A Novel Dual-Membrane Reactor for Continuous Heterogeneous

ARTICLE pubs.acs.org/IECR

A Novel Dual-Membrane Reactor for Continuous Heterogeneous Oxidation Catalysis Hong Jiang,† Lie Meng,† Rizhi Chen,†,‡ Wanqin Jin,*,† Weihong Xing,† and Nanping Xu† †

State Key Laboratory of Materials-Oriented Chemical Engineering, and ‡Jiangsu Key Laboratory of Industrial water-Conservation & Emission Reduction, Nanjing University of Technology, 5 Xinmofan Road, Nanjing 210009, People’s Republic of China

bS Supporting Information ABSTRACT: A novel dual-membrane reactor is proposed for heterogeneous catalysis over ultrafine catalyst particles. In this reactor, one tubular porous ceramic membrane is employed as a distributor controlling the supply of reactant, while the other is employed as a membrane separator for in situ separation of ultrafine catalysts from the products. To evaluate the feasibility and the performance of the dual-membrane reactor, phenol hydroxylation with hydrogen peroxide (H2O2) over TS-1 catalyst was selected as a model reaction. As compared to traditional H2O2 feeding modes, the use of the porous ceramic membrane distributor allows a uniform injection of H2O2 into the reaction system, and as a result significantly promotes the dihydroxybenzene selectivity. The phenol conversion and the DHB selectivity above 15% and about 95%, respectively, can be achieved in a continuous operation of 30 h. TS-1 catalysts can be retained almost completely in the reactor system by the membrane separator. Our study demonstrates the advantages of the novel dual-membrane reactor in enhancing selectivity and catalysts separation in a continuous heterogeneous catalytic reaction.

1. INTRODUCTION Heterogeneous catalysis plays an important role in chemical and petrochemical production processes. Although the use of heterogeneous catalysts with small particle sizes significantly improves the reaction,1 the efficient separation of ultrafine catalyst particles from the reaction slurry, whether by gravity settling or porous tube filtration,2 remains challenging and is still a major concern. As a possible solution, ultrafine catalysts can be immobilized onto a carrier with larger grain size, which allows easier catalyst separation but at the expense of a reduction of the effective surface area of the catalyst. Moreover, catalysts in suspension have been reported with a better catalytic activity than the immobilized counterparts.3,4 Membrane reactor, a process that combines the heterogeneous catalytic reaction with membrane separation, has emerged as a promising approach to solve the problem described above. In general, membrane reactors for heterogeneous catalysis can be divided into two configurations, the side-stream type and submerged type.5 In our previous work, a side-stream membrane reactor was developed2,6 in which the reaction zone was separated from the separation zone. It was found that ultrafine catalysts were easily adsorbed onto the pipeline, the membrane, and into the pumping system, lowering the catalyst concentration in the system and consequently reducing the reactor performance. Recently, the submerged membrane reactor,79 combining the reaction zone and separation zone in a single unit, has received more attention. It shows many advantages over the side-stream membrane reactor, such as a small footprint, reduced catalysts adsorption, and much less energy consumption, due to the absence of a high-flow recirculation pump. Besides catalyst separation, the minimization of byproducts is another important issue in heterogeneous catalysis. The concentration and distribution of reactants play a critical role in the r 2011 American Chemical Society

search for the desired product selectivity,10 which can be achieved by supplying reactants in a controlled way. Membrane distributor is an emerging technology that has been used to control the addition of reactants to the reaction mixture. Many related studies have been carried out.11,12 It has been reported that membrane distributor can be used for a high selectivity in vapor phase reactions.13,14 Coronas et al.13 used ceramic membranes as oxygen distributors for the oxidative coupling of methane, and the selectivity was significantly improved in the membrane reactor when compared to the conventional fixed-bed reactor. Similarly, Tonkovich et al.14 found that the membrane reactor outperformed the tubular reactor at low feed ratios in the oxidative dehydrogenation of ethane. In the present work, a novel dual-membrane reactor has been designed by employing one tubular porous ceramic membrane as a reactant distributor controlling the supply of reactant, thus increasing the product selectivity, and a second tubular porous ceramic membrane as a membrane separator for in situ separation of catalysts from the products. Thus, the two objectives of the ultrafine catalysts separation and product selectivity enhancement could be simultaneously realized in the designed dualmembrane reactor. The scheme is shown in Figure 1. To investigate the feasibility of the continuous operation in this newly developed dual-membrane reactor, the hydroxylation of phenol to dihydroxybenzene (DHB) catalyzed by TS-1 was taken as a model reaction. The effect of the membrane pore size as well as the various operating parameters such as residence time, stirring rate, reaction temperature, catalyst concentration, Received: February 27, 2011 Accepted: August 18, 2011 Revised: May 15, 2011 Published: August 18, 2011 10458

