The Production of Syngas by Dry Reforming in Membrane Reactor ...

I NTERNATIONAL J OURNAL OF C HEMICAL R EACTOR E NGINEERING Volume 6

2008

Article A109

The Production of Syngas by Dry Reforming in Membrane Reactor Using Alumina-Supported Rh Catalyst: A Simulation 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] ∗∗ Indian 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 Rh Catalyst: 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 fixed bed reactor (FBR) and a porous Vycor glass membrane reactor (MR) has been developed. The reactors are packed with 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 analyze the overall performance of MR, a detailed simulation study has been carried out to elucidate the effect of temperature, sweep gas flow rate, dilution ratio, feed ratio on percent conversion of methane, yield of hydrogen and carbon monoxide, and ratio of hydrogen to carbon monoxide in produced syngas. Besides, a comprehensive investigation on the effectiveness of the catalysts for methane reforming with carbon dioxide reaction has been made and presented. KEYWORDS: syngas, dry reforming, membrane reactor, Vycor glass membrane, alumina-supported Rh catalyst

∗

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


Kumar et al.: Production of Syngas by Dry Reforming

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, and also as environmentally clean and easily transportable fuel due to its abundance and enormous surplus in remote areas and underground resources. Natural gas is a methane enriched fuel (contains more than 80% of methane) and its use causes a rise in global concentration of green house gases, CH4 and CO2, in atmosphere, resulting in global warming. The green house effect for CH4 is more pronounced than for CO2. According to the studies of Mackenzie and Mackenzie (1995), the contribution of CH4 and CO2 accounts for three quarters of the total effect. In this regard, therefore, extensive efforts are being made to convert green house gases into valuable products such as syngas. Syngas, a mixture of H2 and CO, forms the feed stock in the chemical and petrochemical industries for the production of methanol, acetic acid, olefins, gasoline, MTBE, oxo-alcohols, and phosgene etc. In some cases either H2 or CO is utilized, for which H2 and CO are acquired from synthesis gas. The hydrogen is then used in fuel cells, in the production of urea and heavy water etc. However, the biggest consumer of H2 from syngas is ammonia synthesis. Recently it is being planned to utilize the hydrogen as a fuel for 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 great interest. Since CO2 is available in large quantities and at low costs, CO2 can be used in place of steam for reforming. Therefore, the dry reforming which is reforming of methane with CO2 seems to be a promising technology for the production of syngas. The synthesis gas produced by steam reforming has high H2/CO ratio which is not suitable for Fischer-Tropsch synthesis in the production of long chain higher hydrocarbons due to the excess hydrogen which suppresses chain growth and decreases the selectivity of higher hydrocarbons (Hou et al., 2006). Conversely, methane reforming with CO2 plays an important role in the industries due to the production of syngas with a low H2/CO ratio (≈ 1.0) which can be preferentially used for production of liquid hydrocarbons in FischerTropsch synthesis network (Luna and Iriarte, 2008). The dry reforming is carried out with excess CO2 to promote reverse water gas shift reaction (RWGS) which results in lower H2/CO ratio (Rostrup-Nielsen et al., 1984). Dry reforming reaction is slightly more endothermic than steam reforming. Therefore, it is favored by low pressure and high temperature (Gadalla and Bower, 1988). The presence of CO2 gives rise to more chances of carbon formation on catalyst surface due to production of CO and consumption of H2 via RWGS reaction. Therefore, the major problem encountered in the application of dry reforming is


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International Journal of Chemical Reactor Engineering

Vol. 6 [2008], Article A109

rapid deactivation of the catalyst. The stability and deactivation constraints render the application of conventional, inexpensive and easily available supported Ni catalysts difficult owing to their high activity for carbon formation. The noble metal catalysts are usually found to be less sensitive to coking. Recently, for dry reforming reaction, excellent stability, high activity, coke resistance ability and activity of Rh/Al2O3 catalyst have been reported in the literature (Richardson and Paripatyadar, 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 based catalysts has been presented in this work in order to justify the applicability of noble metal based catalysts over Ni based catalyst in terms of activity and stability. Keeping this in view, in the present study, Rh/γ-Al2O3 catalyst has been adopted for dry reforming reaction in a membrane reactor. Dry reforming is reversible and endothermic reaction. The conversion is limited by thermodynamic equilibrium. Thus, in order to achieve high conversion, it should be carried out at high temperature. Another possibility to increase the conversion is the continuous and selective removal of one of the products which in turn enhances the forward reaction rate and shifts the equilibrium towards the right side. The inert membrane reactor packed with catalyst implements this concept to improve the conversion by providing the reaction and separation in a single 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 in the range of 2-50 nm. The examples of mesoporous membranes are Vycor glass and composite membrane of γ-alumina. The composite membranes are supported by α-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 gaseous components is governed by Knudsen diffusion. In the present study a one dimensional steady state isothermal mathematical model for dry reforming reaction carried out in a membrane reactor incorporating Vycor glass membrane has been developed. A comprehensive investigation on the effectiveness of the catalyst for dry reforming reaction has been presented. After making the close observation on the brief review on catalyst, no carbon deposition has been observed on the surface of Rh catalyst surface even after a long period of operation. Therefore in the kinetic model all reactions responsible for carbon deposition, for instance CH4 cracking and Boudouard reaction, have been excluded. Thus dry reforming and RWGS reactions have been considered in the present study. Simulation results have been taken out by using ODE solver in MATLAB-7. The performance of two reactor configurations namely commercial fixed bed reactor (FBR) and membrane reactor



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(MR) has been compared. The axial variation in % conversion of CH4 and flow rates of different reaction components have been illustrated at 923 K. The effect of temperature on % conversion of CH4, selectivities and yields of H2 and CO has been investigated. In membrane reactor the sweep gas (Ar) flow rate is varied and its influence on conversion of CH4 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 CO2/CH4 ratios. Another important study has been carried out to find out the effect of various feed compositions on H2/CO ratio in the produced synthesis gas. 2. CATALYST The methane reforming with CO2 has received considerable attention in recent years due to its potential applications in the preservation of the environment and in the production of useful products. However, one major and serious problem of this reaction is the deactivation of catalysts caused by carbon formation and its deposition on the surface of catalyst. The formation of inactive carbon during dry reforming reaction may occur via two reactions; Methane decomposition and Boudouard reaction. These are as follows: CH4 ↔ C + 2H2

(Methane decomposition)

2CO ↔ C + CO2

(Boudouard reaction)

