Zeolite membranes for hydrogen production from natural gas: state of ...

J Porous Mater (2015) 22:635–646 DOI 10.1007/s10934-015-9936-6

Zeolite membranes for hydrogen production from natural gas: state of the art Beata Michalkiewicz1 • Zvi C. Koren2

Published online: 11 March 2015 Ó Springer Science+Business Media New York 2015

Abstract Currently, hydrogen is produced industrially by processes requiring high energy consumption, especially by cracking fossil fuels and by splitting water. In recent years, research has been devoted to the use of membrane catalytic reactors in order to achieve higher hydrogen yield. Zeolite membranes have been shown to be very promising as they are stable to contaminants (such as H2S) and do not have the disadvantages that palladium and silica membranes exhibit. The aim of this work is to summarize the state of the art of hydrogen selective zeolite membranes that can be applied for hydrogen separation after syngas production from reforming streams. In addition, zeolite membrane applications in membrane reactors for hydrogen production are discussed. Keywords Zeolite membranes Membrane reactors Hydrogen separation Hydrogen production

& Beata Michalkiewicz [email protected] Zvi C. Koren [email protected] 1

2

Institute of Chemical and Environment Engineering, West Pomeranian University of Technology, Szczecin, Pułaskiego 10, 70-322 Szczecin, Poland Department of Chemical Engineering, Shenkar College of Engineering and Design, 12 Anna Frank St., 52526 Ramat-Gan, Israel

1 Introduction Hydrogen is one of the most promising energy carriers for the future as it is a high efficiency, zero-emission fuel because when burned, the product is only water [1]. The drawback is that hydrogen is not a primary fuel and always occurs in combined form. Hydrogen must first be produced by chemical processes from fossil fuels or water. The annual world hydrogen production is currently 5 9 1011 Nm3 [2], corresponding to around 2 % of primary energy demand [1]. The main sources for hydrogen production are fossil fuels, namely natural gas—48 %, higher hydrocarbons (mainly oil and naphtha)—30 %, and coal—18 %. Only 3.9 % hydrogen is obtained from water by electrolysis [3]. The conventional methods are equilibrium-limited processes, and hydrogen-rich gas mixtures containing substrates and by-products. This study focuses on the separation and production of hydrogen utilizing zeolite membranes. There are some reviews concerning zeolite membranes [4, 5] and membrane reactors for hydrogen production [6, 7], but there is no systematic and comprehensive review of zeolite membranes for hydrogen technology. The aim of this paper is to show the ability of zeolite membranes as very promising materials for hydrogen production, especially in water–gas shift (WGS) membrane reactors. The preparation, characterization and permeation of the various zeolite membranes will be presented and discussed.

2 Industrial processes for hydrogen production The dominant industrial method to produce hydrogen is the steam methane reforming (SMR) process. The conversion of hydrocarbons into hydrogen in the presence of steam

123

636

J Porous Mater (2015) 22:635–646

was first described by Tessie du Motay and Marechal in 1868 [8]. The SMR process is a mature technology, operating near the theoretical limits of the process. Nearly all the hydrogen produced in the chemical industry is by this method. It is a catalytic process that involves a reaction between natural gas, or other light hydrocarbons, and steam at temperatures of 700–1000 °C and pressures of 3–25 bar, as shown in reaction (1): SMR reaction : CH4 þ H2 O CO þ 3H2 DH 298 ¼ þ206 kJ/mol

ð1Þ

In a second reaction (2), the WGS reaction, one of the products of reaction (1)—carbon monoxide—is also reacted with steam to produce carbon dioxide and more hydrogen: WGS reaction : CO þ H2 O CO2 þ H2 DH 298 ¼ 42 kJ/mol

ð2Þ

The products of these reactions are hydrogen, carbon monoxide and carbon dioxide, and in order to obtain pure hydrogen other steps are needed. Chemical and physical scrubbings are commonly practiced in industry [9]. Both CO and CO2 can be removed simultaneously by methanation or by pressure swing adsorption. The reforming reaction (SMR) is highly endothermic and a large amount of heat is provided for that purpose. Usually the process is performed in multitubular fixed-bed reactors in the presence of a nickel catalyst. SMR efficiencies are currently in the range 60–80 % [10, 11], whereby the efficiency is defined as: nCH4 ;in nCH4 ;out ECH4 ¼ 100 % nCH4 ;in where nCH4 ;in and nCH4 ;out are the number of moles of CH4 at the inlet and at the outlet of the reactor, respectively. Natural gas is the most important and economical feedstock because of its abundant availability. However, steam reforming of higher hydrocarbons can be quite attractive especially at some locations where natural gas is not as abundant, such as, Japan and most European countries (except Norway, Netherlands, and the United Kingdom) [12]. Steam reforming of naphtha is well established in the chemical industry [13], and the reactions associated with steam naphtha reforming are basically the same as those in SMR. Due to the complexity of the heavy hydrocarbons comprising the naphtha fraction, additional complex series of reactions occur. Steam reforming of higher hydrocarbons can be represented by reactions (3)–(5): Cn Hm þ nH2 O ! nCO þ ðn þ m=2ÞH2

ð3Þ

CO þ 3H2 CH4 þ H2 O

ð4Þ

CO þ H2 O CO2 þ H2

ð5Þ

123

The reforming of higher hydrocarbons is an irreversible reaction and the products can be obtained at lower temperatures than those produced from methane. Normally, higher hydrocarbons are not converted to synthesis gas (syngas), but dehydrogenated to more valuable products [7]. Coal gasification is one of the oldest techniques for syngas and hydrogen production, and though the technology is mature, it is less-widely used than SMR. Coal is relatively cheap but the capital investment costs are higher and more variable than SMR. Efficiencies range from 50 to 80 % dependent on the quality of different types of coal. The processing of coal is initiated by first crushing it and feeding it into a gasifier, and then reacting it the gasified coal products with steam and air (or oxygen). The gasification occurs at temperatures as high as 1900 °C, and some processes used very high pressures of 10 MPa [14]. Typical industrial gasification processes today operate in the range 2.5-8 MPa and temperatures above 1250 °C, depending on the application. At these conditions, syngas contains less than 0.5 mol % methane [15]. The mixture of gases produced during gasification mainly consists of CO, H2, CO2, CH4, and N2, as the following reactions (6)–(13) show for the first four gases: coal ! charðCÞ þ coal volatilesðCVÞ

ð6Þ

CV þ H2 O ! CO þ H2

ð7Þ

CV þ H2 ! CH4

ð8Þ

C þ 2H2 ! CH4

ð9Þ

C þ H2 O ! CO þ H2

ð10Þ

C þ CO2 ! 2 CO

ð11Þ

CO þ H2 O ! CO2 þ H2

ð12Þ

C þ O2 ! CO2

ð13Þ

The hydrogen gas stream from coal gasification can be combusted to produce electricity in an integrated gaseous combined cycle plant [16, 17]. This technology allows electricity to be produced at times of peak demand and hydrogen to be produced at other times. In the first step, coal reacts with steam and/or oxygen in a gasifier to produce syngas. The syngas is cooled down to remove its contaminants (for example, H2S). The WGS reaction is exothermic so is favored by decreasing the temperature according to Le Chatelier’s principle. Two reactors are typically applied, where one operates at a high temperature and the other at a lower temperature. This combination produces a higher hydrogen yield [18]. The gas from the WGS reactors must be separated to produce pure hydrogen, and the total process is complex and energy intensive.