dx.doi.org/10.1021/ie200398g | Ind. Eng. Chem. Res. 2011, 50, 10458–10464

Industrial & Engineering Chemistry Research

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Scheme 1. Hydroxylation of Phenol with Hydrogen Peroxide over TS-1 Catalyst

Figure 1. Schematic diagram of the dual-membrane reactor. (a) Membrane distributor. (b) Membrane separator.

Figure 2. Reactor configurations used in this work for hydroxylation of phenol over TS-1.

and phenol to H2O2 molar ratio on the phenol conversion and DHB selectivity was investigated in detail.

2. EXPERIMENTAL SECTION 2.1. Materials. TS-1 catalyst (average particle size, 200 nm; specific surface area, 408 m2 g1; Si/Ti molar ratio, 59) was purchased from Baling Petrochemical Co., SINOPEC. Methanol (>chromatography grade) was supplied by the Yuwang group, and pure water was provided by the Hangzhou Wahaha Group (Co., Ltd., China). Deionized water was homemade and had an electrical conductivity below 12 μs cm1. Phenol, hydroquinone (HQ), and sodium hydroxide were provided by Shantou Xilong Chemical Co., Ltd., China. 30% H2O2 was supplied by the Sinopharm Chemical Reagent Co., Ltd. China, and catechol (CA) was provided by the Shanghai Sansi Reagent Co., Ltd., China. All materials were of analytical grade. 2.2. Apparatus. The experimental setup is shown schematically in Figure 2. This consisted of a slurry stirred reactor, two porous ceramic tubular membrane modules, a feeding system, a productcollecting system, and a heating system. The reactor, with a working volume of 700 mL, was made of glass and equipped with an external

Figure 3. Evolution of phenol conversion and DHB selectivity with time for the different feeding modes of hydrogen peroxide (Cphenol = 1.33 mol L1, CH2O2 = 1.33 mol L1, catalyst concentration is 17.2 g L1, reaction temperature is 353 K, stirring rate is 380 rpm, and flow rate of hydrogen peroxide is 4.5 mL min1).

jacket for temperature control. Single tubular ceramic membranes were provided by Nanjing Jiusi High-Tech Co., Ltd., China. The effective length, outer diameter, and inner diameter were 60, 12, and 8 mm, respectively. Two types of membrane distributors were used, one with a fine layer of α-Al2O3 (nominal pore sizes of 0.2, 0.5, and 0.8 μm, respectively) coated on the outer surface of a tubular αAl2O3 porous support (nominal pore size 2 μm) and the other a tubular α-Al2O3 porous membrane (nominal pore size 2 μm). The membrane separator was composed of a thin layer of α-Al2O3 with a nominal pore size of 0.2 μm coated on the outer surface of a tubular α-Al2O3 porous support (nominal pore size 2 μm). One end of the ceramic tubes was sealed with glazing compound, while the other one was kept open. A constant flow pump (Beijing Chuangxin Tongheng Science & Technology Co., Ltd., China) was employed to feed H2O2 into the phenol solution, which ensured a steady flow of H2O2. Two peristaltic pumps (Baoding Longer Precision Pump Co., Ltd., China) were used to feed phenol solution into the reaction system and take the mixture out of the reactor. 2.3. Hydroxylation Experiment. To investigate the continuous process of the dual-membrane reactor experimentally, the phenol hydroxylation to DHB over TS-1 was taken as a model reaction for the following reasons. First, dihydroxybenzene 10459