The decomposition of carbonaceous species on the catalyst suppresses the active sites available on the surface of catalyst, blocks the catalyst pores and voids leading 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 hot tubes, the uneven flow distribution causes a self accelerating situation with further over heating of the hot tubes (Rostrup-Nielsen, 1993). Therefore, carbon formation can not be tolerated in tubular reformers. Many researchers have explored the approaches, which can improve the coke resistance on the catalyst. These approaches include selection of appropriate catalyst, sulfur passivation of the catalyst, nature of appropriate support, addition of promoters, metal oxide and metal additives with strong Lewis basicity, change of reaction conditions, and the addition of steam or H2 (Luna and Iriarte, 2008; and Pechimuthu et al., 2007). A variety of catalysts have been developed for methane reforming reaction. During the past decade, inexpensive and easily available Ni-based catalyst is more extensively developed (table1). These catalysts have shown very high 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 the deposition of stable and inactive carbon on the active centers of the surface. Recently several studies have been focused on VIII group metals (with the exception of Osmium) specially noble metal catalysts which are known to form less coke under reforming reaction or rather show high selectivity for carbon free operations. A variety of supports on noble metal catalysts for CH4-CO2 reforming have 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 metals especially molybdenum and tungsten have gained considerable interest in various reaction due to abundance of their precursors (Iyer et al., 2003). These carbides are found to be stable at elevated pressure and are moderately resistant to coking. The addition of second metal could result in improvements in activity and stability. For instance, cobalt-tungsten η-carbide is found to be active, stable for at least 150 h and, provide high conversion with H2/CO ratio close to unity. The influencing factors for catalyst behavior include the activity of catalyst towards the conversion, stability, coking resistance and the type of deposited carbon. All these factors depend on the nature of different supports/promoters and interaction between metal and supports/ promoters (Wang and Lu, 1998). There are usually three types of coke formed in dry reforming reaction on supported metal catalyst: polymeric, filamentous and graphitic coke. The polymeric coke may be derived from thermal decomposition of hydrocarbons. Filamentous and graphitic coke is formed on the catalyst. The coke can also be characterized on the basis of its reactivity with H2, H2O and O2 (Guo et al., 2007).For Ni-catalysts, the unreacted carbon residues are dissolved in the metal to generate filamentous carbon causing a significant expansion of the catalyst bed. The contact between metal and support is lost and so it is difficult to regenerate the catalyst system. Further studies indicated that the nature of carbon deposits is a function of support. For instance, on Ni/α-Al2O3 catalyst, high amount of filamentous type of carbon has been observed whereas on Ni/α-Al2O3-ZrO2, smaller amount of graphitic carbon and carbon nanotubes have been observed. The studies on Rh and Ru based catalysts carried out by Rezaei et al. (2006) elucidate the effect of type of carbon on activity and stability of catalyst. According to this study, the higher activity and stability of Ru and Rh catalysts may be due to the formation of highly reactive carbon. This carbon is superficial carbidic carbon and is the reaction intermediate which quickly forms CO. The reactivity of superficial carbidic carbon influences the catalyst activity. The coke deposition analysis on noble metal showed that order of coke deposition is Pt>Pd>Ir>Ru>Rh, indicating minimum on Rh (≈ 0.1%). The observation of Richardson and Paripatyadar (1990) showed that Rh was inactive for Boudouard reaction but Ru was deactivated quickly by same reaction. In the study of Richardson and Paripatyadar (1990), the slight deactivation of Rh was observed probably due to the presence of small amounts



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 the activity of the catalyst was maintained. This type of deactivation is often observed when catalyst loading is very small. The deactivation by impurities is most pronounced at low temperature. Large beds never show this effect and the catalyst can be operated for longer period without any deactivation. The addition of support/promoters greatly suppresses the carbon deposition. A large number of studies have been carried out using the different supports and promoters as mentioned in table 1. The studies of Pompeo et al. (2005) reveal that α - Al2O3 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 to improved resistance to carbon decomposition (Pompeo et al., 2005).The excess steam and promoters control the carbon formation by producing CO. The partial sulfiding also reduces the carbon formation but small amount of steam is still needed (Richardson and Paripatyadar, 1990). Rezaei et al. (2008) have proposed that addition of basic promoters can affect the metal support interaction and enhance 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 rare earth metal oxides as support and promoters such as CeO2 also improves the behavior of alumina based catalysts by increasing the activity, stability and carbon resistance. ZrO2 as a catalyst modifier activates adsorption of CO2 and promotes the gasification of deposited carbon. Many research workers have studied the reforming reaction capacity of Rh and Ru catalysts (table 1). The studies of Rezaei et al. (2006) indicate that Rh and Ru catalysts show a very high stability without any decrease in CH4 conversion with time. Rh and Ru catalysts exhibit highest active metal surface area and there is no change observed in pore size distributions after use, indicating the high stability of catalyst. The activities for Rh and Ru are observed ten times those of Ni, Pt and Pd (Richardson and Paripatyadar, 1990). The conversion and deactivation tests indicate that Rh is much more stable catalyst than Ru and deactivation decreases with increasing temperature. In addition to this, the observation also shows that Rh is inactive for Boudouard reaction. Tsipouriari et al. (1994) in their study also mentioned that Rh and Ru are far more active catalysts than Ni. In all experimental studies mentioned in table 1, no carbon formation by Rh based catalyst occurred in any of the experiments which confirms the ability of Rh to perform the reaction without carbon formation. From the above brief comments and information given in table 1, it can be inferred that it is not necessary to consider carbon formation reactions (methane cracking and Boudouard reaction) with Rh supported catalyst in dry reforming reaction. The most effective support for Rh catalyst is found to be


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Al2O3 which is followed by Rh/TiO2, Rh/SiO2 and Rh/MgO (Richardson and Paripatyadar, 1990; Richardson et al., 2003; and Erdohelyi et al., 1993) in decreasing order. Rh/γ-Al2O3 exhibits higher activity than other Rh/Al2O3 catalysts because active carbon species formed on the catalyst can participate in sequence of steps to form CO (Rezaei et al., 2008).Therefore in the present research work, the Rh/Al2O3 catalyst is selected to carry out the dry reforming reaction in membrane reactor. Table 1: Catalysts for dry reforming reaction Reference

Catalyst

Support

Promoter

Activity/selectivity/ stability/carbon deposition

Vannice and Garten (1979)

Ni Ni Ni Ni

TiO2 SiO2 Al2O3 Graphite

-

Ni/TiO2 showed higher activity than Ni/Al2O3.

Seshan et al. (1994)

Ni Ni Pt Pt

γ-Al2O3 ZrO2 γ-Al2O3 ZrO2

-

Pt/ZrO2 showed high stability in contrast with Pt/γ-Al2O3 and those based on Ni.

