Microreactor Performance Studies of the Cycloaddition of Isoamylene

Ind. Eng. Chem. Res. 2010, 49, 5549–5560

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Microreactor Performance Studies of the Cycloaddition of Isoamylene and r-Methylstyrene Obiefuna C. Okafor,*,† Sunitha Tadepalli,‡ Geatesh Tampy,‡ and Adeniyi Lawal† New Jersey Center for MicroChemical Systems, Department of Chemical Engineering and Materials Science, SteVens Institute of Technology, Castle Point on Hudson, Hoboken, New Jersey 07030, and International FlaVors & Fragrances, 521 West 57th Street, New York, New York 10019

The cycloaddition reaction between isoamylene and R-methylstyrene yields indane compounds 1,1,2,3,3,pentamethylindane and 3-ethyl-1,1,3-trimethylindane, which are intermediate cyclic products used in the synthesis of musk fragrances. This exothermic reaction is usually carried out industrially in large semibatch reactors. The microreactor, with enhanced heat and mass transfer characteristics, was used for the reaction with aqueous sulfuric acid as catalyst. The dependence of reactant conversion, product yield, and average reaction rates on catalyst concentration, temperature, velocity, residence time, and the molar ratio of the reactants in the feed was investigated. A similar study was also performed in the semibatch reactor to compare its performance with that of the microreactor. Higher product yields were obtained in the microreactor, and the average reaction rates in the microreactor were 3 orders of magnitude greater than those obtained in the semibatch reactor. 1. Introduction The cycloaddition reaction between isoamylene (IA) and R-methylstyrene (AMS) yields indane compounds 1,1,2,3,3,pentamethylindane and 3-ethyl-1,1,3-trimethylindane (CP), which are intermediate cyclic products that are used to synthesize musk fragrances used in perfumery. Musks refine, fix, balance, stimulate, and harmonize perfumery compositions.1 Nearly all fragrances contain musk odorants since they basically form the skeleton of a perfume. Of the many polycyclic aromatic musks synthesized over the last century, Galaxolide (1,3,4,6,7,8hexahydro-4,6,6,7,8,8-hexamethylcyclopenta-gamma-2-benzopyran), originally discovered in 1962,2 turned out to be the best in terms of stability, hydrophobicity, odor threshold, and its relatively cheap synthesis. Galaxolide is synthesized by the Friedel-Crafts alkylation of CP with methyloxirane and the acid-catalyzed reaction with paraformaldehyde.1 Today, in spite of the synthesis of a few other macrocyclic musks, Galaxolide is still the most widely used musk with a production capacity of 7000-8000 tonnes per annum.1 CP is also useful in the production of a variety of polycyclic compounds which are used as intermediates in the production of perfumery compositions.3 While a lot of patents have been issued on the production of indane derivatives for use as intermediates for musk fragrances, not much is available elsewhere in open literature. A method for the production of indane derivatives was first reported by Barbier,4 when he discovered that they could be produced by the reaction of p-cumene with isobutylalcohol in the presence of sulfuric acid. Researchers have reported the use of various techniques to improve the yield of the indanes from the cycloaddition reaction between IA and AMS using aqueous sulfuric acid catalyst. Cobb5 discovered that the use of a tetrahydrothiophebe 1,1-dioxide solvent could improve yields, and also suggested that acyclic dimers of AMS could also be employed in the production of the indane compounds. In another patent, Cobb6 also suggested that adding sulfolane to the sulfuric * To whom correspondence should be addressed. Tel.: +1 201 216 8314. E-mail: [email protected]. † Stevens Institute of Technology. ‡ International Flavors & Fragrances.