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(DHB), including hydroquinone (HQ) and catechol (CA), are widely used as antioxidants, medicines, perfumes, polymerization inhibitors, and for organic synthesis, etc.15 Moreover, phenol hydroxylation by H2O2 over the TS-1 catalyst for DHB production is a promising process because of its advantages such as being environmentally benign and having a high atom economy. However, some unexpected byproduct, such as benzoquinone, can be generated because of the strong oxidizing properties of H2O2, resulting in a decrease of the DHB selectivity. The hydroxylation of phenol with hydrogen peroxide over TS-1, depicted in Scheme 1, was performed using the experimental setup described above. In general, the system was heated to a desired temperature under stirring after supplying certain amounts of phenol aqueous solution and TS-1 catalyst into the reactor. The hydrogen peroxide solution was then injected into the lumen of the ceramic membrane distributor and pumped through the micropores of the distributor into the phenol solution to start the reaction. After 60 min, phenol aqueous solution and hydrogen peroxide solution were fed continuously into the reactor. Meanwhile, the product stream was extracted through the membrane separator at a flow rate equal to the total feed rate. Initially, the TS-1 catalyst particles were easily adsorbed onto the clean membrane to form a cake layer, increasing the filtration resistance and decreasing permeation flux with operation time. To keep the liquid level in the reactor constant, the discharge flow rate was increased by increasing the rotational speed of the peristaltic pump. As a result, the transmembrane pressure according to eq 3 discussed below increased. As the equilibrium of deposition and removal of catalyst particles proceeded, the operation pressure and discharge flow rate tended to stabilize, and the system was operated in a steady state. The non-steady-state time was approximately 200 min. Each continuous hydroxylation, lasting 6 h, was used to investigate the optimal conditions; experimental data points were sampled at 6 h from the start of the continuous operation. The catalyst particles were completely retained inside the reactor by the ceramic membranes. The products were collected in a 1000 mL graduated cylinder. The operating pressure of the membrane separator was monitored by a vacuum gauge. After the reaction, the membranes were rinsed with sodium hydroxide solution (1 g/L) before the subsequent experiments. The reaction products were analyzed by high performance liquid chromatography (HPLC, Agilent 1100 Series, U.S.) equipped with a diode array detector (DAD), an autosampler, and a ZORBAX Eclipse XDB-C18 column (5 μm, 4.6 mm 250 mm). The analyses were performed at 308 K by a UV detector at 277 nm using a mobile phase consisting of methanol/pure water (40:60 v/v) at a flow rate of 1.0 mL min1. The automatic injection volume was 5 μL per sample. A calibration curve of oxidation products was drawn using aqueous samples of known composition. Phenol conversion and DHB selectivity were calculated on the basis of the starting amount of phenol,16 according to eqs 1 and 2, respectively. X ¼

CphenolðoÞ CphenolðpÞ CphenolðoÞ

ð1Þ

S¼

CCA þ CHQ CphenolðoÞ CphenolðpÞ

ð2Þ

where X is the phenol conversion, Cphenol(o) is the initial phenol concentration in the feed (mol L1), Cphenol(p) is the phenol

concentration in the outlet of the reactor (mol L1), and CCA and CHQ are the concentrations of CA and HQ in the outlet of the reactor (mol L1). The filtration resistance was obtained from the transmembrane pressure and the membrane flux according to the Darcy law:17 R ¼

Δp Jμ

ð3Þ

where R is the filtration resistance (m1), Δp is the transmembrane pressure (Pa), J is the membrane flux (m s1), and μ is the viscosity of permeate (Pa s). To estimate the retention of the TS-1 catalyst by the membrane separator, the concentrations of silicon and titanium elements in the permeate were analyzed by inductively coupled plasma emission spectroscopy (ICP, Optima 2000DV, U.S.). 2.4. Membrane Characterization. The ceramic porous membranes were characterized using a field emission scanning electron microscope (FESEM, S-4800 II, Hitachi) equipped with an energy dispersive X-ray spectrometer (EDX, INCAXSIGHT, UK).