Luna and Iriarte (2008) Ni Ni Ni Ni

Al2O3 Al2O3 Al2O3 Al2O3

K Ca Mn Sn

Ca, Mn and Sn-modified catalysts showed a dramatic reduction in catalytic activity and significant increase in carbon deposition while Kmodified catalyst showed low carbon deposition and high catalytic activity.

Therdthianwong et al. (2008)

Ni

Al2O3 (15%)

ZrO2

The addition of ZrO2 greatly improved the stability of Ni/Al2O3 in terms of coke inhibition.

Swaan et al. (1994)

Ni Ni Ni Ni Ni Ni Ni Ni

SiO2 La2O3 MgO ZrO2 TiO2 Al2O3-SiO2 SiO2 SiO2

Lemonidou and Vasalos Ni (5%) (2002)

Rate of deactivation of Ni/Al2O3-SiO2 and Ni-Cu/SiO2 was much higher. The activity of Ni/MgO, Ni/TiO2 and NiK/SiO2were found to be close. K Cu

CaO (21.5%) -Al2O3 (78.5%)

The catalyst exhibited high activity and very good stability at stoichiometric CH4/CO2 feed. Despite the amount of coke deposited on the catalyst, no loss of activity was observed. Contd…..



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Table 1 continued…..

Reference

Catalyst

Support

Promoter Activity/selectivity/ stability/carbon deposition

Pechimuthu et al. (2007)

Ni

CeO2-(γ-Al2O3)

K

Appreciable deactivation was not observed for the catalyst at 650, 700 and 750 oC for 60-h run. The carbon formed on catalyst after 60-h at 700 and 750ºC dispersed well and could not be observed while at 650 oC, a significant amount of carbon mainly graphite coke was deposited on the catalyst. However the deposited coke did not affect the high activity of the catalyst. Therefore, the catalyst had a high stability at o 650 C despite the deposition of coke.

Pompeo et al. (2005)

Ni Ni Ni Pt Pt Pt

Al2O3 α-Al2O3-ZrO2 ZrO2 Al2O3 α-Al2O3-ZrO2 ZrO2

-

Pt/ZrO2, Pt/α-Al2O3-ZrO2 and Ni/α-Al2O3-ZrO2 system had lower deactivation levels than Ni/ZrO2, Ni/Al2O3 and Pt/Al2O3. The lowest deactivation level was in Ni/αAl2O3-ZrO2 and Pt/α-Al2O3ZrO2. The highest carbon formation was observed for Ni/Al2O3 and Pt/Al2O3 catalysts.

Sakai et al. (1984)

Ni (10%) Rh (5-10%) Pd (5-10%) Pt (5-10%) Ru (5-10%)

SiO2 Al2O3 Al2O3 Al2O3 Al2O3

-

Ni-SiO2 and Rh-Al2O3 were excellent in selectivity and activity. The selectivity of Rh catalyst was higher than that of Pd, Pt and Ru catalysts. The order of activity was: Rh>Pd>Pt>Ru.

Rostrup-Nielsen and Hansen (1993)

Ni Ru Rh Pt Pd Ir

MgO MgO MgO MgO MgO MgO

-

Ru and Rh catalyst displayed high selectivity and no deposition of carbon was observed. A rapid carbon formation was on Pd catalyst at temperatures > 600 oC. A slow carbon formation occurred on Ir and Pt at temperatures > 750 o C.

Hou et al. (2006)

Ru (5%) Rh (5%) Pt (5%) Pd (5%) Ir (5%) Ni (10%)

α-Al2O3 α-Al2O3 α-Al2O3 α-Al2O3 α-Al2O3 α-Al2O3

-

Noble metals (Ru, Rh, Pd, Ir and, Pt) showed higher coke resistance ability while their activity was lower than that of Ni and Co (10%). Ru dispersed highly on mesoContd…



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Reference

Catalyst

Support

Promoter


Co (10%)

α-Al2O3

Diaz et al. (2007)

Ni (5%)

Activated carbon

Ca (1%)

Ca played a Co-support role inhibiting the deactivation of the catalyst at long periods of reaction. Methane conversion was up to 40% at mild experimental conditions.

Rezaei et al. (2006)

Ru Rh Ir Pt Pd

Spinel Spinel Spinel Spinel Spinel

-

Ru and Rh showed the highest activity for methane dry reforming. The order of activity was found to be Rh≈Ru>Ir>Pt>Pd. The results obtained reported a high stability for Ru, Rh and Pt catalysts and lowest for Pd due to the formation of less reactive deposited carbon on Pd and sintering.

Munera et al. (2003)

Rh (2%) Pt (1%)

La2O3 (100%) La2O3 (100%)

-

The activities of both catalysts were similar. The stability of Rh/La2O3 > Pt/La2O3. Very low amount of carbon formation was observed on both, Rh and Pt catalysts after more than 100-h on stream.

Erdohelyi et al. (1993)

Rh Rh Rh Rh

Al2O3 TiO2 SiO2 MgO

-

The order of activity was found to be: Rh/Al2O3> Rh/TiO2>Rh/SiO2>Rh/MgO. No deactivation and no carbon deposition was observed on Rh/Al2O3. Contd…

porous Al2O3 and exhibited higher coke resistance and higher reforming ability. Supported Pt and Ru catalysts showed poor stability. It might be due to the sintering of these metals at high temperatures. No coke deposition was detected on supported Pt, Ir, and Rh used catalyst while some amount of coke was detected on Pd/α-Al2O3 catalyst. Among coke free catalysts, Rh/α-Al2O3 exhibited the highest reforming activity.



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Reference

Catalyst

Support

Promoter


Qin and Lapszewicz (1994)

Ru Rh Pt Ir Pd

MgO MgO MgO MgO MgO

-

For CO2 reforming, the order of activity was found to be: Rh>Ru>Ir>Pt>Pd. Very less amount of carbon was formed on Rh, Ru and Ir catalyst but some carbon black was observed on test tube for Pt and Pd catalysts.

Sehested et al. (2001)

Mo2C Ru (18%)

MgAl2O4 MgAl2O4

-

Activity of Ru was greater than Mo2C catalyst by more than 2 orders of magnitude. Resistance to carbon of Mo2C was higher than that of Ru. Mo2C was stable at only high conversion.

Richardson et al. (2003) Rh Pt -Re

γ-Al2O3 γ-Al2O3

-

Very low carbon deposition occurred on both of catalysts. Both the catalysts displayed excellent stability but activity loss for Pt-Re catalyst was observed at temperature below 700 oC.