acid catalyst, along with tungsten and molybdenum, or their oxides could be used to improve CP yields. All the processes suggested by patent literature with the use of not just sulfuric acid but other Fiedel-Crafts catalysts such as aluminum chloride and phosphoric acid,3 the use of titanium tetrachloride with cumyl chloride instead of AMS,7 and the use of acid clays and acid ion-exchange resin catalysts8 are performed in the batch and semibatch reactors. However, efficient contacting of the reactant and catalyst phases is a major limitation when aqueous sulfuric acid is used as a catalyst, because of the immiscibility of the aqueous catalyst phase and the organic reactant phase. Ghosh and Chaudhuri9 showed that one of the side reactions, the dimerization of AMS using aqueous sulfuric acid catalyst, takes place in the film of the aqueous layer. Since the cycloaddition reaction is most likely a fast reaction that occurs in the film of the aqueous catalyst layer, the interfacial area between the two phases is of utmost importance. Ende et al.10 proposed that the interfacial area between sulfuric acid and hydrocarbons in a batch reactor was affected by the agitation rate, the volume percent of acid in the reactor, the compositions of both the acid and hydrocarbon phases, and the temperature. Mass transfer in the batch reactor is typically increased by increasing agitator speed after employing other means including the introduction of baffles and other improvements in the impeller design. However, vessel structural stability and agitator drive motors impose limits on the speed of agitators. The use of high agitator speeds increases the possibility of excessive heat generation, which may have an adverse impact on product quality, and it can also cause foaming and vapor entrainment in the headspace.11 Also because of the high demand for these indane products, very large batch and semibatch reactors, with concomitant high mass transfer resistances, are used in its production. Microreactors however offer the opportunity to greatly enhance the mass transfer between the two immiscible phases of aqueous sulfuric acid catalyst and organic reactants. Microchemical systems are sometimes defined as “miniature reaction and other unit operations, possessing specific adVantages over conventional chemical systems”.12 Microreactor technology offers numerous advantages and potential because of its ability to intensify heat and mass transport properties as

10.1021/ie901794p  2010 American Chemical Society Published on Web 05/07/2010

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well as improve flow patterns. Microreactors have been fabricated using various materials of construction such as metals, ceramics, glass, and polymers.13 The use of the microreactor for the cycloaddition reaction has the potential to offer a number of benefits. Better heat and mass transfer in a microreactor can improve selectivity, safety, yield, product quality, and also improve energy efficiency.14-18 The use of the microreactor for the cycloaddition reaction could lead to significant cost savings arising from faster transfer of research to production scale by the numbering-up approach, which increases throughput by parallel operation of a multitude of microreactors, rather than the complex and cost intensive scaling up approach. Complex liquid-liquid reactions like the cycloaddition reaction also present some of the most difficult scale-up problems because of complexities of drop formation and coalescence that change with scale.11 Other advantages of the use of microreactors for this reaction include improved economics by the replacement of batch processes by continuous processes, which has been shown to be economical in the production of bulk chemicals,19 intensification of processing, smaller plant size for mobile applications and distributed production, lower costs for transportation, and improved safety from reduced worker exposure to hazardous chemicals. While the potential exists for every chemical engineering unit operation to be miniaturized, a lot of industrially relevant reactions have already been carried out successfully in microreactors. Specifically, about 50% of reactions in the fine chemicals and pharmaceutical industries could benefit from a continuous process based mainly on microreactor technology,19 and since the 1990s, a lot of work has been done using microreactors for both homogeneous and heterogeneous reactions. In particular, microreactors have also been applied successfully in liquid phase reactions.20-24 Since rapid microfluidic mixing is difficult to achieve with laminar flow on account of the low Reynolds number in microchannels, mixing in microchannels is principally limited to molecular interdiffusion through the fluid interface.14,25,26 Rapid and effective mixing due to the short transverse diffusional time obtained in the microreactor can quickly bring reactants in contact with the aqueous sulfuric acid catalyst to increase conversion.27 Various methods for contacting immiscible liquids have been proposed. Harper28 patented a device that would allow diffusive transfer between immiscible liquids by bringing them in contact with each other without allowing physical mixing of the liquids. However this method of parallel fluid flow with only diffusive transfer requires channels that are below 100 µm to work effectively for rapid reactions. Such tiny channels imply that only low velocities may be used through these channels because of high pressure drops, and consequently the low throughput capabilities of such devices are inadequate for chemical production. Another method of contacting immiscible liquids that would allow greater throughput through the microreactor would be the use of slug flow. Liquid slugs can be used to generate high frequency internal circulation due to the combination of shear within the microchannel and interfacial phenomena. This internal circulation provides a means of rapid transport to and from the interface, thereby increasing the concentration gradients and mass transfer performance.29 Various researchers have studied slug flow experimentally and numerically in microchannels.30-32 Passive mixers that involve the use of rapid mixing based on multilayer flows, such as the use of oriented ridges on the floor of the channel to cause chaotic mixing,33 the use of zigzag microchannels,34 and the repeated dividing and merging of fluids,35 can also be used for high throughput immiscible liquid-liquid reactions. Seong and