J Porous Mater (2015) 22:635–646

637

Technologies to produce hydrogen from biomass are very promising. Efficiencies are higher for biomass-derived biofuels that are processed prior to the hydrogen production plant. The principle advantage of such fuels would be to reduce the fuel transport costs from the plantation to the hydrogen plant. There are no completed industrial-scale demonstrations of any biomass technology for producing hydrogen [19]. The dry reforming of methane has not yet been applied in industry, but it has attracted much attention from both industrial and environmental sectors. The reaction equilibrium for the production of synthesis gas from CH4 and CO2 (reaction 14) is influenced by the simultaneous occurrence of the reverse water–gas shift (RWGS) reaction (15): CH4 þ CO2 2CO þ 2H2 DH 298 ¼ þ248 kJ/mol CO2 þ H2 CO þ H2 O

DH 298 ¼ þ42 kJ=mol

ð14Þ ð15Þ

The main drawbacks of the dry reforming of methane are the deactivation of the catalysts due to coking and the requirement of high temperature, 800 °C [20, 21]. In 1800, Nicholson and Carlisle discovered the ability of electrolytic water decomposition, and by 1902 there were more than 400 industrial water electrolysis units. Despite the discovery that the electrolytic decomposition of water was first observed in an acidic environment, an alkaline medium is preferred in industrial plants because corrosion is more easily controlled and cheaper [22]. Alkaline electrolysis has been used to produce hydrogen since the nineteenth century and is the basis of most commerciallyavailable electrolyzers. Extremely pure hydrogen is produced but at a considerably higher cost than from SMR due to the higher cost of electricity relative to fossil fuel feed stocks.

3 Membranes for hydrogen production Natural gas reforming is currently the most efficient, economical and widely used process for hydrogen production. Because the reforming reaction is highly endothermic, high temperatures need to be applied for hydrogen-rich gas production [23]. The equilibrium hydrogen concentration at 800 °C is far below 100 %. The hydrogen purification process, regardless of the method, requires high energy input [24]. Removal of the main products can drive the reaction towards completion in accordance with Le Chatelier’s principle. Membrane technology has been utilized globally for many decades in the oil refinery and fertilizer industries. The membrane reactor (MR) for SMR has also been demonstrated in small-scale plants [23, 25].

The major advantages of MR are shifting the chemical equilibrium toward the right-hand side of the reaction, improving hydrogen production, and being able to achieve the same methane conversion at relatively low temperatures, such as 650 °C [26]. A membrane is a selective physical barrier that permits the passage of one or more components of a stream through the membrane while retarding the passage of other components. The retentate is that part of the feed that does not pass through the membrane. The permeate is that part of the feed that does pass through the membrane. The selectivity and permeation rate (permeance) are the most basic properties of a membrane. Permeance is defined as a flow rate through the membrane per unit of area and per unit of pressure difference across the material (in units of mol s-1 m-2 Pa-1). Selectivity is a membrane’s ability to separate a desired component from the feed mixture and is often calculated as the ratio of pure gas permeabilities (ideal selectivity). Membranes can be divided into polymeric, metallic and ceramic membranes. Membranes for hydrogen production should be highly selective towards hydrogen, highly permeable, relatively cheap, and possess high thermal, mechanical and chemical stability [7]. Polymeric membranes cannot be used because of the low thermal and mechanical stability and poor chemical fastness. Only dense phase metal and porous ceramic membranes are suitable for hydrogen production [6]. The dense phase metallic membranes, especially palladium, offer very high permeability and exceptional H2 selectivity to produce nearly pure H2. However, Pd membranes have also serious drawbacks that have limited their applications. First, Pd is an expensive material. Additionally, palladium membranes can be poisoned by H2O, CO, and H2S, and even a 1 ppm concentration of H2S adversely affects membrane properties [27, 28]. At the higher temperature and pressure, Pd membranes undergo phase changes resulting in cracks [6]. Dense ceramic membranes with protonic and electronic conductivity were applied for hydrogen separation. High conductivity is necessary to obtain a high hydrogen flux. Dense ceramic membranes are mixed proton-electron conducting oxides, e.g., doped rare earth metal oxides [29] and fluorite-structured metal oxides [30] as well as perovskite-type oxides [31]. A very high purity H2 stream can be obtained but temperatures as high as 900–1400 °C are required. A shortcoming of such membranes is that they can be poisoned by H2S. Silica and silica functionalized ceramic membranes possess high permeability and moderate to high selectivity, and they are chemically and thermally stable. The drawback is instability at high temperatures [32], and especially in the presence of steam, silica membranes lose

123

638

permeability. At high temperatures, reduction in pore volume was observed due to condensation of :S–OH groups and sintering enhanced with the humidity [33]. The new classes of porous crystalline materials, Metal Organic Frameworks (MOFs) and Covalent-OrganicFrameworks (COVs), can also be utilized as membranes. The highly accessible porosity and the wide range of pore sizes of MOFs and COVs would make them possible candidates for the separation of hydrogen from other gases [34]. However, there are certain disadvantages of MOFs and COVs. Successful fabrication of polycrystalline MOF and COV membranes of sufficient quality is quite difficult because of the unfavorable heterogeneous nucleation and weakness of the coordination bonds. Such membranes do not show sufficient interfacial interaction with supports for membrane preparation [35]. The labile nature of the coordination bond on the lattice is a detrimental factor in the fabrication of COV and MOF membranes, and many MOFs show poor stability even under ambient conditions. There are no such problems with zeolite membranes. The zeolite membranes with high silica content, as well as silicalite, are hydrothermally stable and are resistant to sulfur compounds. Such zeolite membranes are relatively cheap and H2-selective over water vapor because of their hydrophobic surface. These properties make them potentially useful for hydrogen production [36, 37]. Membrane reactor applications for hydrogen production have been initially described by Rothfleisch in 1964 [38]. Hydrogen for fuel cells was obtained from an equimolar methanol–water mixture in a single bed catalyst. In the last 15 years, considerable growth in research of MRs for hydrogen production has been achieved (Fig. 1). The continuous polycrystalline zeolite layers for separations were first prepared in the 1990s [39]. In the last several years the growing interest in zeolite MR and zeolite membranes for H2 separation has been observed [40]. The use of membranes with zeolite selective layers for hydrogen production is now only at the early stages of investigation and is still a long way from commercial implementation.

J Porous Mater (2015) 22:635–646

Fig. 1 Number of published papers per year by Scopus Database: light grey—keywords: membrane, reactor, hydrogen, production; white—keywords: membrane, reactor, zeolite; dark grey—keywords: membrane, separation, hydrogen, zeolite; striped—keywords: membrane, reactor, hydrogen, production, zeolite

sieving [47], whereby the smaller molecules permeate preferentially through the membrane in comparison to the larger ones. One of the most important criteria for choosing a zeolite for a separation is the pore size. Figure 2 shows the effective pore size of the different zeolites and the kinetic diameters of gases [48] occurring during the production of hydrogen. Zeolite membranes are synthesized as a thin layer on a ceramic porous support (usually alumina) or a metallic support. The continuous zeolite film is about 2–20 lm thick [49]. The zeolite membranes are synthesized by in situ and ex situ crystallization (secondary growth) methods. In the in situ synthesis method, the zeolite membrane is

3.1 Zeolite membranes Zeolites are ideal membrane materials because they have uniform, molecular-size pores (usually 0.3–1.3 nm) [41]. Nowadays, more than 190 various zeolite framework types are known [42] but \20 have been utilized as membranes. The MFI-type membranes have been the most widely used due to ease of preparation, control of microstructure and versatility of applications [43]. Separation through zeolite membranes proceeds through three mechanisms: molecular sieving [44], diffusion-controlled permeation [45], and adsorption-controlled permeation [46]. The main gas transport mechanism is molecular