Itoh et al. (1992)

Pt Pt Pt

Al2O3 ZrO2 x%-ZrO2- Al2O3 (x=1, 5, 10, 20)

-

Zirconia supported catalyst showed much higher stability even after 60-h run on stream. While Pt/Al2O3 deactivated significantly within 20-h on stream at 1073 K. The amount of carbon deposition on Pt/Al2O3 > on Pt/ZrO2.

Ballarini et al. (2005)

Pt Pt Pt Pt

Al2O3 Na-Al2O3 K-Al2O3 ZrO2

-

Darujati and Thomson (2005)

Mo2C Mo2C Mo2C

γ-Al2O3 ZrO2 MgO

-

Both Pt/Na-Al2O3 and Pt/ZrO2 catalysts showed a good activity and selectivity with a very high catalytic stability at 1073 K. Pt/Al2O3 showed a poor performance and lower stability due to carbon deposition. Pt/K-Al2O3 displayed lower conversion than Pt/Na-Al2O3 catalyst. Surface area and thermal stability and had a much higher activity than a bulk Mo2C catalyst, even though deactivation via oxidation also occurred. The ZrO2 support experienced serious sintering and led to subsequent deactivation. Contd….



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Reference

Catalyst

Support

Promoter


Brungs et al. (2000)

Mo2C Mo2C Mo2C Mo2C

Al2O3 SiO2 ZrO2 TiO2

-

The catalyst stability Mo2C/Al2O3>Mo2C/ZrO2>Mo2 C/SiO2> Mo2C/TiO2. The catalyst with Al2O3 and ZrO2 support did not show the appreciable sign of deactivation for the time period of the study and thus the author concluded as a promising system for methane dry reforming.

3. MEMBRANE In recent years the development of inorganic membranes has paved the way for the application of membrane in the high temperature reactors as these have good mechanical stability and have high resistance to temperature and corrosive environment (Zaman and Chakma, 1994). Conversely, organic polymeric membranes have limited stability at temperatures over 100 ºC (Li et al., 1988). The inorganic membranes can be classified into two types according to mechanism of separation: (a) porous membrane and (b) non -porous dense membrane. Inorganic porous membranes are found to have promising future for industrial applications in catalytic reactor and thereby currently are in great use. In past decades, many valuable reviews on porous inorganic membrane reactors have been published (Coronas and Santamaria, 1999; Dixon, 2003; Julbe et al., 2001). Porous inorganic membranes consist of small pores that are interconnected. Depending upon pore size, these membranes are of three types namely macroporous (dp > 50 nm), mesoporous (50 >dp > 2nm), and microporous (dp < 2nm). Owing to large pore size, macroporous materials such as α-alumina membranes are used only to support the layers of smaller pore size materials to form composite membranes. The mesoporous materials such as Vycor glass and γ-alumina, and microporous materials provide separation function. However, microporous materials such as carbon molecular sieves, porous silicas and zeolites offer very high separation factors (Dixon, 2003). The transport mechanism on porous membranes predominantly is governed by Knudsen diffusion. Therefore, these membranes allow all gases to pass through which results 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



membranes offer high permeability only for hydrogen whereas zirconia and perovskites are highly selective only for oxygen. The transport mechanism follows surface diffusion mechanism. However, on commercial scale their applicators are limited due to their high cost, low permeability because of high wall thickness of 100-150µm (Herman et al., 1997), difficulty in fabrication, sensitivity to poisoning by sulfur species and embrittlement upon aging (Julbe et al., 2001). The performance of dense membranes can be improved by decreasing the membrane thickness, by increasing the surface roughness and by developing the new materials. However, it is difficult to obtain a thin membrane with sufficient strength and satisfactory permeability. Therefore, supported membranes are 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 in terms 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 of withstanding high temperature and harsh chemical environment are made up of ceramic 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 925°C, 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 1000°C. 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 to avoid the shrinkage of the membrane and resulting stresses. In view of above studies on Vycor glass membranes, it is, therefore, worthwhile to use Vycor glass membrane as a potentially high performance membrane for reforming reaction applications upto temperature of 850 °C. In the present study, Vycor glass (7930 glass, Corning) membrane of tubular geometry manufactured by Prabhu et al.(1999) has been employed. The characteristics of this membrane have been evaluated experimentally by Prabhu et al.(1999) and summarized as follows. The normal pore size is in the range of 2-4 nm which is the lowest range of the mesoporous category of porous membrane. The ratio of mean free path / pore size is greater than 180 for all species in the operating reforming reaction temperature range. The effect of temperature and mass of gaseous components on the permeability has been found to be very close to that expected from Knudsen diffusion 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 walls rather than between gas molecules. Since the Knudsen diffusivity is inversely proportional to the molecular weight of the gaseous component, the lighter gaseous components diffuse through the membrane faster than the heavier one. The diffusion mechanism through mesoporous membrane has been presented in detail by Kumar et al. (2006) according to which the permeance of ith gaseous component through Vycor glass porous membrane can be written as: Di =

2 rε 3τRTd

8000 RT πM i

where Mi is molecular weight of ith component. Accordingly, the permeation flux for each component can be written as follows: For CH4:

J CH 4 =

2 rε 3τRTd

8000 RT ' ( PCH 4 − PCH ) 4 πM CH 4

For CO2:

J CO2 =

2 rε 3τRTd

8000 RT ' ( PCO2 − PCO ) 2 πM CO2

For CO:

J CO =

2 rε 3τRTd

8000 RT ' ( PCO − PCO ) πM CO

For H2:

J H2 =

2 rε 3τRTd

8000 RT ( PH 2 − PH' 2 ) πM H 2

For Ar:

J Ar =

2 rε 3τRTd

8000 RT ( PAr − PAr' ) πM Ar

For H2O:

J H 2O =

2 rε 3τRTd

8000 RT ( PH 2O − PH' 2O ) πM H 2O



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4. MATHEMATICAL MODEL

A schematic diagram of a tubular membrane reactor is presented in fig1. The catalytic 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 side zone which is a reaction zone packed with catalyst particles. The reaction occurs Reaction zone

Membrane

feed

product product

Sweep gas

Sweep gas + permeated gas

feed

product Permeation zone

Fig 1: Schematic diagram of a tubular membrane reactor

in this zone. Second is tube side zone, also called permeate zone where the sweep gas is introduced co-currently with respect to feed to carry away the permeated gases from the permeate zone. The feed contains mainly CH4, CO2 and Ar (as diluent) and is fed to the shell side of reactor. The chemically inert sweep gas Ar is introduced into the tube side of reactor. Therefore, in the permeate side (tube side) of the reactor, no chemical reaction occurs. The methane reforming of CO2 to produce synthesis gas can be represented by the following reactions. CH4 + CO2 ↔ 2 CO + 2H2 (Dry reforming)