Crooks25 showed yet another strategy to effective mixing in a microreactor, one that could still yield the high throughput required industrially. This approach involved immobilizing catalysts on microbeads, placing the beads in well-defined microreactor zones, and then passing reactants through these microreactor zones to yield products. This approach effectively mixed the reactants and increased the effective surface area of the microchannel interior. This work reports the use of a microchannel reactor system for the cycloaddition reaction between AMS and IA involving an immiscible liquid-liquid system comprising two miscible reactants and the immiscible sulfuric acid catalyst phase, and compares the microreactor performance with that of the semibatch reactor. The selection of an optimal microreactor design and the use of inert packing material to enhance mixing is discussed. A parametric study was undertaken to optimize the cycloaddition reaction in the microreactor by investigating the dependence of reactant conversion, product yield, and average reaction rates on catalyst concentration, temperature, velocity, residence time, and the molar ratio of the reactants in the feed. A parametric study on the cycloaddition reaction was also performed in the semibatch reactor to compare its performance with that of the microreactor. The dependence of the reactant conversions and product yields on experimental reaction time, temperature, stirrer speed, and catalyst concentration in a semibatch reactor was investigated. 2. Experimental Section 2.1. Materials. The reactants for the cycloaddition reaction, AMS (99%, ACS reagent grade) and IA were obtained from Sigma-Aldrich. The IA obtained was a mixture of two IA isomers: 15% 2-methyl-1-butene and 85% 2-methyl-2-butene, as determined by gas chromatographic analysis. The catalyst was obtained by diluting sulfuric acid (95-98%, Sigma-Aldrich, ACS reagent) with deionized water after standardizing it via titration with oxalic acid dehydrate crystals (J.T. Baker). Sodium hydroxide pellets (97+%, ACS reagent grade) used in sulfuric acid standardization and effluent neutralization were also obtained from Sigma-Aldrich. Spherical glass beads, within a size range of 75-150 µm, used as reactor packing material, were obtained from Greenhill Supply Co., NJ. The glass beads were smooth and nearly without any pores with a surface area of just 0.032 m2/g for glass beads (multipoint BET technique using Quanto-chrome Instruments Autosorb-1). The cyclic products, 1,1,2,3,3,-pentamethylindane and 3-ethyl-1,1,3-trimethylindane (CP), used in GC standardization, were obtained from vacuum distillation of the reactor effluents. 2.2. Experimental Procedure. Figure 1 shows the experimental setup used for the cycloaddition reaction between IA and AMS. Since the reactants, AMS and IA, do not react together in the absence of a catalyst, they were premixed and pumped to the reactor. The catalyst was pumped to the microreactor using a separate pump. HPLC pumps with 10 mL/ min pump heads were used. The various components of the experimental setup were all connected by capillary tubing. Both the reactants and the catalyst were initially contacted in a T-junction before entering the microreactor. The microreactor was constructed of ETFE or Hastelloy C tubing, with an internal diameter in the range of 500 µm to 2.67 mm. The microreactor was packed with inert glass beads and then placed in a circulating water bath, which kept the temperature of the circulating water constant with an accuracy of (0.02 K. The microreactor exit pressure was atmospheric. The volumetric flow rate of the catalyst was the same as that of the reactants into


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Figure 1. Laboratory microreactor setup for cycloaddition reaction.