3. RESULTS AND DISCUSSION 3.1. Comparison of Feeding Modes for the Hydrogen Peroxide. To study the behavior of the porous ceramic mem-

brane distributor, H2O2 was added to the reactor in three different ways for the preliminary reaction of the phenol hydroxylation over the TS-1 catalyst. These were (A) allowing the H2O2 to flow through the ceramic membrane (average pore size of 0.5 μm), (B) adding it directly into the phenol solution drop by drop, and (C) adding in one lot. The experiments of this type were conducted in a batch reactor using the identical operation conditions as the continuous reactor, except for the absence of the membrane separator and products collecting system. The dual-membrane reactor was operated as a batch reactor for this experiment. Figure 3 shows the phenol conversion and DHB selectivity using the different feeding modes for H2O2. At first, the phenol conversion increased quickly initially and then it tended to remain constant, in agreement with a previous report.18 After 90 min, the conversions of modes A and B were similar, and slightly higher than that of mode C. Because of the fact that the decomposition rate of H2O2 increases substantially with its concentration,19 pouring H2O2 directly into the phenol solution using mode C aggravated the decomposition of H2O2 and consequently led to a lower conversion. In contrast, the selectivity of DHB varied markedly with the feeding mode. The DHB selectivity of mode A was higher than that of modes B and C. This is because the porous membrane is a very effective distributor having uniformly sized liquid droplets in the microdiameter range,20 providing the microchannels for a more homogeneous distribution of H2O2 and as a result improving the conversion and selectivity of the phenol hydroxylation reaction. 3.2. Effect of Membrane Pore Size on the Reaction Conversion and Selectivity. The membrane pore size is a key factor for the membrane distributor because it determines the droplet sizes of the reactants21 and therefore affects the phenol hydroxylation reaction. The effect of membrane pore size on the reaction process was investigated using porous ceramic membranes with different pore sizes, that is, 0.2, 0.5, 0.8, and 2 μm. As displayed in Table 1, the DHB selectivity varied significantly with 10460



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Table 1. Effect of Membrane Pore Size on the Conversion and Selectivitya membrane pore size/μm

phenol conversion/%

DHB selectivity/%

0.2

28.7

95.9

0.5 0.8

28.8 28.3

95.5 88.8

2

28.5

88.3

Operation conditions: Cphenol = 3.67 mol L1, CH2O2 = 1.33 mol L1, catalyst concentration is 17.2 g L1, reaction temperature is 353 K, stirring rate is 380 rpm, and residence time is 3.7 h. a

Figure 5. (a) Effect of phenol/H2O2 molar ratio on the conversion, selectivity of phenol hydroxylation. Inset figure plotted the DHB selectivity (y-axis) versus phenol conversion (x-axis). (b) Effect of phenol/H2O2 molar ratio on the filtration resistance (catalyst concentration is 17.2 g L1, reaction temperature is 353 K, stirring rate is 380 rpm, and residence time is 3.7 h).

Figure 4. (a) Effect of catalyst concentration on the conversion, selectivity of phenol hydroxylation. (b) Effect of catalyst concentration on the filtration resistance (Cphenol = 3.67 mol L1, CH2O2 = 1.33 mol L1, residence time is 3.7 h, reaction temperature is 353 K, and stirring rate is 380 rpm).

the membrane pore size, while the phenol conversion remained stable. For instance, the DHB selectivity was improved from 88.3% to 95.9% when the membrane pore size decreased from 2 to 0.2 μm, which indicates that a smaller membrane pore size reduces the droplet sizes of the reactant and enhances the selectivity. This is in agreement with the observation that optimal micromixing at a molecular scale improves the contact between reactants and as a result increases the selectivity, yield, and quality of target products.22 Jia et al.23 reported that membranes with smaller MWCO (molecular weight cutoff) increased the micromixing efficiency. Wu et al.21 also modeled the enhancement of micromixing efficiency with reduced membrane pore size. In this

work, we also found that the smaller membrane pore size results in a more homogeneous distribution of H2O2 in the reactor and, as a result, reduces the further oxidation of DHB to benzoquinone or other heavy byproduct. Taking into account the cost, the membrane with a pore size of 0.5 μm was used in subsequent experiments. 3.3. Optimization of Operation Conditions. Generally, the phenol conversion and product selectivity are strongly affected by operating parameters such as residence time, stirring rate, reaction temperature, catalyst concentration, and molar ratio between the reactants. The effects of these parameters on reactor performance were investigated systematically, as follows, and the corresponding filtration resistances were also analyzed. The effects of residence time, stirring rate, and reaction temperature on phenol conversion, DHB selectivity, and the corresponding filtration resistances were investigated (as shown in Figures S13, Supporting Information). From the results of Figures S13, suitable appropriate residence time, stirring rate, and reaction temperature are, respectively, 225 min, 380 rpm, and 353 K, which are used in subsequent experiments. The effect of catalyst concentration on the oxidation reaction is shown in Figure 4a. As the catalyst concentration was increased from 4.4 to 8.7 g L1, the conversion of phenol was improved from 21.4% to 28.5%, and a further increase of catalyst concentration did not significantly improve the overall conversion. Meanwhile, the selectivity first increased until the catalyst concentration was 17.2 g L1 and then remained constant. As can also be seen in Figure 4b, the filtration resistance kept increasing 10461