Δ Ho298 = 247.4 kJ/ mol

(1)

CO2 + H2 ↔ CO + H2O (RWGS)

Δ Ho298 = 41 kJ/ mol

(2)

The reforming of methane with CO2 is reversible and endothermic in nature (Gallucci et al., 2007). The equation (2) represents reverse water gas shift reaction which occurs in parallel to dry reforming reaction as a side reaction. This reaction is also reversible and less endothermic in nature. There may occur other side reactions responsible for carbon formation via methane cracking and Boudouard reaction as discussed in previous section-2 on catalyst. Here, in present study since Rh/γ-Al2O3 catalyst has been employed in the 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



14


study. For dry reforming and RWGS reactions, the kinetics given by Richardson and Paripatyadar (1990) on Rh/Al2O3 has been utilized. The forward rate of dry reforming reaction ( r1 ) and RWGS reaction ( r2 ) are given by following equations: r1 = k1

K CO2 K CH 4 PCO2 PCH 4

(3)

( 1 + K CO2 PCO2 + K CH 4 PCH 4 ) 2

r2 = k2 ⋅ PCO

(4)

2

The reverse rate expressions are formulated by adding a term to the forward reaction rate so that at equilibrium net rate can become zero (Richardson and Paripatyadar, 1990). As a result, the net rate of dry reforming ( r1' ) and RWGS ( r2' ) reactions can be written as follows:

r = k1 ' 1

K CO2 K CH 4 PCO2 PCH 4 ( 1 + K CO2 PCO2 + K CH 4 PCH 4 ) 2

r2' = k 2 PCO2 ( 1 −

PCO PH 2O K 2 PH 2 PCO2

[1 −

( PCO2 PH 2 ) 2 K 1 K CH 4 PCO2

]

)

(5)

(6)

where k1 and k2 are the rate constants, K1 and K2 are the equilibrium constants for reaction (1) and reaction (2) respectively, Pi is the partial pressure of ith component and K CO2 and K CH 4 are the appropriate adsorption equilibrium constants of carbon dioxide and methane, respectively. The temperature dependence of these constants is as follows:

k1 = 1290e( −102065/ RT )

(7)

k2 = 1.857e( −73105/ RT )

(8)

KCH = 2.60 ⋅10−2 e(40684/ RT )

(9)

4

KCO = 2.61⋅10−2 e(37641/ RT ) 2


(10)


15

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

ln K1 = −6.091ln T − 4.084 ⋅10−3T + 3.0665 ⋅10−7 T 2 +

28623 63050 + + 75.624 T T2

5852.30 0.582 ⋅105 + − 7.1977 ln K 2 = 1.86ln T + 2.70 ⋅10 T − T T2 −4

(11)

(12)

On the basis of stoichiometry of dry reforming and RWGS reactions, the rates of consumption/ formation of reaction species are given below. The net rate of consumption of CH4, rCH 4 = r1'

(13)

The net rate of consumption of CO2, rCO2 = r1' + r2'

(14)

The net rate of production of CO, rCO = 2 r1' + r2'

(15)

The net rate of production of H2, rH 2 = 2 r1' - r2'

(16)

The net rate of production of H2O, rH 2O = r2'

(17)

In order to develop one dimensional mathematical model for membrane reactor, the length of reactor is divided into small elemental length segments of size dz keeping cross sectional area of reactor constant. The material balance for each component has been taken around this control volume and equations have been formulated. These model equations may also be directly formulated by using comprehensive model for membrane reactor given by Kumar et al. (2006). The balance equations rely on the following assumptions: 1) 2) 3) 4) 5) 6)

Steady state operation Isobaric condition Isothermal operation Ideal gas law holds true Plug flow behavior (no radial concentration gradient) Resistances by gas film on both sides of membrane and by membrane support to permeation are negligible.



16


Shell side material balance equation

The material balance for ith component is as follows:

−

dFi ± (ri ) ⋅ area − Ji ⋅ 2π R1' = 0 dz

(18)

There are 6 components in shell side of reactor namely CH4, CO2, CO, H2, H2O, and Ar. The equation (18) can be written for each component as follows:

dFCH 4 2 + π (rCH 4 )( R22 − R '1 ) + J CH 4 (2π R '1 ) = 0 dz

(19)

dFCO2 2 + π (rCO2 )( R22 − R '1 ) + J CO2 ⋅ (2π R '1 ) = 0 dz

(20)

2 dFCO − π (rCO )( R22 − R '1 ) + J CO ⋅ (2π R '1 ) = 0 dz

(21)

dFH 2 2 − π (rH 2 )( R22 − R '1 ) + J H 2 ⋅ (2π R '1 ) = 0 dz

(22)

dFH 2O 2 − π (rH 2O )( R22 − R '1 ) + J H 2O ⋅ (2π R '1 ) = 0 dz

(23)

dFAr + J Ar ⋅ (2π R '1 ) = 0 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 contain only permeation term and no reaction term. The material balance equation for ith component can be written as:

dFi ' = J i ⋅ 2π R1 dz

(25)



17

The membrane is porous, therefore, all components present in the reaction side get permeated to tube side. Since Ar is used as sweep gas as well as diluent to feed, there are also 6 components viz. CH4, CO2, CO, H2, H2O, and Ar in tube side of reactor. The equation (25) can be written for each component as follows:

dF 'CH4 − JCH4 ⋅ (2π R1 ) = 0 dz

(26)

dF 'CO2 − JCO2 ⋅ (2π R1 ) = 0 dz

(27)

dF 'CO − J CO ⋅ (2π R1 ) = 0 dz

(28)

dF ' H 2 − J H 2 ⋅ (2π R1 ) = 0 dz

(29)

dF ' H 2O − J H 2O ⋅ (2π R1 ) = 0 dz

(30)

dF ' Ar − J Ar ⋅ (2π R1 ) = 0 dz

(31)

5. SOLUTION PROCEDURE

The mathematical model developed for membrane reactor (MR) consists of a set of twelve differential equations. These equations contain rate of reaction term and rate of permeation term for each component. In order to modify the model equations 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 terms are set to zero for all components. Thus, model for FBR consists of only six ordinary differential equations. The operating conditions combined with initial conditions are listed in table 2. For the simulation of reactor, the model equations are solved simultaneously in MATLAB-7 using ODE solver tool. The following definitions have been used for describing the performance of FBR and MR.