Scheme 1. Mechanism of the Cycloaddition Reaction39

the microreactor for all the experiments. An equimolar ratio of the reactants was also used in all the experiments, except for that described in section 3.2.3, where the effect of varying the molar feed ratio of the reactants was reported. The total volumetric flow rate through the microreactor ranged from 0.1 to 3 mL/min, except for in the experiment described in section 3.2.5, where higher volumetric flow rates were also used. For all the microreactor experiments, the values reported for each experimental point are an average of at least three product samples measured after the microreactor performance had been stabilized for at least 30 min. Samples were taken at regular intervals of at least 20 min. The experimental runs were also repeated successfully to ensure repeatability. Experiments to compare the performance of the microreactor with the semibatch reactor were conducted in a 25 mL PARR 5500 compact batch reactor, and reported in section 3.3. The reaction was carried out in the semibatch mode, rather than in the batch mode because the cycloaddition reaction and all the other major reactions occurring concurrently are exothermic reactions as shown by the heat of reaction values in Schemes 1-3. At the beginning of each experiment, the reactor was charged with 4 mL of the catalyst before being properly sealed, since the reactor was operated under autogenous pressure. After the desired reactor temperature was attained, the stirrer was

switched on to the desired speed, the reactant was fed to the reactor at a flow rate of 0.08 mL/min for 3 h (except for some of the experiments investigating the effect of experimental run time in section 3.3.1), and the stirring continued for one more hour. The reaction temperature for all the semibatch experiments was 65 °C, and the stirrer speed was 1500 rpm, except where otherwise stated. 2.3. Analytical Procedure. The reactor effluents/product sample contained two immiscible liquid phases: the aqueous catalyst phase and the organic phase which contained the reaction products and unreacted reactants. Because these two layers were immiscible, they were readily separated by gravity to produce two distinct layers. The upper organic phase was collected and first mixed in 25 wt % sodium hydroxide to neutralize any traces of sulfuric acid, then extracted again and mixed in water to remove any traces of sodium hydroxide that might contaminate the GC column. A Varian CP-3800 gas chromatograph equipped with an Agilent DB-5, 30 m column with 0.53 mm ID and 1 µm film thickness, a Varian 8410 autosampler, and a thermal conductivity detector was used for the analysis. Helium (ultra-high-purity gas, Praxair) was used as the carrier gas. A 1 µL aliquot of each sample was vaporized at 250 °C at the injector, and the temperature program started

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Scheme 2. Major Products of IA Dimerization Identified via GC/MS40

Scheme 3. Reaction Network for AMS Dimerization41 (Heat of Reactions Were Obtained from the Heat of Formation of the Products That Were Calculated Based on the Cohen and Benson42 Method)

molecular weight of CP, R is the proportion of catalyst flow rate compared to the total inlet flow (0.5), and V is the volume of the microreactor (L). Since the aqueous catalyst phase and organic phases are immiscible, it is assumed that the organic mass flow rate into the reactor is equal to the organic flow rate out of the reactor, even though it is likely that a very small portion of the organic phase ends up in the aqueous catalyst phase. A fractional void space of 0.4 was assumed for all the packed microreactors used in this work.18 The average reaction rate in the semibatch reactor was calculated as ARR )

at 75 °C, was ramped up to 250 °C at 8 °C/min, and then kept at 250 °C for 10 more minutes. Standards containing pure samples of IA, AMS, and CP were used in standardizing the GC using the external standardization method. All the calibration curves generated were straight lines with regression coefficients very close to one (>0.97). All calibration curves also passed through the origin, which suggested that neither the column nor the detector was overloaded, that the electronics were responding linearly, and that there was no apparent component adsorption in the injection port, the column, the detector, or the associated plumbing.36 Productivity of CP was expressed as the average reaction rate (ARR) and the yield37 calculated as ARR ) yield )