7

26

95.0

90.0

4.66 3.5:1 8 reactor

reactor

dual-membrane

TS-1 membrane

submerged

TS-1

5

80

2.64:1 80

60 microreactor

microreactor

catalytic wall

TS-1

0.02 catalytic

packed bed

TS-1

0.35

60

0.55:1

3:1 70 2 FeCoNaY reactor

fixed-bed flow

0.55:1

1:4.5

2.5

30

20

2

3.66

225

144

15.0

17.0

44.6 0.16

1 2

0.16

2

47.8

21.8 44

52 8.46 1.25:1 3:1 84 7 TS-1/diatomite reactor

fixed-bed

catalyst reactor

98.4

26 98.5

25 61.3

67.4 12.0

% % mol L1 H2O2(mol) C g

phenol/acetone phenol/ temperature/°

reaction catalyst

concentration/

or phenol/water (wt)

reaction time/h h1

time/min

DHB phenol 10462

operation conditions

Table 2. Comparison of Phenol Conversion and DHB Selectivity in Various Reactors

with higher catalyst concentration. On the one hand, higher catalyst concentrations led to more active titanium moieties available for the phenol hydroxylation, contributing to an increase of the conversion and selectivity. However, it also aggregated the catalyst particles depositing on the membrane surface, and consequently resulted in a higher filtration resistance. Therefore, an optimal catalyst concentration value of 17.2 g L1 is employed for easier control of the reaction. Figure 5a exhibits the effect of the molar ratio of phenol to H2O2 on the phenol hydroxylation reaction, in which the H2O2 concentration was kept at 1.33 mol L1. With increasing molar ratio, there was a gradual decrease in the conversion and increase in selectivity. To make the phenomena clearer, we plotted the DHB selectivity (y-axis) versus phenol conversion (x-axis) in Figure 5a. As shown in the inset of Figure 5a, the low phenol conversion was corresponded to high DHB selectivity. The phenol excessive greatly compared to H2O2 could facilitate the enhancement of DHB selectivity. This result was in good agreement with those reported in the literature.19 It is noted from Figure 5b that the filtration resistance decreased with the increase of the molar ratio. Organic substances can be adsorbed on to the surfaces of the catalyst particles in the cake layer and subsequently change the permeability of the cake layer.3 In this work, less byproduct was formed and absorbed on the catalyst particles in the cake layer, resulting in a loose cake layer and a lower filtration resistance. In the present work, low phenol/H2O2 molar ratio could produce high phenol conversion, but low DHB selectivity and high filtration resistance. For example, when the molar ratio of phenol to H2O2 was 1, the phenol conversion, DHB selectivity, and filtration resistance were 61.0%, 84.3%, and 23.93 1011 m1, respectively. In addition, at this mole ratio, a thick cake layer was formed on the membrane surface, according to visual inspection after the experiment. To keep the liquid level in the reactor constant, the discharge flow rate was adjusted by changing the rotational speed of the peristaltic pump. The thicker was the cake layer, the higher was the rotational speed required for the experiment, which made continuous operation harder to maintain. Therefore, an optimum phenol/H2O2 molar ratio in the feed steam for the hydroxylation reaction was fixed at 3.5. 3.4. Feasibility of the Continuous Reaction Process. The reaction process was run for 30 h, and the results are presented in

phenol concentration/

Figure 6. Operation stability of the coupling system (pore size of membrane distributor is 0.5 μm, pore size of membrane separator is 0.2 μm, residence time is 3.7 h, reaction temperature is 353 K, stirring rate is 380 rpm, catalyst concentration is 17.2 g L1, phenol/H2O2 molar ratio is 3.5, and CH2O2 = 1.33 mol L1).

