 2, in atmosphere,resulting in global warming. The green house effect for CH 4 is more pronouncedthan for CO 2. According to the studies of Mackenzie and Mackenzie (1995), thecontribution of CH 4 and CO 2 accounts for three quarters of the total effect. In thisregard, therefore, extensive efforts are being made to convert green house gasesinto valuable products such as syngas. Syngas, a mixture of H 2 and CO, forms thefeed stock in the chemical and petrochemical industries for the production ofmethanol, acetic acid, olefins, gasoline, MTBE, oxo-alcohols, and phosgene etc.In some cases either H 2 or CO is utilized, for which H 2 and CO are acquired fromsynthesis gas. The hydrogen is then used in fuel cells, in the production of ureaand heavy water etc. However, the biggest consumer of H 2 from syngas isammonia synthesis. Recently it is being planned to utilize the hydrogen as a fuelfor non-polluting vehicle. The carbon monoxide is used in the production of

paints, plastics, pesticides, insecticides, acetic acid and ethylene glycol etc.In the past decades, the synthesis gas production via steam reforming,

partial oxidation of methane, and dry reforming reaction has received a greatinterest. Since CO 2 is available in large quantities and at low costs, CO 2 can beused in place of steam for reforming. Therefore, the dry reforming which isreforming of methane with CO 2 seems to be a promising technology for the

production of syngas. The synthesis gas produced by steam reforming has highH2/CO ratio which is not suitable for Fischer-Tropsch synthesis in the productionof long chain higher hydrocarbons due to the excess hydrogen which suppresseschain growth and decreases the selectivity of higher hydrocarbons (Hou et al.,2006). Conversely, methane reforming with CO 2 plays an important role in theindustries due to the production of syngas with a low H 2/CO ratio ( 1.0) whichcan be preferentially used for production of liquid hydrocarbons in Fischer-Tropsch synthesis network (Luna and Iriarte, 2008). The dry reforming is carriedout with excess CO 2 to promote reverse water gas shift reaction (RWGS) whichresults in lower H 2/CO ratio (Rostrup-Nielsen et al., 1984). Dry reformingreaction is slightly more endothermic than steam reforming. Therefore, it isfavored by low pressure and high temperature (Gadalla and Bower, 1988). The

presence of CO 2 gives rise to more chances of carbon formation on catalystsurface due to production of CO and consumption of H 2 via RWGS reaction.Therefore, the major problem encountered in the application of dry reforming is

Kumar et al.: Production of Syngas by Dry Reforming




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rapid deactivation of the catalyst. The stability and deactivation constraints renderthe application of conventional, inexpensive and easily available supported Nicatalysts difficult owing to their high activity for carbon formation. The noblemetal catalysts are usually found to be less sensitive to coking. Recently, for dryreforming reaction, excellent stability, high activity, coke resistance ability andactivity of Rh/Al 2O3 catalyst have been reported in the literature (Richardson andParipatyadar, 1990; Hou et al., 2006; Sakai et al., 1984; Qin and Lapszewicz,1994; and Erdohelyi et al., 1993).A brief review on Ni and noble metal basedcatalysts has been presented in this work in order to justify the applicability ofnoble metal based catalysts over Ni based catalyst in terms of activity andstability. Keeping this in view, in the present study, Rh/ -Al 2O3 catalyst has beenadopted for dry reforming reaction in a membrane reactor.

Dry reforming is reversible and endothermic reaction. The conversion islimited by thermodynamic equilibrium. Thus, in order to achieve high conversion,it should be carried out at high temperature. Another possibility to increase theconversion is the continuous and selective removal of one of the products whichin turn enhances the forward reaction rate and shifts the equilibrium towards theright side. The inert membrane reactor packed with catalyst implements thisconcept to improve the conversion by providing the reaction and separation in asingle unit. The development of the porous inorganic membrane based separation

process has gained a great interest to enhance the conversion (Kumar et al., 2006).The mesoporous membrane made of silica or alumina, have generally pore size inthe range of 2-50 nm. The examples of mesoporous membranes are Vycor glassand composite membrane of -alumina. The composite membranes are supported

by -alumina. The most important characteristics of the membrane are permeability and selectivity. Mesoporous membranes exhibit high permeability but relatively low selectivity and, therefore, low separation factor since available pore sizes are sufficient for molecular sieving. The separation of gaseouscomponents is governed by Knudsen diffusion.

In the present study a one dimensional steady state isothermalmathematical model for dry reforming reaction carried out in a membrane reactorincorporating Vycor glass membrane has been developed. A comprehensiveinvestigation on the effectiveness of the catalyst for dry reforming reaction has

been presented. After making the close observation on the brief review oncatalyst, no carbon deposition has been observed on the surface of Rh catalystsurface even after a long period of operation. Therefore in the kinetic model allreactions responsible for carbon deposition, for instance CH 4 cracking andBoudouard reaction, have been excluded. Thus dry reforming and RWGSreactions have been considered in the present study. Simulation results have beentaken out by using ODE solver in MATLAB-7. The performance of two reactorconfigurations namely commercial fixed bed reactor (FBR) and membrane reactor

2 International Journal of Chemical Reactor Engineering Vol. 6 [2008], Article A109




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(MR) has been compared. The axial variation in % conversion of CH 4 and flowrates of different reaction components have been illustrated at 923 K. The effectof temperature on % conversion of CH 4, selectivities and yields of H 2 and CO has

been investigated. In membrane reactor the sweep gas (Ar) flow rate is varied andits influence on conversion of CH 4 at various diluent flow rates has been studied.Ar is also used as a diluent in feed. The effect of diluent flow rate in feed on %conversion of methane has been investigated at three CO 2/CH 4 ratios. Anotherimportant study has been carried out to find out the effect of various feedcompositions on H 2/CO ratio in the produced synthesis gas.

2. CATALYST

The methane reforming with CO 2 has received considerable attention in recentyears due to its potential applications in the preservation of the environment andin the production of useful products. However, one major and serious problem ofthis reaction is the deactivation of catalysts caused by carbon formation and itsdeposition on the surface of catalyst. The formation of inactive carbon during dryreforming reaction may occur via two reactions; Methane decomposition andBoudouard reaction. These are as follows:

CH 4 C + 2H 2 (Methane decomposition)

2CO C + CO 2 (Boudouard reaction)

The decomposition of carbonaceous species on the catalyst suppresses theactive sites available on the surface of catalyst, blocks the catalyst pores and voidsleading to pressure rise in the reformer tubes (Quiroga and Luna, 2007). Besides,the deposition may also cause the break down of catalyst which also results in

plugging of reformer tubes. As a consequence, hot spots are developed on hottubes, the uneven flow distribution causes a self accelerating situation with furtherover heating of the hot tubes (Rostrup-Nielsen, 1993). Therefore, carbonformation can not be tolerated in tubular reformers. Many researchers haveexplored the approaches, which can improve the coke resistance on the catalyst.These approaches include selection of appropriate catalyst, sulfur passivation ofthe catalyst, nature of appropriate support, addition of promoters, metal oxide andmetal additives with strong Lewis basicity, change of reaction conditions, and theaddition of steam or H 2 (Luna and Iriarte, 2008; and Pechimuthu et al., 2007).