123

Fig. 2 Pore sizes of zeolites and kinetic diameters of small gaseous molecules

J Porous Mater (2015) 22:635–646

crystallized from gel on the surface of the support. The gel is composed of water, amorphous silica, a source for tetrahedral framework atoms other than Si, an organic template, and sometimes a mineralizing agent, such as NaOH. The crystallization occurs in an autoclave under elevated temperatures and under autogenous pressure. When a template is used, post-synthesis treatment, such as, calcination or burn out of the template is necessary [50–52]. In the ex situ synthesis technique, the support is pre-seeded with zeolite crystals (a seed layer), which are then sealed by a secondary hydrothermal growth [50, 51, 53]. The secondary growth method is considered to be the best in preparing high quality, defect-free membranes with high separation performance [54, 55]. 3.1.1 Zeolite membranes for H2 separation 3.1.1.1 The MFI zeolite membranes The ZSM-5 possesses two types of pore structures, and both of them are composed of ten tetrahedral-membered rings. One pore system is sinusoidal with a nearly circular cross Sect. (0.54 9 0.56 nm). The other is straight and perpendicular to the sinusoidal system with elliptical pores (0.51 9 0.54 nm). MFI channel sizes are similar to the sizes of hydrocarbons, and therefore, MFI membranes show very good selectivity in separating hydrocarbon isomers, such as n-butane from isobutane and n-hexane from 2,2-dimethylbutane [56]. There are a few reports about high ideal selectivities for H2 over gases such as N2, CO2, CO, CH4. Lovallo and Tsapatsis [57] prepared an oriented submicron silicalite membrane by growing a layer of oriented silicalite crystals on a composite precursor nanocrystalline silicalite/ alumina film using controlled secondary growth. This zeolite membrane showed a H2/N2 ideal selectivity of 61 at 150 °C (permeance of H2: 1230 9 10-10 mol s-1 m-2 Pa-1). Lai and Gavalas [58] investigated ZSM-5 membranes prepared by a hydrothermal reaction on asymmetric a-Al2O3 tubular supports. Tetrapropylammonium (TPA)-free synthesis gel was applied, and three different Si:Al ratios were investigated. They observed a H2/CH4 ideal selectivity of 1000 at 25 °C and a H2/CO2 ideal selectivity of 15 at 150 °C (permeance of H2 at these respective selectivities: 67 and 1200, both in units of 10-10 mol s-1 m-2 Pa-1). Most investigations show that because of pore size, MFI membranes have not been effective for H2 separation from light gas mixtures [59]. The high selectivities reported by some authors may be caused by the orientation of the zeolite crystals at the intergrown surface layer [57] and occlusion of the amorphous species inside the channel [58]. Diffusion of small gas molecules, such as H2, N2, CO, and CO2, at high temperatures in MFI zeolites can be described by the Knudsen-type diffusion mode. The

639

permeance and the permselectivities of these gas mixtures are limited by their Knudsen selectivities [60]. Xiao and Wei [61] proposed a translational diffusion model based on the relative size of the gas molecule as compared with the zeolitic pore size and gas molecular loading in the material. Gas molecules inside the zeolitic pores retain their gaseous characteristics, and the movement of molecules between the sites has to overcome an energy barrier that comes from the zeolitic pore channels. The diffusion activation energy can be estimated on the basis of Lennard-Jones potentials of a molecule at an intersection and in a channel. The transition diffusion behavior depends on the temperature, the ratio of molecule diameter to zeolite pore size, Lennard-Jones length constant for the molecule, and the pore structure of the individual zeolite. At the high temperature, the type of diffusion of nonadsorbing gases (e.g., H2, He, CO, CO2) depends on the ratio of the kinetic diameter of a gas molecule to the zeolitic pore size, L. When L \ 0.6, transport through zeolite membrane follows a Knudsen diffusion. When L [ 0.8, an activated diffusion behavior is observed. Between 0.6 and 0.8, the permeation of small gas molecules through zeolite membranes is controlled by both diffusion types [61]. The higher selectivity of H2 over other small gas molecules can be achieved by reducing the diameter of the zeolitic pores. The possibility of control of zeolite pore size without affecting the acid site was first described by Niwa et al. [62, 63]. Chemical vapor deposition (CVD) of an organo-silicon compound (tetramethyl orthosilicate) on the H mordenite [62, 63] and Y zeolites [64] has been applied to reduce pore diameter. The idea of chemical cracking of silanes for zeolitic pore size decrease is presented in Fig. 3. The silane molecule was adsorbed on the zeolite pore surface by interaction with oxygen ions of the zeolite framework. At high temperatures, amorphous SiO2 was formed. The diameter of a Si?4 is about 0.08 nm, and the MFI pore was reduced from 0.55 nm to about 0.36–0.47 nm, depending on the number of acidic sites at a local area. Masuda et al. [65] applied catalytic cracking of silanes (CCS) for modifying the effective pore of ZSM-5 with the use of methyldiethoxysilane (MDES) and methyldimethoxysilane. In this method, silane compounds small enough to fit

Fig. 3 Modification of the MFI membrane by catalytic cracking of silanes (CCS)

123

640

into the pores of ZSM-5 were adsorbed on acid sites and metal cations inside the zeolite. After catalytic cracking and calcination, SiO2 species resulted and reduced the size of the pores. The CCS method was applied for modifying the pore size of a ZSM-5 membrane used for H2 separation from a mixture of H2, N2, and O2 [65]. The permeance of H2 of the modified membrane decreased by 1/10. The permeances of the other gases for the treated membrane were reduced to about 1/500 of that of a fresh membrane. The treated membrane showed a H2 separation factor of about 90–140 (1.4–4.5 for the pristine membrane). ZSM-5 membranes containing Si and B atoms (instead of Al) in the framework were silylated using the catalytic cracking of MDES [59]. Boron atoms create acid sites weaker than those of Al-ZSM-5. The adsorption and reaction sites for MDES are the acid sites. The B-ZSM-5 membranes (Si/B = 12.5) were synthesized on the inner surface of tubular a-alumina supports. Single-gas permeance through the membrane and fifty/fifty gas mixtures of H2/CO2, H2/CH4, H2/N2 separations were investigated between 22 and 300 °C. The permeate pressure was 84 kPa, and the pressure difference across the membrane was 138 kPa. After CCS, H2 permeances decreased by more than one order of magnitude because MDES forms SiO2 in the zeolite pores, which reduced their pore size. All gases diffuse slower in the smaller pores. The permeance of the smaller H2 molecule decreased less than that of the larger molecules, and so the H2 selectivities increased. The H2 permeance and H2/CO2 and H2/CH4 separation selectivities increased with temperature. At 300 °C, the H2 permeance was 1.0 9 10-7 mol s-1 m-2 Pa-1 and the H2/CO2 and H2/CH4 separation selectivities were 47 and 45, respectively. The a-alumina-supported ZSM-5 membranes were modified by on-stream CCS with the use of the silane, MDES [66]. After modification, a significant increase in H2 selectivity over CO2 with a moderate decrease in H2 permeance was observed. At 450 °C, H2/CO2 permselectivity was equal to 17.5 (2.78 before CCS) and H2 single gas permeance of 1.86 9 10-7 mol s-1 m-2 Pa-1 (2.75 9 10-7 mol s-1 m-2 Pa-1 before CCS). Good performance and stability in separation of moisture from H2/CO2 (up to 28.4 % water vapor) at 450 °C was found. ZSM-5 membranes were prepared on a-alumina disks by the template-free secondary growth method followed by on-stream counter-diffusion or one-side CVD by tetraethoxysilane (TEOS) to eliminate intercrystalline pores [67]. Gas permeation and separation of H2, CO, and CO2 experiments were performed in the range of 25–500 °C. At 500 °C, H2/CO separation selectivity for one-side CVD modified membrane was about 16 times higher than the pristine membrane, and H2/CO2 separation selectivity for