WHSV/

residence

conversion/

selectivity/

24

ref

this work

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Figure 7. FESEM micrograph of the surface of membrane distributor. (a) Fresh. (b) Used.

Figure 8. FESEM micrograph of the surface of membrane separator. (a) Fresh. (b) Used.

Figure 9. Performance of different usage state catalysts for hydroxylation of phenol (reaction temperature is 353 K, stirring rate is 380 rpm, catalyst concentration is 17.2 g L1, phenol/H2O2 molar ratio is 2.8, and CH2O2 = 1.33 mol L1).

Figure 6. We found that the conversion decreased significantly within the first 4 h and then decreased gradually to ∼15% in the following 26 h, while the selectivity remained at ∼95%. The filtration resistance leveled off through the continuous reaction. Table 2 lists the typical results from this work and compares them with data from different reactors under their respective optimal operation conditions as found in the literature.7,2426 It is noted that the DHB selectivity in this novel dual-membrane reactor was higher than those in the fixed-bed reactor but lower than those from a packed bed catalytic microreactor and catalytic wall microreactor, possibly because of the shorter residence time. A higher selectivity could be obtained with comparable conversion in the novel dual-membrane reactor in contrast to the submerged membrane reactor reported in our previous work,7 indicating that the membrane distributor promotes the selectivity. According to the ICP analysis, the concentrations of silicon and titanium elements in the permeate were close to zero, showing that the TS-1 catalyst particles were almost completely retained in the reactor by the membrane separator. The decline of the conversion at the beginning of the operation is likely to be due to the adsorption of TS-1 catalyst particles

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on the clean membrane surface, leading to the decrease of the catalyst concentration in the reactor. This explanation was further confirmed by FESEM analysis of the fresh membrane and the one used for the hydroxylation of phenol. As can be seen in Figure 7, there were some particles smaller than 400 nm aggregated on the surface of membrane distributor. EDX data of the fouled surface of the membrane distributor indicated that the foulants absorbed on the membrane surface were composed of C, Si, Ti, O, and Al. The C element was attributed to the organic matter in the reaction slurry adsorbed onto the membrane surface. The Si, Ti, and O came from TS-1 particles. Aluminum was one of the components of the membrane material. Figure 8 shows FESEM images of the surfaces of the fresh and used membrane separator. There were almost no catalyst particles deposited on the surface, in agreement with the EDX analysis, resulting in a stable membrane filtration. The H2O2 concentration on the membrane surface was higher than in other regions of the reactor, leading to the generation of large molecular byproducts or tar. Thus, it is easier for the TS-1 catalyst particles to adhere to the surface of membrane distributor when compared to the membrane separator. These results showed that TS-1 particle deposition is one of the main reasons for the decrease of phenol conversion. Zhong et al.27 found that the membrane fouling was mainly caused by the deposition of TS-1 particles, silica additive adsorption, and iron precipitation, possibly due to the construction of the membrane reactor and the reaction in their experiments. The further decrease of phenol conversion might be caused by the adsorption of organic species on the TS-1 surface, which might hinder the diffusion of the reactant into the TS-1 channels. To confirm these assumptions, the performance of the regenerated catalysts for hydroxylation of phenol was compared to that of the fresh TS-1 and the used TS-1, and the results are shown in Figure 9. As can be seen, the conversion of fresh catalyst decreased from 29.9% to 18% after 30 h of operation. After catalyst regeneration at 500 °C in air for 6 h, the performance was largely recovered.

4. CONCLUSIONS In this work, a novel dual-tubular membrane reactor was developed for continuous heterogeneous oxidation catalysis, and its performance was evaluated using the hydroxylation of phenol over ultrafine TS-1 catalyst as the model reaction. It was found that a smaller pore size for the membrane distributor was more efficient and enhanced the DHB selectivity. In a 30 h continuous operation, the selectivity remained at ∼95%, while the conversion rapidly decreased within the first 4 h and then gradually to ∼15%. The decrease of conversion was mainly caused by the adsorption of TS-1 catalysts onto the membrane distributor surface and the catalyst deactivation due to the adsorption of reaction mixture on the catalysts surface. The membrane separator was able to retain almost all of the TS-1 catalyst. The present work demonstrated that the dual-membrane reactor is feasible for heterogeneous oxidation catalysis and is effective for the enhancement of selectivity and separation of ultrafine catalysts. Further research into the mechanism and modeling of the reaction system is underway. ’ ASSOCIATED CONTENT

bS

Supporting Information. Effect of residence time, effect of stirring rate, effect of reaction temperature, and particle size of

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Industrial & Engineering Chemistry Research TS-1 catalyst. This material is available free of charge via the Internet at http://pubs.acs.org.