XprodFo,o εVMCPR

(1)

moles of CP formed 100 moles of CP that would have been formed if there were no side reactions and the limiting reactant had reacted completely (2)

Conversion37 of the reactants was expressed as conversion )

Xr,iFo,i - Xr,oFo,o Xr,iFo,i

(3)

where Xprod is the mass fraction of the product in the organic phase outlet, Xr is the mass fraction of the reactant in the organic phase, Fo is the mass flow rate of the organic phase (g/h), while the subscripts i and o stand for inlet and outlet, respectively, ε is the fractional void space of the packed bed, MCP is the

XprodWreactants VcatMCPt

(4)

where Xprod is the weight fraction of CP, Wreactants is the mass of reactants fed to the reactor (g), Vcat is the volume of the catalyst in the semibatch reactor (L), and t is the reaction time (h). 3. Results and Discussion 3.1. Reaction Pathway. A mechanism for the cycloaddition reaction (Scheme 1) has been suggested by Shah et al.,38 in which AMS is protonated to give a cumyl cation which reacts with isoamylene to give the two cyclic product isomers of interest: 1,1,2,3,3,-pentamethylindane and 3-ethyl-1,1,3-trimethylindane (CP). In this work, the yield and ARR refer to the combined values for these two product isomers of interest. Apart from these major products, a plethora of other byproducts are also produced during the cycloaddition of AMS and IA with H2SO4 as a catalyst. The byproducts of the cycloaddition reaction are dimers of IA (DIA), dimers of AMS (DAMS), and other isomers of CP. Scheme 2 shows the four major isomers of DIA (C10H20), 3,4,5,5-tetramethyl-2-hexene, 2,3,4,4-tetramethyl-1-hexene, 3,4,4,5-tetramethyl-2-hexene, and 3,5,5-trimethyl-2-heptene, that were detected as byproducts of the cycloaddition reaction via GC/MS analysis. These dimers of IA were also obtained as reaction products by Cruz et al.39 in the dimerization and etherification of isoamylenes using ion exchange resins in the presence of ethanol and methanol, and also by Shah and Sharma40 in their study on the dimerization of isoamylene using a cation exchange resin and Filtrol-24 as catalysts. The DAMS that were identified by GC/MS as byproducts of the cycloaddition reaction are the unsaturated DAMS, 2,4-diphenyl-4-methyl-1-pentene and 2,4-diphenyl-4methyl-2-pentene, and the saturated DAMS, 1,1,3-trimethyl-3phenylindane. Scheme 3 shows a reaction network for the dimerization of AMS proposed by Chaudhuri and Sharma.41 These byproduct have important industrial relevance with DIA being used extensively as an important intermediate in the


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Figure 2. 2-D image of a packed-bed entrance and schematic of possible mixing of two streams.43