123

J Porous Mater (2015) 22:635–646

on-stream counter-diffusion modified membrane was about 12 times higher than the pristine membrane. Considerable improvement in H2/CO2 separation through CVD modification can be achieved only on very good quality ZSM-5 membranes with few intercrystalline defects [68, 69]. The pore size reduction was performed by MDES deposition. The best H2 permeance (1.1 9 10-7 mol s-1 m-2 Pa-1) was observed at 450 °C and decreased with the temperature drop to 0.22 9 10-7 mol s-1 m-2 Pa-1 at 20 °C. The CO2 permeance decreases from 0.42 to 0.13 9 10-7 mol s-1 m-2 Pa-1 during increasing temperature. The ideal H2/CO2 separation factor at 450 °C for the CVD-modified good quality ZSM-5 membranes was about 20, and 8.6 for equimolar H2/CO2 mixtures. The improvement of H2/CO2 separation factors of MFI zeolite membranes modified by CCS of silanes accompanies reduction in H2 permeance (usually more than one order of magnitude). The thickness of the ZSM-5 layer determined the reduction in H2 permeance. In order to obtain a thin, high quality ZSM-5 layer on the a-alumina support, the yttrium stabilized zirconium (YSZ) intermediate barrier layer was applied [70]. On the YSZ layer, thick Al-free MFI (silicalite) and then thin Al-containing MFI zeolite layers were coated. The CCS modification of such a membrane resulted in an improvement of H2/CO2 separation factor from 4.95 to 25.3 and reduction in H2 permeance from 1.85 to 1.28 (10-7 mol s-1 m-2 Pa-1) at 450 °C. To obtain high selectivity for H2 separation, a high temperature air treatment on MFI zeolite membranes was applied [71]. MFI zeolite membranes were synthesized on porous a-alumina hollow fibers by in situ hydrothermal synthesis. Activation by air at 500 °C effectively promoted CCS of MDES in the zeolitic pores. The H2/CO2 separation factor of 45.6 with H2 permeance of 1 9 10-7 mol s-1 m-2 Pa-1 was observed at 500 °C. It was found that the H? ion-exchange of MFI zeolite membranes was beneficial to CCS of MDES and improved the deposited amount of SiO2 [72]. H? ion exchange of zeolite membranes increased acid sites in the zeolite membrane. The modified membrane showed a separation factor of 42.6 and H2 permeance of 2.82 9 10-7 mol s-1 m-2 Pa-1 at 500 °C. Additionally, the activated diffusion for H2 permeation was found. 3.1.1.2 Non-MFI zeolite membranes Separation by differences in size has a greater potential to be effective at high temperature than competitive adsorption, but pores smaller than those in MFI zeolites are required. An A-type zeolite membrane (0.41 nm pore diameter) was synthesized on an a-alumina tube by a hydrothermal synthesis [73]. The values of the ideal separation factor for H2/CO2 were 3.6 at 35 °C and 6.4 at 300 °C. The permeance of hydrogen was about 1 9 10-7 mol s-1 m-2 Pa-1.

J Porous Mater (2015) 22:635–646

Microwave heating instead of conventional heating in an autoclave was applied to obtain high selectivity for H2 and good permeability for an NaA zeolite membrane on aAl2O3 disk [74]. The crystallization of zeolite took only 15 min. The permeance of H2 was equal to 2.13 9 10-6 mol s-1 m-2 Pa-1 whereas the permeances of N2 and n-C4H10 were 0.67 and 0.18 (10-6 mol s-1 m-2 Pa-1), respectively. The synthesis by microwave heating was 8–12 times shorter than conventional. The permeance of the zeolite membrane synthesized by this method was four times higher. The SAPO-34 (0.38 nm pore diameter) zeolite membranes supported on a-alumina disk were prepared by the in situ method [75]. The ideal selectivities of H2/N2 and H2/CO2 were 11.2 and 1.6, respectively, at 100 °C and a feed pressure of 80 kPa. SAPO-34 membranes obtained by crystallization onto tubular stainless steel supports were applied by another group to separate CO2/H2 and H2/CH4 mixtures at feed pressures up to 1.7 MPa and at the temperature range between -20 and 35 °C [76]. The results showed strong adsorption of CO2 and inhibited H2 adsorption at a temperature of -20 °C. As a consequence, SAPO-34 membranes removed CO2 from CO2/H2 mixtures and the CO2/ H2 selectivity was greater than 100. The CO2 permeance was 3 9 10-8 mol s-1 m-2 Pa-1, and the CO2/H2 selectivity dropped strongly with increasing temperature. The H2/CH4 separation selectivity was about 20 and was a weak function of temperature. SAPO-34 zeolite membranes were deposited as a continuous layer on the inner wall of a-alumina tubular supports [77, 78]. The permeance of H2 at 27 °C was 2.4 9 10-8 mol s-1 m-2 Pa-1. The H2/CH4, H2/N2 and H2/CO2 ideal selectivities at 27 °C and 270 kPa feed pressure with a 138 kPa pressure drop were 25, 7.4, and 1.3 respectively. The selectivities decreased with increasing temperature and increased with increasing pressure. The pores in SAPO-34 are too small for MDES to penetrate but CCS was used to determine if the pore entrance could be modified by the same method as ZSM-5 [59]. Moreover the SAPO-34 structure has a large concentration of acid sites. The SAPO-34 membranes were obtained by crystallization onto tubular stainless steel supports. After silylation of SAPO-34 and B-ZSM-5 membranes, H2 permeances did not decrease; however, after silylation of ZSM-5 the H2 permeances did decrease. The CH4 permeance decreased after CCS, so the H2/CH4 separation selectivity was 59 at 22 °C (35 unsilylated) and decreased with the temperature. The H2/CO2 separation on the treated SAPO-34 membrane was slightly higher than on a fresh one. The size of SAPO-34 pores is almost the same as the CH4 kinetic diameter. The CH4 molecules probably permeate through nonzeolite pores. Because of changing