18

' FCH 4 ,0 − ( FCH 4 ,1 + FCH ) 4 ,1

Methane conversion (%) =

Selectivity of hydrogen =

Selectivity of CO =

Yield of CO =

Yield of H2 =


(32)

FCH 4 ,0 FH 2 ,1 + FH' 2 ,1

(33)

' FH 2 ,1 + FH' 2 ,1 + FH 2O ,1 + FH' 2O ,1 + FCO ,1 + FCO ,1 ' FCO ,1 + FCO ,1

(34)

' FH 2 ,1 + FH' 2 ,1 + FH 2O ,1 + FH' 2O ,1 + FCO ,1 + FCO ,1

' FCO ,1 + FCO ,1

(35)

FCH 4 ,0 FH 2 ,1 + FH' 2 ,1

(36)

FCH 4 ,0 Table 2: Operating and boundary conditions

Parameter L Pt

Value 0.04 m 1.0 atm

Parameter

r

Value 3.0 4.0 nm

P

1.0 atm

d

1.0 mm

R

8.314 J/mol K

FCH 4 ,o

24 μmol/s

R1

0.004 m

FCO2 ,o

24 μmol/s

R1'

0.005 m

FAr ,o

27 μmol/s

R2

0.007 m

FAr' ,o

7 μmol/s

T

923 K

0 μmol/s

ε

' FCH 4 ,o

0.6

' FCO 2 ,o

0 μmol/s

' t

τ

6. SIMULATION RESULTS

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



19

model, the experimental studies of Prabhu et al. (2000) on dry reforming reaction using Vycor glass membrane have been considered. The model equations have been solved at the same operating conditions used by Prabhu et al.(2000). The comparison of model results with experimental results has been made in terms of % conversion of CH4 at various temperatures for FBR and MR in tables 3 and 4 respectively. The conversion in MR appears to be higher than in FBR due to simultaneous removal of gaseous product (H2, CO and H2O) from feed side of membrane to the permeate side. Further, it is noteworthy from these tables that the difference between two results (experimental and model) is very small in the whole temperature range for FBR as well as for MR. In case of FBR, the % error varies 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 both FBR and MR. These results clearly show that the model predictions are in good agreement with experimental predictions. Thus model simulates the laboratory reactor very well. Table 3: Validation of model results of fixed bed reactor with experimental results Temperature (K) 848 873 898 923 948

Experimental results

Model results

Error (%)

36.7 45.2 54.3 62.6 70.2

36.5 45.05 53.48 62.75 68.65

0.54 0.33 1.5 - 0.24 2.2

Table 4: Validation of model results of membrane reactor with experimental results Temperature (K) 848 873 898 923 948

Experimental results

Model results

Error (%)

42.7 51.3 59.9 65.3 76.1

41.68 50.72 58.72 64.57 72.13

2.3 1.10 1.96 1.11 5.2




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. In fixed bed reactor (fig. 2), the flow rates of reactants CH4 and CO2 decrease continuously down the length of the reactor, until the equilibrium is achieved. The flow rate of CO2 is lower than the flow rate of CH4 because of the consumption of CO2 in dry reforming reaction as well as in RWGS reaction, and consumption of CH4 only in dry reforming reaction. The flow rate of inert diluent Ar in feed remains constant. Conversely, the flow rates of products H2 and CO increase down the length of reactor significantly as these are produced in more molar amounts as compared to molar consumption of reactants. The flow rate of H2 is lower than that of CO because some of the H2 produced in dry reforming reaction, gets consumed in (RWGS) to yield more CO and H2O whereas CO is produced from both reactions. The flow rate of water is very small and increases slowly along the length of reactor because it is produced in small quantity only by RWGS reaction. The end flow rate of water is about 2.2 μmol/sec. Flow rate of components in fixed bed reactor (μmol/s)

20

45 CO

40 35

H2

30 25

Ar

20 15

CH4

10

CO2

5

H2 O

0 0

0.01

0.02

0.03

0.04

0.05

Length of reactor (m)

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

For the membrane reactor utilizing the Vycor glass membrane, two figures are 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



21

Flow rate of components in shell side of membrane reactor (μmol/s)

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 reactor except of Ar. The profiles differ only in values of flow rates. In membrane reactor the flow rates of reactants CH4 and CO2 are lower while the flow rates of products H2, CO and H2O are higher than that obtained in fixed bed reactor. The valid reason for these trends is as follows. The reforming reaction is a reversible reaction and equilibrium conversion is thermodynamically limited. The reversibility 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 overcomes equilibrium limitations. In dry reforming, the removal of H2, CO, and H2O from reaction side by the membrane in the reactor shifts the equilibrium towards product formation. As a result, the conversion of CH4 further increases and correspondingly the flow rates of products increase and of reactants decrease. Therefore, the exit values of the flow rates of components differ from those obtained in fixed bed reactor. The exit values of flow rates of CH4 and CO2 in 40

CO

35 30

Ar

H2

25 20 15

CH4

10 5

H2O

CO2

0 0

0.01

0.02

0.03

0.04

0.05


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

shell side of the membrane reactor are 9.4405 μmol/sec and 6.3533 μmol/sec and in fixed bed reactor are 10.7092 μmol/sec and 8.3082 μmol/sec respectively. The flow rate of water in membrane reactor is increased up to 3.051 μmol/sec. The difference between the flow rates of H2 and CO is larger as shown in fig. 3, because more CO is produced and in parallel more H2 is consumed in RWGS reaction. The flow rate of Ar increases in the beginning of reaction (from 27 to



22


30.9µmol/sec) and then remains almost constant. In the present study Ar is also used 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 of Ar 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 which makes 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 flow rates of the products (H2, CO and H2O) increase along the length of reactor. The permeation of gaseous species through the porous Vycor membrane proceeds by the mechanism of Knudsen diffusion. This mechanism predicts that permeance of gaseous component is inversely proportional to the square root of its molecular weight, √(1/Mi). Since molecular weight of the hydrogen is small as compared to CO and H2O, it can permeate through the membrane at a higher rate than other components. But as it is getting consumed in RWGS reaction, the driving force for permeation which is trans - membrane partial pressure difference of H2, is

Flow rate of components in tube side of membrane reactor (μmol/s)