A variety of catalysts have been developed for methane reformingreaction. During the past decade, inexpensive and easily available Ni-basedcatalyst is more extensively developed (table1). These catalysts have shown veryhigh activity from the industrial point of view. However, the Ni based catalysts





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are completely deactivated within a few hours of the reaction due to thedeposition of stable and inactive carbon on the active centers of the surface.Recently several studies have been focused on VIII group metals (with theexception of Osmium) specially noble metal catalysts which are known to formless coke under reforming reaction or rather show high selectivity for carbon freeoperations. A variety of supports on noble metal catalysts for CH 4-CO 2 reforminghave been studied. Since the noble metals are costly and not easily available, the

bimetallic type catalysts (e.g. Ni-Rh) have been developed. The carbide of metalsespecially molybdenum and tungsten have gained considerable interest in variousreaction due to abundance of their precursors (Iyer et al., 2003). These carbidesare found to be stable at elevated pressure and are moderately resistant to coking.The addition of second metal could result in improvements in activity andstability. For instance, cobalt-tungsten -carbide is found to be active, stable for atleast 150 h and, provide high conversion with H 2/CO ratio close to unity. Theinfluencing factors for catalyst behavior include the activity of catalyst towardsthe conversion, stability, coking resistance and the type of deposited carbon. Allthese factors depend on the nature of different supports/promoters and interaction

between metal and supports/ promoters (Wang and Lu, 1998). There are usuallythree types of coke formed in dry reforming reaction on supported metal catalyst:

polymeric, filamentous and graphitic coke. The polymeric coke may be derivedfrom thermal decomposition of hydrocarbons. Filamentous and graphitic coke isformed on the catalyst. The coke can also be characterized on the basis of itsreactivity with H 2, H 2O and O 2 (Guo et al., 2007).For Ni-catalysts, the unreactedcarbon residues are dissolved in the metal to generate filamentous carbon causinga significant expansion of the catalyst bed. The contact between metal and supportis lost and so it is difficult to regenerate the catalyst system. Further studiesindicated that the nature of carbon deposits is a function of support. For instance,on Ni/ -Al 2O3 catalyst, high amount of filamentous type of carbon has beenobserved whereas on Ni/ -Al 2O3-ZrO 2, smaller amount of graphitic carbon andcarbon nanotubes have been observed. The studies on Rh and Ru based catalystscarried out by Rezaei et al. (2006) elucidate the effect of type of carbon onactivity and stability of catalyst. According to this study, the higher activity andstability of Ru and Rh catalysts may be due to the formation of highly reactivecarbon. This carbon is superficial carbidic carbon and is the reaction intermediatewhich quickly forms CO. The reactivity of superficial carbidic carbon influencesthe catalyst activity. The coke deposition analysis on noble metal showed thatorder of coke deposition is Pt>Pd>Ir>Ru>Rh, indicating minimum on Rh ( 0.1%). The observation of Richardson and Paripatyadar (1990) showed that Rhwas inactive for Boudouard reaction but Ru was deactivated quickly by samereaction. In the study of Richardson and Paripatyadar (1990), the slightdeactivation of Rh was observed probably due to the presence of small amounts





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of impurities in feed which got adsorbed on the surface of the catalyst. However,at high temperature these impurities were found to be destroyed so that theactivity of the catalyst was maintained. This type of deactivation is often observedwhen catalyst loading is very small. The deactivation by impurities is most

pronounced at low temperature. Large beds never show this effect and the catalystcan be operated for longer period without any deactivation.

The addition of support/promoters greatly suppresses the carbondeposition. A large number of studies have been carried out using the differentsupports and promoters as mentioned in table 1. The studies of Pompeo et al.(2005) reveal that - Al 2O3 is most suitable support due to its chemical and

physical stability but shows poor sintering tolerance. The addition of metallic promoters and supports modified with alkaline metal such as Li or K, leads toimproved resistance to carbon decomposition (Pompeo et al., 2005).The excesssteam and promoters control the carbon formation by producing CO. The partialsulfiding also reduces the carbon formation but small amount of steam is stillneeded (Richardson and Paripatyadar, 1990). Rezaei et al. (2008) have proposedthat addition of basic promoters can affect the metal support interaction andenhance the basicity of catalysts, leading to improve both activity and stability of

Ni-catalyst. Results also indicate that despite the amount of coke formation,activity of Ni catalyst is not decreased for promoted catalysts. The addition of rareearth metal oxides as support and promoters such as CeO 2 also improves the

behavior of alumina based catalysts by increasing the activity, stability and carbonresistance. ZrO 2 as a catalyst modifier activates adsorption of CO 2 and promotesthe gasification of deposited carbon. Many research workers have studied thereforming reaction capacity of Rh and Ru catalysts (table 1). The studies ofRezaei et al. (2006) indicate that Rh and Ru catalysts show a very high stabilitywithout any decrease in CH 4 conversion with time. Rh and Ru catalysts exhibithighest active metal surface area and there is no change observed in pore sizedistributions after use, indicating the high stability of catalyst. The activities forRh and Ru are observed ten times those of Ni, Pt and Pd (Richardson andParipatyadar, 1990). The conversion and deactivation tests indicate that Rh ismuch more stable catalyst than Ru and deactivation decreases with increasingtemperature. In addition to this, the observation also shows that Rh is inactive forBoudouard reaction. Tsipouriari et al. (1994) in their study also mentioned that Rhand Ru are far more active catalysts than Ni. In all experimental studiesmentioned in table 1, no carbon formation by Rh based catalyst occurred in any ofthe experiments which confirms the ability of Rh to perform the reaction withoutcarbon formation. From the above brief comments and information given in table1, it can be inferred that it is not necessary to consider carbon formation reactions(methane cracking and Boudouard reaction) with Rh supported catalyst in dryreforming reaction. The most effective support for Rh catalyst is found to be





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Al2O3 which is followed by Rh/TiO 2, Rh/SiO 2 and Rh/MgO (Richardson andParipatyadar, 1990; Richardson et al., 2003; and Erdohelyi et al., 1993) indecreasing order. Rh/ -Al 2O3 exhibits higher activity than other Rh/Al 2O3catalysts because active carbon species formed on the catalyst can participate insequence of steps to form CO (Rezaei et al., 2008).Therefore in the presentresearch work, the Rh/Al 2O3 catalyst is selected to carry out the dry reformingreaction in membrane reactor.

Table 1: Catalysts for dry reforming reaction

Contd..

Reference Catalyst Support Promoter Activity/selectivity/stability/carbondeposition

Vannice and Garten(1979)

Ni Ni Ni Ni

TiO 2 SiO 2Al2O3Graphite

- Ni/TiO 2 showed higher activitythan Ni/Al 2O3.

Seshan et al. (1994) Ni NiPtPt

-Al 2O3ZrO 2-Al 2O3ZrO 2

- Pt/ZrO 2 showed high stability incontrast with Pt/ -Al 2O3 andthose based on Ni.

Luna and Iriarte (2008) Ni Ni Ni Ni

Al2O3Al2O3Al2O3Al2O3

KCaMnSn

Ca, Mn and Sn-modifiedcatalysts showed a dramaticreduction in catalytic activityand significant increase incarbon deposition while K-modified catalyst showed lowcarbon deposition and high

catalytic activity.

Therdthianwong et al.(2008)

Ni Al 2O3 (15%) ZrO 2 The addition of ZrO 2 greatlyimproved the stability of

Ni/Al 2O3 in terms of cokeinhibition.