641

nonzeolite pore size, H2/CH4 separation selectivities increased after CCS. The results indicated that silylation did not significantly affect the pores or the pore entrance of the SAPO-34 but partially blocked nonzeolite pores. P-type zeolite membranes hydrothermally synthesized on porous a-Al2O3 support were used for H2 separation from SF6 and Ar [79]. The typical permeances of H2, Ar, and SF6 for the double coated membranes at room temperature were equal to 5.71, 1.08, and 0.05 (10-7 mol s-1 m-2 Pa-1), respectively. Permselectivity of 102 was achieved for H2/SF6 and 5.29 for H2/Ar. The permeances of the single gases H2, CH4, N2, and CO2 through composite membranes that consist of thin faujasite layers and a seed layer of Na-Y nanocrystals on the surfaces of porous a-Al2O3 support disks were investigated [80]. The zeolite membranes were grown hydrothermally. Two different kinds of heating during hydrothermal synthesis of polycrystalline faujasite layer were utilized, either with a conventional oven or a microwave oven. The resulting permeances of H2, CH4, N2, and CO2 were 18.7, 7.6, 5.8, and 5.3 (10-10 mol s-1 m-2 Pa-1), respectively for the conventional heating and 8.5, 3. 5, 2.5, and 2.1 (10-10 mol s-1 m-2 Pa-1), respectively, for the microwave heating. The permselectivities for the H2/CH4, H2/N2, H2/CO2 are slightly higher (over a dozen per cent) for the microwave than for conventional heating. The permeances for conventional heating are considerably higher than for microwave heating so the traditional method is better for this membrane. The authors concluded that the separation of small gas molecules by the large and heteropolar FAU (Na-X) pore system with 0.74 nm pore opening diameters is mainly based on differences of adsorption or diffusion rates, rather than on size exclusion like in LTA- or MFI-type zeolite membranes. AlPO4-type membranes were prepared on porous a-alumina tubes by a hydrothermal synthesis [81]. The ideal separation factors for single-component H2/N2 and H2/CO2 at 37 °C were equal to 13.3 and 23.9, respectively, and a permeance of H2 of 0.5 9 10-8 mol s-1 m-2 Pa-1. With increasing temperature, permeances of the gases increased, and separation factors of H2 with respect to the other gases decreased. Single gas permeation of He, H2, CO2, and CO at 25 – 500 °C through a highly siliceous DDR (Deca-Dodecasil 3R) zeolite membrane was investigated [43]. DDR membranes were grown on a porous a-alumina disk. To eliminate the intercrystalline micropores on-stream, counter diffusion CVD modification was applied. The DDR-type zeolite has a high thermal stability and small pore openings (0.36 9 0.44 nm) [42]. The permeance of H2 was about 1.1 9 10-10 mol s-1 m-2 Pa-1 at 500 °C and decreased linearly with lowering the temperature. The permselectivity of H2/CO and H2/CO2 at 500 °C are about 11 and 9, respectively.

123

642

The main step in fabrication of zeolite membranes is usually hydrothermal deposition. This traditional way of zeolite synthesis via the hydrothermal deposition cannot be easily scaled up in a cost-effective way needed for large scale separation applications [82, 83]. Choi and Tsapatsis [84] described production of zeolite MCM-22 on an aAl2O3 membrane for H2 separations without the hydrothermal growth. The preparation of zeolite composite films using layer-by-layer deposition of thin, plate-like crystals of MCM-22 on a-Al2O3 support has been proposed. The H2/N2 ideal selectivities approached or even exceeded 100 and remained high over the temperature range of 20–200 °C. The permeance of H2 of these selective membranes was higher than 1 9 10-8 mol s-1 m-2 Pa-1. Figure 4 summarizes the substantial amount of published research performed towards the goal of producing more effective H2/CO2 separations with zeolite membranes. Zeolite membrane systems generally show H2 permeances lower than 1 9 10-6 mol s-1 m-2 Pa-1 and H2/CO2 selectivity below 100. More intense research is needed to provide better insight into the synthesis of zeolite membranes with precise pore structures. 3.2 Membrane reactors for H2 production Membrane catalytic reactors combine a catalytic reaction and separation, which results in a high degree of process intensification. MRs enable cost-reduction as a result of the lessening in the number of process units and eliminate the separation stage as well as producing improvement in yields due to a shift in the equilibrium of the reaction. The packed bed MR is usually used for hydrogen production in MR investigations. In a packed bed MR, the catalyst is placed in a fixed bed close to a perm-selective membrane. The two configurations of packed bed MR – catalyst in tube and catalyst in shell—are shown in

Fig. 4 Permeability and selectivity data for H2/CO2 separation using zeolite membranes

123

J Porous Mater (2015) 22:635–646

Fig. 5. Carbon monoxide and water are present in the feed. The permeate contains mostly hydrogen whereas carbon dioxide is present in the retentate stream. The other configurations of zeolites in membrane reactors are: a catalytic membrane reactor, a catalytic nonpermselective membrane reactor, a non-permselective membrane, and a reactant-selective packed bed reactor. More detailed descriptions can be found in Ref. [5]. The WGS reaction was performed by applying the modified MFI zeolite membrane tube packed with a nanocrystalline cerium-doped ferrite catalyst at temperatures of 400–550 °C [85]. The MFI-type zeolite membrane was modified by CCS of MDES, and the aalumina tube was used as a support. The H2/CO2 permselectivity of modified MFI membranes was equal to 68.3 and H2 permeance of 2.94 9 10-7 mol s-1 m-2 Pa-1 at 550 °C. The conversion of CO in the MR was much higher than that in the traditional reactor. The MFI membrane enhanced the WGS reaction especially at high temperature, and by increasing the weight hourly space velocity (WHSV), and at low H2O:CO ratios. The 81.7 % CO conversion was achieved at 550 °C, with a WHSV of 60,000 h-1 and H2O:CO of 1.0. Such a value is considerably higher than the CO conversion obtained in the traditional packed-bed reactor and equilibrium conversion, i.e., 65 %. The increase of CO conversion increased with reaction pressure from 2 to 6 atm at all temperatures [86]. At 6 atm and at 550 °C, CO conversion of 98 % was observed. Hydrogen recovery (amount of H2 in permeate/ amount of H2 generated) of 64 % was obtained. It was also found that both the membrane and catalyst have good resistances to high H2S concentration. The model calculations have indicated that CO conversion higher than 99 % could be achieved using MFI zeolite membrane tube packed with a nanocrystalline cerium-doped ferrite catalyst at temperatures higher than 500 °C and a pressure of 30 atm, with a H2O:CO ratio of 3.5 [87]. A ZSM-5/silicalite bilayer membrane was synthesized on porous a-alumina supports coated with a YSZ barrier layer modified by CCS of MDES applied in a WGS reactor [88]. The reaction was performed over cerium-doped iron oxide catalyst at 500 °C, with a H2O:CO ratio of three, and gas hourly space velocity (GHSV) of 60,000 h-1. The CO conversion was about 84 % and hydrogen recovery about 23 %. These values remained unchanged even after 24 days. The membranes were stable after 24 days treatment with 400 ppm H2S. CuO/ZnO/Al2O3 commercial catalysts with MFI zeolite membranes were utilized for the low-temperature WGS reaction [89]. The zeolite membranes were synthetized on disk-shaped porous a-alumina by in situ hydrothermal crystallization. Zeolites were then modified by CCS of MDES. Although H2 permeance through modified

J Porous Mater (2015) 22:635–646

643

Fig. 5 Packed bed membrane reactors configurations: a catalyst in tube b catalyst in shell

Table 1 Hydrogen/carbon dioxide permselectivity and hydrogen permeance values for zeolite membranes Membrane

Temperature (oC)

Permselectivity H2/CO2

Hydrogen permeance (mol s-1 m-2 Pa-1107)

Ref

ZSM-5 modified by MDES

550

68.3

2.94

85


300

47

1

59


500

45.6

1

71


450

20

1.1

68


450

17.5

1.86

66


20

8.6

0.22

68

ZSM-5 modified by MDES ? YSZ layer H-ZSM-5 modified by MDES

450 500

25.3 42.6

1.28 2.82

70 72

ZSM-5

450

2.78

2.75

66

ZSM-5

150

15

1.2

58

SAPO-34

27

1.3

0.24

77

Na-Y nanocrystals

no data

3.5

0.0187

80

DDR

500

9

0.0011

43

AlPO4-type

37

23.9

0.05

81

A type zeolite

35

3.6

1

73

A type zeolite

300

6.4

1

73

membranes decreased about 15 %, CO2 permeance decreased about 1 order of magnitude. The H2/CO2 separation factor was considerably high and was equal to 42.6 and 10 at 500 °C and 200 °C, respectively. The good selectivity and high H2 permeance was due to a precise pore size for molecular sieving of H2 over CO2. The CO conversion of 95.4 % was obtained in the WGS process in membrane reactors at 300 °C. This value exceeded that for the calculated equilibrium conversion. The dry reforming of methane was performed using a La2NiO4/NaA zeolite composite membrane on a-Al2O3-c-