’ AUTHOR INFORMATION Corresponding Author

*Tel.: +86-25-83172266. Fax: +86-25-83172292. E-mail: wqjin@ njut.edu.cn.

’ ACKNOWLEDGMENT Financial support from the National Natural Science Foundation (20990222), the National Basic Research Program (2009CB623406), the Natural Science Foundation of Jiangsu Province (BK2010549, BK2009021), and the Natural Science Foundation of the Higher Education Institutions of Jiangsu Province (09KJB530006) of China is gratefully acknowledged. ’ REFERENCES (1) Hutchings, G. J. New approaches to rate enhancement in heterogeneous catalysis. Chem. Commun. 1999, 4, 301. (2) Zhong, Z. X.; Xing, W. H.; Jin, W. Q.; Xu, N. P. Adhesion of nanosized nickel catalysts in the nanocatalysis/UF system. AIChE J. 2007, 53, 1204. (3) Lee, S. A.; Choo, K. H.; Lee, C. H.; Lee, H. I.; Hyeon, T.; Choi, W.; Kwon, H. H. Use of ultrafiltration membranes for the separation of TiO2 photocatalysts in drinking water treatment. Ind. Eng. Chem. Res. 2001, 40, 1712. (4) Meng, Y. B.; Huang, X.; Yang, Q. H.; Qian, Y.; Kubota, N.; Fukunaga, S. Treatment of polluted river water with a photocatalytic slurry reactor using low-pressure mercury lamps coupled with a membrane. Desalination 2005, 181, 121. (5) Yang, W. B.; Cicek, N.; Ilg, J. State-of-the-art of membrane bioreactors: Worldwide research and commercial applications in North America. J. Membr. Sci. 2006, 270, 201. (6) Chen, R. Z.; Bu, Z.; Li, Z. H.; Zhong, Z. X.; Jin, W. Q.; Xing, W. H. Scouring-ball effect of microsized silica particles on operation stability of the membrane reactor for acetone ammoximation over TS-1. Chem. Eng. J. 2010, 156, 418. (7) Lu, C. J.; Chen, R. Z.; Xing, W. H.; Jin, W. Q.; Xu, N. P. A submerged membrane reactor for continuous phenol hydroxylation over TS-1. AIChE J. 2008, 54, 1842. (8) Tian, Y.; Chen, L.; Jiang, T. L. Simulation of a membrane bioreactor system for wastewater Organic removal: biological treatment and cake layer- membrane filtration. Ind. Eng. Chem. Res. 2011, 50, 1127. (9) Vyrides, I.; Stuckey, D. C. Saline sewage treatment using a submerged anaerobic membrane reactor (SAMBR): Effects of activated carbon addition and biogas-sparging time. Water Res. 2009, 43, 933. (10) Lu, Y. P.; Dixon, A. G.; Moser, W. R.; Ma, Y. H. Analysis and optimization of cross-flow reactors with distributed reactant feed and product removal. Catal. Today 1997, 35, 443. (11) Wang, Y. J.; Zhang, C. L.; Bi, S. W.; Luo, G. S. Preparation of ZnO nanoparticles using the direct precipitation method in a membrane dispersion micro-structured reactor. Powder Technol. 2010, 202, 130. (12) Xu, J. H.; Luo, G. S.; Chen, G. G.; Tan, B. Mass transfer performance and two-phase flow characteristic in membrane dispersion mini-extractor. J. Membr. Sci. 2005, 249, 75. (13) Coronas, J.; Menendez, M.; Santamaria, J. Methane oxidative coupling using porous ceramic membrane reactors-II. Reaction studies. Chem. Eng. Sci. 1994, 49, 2015. (14) Tonkovich, A. L. Y.; Zilka, J. L.; Jimenez, D. M.; Roberts, G. L.; Cox, J. L. Experimental investigations of inorganic membrane reactors: A distributed feed approach for partical oxidation reactions. Chem. Eng. Sci. 1996, 51, 789.

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