Swaan et al. (1994) Ni Ni Ni Ni Ni Ni Ni Ni

SiO 2 La2O3MgOZrO 2 TiO 2Al2O3-SiO 2SiO 2SiO 2

KCu

Rate of deactivation of Ni/Al 2O3-SiO 2 and Ni-Cu/SiO 2was much higher. The activityof Ni/MgO, Ni/TiO 2 and Ni-K/SiO 2were found to be close.

Lemonidou and Vasalos(2002)

Ni (5%) CaO (21.5%) -Al 2O3 (78.5%)

- The catalyst exhibited highactivity and very good stabilityat stoichiometric CH 4/CO 2 feed.Despite the amount of cokedeposited on the catalyst, noloss of activity was observed.





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Table 1 continued..

Contd


Pechimuthu et al. (2007) Ni CeO 2-(-Al 2O3) K Appreciable deactivation wasnot observed for the catalyst at650, 700 and 750 oC for 60-hrun. The carbon formed oncatalyst after 60-h at 700 and750C dispersed well andcould not be observed while at650 oC, a significant amountof carbon mainly graphitecoke was deposited on thecatalyst. However thedeposited coke did not affectthe high activity of thecatalyst. Therefore, thecatalyst had a high stability at

650o

C despite the depositionof coke.

Pompeo et al. (2005) Ni Ni NiPtPtPt

Al2O3-Al 2O3-ZrO 2ZrO 2 Al2O3-Al 2O3-ZrO 2ZrO 2

- Pt/ZrO 2, Pt/ -Al 2O3-ZrO 2 and Ni/-Al 2O3-ZrO 2 system hadlower deactivation levels than

Ni/ZrO 2, Ni/Al 2O3 andPt/Al 2O3. The lowestdeactivation level was in Ni/ -Al2O3-ZrO 2 and Pt/ -Al 2O3-ZrO 2. The highest carbonformation was observed for

Ni/Al 2O3 and Pt/Al 2O3catalysts.

Sakai et al. (1984) Ni (10%)Rh (5-10%)Pd (5-10%)Pt (5-10%)Ru (5-10%)

SiO 2Al2O3Al2O3Al2O3Al2O3

- Ni-SiO 2 and Rh-Al 2O3 wereexcellent in selectivity andactivity. The selectivity of Rhcatalyst was higher than that of Pd, Pt and Ru catalysts. Theorder of activity was:Rh>Pd>Pt>Ru.

Rostrup-Nielsen andHansen (1993)

NiRuRhPtPdIr

MgOMgOMgOMgOMgOMgO

- Ru and Rh catalyst displayedhigh selectivity and nodeposition of carbon wasobserved. A rapid carbonformation was on Pd catalyst attemperatures > 600 oC. A slowcarbon formation occurred onIr and Pt at temperatures > 750oC.

Hou et al. (2006) Ru (5%)Rh (5%)Pt (5%)Pd (5%)Ir (5%)

Ni (10%)

-Al 2O3-Al 2O3-Al 2O3-Al 2O3-Al 2O3-Al 2O3

- Noble metals (Ru, Rh, Pd, Ir and, Pt) showed higher cokeresistance ability while their activity was lower than that of

Ni and Co (10%). Rudispersed highly on meso-





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Table 1 continued..

Contd


Co (10%) -Al 2O3 porous Al 2O3 and exhibitedhigher coke resistance andhigher reforming ability.Supported Pt and Ru catalystsshowed poor stability. It might

be due to the sintering of thesemetals at high temperatures.

No coke deposition wasdetected on supported Pt, Ir,and Rh used catalyst whilesome amount of coke wasdetected on Pd/ -Al 2O3catalyst. Among coke freecatalysts, Rh/ -Al 2O3 exhibitedthe highest reforming activity.

Diaz et al. (2007) Ni (5%) Activated carbon Ca (1%) Ca played a Co-support roleinhibiting the deactivation of the catalyst at long periods of reaction. Methane conversionwas up to 40% at mildexperimental conditions.

Rezaei et al. (2006) RuRhIrPtPd

SpinelSpinelSpinelSpinelSpinel

- Ru and Rh showed the highestactivity for methane dryreforming. The order of activity was found to beRhRu>Ir>Pt>Pd. The resultsobtained reported a highstability for Ru, Rh and Ptcatalysts and lowest for Pd dueto the formation of lessreactive deposited carbon onPd and sintering.

Munera et al. (2003) Rh (2%)Pt (1%)

La2O3 (100%)La2O3 (100%)

- The activities of both catalystswere similar. The stability of Rh/La 2O3 > Pt/La 2O3. Verylow amount of carbonformation was observed on

oth, Rh and Pt catalysts after more than 100-h on stream.

Erdohelyi et al. (1993) RhRhRhRh

Al2O3TiO 2SiO 2MgO

- The order of activity wasfound to be: Rh/Al 2O3>Rh/TiO 2>Rh/SiO 2>Rh/MgO.

No deactivation and no carbondeposition was observed on

Rh/Al 2O3.





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Table 1 continued..

Contd.


Qin and Lapszewicz(1994)

RuRhPtIrPd

MgOMgOMgOMgOMgO

- For CO 2 reforming, the order of activity was found to be:Rh>Ru>Ir>Pt>Pd. Very lessamount of carbon was formedon Rh, Ru and Ir catalyst butsome carbon black wasobserved on test tube for Ptand Pd catalysts.

Sehested et al. (2001) Mo 2CRu (18%)

MgAl 2O4MgAl 2O4

- Activity of Ru was greater thanMo 2C catalyst by more than 2orders of magnitude.Resistance to carbon of Mo 2Cwas higher than that of Ru.Mo 2C was stable at only high

conversion.

Richardson et al. (2003) RhPt -Re

-Al 2O3-Al 2O3

- Very low carbon depositionoccurred on both of catalysts.Both the catalysts displayedexcellent stability but activityloss for Pt-Re catalyst wasobserved at temperature below700 oC.

Itoh et al. (1992) PtPtPt

Al2O3ZrO 2

x%-ZrO 2- Al 2O3( x=1, 5, 10, 20 )

- Zirconia supported catalystshowed much higher stabilityeven after 60-h run on stream.While Pt/Al 2O3 deactivatedsignificantly within 20-h onstream at 1073 K. The amountof carbon deposition onPt/Al 2O3 > on Pt/ZrO 2.

Ballarini et al. (2005) PtPtPtPt

Al2O3 Na-Al 2O3K-Al 2O3ZrO 2

- Both Pt/Na-Al 2O3 and Pt/ZrO 2catalysts showed a goodactivity and selectivity with avery high catalytic stability at1073 K. Pt/Al 2O3 showed a

poor performance and lower stability due to carbondeposition. Pt/K-Al 2O3displayed lower conversionthan Pt/Na-Al 2O3 catalyst.

Darujati and Thomson(2005)

Mo 2CMo 2CMo 2C

-Al 2O3ZrO 2MgO

- Surface area and thermalstability and had a much higher activity than a bulk Mo 2Ccatalyst, even though

deactivation via oxidation alsooccurred. The ZrO 2 supportexperienced serious sinteringand led to subsequentdeactivation.





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Table 1 continued..