8 7 6 5 4 CO

Ar

3 2

H2 CO2

1

CH4 H2O

0 0

0.01

0.02

0.03

0.04

0.05


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



23

lower than that of CO. As a consequence, the permeation of H2 decreases and thereby flow rate becomes lower than that of CO. The reactants CH4 and CO2 get permeated with high rates in the starting of the reactor as the conversion is low and 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 CH4 and CO2 decrease along the length of reactor. The flow rate of Ar decreases by 3.9 μmol/ sec as it is diffusing through the membrane to the shell side of the membrane reactor. Fig. 5 shows the conversion profiles of methane in axial direction at 923 K in fixed bed reactor and membrane reactor. This figure shows that the conversion increases 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 suggests that the residence time corresponding to the reactor length of 0.04 m is sufficient to achieve the equilibrium in the reactor at prevailing operating conditions. Further, the equilibrium conversion of methane in membrane reactor (64.57%) is 2.9% higher than that in fixed bed reactor (62.75%) at 923 K. The reason of getting higher conversion in membrane reactor is similar as discussed above with figures 3 and 4. 70

% conversion of methane

60 50 40 30 20 MR

10

FBR

0 0

0.01

0.02

0.03

0.04

0.05


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



24


Effect of temperature

Reaction temperature plays an important role in reactor performance. Fig. 6 shows the variation in conversion of CH4 with the temperature for fixed bed reactor and membrane reactor. Owing to the endothermic nature of reaction, the percent conversion increases with increase in temperature in both reactors. In membrane reactor, the % conversion is higher than corresponding % conversion value obtained in the fixed bed reactor under the same condition caused by the permeation of the products through membrane and thereby shifting of equilibrium towards right of reaction. The results also show that the influence of the temperature of membrane reactor is less at high temperature as the deviation between the fixed bed conversion and membrane reactor conversion is greater at lower temperature than the deviation at high temperature. This may be mainly because of the fact that the permeation rates of reactant through porous Vycor glass membrane also increases with temperature in conjunction with products leading to the loss of reactants. 80


70 60 50 40 30 20

FBR

MR

10 0 840

860

880

900

920

940

960

Temperature (K) Fig 6: Effect of temperature on conversion

The influence of temperature on yield of CO and H2 is illustrated in fig. 7. The yields of CO and H2 increase with the increase in temperature. The yield of



25

Yield of components in membrane reactor

CO is higher than that of H2 because of getting produced by both the reactions and not consumed in any reaction. On the other hand, H2 is produced via dry reforming reaction and consumed in RWGS reaction. The yield of H2 slightly decreases at higher temperature of 948 K. This decrease could be due to the 1.8 1.6 1.4

CO

1.2 1

H2

0.8 0.6 0.4 0.2 0 840

860

880

900

920

940

960

Temperature (K)

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

higher value of equilibrium constant at high temperature for RWGS reaction. The deviation in the yield of CO and H2 is less at low temperature and is high at high temperature. This indicates that the H2/CO ratio decreases with temperature. As far as the selectivity is concerned, the selectivity of CO is higher than the selectivity of H2 at all temperatures (fig. 8). There is no pronounced effect of temperature on selectivity of CO and H2. The selectivity of CO and H2 varies around 0.52 and 0.45 respectively.



Selectivity of components in membrane reactor

26


0.52

CO

0.51 0.5 0.49 0.48 0.47 0.46

H2

0.45 0.44 840

860

880

900

920

940

960

Temperature (K) 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 much appreciable due to permeation of both reactants and products through porous membrane (Gallucci et al., 2008). However, the studies of (Aparicio et al., 2002) reveal that even for porous membrane, high sweep gas flow rates induce large changes in the distribution of all gaseous components at both sides of the membrane. This change depends on the membrane permeance and not only provides moderate conversion enhancement but also maximizes the selectivity of H2 by suppressing the RWGS reaction. On this ground, in the present study, the effect of sweep gas flow on the performance of MR has been investigated. The inert 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 the large amount of permeation of Ar through porous membrane from either side of membrane can be avoided. Fig 9 illustrates the CH4 conversion in MR as a function of sweep gas flow rate for three different flow rates of diluent Ar ranging from 27 – 47 µmol/s using feed with CO2/CH4 ratio close to unity (24 µmol/s : 24 µmol/s) at 923 K. It is noteworthy from this figure that % conversion of CH4



increases with increase in the flow rate of sweep gas and diluent Ar. The increase in conversion becomes insignificant at sufficiently high sweep gas flow rate. This observation 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 (CH4, CO2) and products (CO, H2, H2O) on permeate side and thereby increases the driving force for the permeation of these gases resulting in the enhancement of methane conversion. However, as the sweep gas flow rate keeps on increasing, the reduction in partial pressure of gaseous components in permeate side becomes pronounced and finally the partial pressures become negligibly small. The resulting permeation flux remains almost constant which in turn 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 in Table 5. It can be seen that conversion of CH4 increases with increase in flow Table 5: Percent CH4 conversion at various sweep gas and diluent flow rates

Diluent % conc. of With sweep gas % increase flow rate CH4 without Max. conv. Optimum sweep gas in µmol/sec sweep gas conversion of CH4 flow rate µmol/sec 27 65.1 66.6 67 2.3 37 66.2 67.3 57 1.66 47 66.9 67.7 47 1.20 rate of diluent. The optimum value of sweep gas flow rate indicates the maximum flow rate above which improvement in percent conversion of CH4 becomes insignificant at given conditions. It can be seen from table 5 that conversion of CH4 increases with increase in sweep gas flow rate and diluent flow rate. However, the percent increase in conversion and optimum value of sweep gas flow rate decrease with increase in diluent flow rate. From these observations it, can be concluded that there is only moderate enhancement in the conversion of CH4 on increasing the sweep gas flow rate.


27


28


% conversion of m ethane

68 47 µmol/s

67.5

37 µmol/s

67 66.5

27 µmol/s

66 65.5 65 64.5 64 63.5 0

7

17

27

37

47

57

67

Sweep gas flow rate( µmol/s) Fig 9: Effect of sweep gas flow rate on conversion of methane Effect of dilution ratio

The feed, a mixture of CH4 and CO2, is diluted by addition of Ar. Fig 10 demonstrates the effect of dilution ratio on the conversion of methane at different values of CO2/CH4 ratios in the feed. Three CO2/CH4 ratios, 1, 2, and 3, have been considered by keeping the flow rate of CH4 constant at 24 μmol/sec and accordingly varying the flow rate of CO2. The flow rate of inert diluent Ar is varied from 1 μmol/sec to 75 μmol/sec. The conversions of CH4 at no diluent in feed are 59%, 69% and 87% for feed ratios of 1, 2 and 3 respectively. Fig. 10 depicts that the effect of flow rate of Ar on the conversion of methane is insignificant at all values of CO2/CH4 ratios. However at feed ratios of unity, there exists slight increase in conversion of CH4 with increase in diluent.