Al2O3 tube [90]. An NaA zeolite membrane was prepared by in situ hydrothermal synthesis on an alumina support. The H2/CH4 separation factor at room temperature was 4.2 (Knudsen value = 2.83), and the WGS reaction took place at 500–700 °C. The conversions of CH4 and CO2 and the selectivities of CO and H2 were remarkably higher than those over a traditional fixed-bed reactor. Moreover, the carbon deposition was lower. The dry reforming of methane was investigated over two composite membranes: Pd and a Pd–Ag alloy [91]. The porous stainless steel support was modified with an

123

644

NaA-zeolite by dip coating and hydrothermal synthesis, and the metal films were deposited by electroless plating. The NaA penetrates into the pore of the stainless steel support and makes the deposition of Pd and Pt layers easier. The Rh/La2O3 was used as the catalyst. Both membranes showed good performances, but the Pd membrane was better. The enhancement of the methane conversion was about 80 % at 450 °C. Hydrogen purity of 99.5 % was obtained. After 570 h work at 450–500 °C, the membrane was stable and no changes of the permeation were observed. The zeolite Y/a-Al2O3/RuO2 system was used to produce H2 from water [92]. The Y zeolite membrane was supported on a porous alumina support. The nanofingers of RuO2 have been assembled within a-Al2O3. Visible light initiated an electron-transfer reaction on one side of the membrane, followed by a reduction of water to H2 on the opposite side of the membrane. Propyl viologen sulfonate was applied as the electron acceptor (on the alumina side), and EDTA was used as the sacrificial donor (on the zeolite side). After 300 min of photolysis 90 9 10-7 mol of H2 was formed. Table 1 compares the H2/CO2 permselectivities and H2 permeances for various zeolite membranes. The ZSM-5 zeolites were widely studied as membranes, and the ZSM-5 modified by MDES membranes exhibit the high permselectivity and hydrogen permeance values. The highest reported permselectivity and hydrogen permeance were obtained by Tang et al. [87]. This was accomplished by using a zeolite-membrane reactor and employing the WGS reaction in the membrane’s inner tube, which was packed with a nanocrystalline cerium-doped ferrite catalyst.

4 Conclusions In this study, an overview of the zeolite membranes for hydrogen separation is presented. Although dense metal membranes, and among them the Pd-based membranes were most widely studied for hydrogen separation, the zeolite membranes have been shown to be more promising because of their high thermal and chemical stability. There have been only a limited number of investigations on zeolite membranes for hydrogen separation and zeolite membrane reactors. A closer look at zeolite synthesis in order to produce high quality membranes is needed. In zeolite membrane separations, molecular size exclusion sieving is the dominant mechanism. Because of the H2 kinetic diameter, 0.3 nm pore diameter in zeolite membranes is desired for high H2 selectivity. The most recent successes were exhibited by the pore size reduction achieved by CCS. This technique is a key approach to obtain high hydrogen selectivity, and the CCS method of

123

J Porous Mater (2015) 22:635–646

tailoring the pore sizes needs to be developed. The challenge is to prepare the high quality, large surface, reproducible zeolite membranes in a cost-effective way. More studies on the economic feasibility of zeolite membrane reactor processes, long-term stability of membranes, and scaling up are needed.

References 1. National Research Council, The Hydrogen Economy: Opportunities, Costs, Barriers, and R&D Needs (The National Academies Press, Washington, 2004) 2. Z. Chen, R. John, J.R. Grace, J. Lim, CO2 capture and hydrogen production in an integrated fluidized bed reformer-regenerator system. Ind. Eng. Chem. Res. 50, 4716–4721 (2011) 3. B.C.R. Ewan, R.W.K. Allen, A figure of merit assessment of the routes to hydrogen. Int. J. Hydrogen Energy 30, 809–819 (2005) 4. J. Caro, M. Noack, P. Kolsch, R. Schafer, Zeolite membranes– state of their development and perspective. Microporous Mesoporous Mater. 38, 3–24 (2000) 5. E.E. McLeary, J.C. Jansen, F. Kapteijn, Zeolite based films, membranes and membrane reactors: progress and prospects. Microporous Mesoporous Mater. 90, 198–220 (2006) 6. G.Q. Lu, J.C. da Costa, M. Duke, S. Giessler, R. Socolow, R.H. Williams, T. Kreutz, Inorganic membranes for hydrogen production and purification: a critical review and perspective. J. Colloid Interface Sci. 314, 589–603 (2007) 7. F. Gallucci, E. Fernandez, P. Corengia, M. van Sint, Annaland, Recent advances on membranes and membrane reactors for hydrogen production. Chem. Eng. Sci. 92, 40–66 (2013) 8. M. Du Tessie, M. Motay, M. Marechal, Bull. Soc. Chim. France 9, 334 (1868) 9. B. Gaudernack, S. Lynum, Hydrogen from natural gas without release of CO2 to the Atmosphere. Int. J. Hydrogen Energy 23, 1087 (1998) 10. C.N. Avila-Neto, S.C. Dantas, F.A. Silva, T.V. Franco, L.L. Romanielo, C.E. Hori, A.J. Assis, Hydrogen production from methane reforming. Thermodynamic assessment and auto thermal reactor design. J. Nat. Gas Sci. Eng. 1, 205–215 (2009) 11. A.L. da Silva, I.L. Mu¨ller, Towards H2-rich gas production from unmixed steam reforming of methane: thermodynamic modelling. J. Power Sources 196, 8568–8582 (2011) 12. www.eia.gov, file: Proved_Reserves_of_Natural_Gas_(Trillion_Cubic_Feet).xls 13. Y. Chen, H. Xu, Y. Wang, G. Xiong, Hydrogen production from the steam reforming of liquid hydrocarbons in membrane reactor. Catal. Today 118, 136–143 (2006) 14. G. Jinsheng (ed.) Coal, Oil, Shale, Natural Bitumen, Heavy Oil and Peat, East China University of Science and Technology (ECUST), China, ISBN: 978-1-84826-467-0, 978-1-84826-468-7 15. Ch. Higman, S. Tam, Advances, coal gasification, hydrogenation, and gas treating for the production of chemicals and fuels. Chem. Rev. 114, 1673–1708 (2014) 16. P. Chiesa, S. Consonni, T. Kreutzb, R. Williams, Co-production of hydrogen, electricity and CO2 from coal with commercially ready technology. Part A: performance and emissions. Int. J. Hydrogen Energy 30, 747–767 (2005) 17. C. Corte´s, E. Tzimas, S.D. Peteves, Technologies for coal based hydrogen and electricity co-production power plants with CO2 capture (Institute for Energy, European Commission, Luxembourg, 2009)