3. MEMBRANE

In recent years the development of inorganic membranes has paved the way forthe application of membrane in the high temperature reactors as these have goodmechanical stability and have high resistance to temperature and corrosiveenvironment (Zaman and Chakma, 1994). Conversely, organic polymericmembranes have limited stability at temperatures over 100 C (Li et al., 1988).The inorganic membranes can be classified into two types according tomechanism of separation: (a) porous membrane and (b) non -porous densemembrane. Inorganic porous membranes are found to have promising future forindustrial applications in catalytic reactor and thereby currently are in great use.In past decades, many valuable reviews on porous inorganic membrane reactorshave been published (Coronas and Santamaria, 1999; Dixon, 2003; Julbe et al.,2001). Porous inorganic membranes consist of small pores that areinterconnected. Depending upon pore size, these membranes are of three typesnamely macroporous (d p > 50 nm), mesoporous (50 >d p > 2nm), and microporous(d p < 2nm). Owing to large pore size, macroporous materials such as -aluminamembranes are used only to support the layers of smaller pore size materials toform composite membranes. The mesoporous materials such as Vycor glass and-alumina, and microporous materials provide separation function. However,microporous materials such as carbon molecular sieves, porous silicas andzeolites offer very high separation factors (Dixon, 2003). The transportmechanism on porous membranes predominantly is governed by Knudsendiffusion. Therefore, these membranes allow all gases to pass through whichresults in high values of permeability and low selectivity. On the other hand,nonporous dense inorganic membranes are made of either metals such as Pd,silver or their alloys or solid oxide electrolytes such as modified zirconias and

perovskites (Dixon, 2003; Zaman and Chakma, 1994). Dense membranes are permeable to atomic or ionic forms of hydrogen or oxygen. Pd-Pd alloy


Brungs et al. (2000) Mo 2CMo 2CMo 2CMo 2C

Al2O3SiO 2ZrO 2TiO 2

- The catalyst stabilityMo 2C/Al 2O3>Mo 2C/ZrO 2>Mo 2C/SiO 2> Mo 2C/TiO 2. Thecatalyst with Al 2O3 and ZrO 2support did not show theappreciable sign of deactivation for the time

period of the study and thusthe author concluded as a

promising system for methanedry reforming.





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membranes offer high permeability only for hydrogen whereas zirconia and perovskites are highly selective only for oxygen. The transport mechanismfollows surface diffusion mechanism. However, on commercial scale theirapplicators are limited due to their high cost, low permeability because of highwall thickness of 100-150m (Herman et al., 1997), difficulty in fabrication,sensitivity to poisoning by sulfur species and embrittlement upon aging (Julbe etal., 2001). The performance of dense membranes can be improved by decreasingthe membrane thickness, by increasing the surface roughness and by developingthe new materials. However, it is difficult to obtain a thin membrane withsufficient strength and satisfactory permeability. Therefore, supported membranesare being manufactured and used. Shu et al.(1991) and Dittmeyer et al. (2001)have provided an extensive coverage of the Pd based membranes and reactors interms of their properties, preparation techniques and applications, and effect on

performance of reactor.Currently there is a growing interest in the development of novel high

temperature inorganic membrane reactors. The membranes which are capable ofwithstanding high temperature and harsh chemical environment are made up ofceramic materials such as porous Vycor glass. The membranes developed from

porous Vycor glass have been made successful at a commercial scale and have been reported to have good chemical and thermal stability upto 850 C with a pore size of 4 nm [Phair and Badwal (2006); Shelekhin et al. (1995); Prabhu et al.(1999), Li et al. (1988); Qiu and Hwang (1991)]. The studies of Shelekhin et al.(1995) reveal that at temperature higher than 925C, the change in the internal

pore structure occurs due to the collapse of few pores rather than change in the pore size. This results in the reduction in membrane permeability. No gas permeability is observed in membrane heat treated at temperature above 1000C.In addition to this, the same study suggests that Vycor glass membrane should be

preheated at temperature higher than the operating temperature of the process toavoid the shrinkage of the membrane and resulting stresses. In view of abovestudies on Vycor glass membranes, it is, therefore, worthwhile to use Vycor glassmembrane as a potentially high performance membrane for reforming reactionapplications upto temperature of 850 C. In the present study, Vycor glass (7930glass, Corning) membrane of tubular geometry manufactured by Prabhu etal.(1999) has been employed. The characteristics of this membrane have beenevaluated experimentally by Prabhu et al.(1999) and summarized as follows. Thenormal pore size is in the range of 2-4 nm which is the lowest range of themesoporous category of porous membrane. The ratio of mean free path / pore sizeis greater than 180 for all species in the operating reforming reaction temperaturerange. The effect of temperature and mass of gaseous components on the

permeability has been found to be very close to that expected from Knudsendiffusion equation. These results indicate that the mean free path of a gaseous





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molecule is much larger than the pore size of the membrane and the transport processes involve interaction of gaseous molecules with the membrane pore wallsrather than between gas molecules. Since the Knudsen diffusivity is inversely

proportional to the molecular weight of the gaseous component, the lightergaseous components diffuse through the membrane faster than the heavier one.The diffusion mechanism through mesoporous membrane has been presented indetail by Kumar et al. (2006) according to which the permeance of i th gaseouscomponent through Vycor glass porous membrane can be written as:

ii M

RT RTd

r D

80003

2=

where M i is molecular weight of i th component. Accordingly, the permeation fluxfor each component can be written as follows:

For CH 4: ) P P ( M

RT RTd

r J ' CH CH

CH CH 44

4

4

80003

2=

For CO 2: ) P P ( M

RT RTd

r J ' COCO

COCO 22

2

2

80003

2=

For CO: ) P P ( M RT

RTd r

J '

COCOCO

CO =

80003

2

For H 2: ) P P ( M

RT RTd

r J ' H H

H H 22

2

2

80003

2=

For Ar: ) P P ( M

RT RTd

r J ' Ar Ar

Ar Ar

=

80003

2

For H 2O: ) P P ( M RT

RTd r

J ' O H O H O H

O H 222

2

80003

2=





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product

Sweep gas + permeated gas

Reaction zone Membrane

product

product

feed

feed

Sweepgas

Permeation zone

4. MATHEMATICAL MODEL

A schematic diagram of a tubular membrane reactor is presented in fig1. Thecatalytic membrane reactor is a cylindrical reactor equipped with a membrane.This membrane is inert with respect to chemical reaction and tubular in shape.The tubular membrane divides the reactor in two zones. First zone is shell sidezone which is a reaction zone packed with catalyst particles. The reaction occurs

Fig 1: Schematic diagram of a tubular membrane reactor

in this zone. Second is tube side zone, also called permeate zone where the sweepgas is introduced co-currently with respect to feed to carry away the permeatedgases from the permeate zone. The feed contains mainly CH 4, CO 2 and Ar (asdiluent) and is fed to the shell side of reactor. The chemically inert sweep gas Aris introduced into the tube side of reactor. Therefore, in the permeate side (tubeside) of the reactor, no chemical reaction occurs. The methane reforming of CO 2 to produce synthesis gas can be represented by the following reactions.

CH 4 + CO 2 2 CO + 2H 2 (Dry reforming) Ho298 = 247.4 kJ/ mol (1)

CO 2 + H 2 CO + H 2O (RWGS) Ho298 = 41 kJ/ mol (2)

The reforming of methane with CO 2 is reversible and endothermic innature (Gallucci et al., 2007). The equation (2) represents reverse water gas shiftreaction which occurs in parallel to dry reforming reaction as a side reaction. Thisreaction is also reversible and less endothermic in nature.