29


100 90

CO2/CH4=3

80

CO2/CH4=2

70 60

CO2/CH4=1

50 40 30 20 10 0 0

20

40

60

80

Flow rate of Ar in shell side of reactor (μmol/s) Fig 10: Effect of dilution ratio on conversion of methane Effect of CO2 / CH4 ratio

The influence of CO2/CH4 ratio on the yield, selectivity and H2/CO ratio has been shown in table 6. The ratio of CO2/CH4 in feed has been varied in two manners. In the first manner, the total flow rate of CO2 and CH4 (48 μmol/sec) is kept constant and individual flow rates are varied accordingly. In the second manner, the flow rate of methane is kept constant at 24 μmol/sec and the flow rate of CO2 is varied. In the first case, the yield of H2 and CO increases with increase in feed ratio. On the contrary, the selectivity of H2 decreases, while, selectivity of CO increases. The H2/CO ratio also decreases with increase in feed ratio. The yield of H2 and CO is high due to high conversion at high CO2/CH4 ratio (fig.10). In the second case, the total flow rate of feed increases on increasing the CO2/CH4 ratio which reduces the residence time for the reaction. The increase in conversion and so the yield of H2 and CO with feed ratio, may be the result of two opposite effects. The reduction in residence time decreases the conversion of methane. On the other hand, high flow rate of CO2 enhances the rate of RWGS reaction which as a result, increases the consumption of product H2. More consumption of H2 together with its removal through membrane increases the conversion of methane which


30



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 CO2 in feed. This fact leads to the reduction in H2/CO ratio with increasing CO2/CH4 ratio. The results demonstrated in table 6 leads to the conclusion that although at fixed total flow rate of feed (1st case), the yield of H2 and CO is higher than the yield at varying flow rate (2nd case). H2/CO ratio increases from 0.6527 to 0.9312 (1st case) and from 0.7063 to 0.9164 (2nd case) with decreasing CO2/CH4 ratio. Thus, the first case, where the total feed rate is kept constant shows favorable results in the present study. Table 6: Variation of H2/CO ratio with feed ratio Feed ratio Yield of Yield of Selectivity Selectivity H2/CO (CO2/CH4) H2 CO of H2 of CO ratio 1. Keeping total flow rate of feed (CH4 + CO2) constant (= 48 μmol/s) 0.33 0.5558 0.5968 0.4737 0.5087 0.9312 0.50 0.7747 0.8420 0.4694 0.5101 0.9200 1 1.1825 1.3549 0.4505 0.5165 0.8715 2 1.4557 1.9148 0.4043 0.5318 0.7602 3 1.4598 2.2363 0.3573 0.5475 0.6527 2. Keeping flow rate of methane (CH4) constant (= 24μmol/s) 0.33 0.4860 0.5303 0.4860 0.5106 0.9164 0.50 0.7349 0.8037 0.4671 0.5109 0.9143 1 1.1825 1.3549 0.4505 0.5165 0.8715 2 1.4361 1.8399 0.4129 0.5290 0.7805 3 1.4462 2.0475 0.3811 0.5396 0.7063 7. CONCLUSIONS

A one dimensional isothermal mathematical model has been presented to analyze the performance of membrane reactor (MR) packed with Rh/γ-Al2O3 catalyst. Porous Vycor glass membrane has been used. The performance of MR is compared with conventional fixed bed reactor (FBR). The simulation study shows that conversion of CH4 is higher in MR than that of FBR at all temperatures due to the continuous removal of products from the reaction side of reactor. The yield of CO is higher than the yield of H2 and both yields increase with increase in temperatures. The yield results indicate that H2/CO ratio decreases with the increase in temperature. The selectivity of CO is higher than the selectivity of H2



at all temperatures. However, there is no pronounced effect of temperature on selectivity of CO and H2. The percent conversion of CH4 increases with the increase in flow rates of sweep gas and diluent in feed and attains a maximum value. The percent increase in conversion with and without sweep gas decreases with increase in diluent flow. However, there is no significant enhancement in conversion of CH4 on applying sweep gas. The flow rate of inert diluent Ar is varied from 1 μmol/sec to 75 μmol/sec at three values of CO2/CH4 ratios. The results indicate that the effect of diluent flow rate on the conversion is insignificant at all CO2/CH4 ratios. The ratio CO2/CH4 is varied from 0.33 to 3 keeping total flow rate constant at 48 μmol/sec. The highest H2/CO ratio is found to be close to unity at CO2/CH4 ratio of 0.33. The selectivity of CO increases and of H2 decreases with the increase in feed ratio. The feed ratio is also varied from 0.33 to 3 by keeping the flow rate of CH4 constant at 24 μmol/sec. The increasing and decreasing trends of the results are same as found in previous case. However, the yields of H2 and CO and H2/CO ratio are found to be higher in previous case. Besides, a critique on the effectiveness of the catalysts for dry reforming reaction has been presented, which supports the use of Rh/γ-Al2O3 catalyst in the present study. NOTATIONS

d dp Di Fi Fi ,o Fi ,1

Thickness of the membrane, m Membrane pore diameter, m Effective permeability of ith component, mol/m2.atm.s Molar flow rate of ith component in shell side of the reactor, mol/s Initial molar flow rate of ith component in shell side, mol/s Exit molar flow rate of ith component in shell side, mol/s

Fi '

Molar flow rate of ith component in tube side, mol/s

Fi ,'o

Initial molar flow rate of ith component in tube side, mol/s

Fi ,1' Ji k1 k2

Exit molar flow rate of ith component in tube side, mol/s Molar flux of ith component through membrane, mol/m2.s Rate constant for dry reforming reaction, gmol/gcat.s Rate constant for RWGS reaction, gmol/gcat.s.atm Equilibrium constants for dry reforming reaction, (-) Equilibrium constants for RWGS reaction, (-) Adsorption equilibrium constant of methane, atm-1

K1 K2

K CH K CO

4

2

Adsorption equilibrium constant of carbon dioxide, atm-1


31


32

Mi Pi Pt

Molecular weight of ith component, g/mol Partial pressure of ith component on shell side, atm Total pressure on shell side, atm

Pi '

Partial pressure of ith component on tube side, atm

Pt ' r r1 r2 r1' r2'

Total pressure on tube side, atm Pore radius of membrane, m Forward rate of dry reforming reaction, mol/m3.s Forward rate of RWGS reaction, mol/m3.s Net rate of dry reforming reaction, mol/m3.s Net rate of RWGS reaction, mol/m3.s Net rate of consumption/production of ith component Universal gas constant, J/mol K Inner radius of the tube, m Outer radius of the tube, m

ri

R R1 R2 R2' T L


Inner radius of the shell, m Reactor temperature, K Reactor length, m

Greek letters

ε

τ

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

REFERENCES

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