J Porous Mater (2015) 22:635–646 18. J.R. Ladebeck, J.P. Wagner, Handbook of fuel cells-fundamentals, technology and applications (Wiley, Chichester, 2003) 19. Y. Kalinci, A. Hepbasli, I. Dincerc, Biomass-based hydrogen production: a review and analysis. Int. J. Hydrogen Energy 34, 8799–8817 (2009) 20. M.C.J. Bradford, M.A. Vannice, CO2 reforming of CH4. Catal. Rev. Sci. Eng. 41, 1–42 (1999) 21. P. Gangadharan, K.C. Kanchi, H.H. Lou, Evaluation of the economic and environmental impact of combining dry reforming with steam reforming of methane. Chem. Eng. Res. Des. 90, 1956–1968 (2012) 22. W. Kreuter, H. Hofmann, Electrolysis: the important energy transformer in a world of sustainable energy. Int. J. Hydrogen Energy 23, 661–666 (1998) 23. J.R. Rostrup-Nielsen, Catalytic steam reforming in catalysis. Science and technology, vol. 5 (Springer, Berlin, 1984) 24. Y. Shirasaki, T. Tsuneki, Y. Ota, I. Yasuda, S. Tachibana, H. Nakajima, K. Kobayashi, Development of membrane reformer system for highly efficient hydrogen production from natural gas. Int. J. Hydrogen Energy 34, 4482–4487 (2009) 25. G. Iaquaniello, F. Giacobbe, B. Morico, S. Cosenza, A. Farace, Membrane reforming in converting natural gas to hydrogen: production costs, Part II. Int. J. Hydrogen Energy 33, 6595–6601 (2008) 26. T. Tsuru, K. Yamaguchi, T. Yoshioka, M. Asaeda, In the recent years, research has been devoted to the use catalytic membrane reactors in overcoming the equilibrium limitation. AIChE J. 50, 2794–2805 (2004) 27. R.C. Hurlbert, J.O. Konecny, Diffusion of hydrogen through palladium. J. Chem. Phys. 34, 655–658 (1961) 28. A. Kulprathipanja, G. Alptekin, J. Falconer, D. Way, Pd and Pd– Cu membranes: inhibition of H2 permeation by H2S. J. Membr. Sci. 254, 49–62 (2005) 29. S. Escola´stico, C. Solı´s, J.M. Serra, Hydrogen separation and stability study of ceramic membranes based on the system Nd5LnWO12. Int. J. Hydrogen Energy 36, 1194–1954 (2011) 30. E. Serra, A.C. Bini, G. Cosoli, L. Pilloni, Hydrogen permeation measurements on alumina. J. Am. Ceram. Soc. 88, 15–18 (2005) 31. W.H. Yuan, L.L. Mao, L. Li, Novel SrCe0.75Zr0.20Tm0.05O3-d membrane for hydrogen separation. Chin. Chem. Lett. 21, 369–372 (2010) 32. J. Dong, Y.S. Lin, M. Kanezashi, Z. Tang, Microporous inorganic membranes for high temperature hydrogen purification. J. Appl. Phys. 104, 121301 (2008) 33. Y.S. Lin, Microporous and dense inorganic membranes: current status and prospective. Sep. Purif. Technol. 25, 39–55 (2001) 34. J. Gascon, F. Kapteijn, Metal-organic framework membranes— high potential, bright future? Angew. Chem. 49, 1530–1532 (2010) 35. M. Shah, M.C. McCarthy, S. Sachdeva, A.K. Lee, H.-K. Jeong, Current status of metal-organic framework membranes for gas separations: promises and challenges. Ind. Eng. Chem. Res. 51, 2179–2219 (2012) 36. N.W. Ockwig, T.M. Nenoff, Membranes for hydrogen separation. Chem. Rev. 107, 4078–4110 (2007) 37. H. Verweij, Y.S. Lin, J.H. Dong, Microporous silica and zeolite membranes for hydrogen purification. MRS Bull. 31, 756–764 (2006) 38. J. Rothfleisch, Hydrogen from Methanol for Fuel Cells. SAE Technical Paper, 640377 (1964) 39. S. Albert Jr, Large zeolites: why and how to grow in space. Proc. SPIE Int. Soc. Opt. Eng. Proc. SPIE. 1557, 6–9 (1991) 40. J. Coronas, J. Santamaria, State-of-the-art in zeolite membrane reactors. Top. Catal. 29, 29–44 (2004) 41. M. Yu, R.D. Noble, J.L. Falconer, Zeolite membranes: microstructure characterization and permeation mechanisms. Acc. Chem. Res. 44, 1196–1206 (2011)

645 42. C. Baerlocher, W.M. Meier, D.H. Olson, Atlas of zeolite framework types (Elsevier, Amsterdam, 2001) 43. Y.S. Lin, I. Kumakiri, B.N. Nair, H. Alsyouri, Microporous inorganic membranes. Sep. Purif. Meth. 31, 229–379 (2002) 44. S. Mohanty, A.V. McCormick, Prospects for principles of size and shape selective separations using zeolites. Chem. Eng. J. 74, 1–14 (1999) 45. J. Caro, M. Noack, P. Kolsch, Zeolite membranes: from the laboratory scale to technical applications. Adsorption 11, 215–227 (2005) 46. M. Kanezashi, J. O’Brien-Abraham, Y.S. Lin, K. Suzuki, Gas permeation through DDR - type zeolite membranes at high temperatures. AIChE J. 54, 1478–1486 (2008) 47. S.T. Oyama, S.M. Stagg-Williams, Inorganic polymeric and composite membranes: structure function and other correlations (Elsevier, Amesterdam, 2011) 48. N. Mehio, S. Dai, D. Jiang, Quantum mechanical basis for kinetic diameters of small gaseous molecules. J. Phys. Chem. A 118, 1150–1154 (2014) 49. S. Kulprathipanja, Zeolites in Industrial Separation and Catalysis (Wiley, Weinheim, 2010) 50. A.K. Pabby, S.S.H. Rizvi, A.M.S. Requena, Handbook of Membrane Separations: Chemical, Pharmaceutical, Food, and Biotechnological Applications (CRC Press, Boca Raton, 2008) 51. E. van Steen, L.H. Callanan, M. Claeys, Recent Advances in the Science and Technology of Zeolites and Related Materials Series: Studies in Surface Science and Catalysis (Book 154) (Elsevier, New York, 2005) 52. T.C. Bowen, R.D. Noble, J.L. Falconer, Fundamentals and applications of pervaporation through zeolite membranes. J. Membr. Sci. 245, 1–33 (2004) 53. M.C. Lovallo, A. Gouzinis, M. Tsapatsis, Synthesis and characterization of oriented MFI membranes prepared by secondary growth. AIChE J. 44, 1903–1913 (1998) 54. J. Ma, J. Shao, Z. Wang, Y. Yan, Preparation of zeolite naa membranes on macroporous alumina supports by secondary growth of gel layers. Ind. Eng. Chem. Res. 53, 6121–6130 (2014) 55. Z. Wang, Q. Ge, J. Shao, Y. Yan, High performance zeolite LTA pervaporation membranes on ceramic hollow fibers by dipcoating - wiping seed deposition. J. Am. Chem. Soc. 131, 6910–6911 (2009) 56. C.J. Gump, X. Lin, J.L. Falconer, R.D. Noble, Experimental configuration and adsorption effects on the permeation of C4 isomers through ZSM-5 zeolite membranes. J. Membr. Sci. 173, 35–52 (2000) 57. M.C. Lovallo, M. Tsapatsis, Preferentially oriented submicron silicalite membranes. AIChE J. 42, 3020–3029 (1996) 58. R. Lai, G.R. Gavalas, ZSM-5 membrane synthesis with organicfree mixtures. Microporous Mesoporous Mater. 38, 239–245 (2000) 59. M. Hong, J.L. Falconer, R.D. Noble, Modification of zeolite membranes for H2 separation by catalytic cracking of methyldiethoxysilane. Ind. Eng. Chem. Res. 44, 4035–4041 (2005) 60. A.J. Burggraaf, Single gas permeation of thin zeolite (MFI) membranes: theory and analysis of experimental observations. J. Membr. Sci. 155, 45–65 (1999) 61. J. Xiao, J. Wei, Diffusion mechanism of hydrocarbons in zeolitess. Theory Chem. Eng. Sci. 47, 1123–1141 (1992) 62. M. Niwa, H. Itoh, S. Kato, T. Hattori, Y. Murakami, Modification of H-mordenite by a vapour-phase deposition method. J. Chem. Soc. Chem. Commun. 819–820 (1982) 63. M. Niwa, S. Kato, T. Hattori, Y. Murakami, Fine control of the pore-opening size of the zeolite mordenite by chemical vapour deposition of silicon alkoxide. J. Chem. Soc. Faraday Trans. I 80, 3135–3145 (1984)