There may occur other side reactions responsible for carbon formation viamethane cracking and Boudouard reaction as discussed in previous section-2 oncatalyst. Here, in present study since Rh/ -Al 2O3 catalyst has been employed inthe reactor, the coke formation and deactivation of the catalyst may be neglected(section-2). Thus all side reactions other than RWGS reaction are excluded in this





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study. For dry reforming and RWGS reactions, the kinetics given by Richardsonand Paripatyadar (1990) on Rh/Al 2O3 has been utilized.

The forward rate of dry reforming reaction ( 1r ) and RWGS reaction ( 2r )are given by following equations:

211

4422

4242

1 ) P K P K (

P P K K k r

CH CH COCO

CH COCH CO

++= (3)

222 CO P r k = (4)

The reverse rate expressions are formulated by adding a term to theforward reaction rate so that at equilibrium net rate can become zero (Richardson

and Paripatyadar, 1990). As a result, the net rate of dry reforming ('

1r ) and RWGS( '2r ) reactions can be written as follows:

] P K K

) P P ( [

) P K P K (

P P K K k r

COCH

H CO

CH CH COCO

CH COCH CO'

24

22

4422

4242

1

2

211 11

++= (5)

) P P K

P P ( P k r

CO H

O H COCO

'

22

2

2

222 1 = (6)

where k 1 and k 2 are the rate constants, K 1 and K 2 are the equilibrium constants forreaction (1) and reaction (2) respectively, P i is the partial pressure of i th

component and2CO

K and4CH

K are the appropriate adsorption equilibrium

constants of carbon dioxide and methane, respectively. The temperaturedependence of these constants is as follows:

( 102065/ )1 1290

RT k e = (7)

( 73105/ )2 1.857

RT k e = (8)

42 (40684/ )2.60 10CH RT K e= (9)

2

2 (37641/ )2.6110CO RT K e= (10)





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Chemical equilibrium constants for dry reforming reaction and RWGSreaction have been derived by using thermodynamic properties (Van Ness et al.,2004). The final expressions are as follows:

3 7 21 2

28623 630506.091ln 4.084 10 3.0665 10 75.624ln T T T

T T K = + + + + (11)

54

2 2

5852.30 0.582 10ln 1.86ln 2.70 10 7.1977 K T T

T T = + + (12)

On the basis of stoichiometry of dry reforming and RWGS reactions, the rates ofconsumption/ formation of reaction species are given below.

The net rate of consumption of CH 4, 4CH r ='

1r (13)

The net rate of consumption of CO 2, 2COr ='

1r +'

2r (14)

The net rate of production of CO, COr = 2'

1r +'

2r (15)

The net rate of production of H 2, 2 H r = 2'

1r -'

2r (16)

The net rate of production of H 2O, 2 H Or ='

2r (17)

In order to develop one dimensional mathematical model for membranereactor, the length of reactor is divided into small elemental length segments ofsize dz keeping cross sectional area of reactor constant. The material balance foreach component has been taken around this control volume and equations have

been formulated. These model equations may also be directly formulated by usingcomprehensive model for membrane reactor given by Kumar et al. (2006). The

balance equations rely on the following assumptions:

1) Steady state operation2) Isobaric condition3) Isothermal operation4) Ideal gas law holds true5) Plug flow behavior (no radial concentration gradient)6) Resistances by gas film on both sides of membrane and by membrane

support to permeation are negligible.





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Shell side material balance equation

The material balance for i th component is as follows:

'1( ) 2 0

ii i

dF r area J R

dz = (18)

There are 6 components in shell side of reactor namely CH 4, CO 2, CO, H 2,H2O, and Ar. The equation (18) can be written for each component as follows:

4

4 4

22 ' '1 12( )( ) (2 ) 0

CH CH CH

dF r R R J R

dz + + = (19)

2

2 2

22 ' '1 12( )( ) (2 ) 0

COCO CO

dF r R R J Rdz

+ + = (20)

22 ' '1 12( )( ) (2 ) 0

COCO CO

dF r R R J R

dz + = (21)

2

2 2

22 ' '1 12( )( ) (2 ) 0

H H H

dF r R R J R

dz + = (22)

2

2 2

22 ' '1 1

2( )( ) (2 ) 0 H O

H O H O

dF r R R J R

dz + = (23)

'1(2 ) 0 Ar Ar

dF J R

dz + = (24)

Tube side material balance equations

Tube side is permeating side of reactor where inactive sweep gas flows. Thus,there is no chemical reaction. As a result, the material balance equations containonly permeation term and no reaction term. The material balance equation for i th component can be written as:

'

12i

idF J Rdz

= (25)





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The membrane is porous, therefore, all components present in the reactionside get permeated to tube side. Since Ar is used as sweep gas as well as diluent tofeed, there are also 6 components viz. CH 4, CO 2, CO, H 2, H 2O, and Ar in tubeside of reactor. The equation (25) can be written for each component as follows:

4

4

'

1(2 ) 0CH

CH dF J R

dz = (26)

2

2

'

1(2 ) 0CO

COdF J R

dz = (27)

'

1(2 ) 0CO

COdF J R

dz = (28)

2

2

'

1(2 ) 0 H

H dF J R

dz = (29)

2

2

'

1(2 ) 0 H O

H OdF J R

dz = (30)

'

1(2 ) 0 Ar

Ar dF J R

dz = (31)

5. SOLUTION PROCEDURE

The mathematical model developed for membrane reactor (MR) consists of a setof twelve differential equations. These equations contain rate of reaction term andrate of permeation term for each component. In order to modify the modelequations for fixed bed reactor (FBR), six model equations representing the

permeation zone of the reactor, are excluded from the model as there is no permeation zone in FBR. In reaction side model equations, the permeation termsare set to zero for all components. Thus, model for FBR consists of only sixordinary differential equations. The operating conditions combined with initialconditions are listed in table 2. For the simulation of reactor, the model equationsare solved simultaneously in MATLAB-7 using ODE solver tool. The followingdefinitions have been used for describing the performance of FBR and MR.





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Methane conversion (%) = 4 4 4

4

',0 ,1 ,1

,0

( )CH CH CH

CH

F F F

F

+ (32)

Selectivity of hydrogen = 2 2

2 2 2 2

',1 ,1

' ' ',1 ,1 ,1 ,1 ,1 ,1

H H

H H H O H O CO CO

F F

F F F F F F

+

+ + + + + (33)

Selectivity of CO =2 2 2 2

',1 ,1

' ' ',1 ,1 ,1 ,1 ,1 ,1

CO CO

H H H O H O CO CO

F F

F F F F F F

+

+ + + + + (34)

Yield of CO =4

',1 ,1

,0

CO CO

CH

F F

F

+ (35)

Yield of H 2 = 2 2

4

',1 ,1

,0

H H

CH

F F

F

+ (36)

Table 2: Operating and boundary conditions

Parameter Value Parameter ValueL 0.04 m 3.0

tP 1.0 atm r 4.0 nm'tP 1.0 atm d 1.0 mm

R 8.314 J/mol K4 ,CH o

F 24 mol/s

1 R 0.004 m 2 ,CO o F 24 mol/s'

1 R 0.005 m , Ar o F 27 mol/s

2 R 0.007 m'

, Ar o F 7 mol/sT 923 K

4

',CH o F 0 mol/s

0.62

',CO o F 0 mol/s

6. SIMULATION RESULTS

The fixed bed reactor (FBR) and membrane reactor (MR) are analyzed by usingsame operating conditions of feed, temperature and pressure for the sake ofcomparison of two reactor configurations via simulation. In order to validate the





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model, the experimental studies of Prabhu et al. (2000) on dry reforming reactionusing Vycor glass membrane have been considered. The model equations have

been solved at the same operating conditions used by Prabhu et al.(2000). Thecomparison of model results with experimental results has been made in terms of% conversion of CH 4 at various temperatures for FBR and MR in tables 3 and 4respectively. The conversion in MR appears to be higher than in FBR due tosimultaneous removal of gaseous product (H 2, CO and H 2O) from feed side ofmembrane to the permeate side. Further, it is noteworthy from these tables thatthe difference between two results (experimental and model) is very small in thewhole temperature range for FBR as well as for MR. In case of FBR, the % errorvaries from 0.24 to 2.2, while it varies from 1.1 to 5.2 for MR. In particular, it can

be noticed that the % error is minimum at 923 K and maximum at 948 K for bothFBR and MR. These results clearly show that the model predictions are in goodagreement with experimental predictions. Thus model simulates the laboratoryreactor very well.

Table 3: Validation of model results of fixed bed reactor with experimentalresults

Temperature(K)

Experimental results Model results Error (%)

848 36.7 36.5 0.54873 45.2 45.05 0.33898 54.3 53.48 1.5

923 62.6 62.75 - 0.24948 70.2 68.65 2.2

Table 4: Validation of model results of membrane reactor with experimentalresults

Temperature(K)

Experimental results Model results Error (%)

848 42.7 41.68 2.3873 51.3 50.72 1.10

898 59.9 58.72 1.96923 65.3 64.57 1.11948 76.1 72.13 5.2





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The calculated axial flow rate profiles of different gaseous reactants and products in FBR and MR at 923K are shown in figs. 2, 3, and 4, respectively. Infixed bed reactor (fig. 2), the flow rates of reactants CH 4 and CO 2 decreasecontinuously down the length of the reactor, until the equilibrium is achieved. Theflow rate of CO 2 is lower than the flow rate of CH 4 because of the consumption ofCO 2 in dry reforming reaction as well as in RWGS reaction, and consumption ofCH 4 only in dry reforming reaction. The flow rate of inert diluent Ar in feedremains constant. Conversely, the flow rates of products H 2 and CO increasedown the length of reactor significantly as these are produced in more molaramounts as compared to molar consumption of reactants. The flow rate of H 2 islower than that of CO because some of the H 2 produced in dry reforming reaction,gets consumed in (RWGS) to yield more CO and H 2O whereas CO is producedfrom both reactions. The flow rate of water is very small and increases slowlyalong the length of reactor because it is produced in small quantity only byRWGS reaction. The end flow rate of water is about 2.2 mol/sec.

0

5

10

15

20

25

30

35

40

45

0 0.01 0.02 0.03 0.04 0.05

F l o w r a

t e o f c o m p o n e n t s

i n f i x e

d b e d

r e a c t o r

( m o l

/ s )

Length of reactor (m)

H 2

CO

Ar

H 2O

CH 4

CO 2

Fig 2: Axial variation of flow rate of components in FBR

For the membrane reactor utilizing the Vycor glass membrane, two figuresare given, the reaction side (shell side, fig.3) and the permeate side (tube side,fig.4). The Vycor glass membrane is permeable to all gaseous components,therefore it allows the reactants also to pass through along with all products and





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diluent. From fig.3 it can be noted that in reaction side of the membrane reactor,the trends of flow rate profiles of all components are similar to fixed bed reactorexcept of Ar. The profiles differ only in values of flow rates. In membrane reactorthe flow rates of reactants CH 4 and CO 2 are lower while the flow rates of productsH2, CO and H 2O are higher than that obtained in fixed bed reactor. The validreason for these trends is as follows. The reforming reaction is a reversiblereaction and equilibrium conversion is thermodynamically limited. Thereversibility of the reaction limits the maximum conversion of methane in fixed

bed reactor. In such reactions, the preferential removal of one or more of the products during reaction induces reaction enhancements and overcomesequilibrium limitations. In dry reforming, the removal of H 2, CO, and H 2O fromreaction side by the membrane in the reactor shifts the equilibrium towards

product formation. As a result, the conversion of CH 4 further increases andcorrespondingly the flow rates of products increase and of reactants decrease.Therefore, the exit values of the flow rates of components differ from thoseobtained in fixed bed reactor. The exit values of flow rates of CH 4 and CO 2 in

0

5

10

15

20

25

30

35

40

0 0.01 0.02 0.03 0.04 0.05


F l o w r a

t e o f c o m

p o n e n t s

i n s h e l

l s i d e o f

m e m

b r a n e r e a c

t o r

( m o l

/ s )

H2

CO

Ar

CH 4

CO 2H2O

Fig 3: Axial variation of flow rate of components in shell side of MR

shell side of the membrane reactor are 9.4405 mol/sec and 6.3533 mol/sec andin fixed bed reactor are 10.7092 mol/sec and 8.3082 mol/sec respectively. Theflow rate of water in membrane reactor is increased up to 3.051 mol/sec. Thedifference between the flow rates of H 2 and CO is larger as shown in fig. 3,

because more CO is produced and in parallel more H 2 is consumed in RWGSreaction. The flow rate of Ar increases in the beginning of reaction (from 27 to





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30.9mol/sec) and then remains almost constant. In the present study Ar is alsoused as sweep gas in the permeate side of reactor. The high flow rate in the

beginning of reaction can be attributed to the high partial pressure of Ar at permeate side as compared to the reaction side. This results in the permeation ofAr from permeate side to reaction side of reactor. As the reaction proceeds, the

production of gaseous products and their permeation along with reactants through porous membrane, reduces the partial pressure of Ar at permeate side whichmakes the flow rate of Ar almost constant in each side of the reactor.

In permeate side (tube side) of the membrane reactor (fig. 4), the flowrates of the products (H 2, CO and H 2O) increase along the length of reactor. The

permeation of gaseous species through the porous Vycor membrane proceeds bythe mechanism of Knudsen diffusion. This mechanism predicts that permeance ofgaseous component is inversely proportional to the square root of its molecularweight, (1/Mi). Since molecular weight of the hydrogen is small as compared toCO and H 2O, it can permeate through the membrane at a higher rate than othercomponents. But as it is getting consumed in RWGS reaction, the driving forcefor permeation which is trans - membrane partial pressure difference of H 2, is

0

1

2

3

4

5

6

7

8

0 0.01 0.02 0.03 0.04 0.05

F l o w r a

t e o f c o m p o n e n t s i n

t u b e s i

d e o f

m e m

b r a n e r e a c

t o r ( m o l

/ s )


H 2

COAr

H 2O

CH 4CO 2

Fig 4: Axial variation of flow rate of components in tube side of MR





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0

10

20

30

40

50

60

70

0 0.01 0.02 0.03 0.04 0.05

% c

o n v e r s

i o n o f m e t

h a n e


MR FBR

lower than that of CO. As a consequence, the permeation of H 2 decreases andthereby flow rate becomes lower than that of CO. The reactants CH 4 and CO 2 get

permeated with high rates in the starting of the reactor as the conversion is lowand they are available in large amounts. As the reaction proceeds, the partial

pressures of reactants start decreasing, leading to lower permeation of reactants.As a result, the flow rates of CH 4 and CO 2 decrease along the length of reactor.The flow rate of Ar decreases by 3.9 mol/ sec as it is diffusing through themembrane to the shell side of the membrane reactor.