123

646 64. J.H. Kim, Y. Ikoma, M. Niwa, Control of the pore-opening size of HY zeolite by CVD of silicon alkoxide. Microporous Mesoporous Mater. 32, 37–44 (1999) 65. T. Masuda, N. Fukumoto, M. Kitamura, S.R. Mukai, K. Hashimoto, T. Tanaka, T. Funabiki, Modification of pore size of MFI-type zeolite by catalytic cracking of silane and application to preparation of H2-separating zeolite membrane. Microporous Mesoporous Mater. 48, 239–245 (2001) 66. X. Gu, Z. Tang, Dong J. On-stream modification of MFI zeolite membranes for enhancing hydrogen separation at high temperature. Microporous Mesoporous Mater. 111, 441–448 (2008) 67. M. Kanezashi, Y.S. Lin, Gas permeation and diffusion characteristics of MFI-type zeolite membranes at high temperatures. J. Phys. Chem. C 113, 3767–3774 (2009) 68. X. Zhu, H. Wang, Y.S. Lin, Effect of the membrane quality on gas permeation and chemical vapor deposition modification of MFI-type zeolite membranes. Ind. Eng. Chem. Res. 49, 10026–10033 (2010) 69. H. Wang, Y.S. Lin, Effects, synthesis conditions on MFI zeolite membrane quality and catalytic cracking deposition modification results. Microporous Mesoporous Mater. 142, 481–488 (2011) 70. H. Wang, Y.S. Lin, Synthesis and modification of ZSM-5/silicalite bilayer membrane with improved hydrogen separation performance. J. Membr. Sci. 396, 128–137 (2012) 71. Z. Hong, F. Sun, D. Chen, C. Zhang, X. Gu, N. Xu, Improvement of hydrogen-separating performance by on-stream catalytic cracking of silane over hollow fiber MFI zeolite membrane, nt. J. Hydrogen Energy 39, 4739–4748 (2014) 72. Z. Hong, Z. Wu, Y. Zhang, X. Gu, Catalytic cracking deposition of methyldiethoxysilane for modification of zeolitic pores in MFI/ a-Al2O3 zeolite membrane with H ? ion exchange pretreatment. Ind. Eng. Chem. Res. 52, 13113–13119 (2013) 73. K. Aoki, K. Kusakabe, S. Morooka, Gas permeation properties of A-type zeolite membrane formed on porous substrate by hydrothermal synthesis. J. Membr. Sci. 141, 197–205 (1998) 74. X. Xu, W. Yang, J. Liu, L. Lin, Synthesis of a high-permeance NaA zeolite membrane by microwave heating. Adv. Mater. 12, 195–198 (2000) 75. Z. Lixiong, J.M. Dong, M. Enze, Synthesis of SAPO-34/ceramic composite membranes. Stud. Surf. Sci. Catal. 105, 2211–2216 (1997) 76. M. Hong, S. Li, J.L. Falconer, R.D. Noble, Hydrogen purification-34 membrane. J. Membr. Sci. 307, 277–283 (2008) 77. J.C. Poshusta, V.A. Tuan, E.A. Pape, R.D. Noble, J.L. Falconer, Separation of light gas mixtures using SAPO-34 membranes. AIChE J. 46, 779–789 (2000) 78. J.C. Poshusta, V.A. Tuan, J.L. Falconer, R.D. Noble, Synthesis and permeation properties of SAPO-34 tubular membranes. Ind. Eng. Chem. Res. 37, 3924–3929 (1998)

123

J Porous Mater (2015) 22:635–646 79. J. Dong, Y.S. Lin, In situ synthesis of P-type zeolite membranes on porous a-alumina supports. Ind. Eng. Chem. Res. 37, 2404–2409 (1998) 80. K. Weh, M. Noack, I. Sieber, J. Caro, Permeation of single gases and gas mixtures through faujasite-type molecular sieve membranes. Microporous Mesoporous Mater. 54, 27–36 (2002) 81. G. Guan, T. Tanaka, K. Kusakabe, K.I. Sotowa, S. Morooka, Characterization of AlPO4-type molecular sieving membranes formed on a porous a-alumina tube. J. Membr. Sci. 214, 191–198 (2003) 82. W.J. Koros, R. Mahajan, Pushing the limits on possibilities for large scale gas separation: which strategies? J. Membr. Sci. 175, 181–196 (2000) 83. J. Coronas, J. Santamaria, The use of zeolite films in small-scale and micro-scale applications. Chem. Eng. Sci. 59, 4879–4885 (2004) 84. J. Choi, M. Tsapatsis, MCM-22/silica selective flake nanocomposite membranes for hydrogen separations. J. Am. Chem. Soc. 132, 448–449 (2010) 85. Z. Tang, S.J. Kim, G.K. Reddy, J. Dong, P. Smirniotis, Modified zeolite membrane reactor for high temperature water gas shift reaction. J. Membr. Sci. 354, 114–122 (2010) 86. S.J. Kim, Z. Xu, G.K. Reddy, P. Smirniotis, J. Dong, Effect of pressure on high-temperature water gas shift reaction in microporous zeolite membrane reactor. Ind. Eng. Chem. Res. 51, 1364–1375 (2012) 87. S.J. Kim, S. Yang, G.K. Reddy, P. Smirniotis, J. Dong, Zeolite membrane reactor for high-temperature water-gas shift reaction: effects of membrane properties and operating conditions. Energy Fuels 27, 4471–4480 (2013) 88. H. Wang, X. Dong, Y.S. Lin, Highly stable bilayer MFI zeolite membranes for high temperature hydrogen separation. J. Membr. Sci. 450, 425–432 (2014) 89. Y. Zhang, Z. Wu, Z. Hong, X. Gu, N. Xu, Hydrogen-selective zeolite membrane reactor for low temperature water gas shift reaction. Chem. Eng. J. 197, 314–321 (2012) 90. B.S. Liu, L.Z. Gao, C.T. Au, Preparation, characterization and application of a catalytic NaA membrane for CH4/CO2 reforming to syngas. Appl. Catal. A Gener. 235, 193–206 (2002) 91. M.L. Bosko, J.F. Mu´nera, E.A. Lombardo, L.M. Cornaglia, Dry reforming of methane in membrane reactors using Pd and Pd–Ag composite membranes on a NaA zeolite modified porous stainless steel support. J. Membr. Sci. 364, 17–26 (2010) 92. Y. Kim, P.K. Dutta, An integrated zeolite membrane/RuO2 photocatalyst system for hydrogen production. J. Phys. Chem. C 111, 10575–10581 (2007)