Fig. 5 shows the conversion profiles of methane in axial direction at 923 Kin fixed bed reactor and membrane reactor. This figure shows that the conversionincreases along the length of reactor. At 923 K, the equilibrium conversion has

been achieved at a reactor length of 0.04 m for both reactors. This fact suggeststhat the residence time corresponding to the reactor length of 0.04 m is sufficientto achieve the equilibrium in the reactor at prevailing operating conditions.Further, the equilibrium conversion of methane in membrane reactor (64.57%) is2.9% higher than that in fixed bed reactor (62.75%) at 923 K. The reason ofgetting higher conversion in membrane reactor is similar as discussed above withfigures 3 and 4.

Fig 5: Axial variation of % conversion of methane in MR and FBR





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Effect of temperature

Reaction temperature plays an important role in reactor performance. Fig. 6shows the variation in conversion of CH 4 with the temperature for fixed bedreactor and membrane reactor. Owing to the endothermic nature of reaction, the

percent conversion increases with increase in temperature in both reactors. Inmembrane reactor, the % conversion is higher than corresponding % conversionvalue obtained in the fixed bed reactor under the same condition caused by the

permeation of the products through membrane and thereby shifting of equilibriumtowards right of reaction. The results also show that the influence of thetemperature of membrane reactor is less at high temperature as the deviation

between the fixed bed conversion and membrane reactor conversion is greater atlower temperature than the deviation at high temperature. This may be mainly

because of the fact that the permeation rates of reactant through porous Vycorglass membrane also increases with temperature in conjunction with productsleading to the loss of reactants.

0

10

20

3040

50

60

70

80

840 860 880 900 920 940 960

Temperature (K)

% c

o n v e r s i o n o f m e t

h a n e

FBR MR

Fig 6: Effect of temperature on conversion

The influence of temperature on yield of CO and H 2 is illustrated in fig. 7.The yields of CO and H 2 increase with the increase in temperature. The yield of





27/39

CO is higher than that of H 2 because of getting produced by both the reactionsand not consumed in any reaction. On the other hand, H 2 is produced via dryreforming reaction and consumed in RWGS reaction. The yield of H 2 slightlydecreases at higher temperature of 948 K. This decrease could be due to the

0

0.2

0.4

0.6

0.8

1

1.2

1.4

1.6

1.8

840 860 880 900 920 940 960

Temperature (K)

Y i e l d o f c o m p o n e n

t s i n m e m

b r a n e r e a c

t o r

H 2

CO

Fig 7: Effect of temperature on yield of components in MR

higher value of equilibrium constant at high temperature for RWGS reaction. Thedeviation in the yield of CO and H 2 is less at low temperature and is high at hightemperature. This indicates that the H 2/CO ratio decreases with temperature. Asfar as the selectivity is concerned, the selectivity of CO is higher than theselectivity of H 2 at all temperatures (fig. 8). There is no pronounced effect oftemperature on selectivity of CO and H 2. The selectivity of CO and H 2 variesaround 0.52 and 0.45 respectively.





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0.44

0.45

0.46

0.47

0.48

0.49

0.50.51

0.52

840 860 880 900 920 940 960

Temperature (K)

S e l e c t i v i

t y o f c o m p o n e n t s

i n m e m b r a n e

r e a c

t o r

H 2

CO

Fig 8: Effect of temperature on selectivity of components in MR

Effect of sweep gas flow rate

When performing the reaction in MR by using the porous Vycor glass membrane,the benefit of preferential product removal by using sweep gas is not so muchappreciable due to permeation of both reactants and products through porousmembrane (Gallucci et al., 2008). However, the studies of (Aparicio et al., 2002)reveal that even for porous membrane, high sweep gas flow rates induce largechanges in the distribution of all gaseous components at both sides of themembrane. This change depends on the membrane permeance and not only

provides moderate conversion enhancement but also maximizes the selectivity ofH2 by suppressing the RWGS reaction. On this ground, in the present study, theeffect of sweep gas flow on the performance of MR has been investigated. Theinert gas Ar is used as diluent as well as sweep gas for maintaining the sufficient

partial pressure of Ar at permeate and reaction sides of the reactor so that thelarge amount of permeation of Ar through porous membrane from either side ofmembrane can be avoided. Fig 9 illustrates the CH 4 conversion in MR as afunction of sweep gas flow rate for three different flow rates of diluent Ar rangingfrom 27 47 mol/s using feed with CO 2/CH 4 ratio close to unity (24 mol/s : 24mol/s) at 923 K. It is noteworthy from this figure that % conversion of CH 4





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increases with increase in the flow rate of sweep gas and diluent Ar. The increasein conversion becomes insignificant at sufficiently high sweep gas flow rate. Thisobservation could be explained by examining the gaseous component partial

pressure as follows. The increase of sweep gas flow rate depletes the partial pressure of reactants (CH 4, CO 2) and products (CO, H 2, H 2O) on permeate sideand thereby increases the driving force for the permeation of these gases resultingin the enhancement of methane conversion. However, as the sweep gas flow ratekeeps on increasing, the reduction in partial pressure of gaseous components in

permeate side becomes pronounced and finally the partial pressures becomenegligibly small. The resulting permeation flux remains almost constant which inturn keeps the % conversion of methane almost constant.

Simulation has also been carried out on MR operating without sweep gas.The results obtained at various sweep gas and diluent flow rates are compiled inTable 5. It can be seen that conversion of CH 4 increases with increase in flow

Table 5: Percent CH 4 conversion at various sweep gas and diluent flowrates

Diluentflow ratemol/sec

% conc. ofCH 4 withoutsweep gas

With sweep gas % increaseinconversion

Max. conv.of CH 4

Optimum sweep gasflow rate mol/sec

27 65.1 66.6 67 2.337 66.2 67.3 57 1.6647 66.9 67.7 47 1.20

rate of diluent. The optimum value of sweep gas flow rate indicates the maximumflow rate above which improvement in percent conversion of CH 4 becomesinsignificant at given conditions. It can be seen from table 5 that conversion ofCH 4 increases with increase in sweep gas flow rate and diluent flow rate.However, the percent increase in conversion and optimum value of sweep gasflow rate decrease with increase in diluent flow rate. From these observations it,can be concluded that there is only moderate enhancement in the conversion ofCH 4 on increasing the sweep gas flow rate.





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63.5

64

64.5

65

65.5

66

66.5

67

67.5

68

0 7 17 27 37 47 57 67

Sweep gas flow rate( mol/s)

% c o n v

e r s i o n o f m e t h a n e

Fig 9: Effect of sweep gas flow rate on conversion of methane

Effect of dilution ratio

The feed, a mixture of CH 4 and CO 2, is diluted by addition of Ar. Fig 10demonstrates the effect of dilution ratio on the conversion of methane at differentvalues of CO 2/CH 4 ratios in the feed. Three CO 2/CH 4 ratios, 1, 2, and 3, have

been considered by keeping the flow rate of CH 4 constant at 24 mol/sec andaccordingly varying the flow rate of CO 2. The flow rate of inert diluent Ar isvaried from 1 mol/sec to 75 mol/sec. The conversions of CH 4 at no diluent infeed are 59%, 69% and 87% for feed ratios of 1, 2 and 3 respectively. Fig. 10depicts that the effect of flow rate of Ar on the conversion of methane isinsignificant at all values of CO 2/CH 4 ratios. However at feed ratios of unity,there exists slight increase in conversion of CH 4 with increase in diluent.

47 mol/s

37 mol/s

27 mol/s





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0

10

20

30

40

50

60

70

8090

100

0 20 40 60 80

% c

o n v e r s

i o n o f m e t

h a n e

Flow rate of Ar in shell side of reactor ( mol/s)

CO 2/CH 4=1

CO 2/CH 4=2

CO 2/CH 4=3

Fig 10: Effect of dilution ratio on conversion of methane

Effect of CO 2 / CH 4 ratio

The influence of CO 2/CH 4 ratio on the yield, selectivity and H 2/CO ratio has beenshown in table 6. The ratio of CO 2/CH 4 in feed has been varied in two manners. Inthe first manner, the total flow rate of CO 2 and CH 4 (48 mol/sec) is kept constantand individual flow rates are varied accordingly. In the second manner, the flowrate of methane is kept constant at 24 mol/sec and the flow rate of CO 2 is varied.In the first case, the yield of H 2 and CO increases with increase in feed ratio. Onthe contrary, the selectivity of H 2 decreases, while, selectivity of CO increases.The H 2/CO ratio also decreases with increase in feed ratio. The yield of H 2 andCO is high due to high conversion at high CO 2/CH 4 ratio (fig.10). In the secondcase, the total flow rate of feed increases on increasing the CO 2/CH 4 ratio whichreduces the residence time for the reaction. The increase in conversion and so theyield of H 2 and CO with feed ratio, may be the result of two opposite effects. Thereduction in residence time decreases the conversion of methane. On the otherhand, high flow rate of CO 2 enhances the rate of RWGS reaction which as aresult, increases the consumption of product H 2. More consumption of H 2 togetherwith its removal through membrane increases the conversion of methane which





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overcomes the reduction in conversion caused by the lowering of residence time.Further, the selectivity of CO continuously increases due to continuous higher

production of CO as a result of increasing flow rate of CO 2 in feed. This fact leadsto the reduction in H 2/CO ratio with increasing CO 2/CH 4 ratio. The resultsdemonstrated in table 6 leads to the conclusion that although at fixed total flowrate of feed (1 st case), the yield of H 2 and CO is higher than the yield at varyingflow rate (2 nd case). H 2/CO ratio increases from 0.6527 to 0.9312 (1 st case) andfrom 0.7063 to 0.9164 (2 nd case) with decreasing CO 2/CH 4 ratio. Thus, the firstcase, where the total feed rate is kept constant shows favorable results in the

present study.

Table 6: Variation of H 2/CO ratio with feed ratio

Feed ratio(CO 2/CH 4)

Yield ofH 2

Yield ofCO

Selectivityof H 2

Selectivityof CO

H 2/COratio

1. Keeping total flow rate of feed (CH 4 + CO 2) constant (= 48 mol/s)0.33 0.5558 0.5968 0.4737 0.5087 0.93120.50 0.7747 0.8420 0.4694 0.5101 0.9200

1 1.1825 1.3549 0.4505 0.5165 0.87152 1.4557 1.9148 0.4043 0.5318 0.76023 1.4598 2.2363 0.3573 0.5475 0.6527

2. Keeping flow rate of methane (CH 4) constant (= 24 mol/s) 0.33 0.4860 0.5303 0.4860 0.5106 0.91640.50 0.7349 0.8037 0.4671 0.5109 0.9143

1 1.1825 1.3549 0.4505 0.5165 0.87152 1.4361 1.8399 0.4129 0.5290 0.78053 1.4462 2.0475 0.3811 0.5396 0.7063

7. CONCLUSIONS

A one dimensional isothermal mathematical model has been presented to analyzethe performance of membrane reactor (MR) packed with Rh/ -Al 2O3 catalyst.Porous Vycor glass membrane has been used. The performance of MR iscompared with conventional fixed bed reactor (FBR). The simulation study shows

that conversion of CH 4 is higher in MR than that of FBR at all temperatures dueto the continuous removal of products from the reaction side of reactor. The yieldof CO is higher than the yield of H 2 and both yields increase with increase intemperatures. The yield results indicate that H 2/CO ratio decreases with theincrease in temperature. The selectivity of CO is higher than the selectivity of H 2





33/39

at all temperatures. However, there is no pronounced effect of temperature onselectivity of CO and H 2. The percent conversion of CH 4 increases with theincrease in flow rates of sweep gas and diluent in feed and attains a maximumvalue. The percent increase in conversion with and without sweep gas decreaseswith increase in diluent flow. However, there is no significant enhancement inconversion of CH 4 on applying sweep gas. The flow rate of inert diluent Ar isvaried from 1 mol/sec to 75 mol/sec at three values of CO 2/CH 4 ratios. Theresults indicate that the effect of diluent flow rate on the conversion isinsignificant at all CO 2/CH 4 ratios. The ratio CO 2/CH 4 is varied from 0.33 to 3keeping total flow rate constant at 48 mol/sec. The highest H 2/CO ratio is foundto be close to unity at CO 2/CH 4 ratio of 0.33. The selectivity of CO increases andof H 2 decreases with the increase in feed ratio. The feed ratio is also varied from0.33 to 3 by keeping the flow rate of CH 4 constant at 24 mol/sec. The increasingand decreasing trends of the results are same as found in previous case. However,the yields of H 2 and CO and H 2/CO ratio are found to be higher in previous case.Besides, a critique on the effectiveness of the catalysts for dry reforming reactionhas been presented, which supports the use of Rh/ -Al 2O3 catalyst in the presentstudy.

NOTATIONS

d Thickness of the membrane, md p Membrane pore diameter, m

i D Effective permeability of ith component, mol/m 2.atm.s

i F Molar flow rate of ith

component in shell side of the reactor, mol/s,i o F Initial molar flow rate of i

th component in shell side, mol/s

,1i F Exit molar flow rate of ith component in shell side, mol/s

'i F Molar flow rate of i

th component in tube side, mol/s',i o F Initial molar flow rate of i

th component in tube side, mol/s',1i F Exit molar flow rate of i

th component in tube side, mol/s

i J Molar flux of ith component through membrane, mol/m 2.s

1k Rate constant for dry reforming reaction, gmol/gcat.s

2k Rate constant for RWGS reaction, gmol/gcat.s.atm

1 K Equilibrium constants for dry reforming reaction, (-)

2 K Equilibrium constants for RWGS reaction, (-)

4CH K Adsorption equilibrium constant of methane, atm -1

2CO K Adsorption equilibrium constant of carbon dioxide, atm -1





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i Molecular weight of ith component, g/mol

i P Partial pressure of ith component on shell side, atm

t P Total pressure on shell side, atm'

i P Partial pressure of ith component on tube side, atm

't P Total pressure on tube side, atm

r Pore radius of membrane, m1r Forward rate of dry reforming reaction, mol/m

3.s

2r Forward rate of RWGS reaction, mol/m3.s

'1r Net rate of dry reforming reaction, mol/m

3.s'

2r Net rate of RWGS reaction, mol/m3.s

ir Net rate of consumption/production of ith component

R Universal gas constant, J/mol K1 R Inner radius of the tube, m

2 R Outer radius of the tube, m'2 R Inner radius of the shell, m

T Reactor temperature, KL Reactor length, m

Greek letters

Porosity of the membrane, (-) Tortuosity of the membrane, (-)

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