perfumery and flavor industry,40 while the unsaturated dimers of AMS are used industrially as chain-transfer agents or molecular weight regulators in the production of polymers such as polystyrene, AS resin, and ABS resin.41 3.2. Microreactor Studies. 3.2.1. Fluid Flow Characterization. The packed-bed microreactor was chosen as the microreactor for this cycloaddition reaction because of the ability of the reactor packing to introduce tortuosity, reduce channel size, and increase contact areas between fluid elements. Figure 243 shows a 2-D image of a packed bed and the schematic of possible mixing by the splitting (or bifurcation) and recombination of flow streams via a multilayered approach. This multilayered approach enhances contacting/mixing between the two immiscible fluid layers and hence increases the average reaction rate. While the diagram basically shows the bifurcation and recombination of the fluid streams caused by fixed pillars in a microreactor, a packed bed microreactor offers the same advantages with a greatly increased number of “pillars” in 3-D to cause more bifurcation and recombination of the flow streams. The packed bed microreactor approach involves placing inert glass beads into a well-defined microreactor zone, and then passing reactants through the reactor zone to yield products. The inert glass beads also increase the effective surface area of the channel interior, which improves reaction velocities compared to open channels. Multilayered flow, rather than slug flow, was verified by flow visualization experiments. This multilayered approach in a packed bed microreactor, that promotes both laminar and turbulent mixing, was also demonstrated by Seong and Crook25 when they performed two sequential reactions catalyzed by glucose oxidase and horseradish peroxidase using microbead supported catalysts as a general route to chemical synthesis within microfluidic systems. 3.2.2. Dependence of Conversion, Yield, and ARR on Catalyst Concentration and Residence Time. The effect of the concentration of sulfuric acid catalyst was investigated to obtain the optimum catalyst concentration. The results from Figure 3 show that at the operating conditions used, the cycloaddition reaction does not occur until a catalyst concentration of at least 72 wt % is attained in the microreactor. The yield of CP continued to increase with an increase in catalyst concentration until a catalyst concentration of 90 wt %, above which the effluent became difficult to separate because of the probable formation of sulphonates. However, even at a low catalyst concentration of 65 wt % the reactants readily dimerized to form DIA and DAMS with a reactant conversion of 26 and 35% for AMS and IA, respectively. Depending on the speed of a reaction, the residence time of the reactants in a reactor can play a vital role in conversion.44 The low reactant conversions at low concentrations of the sulfuric acid catalyst suggests that the activity of the catalyst at those concentrations is too low

Figure 3. Dependence of yield, ARR, and conversion on catalyst concentration (reaction conditions: total flow rate, 3 mL/min; temp, 35 °C; reactor length, 6 cm).

Figure 4. Dependence of yield, ARR, and conversion on residence time (reaction conditions: velocity, 1.11 cm/s; total flow rate, 3 mL/min; 90 wt % acid; temp, 35 °C).

for complete reactant conversion within the low residence time (2.2 s) available in the microreactor at the operating conditions used. Ghosh and Chaudhuri9 in their study of the dimerization of AMS using an aqueous sulfuric acid catalyst concentration of 40 wt % in the batch reactor observed that the AMS conversion increased from 6.9% at a reaction time of 1.83 h to 41% after 15 h. However, they also observed an increase in the average reaction rate shown by a reactant conversion of 96% after a reaction time of just 30 min using a sulfuric acid catalyst concentration of 98 wt %. As in the dimerization of AMS, the cycloaddition reaction requires a more concentrated sulfuric acid catalyst to improve reaction rates within small residence times. To study the effect of residence time, all the experiments were conducted at the same velocity, which necessitated adjusting the reactor lengths so as to separate the effects of velocity and residence time. The small dimensions of the microreactor imply that the residence times of the reactants and catalyst in the reactor are much smaller than those obtained in conventional reactors. However, the residence time of the reactants in the reactor must be at least as much as the reaction time needed for complete reactant conversion. Figure 4 shows that product yields and reactant conversions were not increased with increased residence time, suggesting that the reaction time for the cycloaddition reaction was less than 5.43 s when using

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Figure 5. Dependence of (a) yield, (b) IA conversion, (c) AMS conversion on residence time at 35 °C using 72 wt % sulfuric acid catalyst at different volumetric flow rates through the microreactor.

a catalyst concentration of 90 wt %. Increasing the residence time of the reactants in the microreactor by increasing the reactor length means that the volume of the catalyst in the microreactor was increased. This increase in catalyst volume at a high catalyst concentration means that the reactant concentration in the downstream of the microreactor becomes small due to the high conversion of the reactants in the upstream of the microreactor, leading to a reduction in average reaction rates, which is also shown in Figure 4. However, with increased residence time it is possible to increase the reactant conversions and product yields at lower catalyst concentrations. This is shown in Figure 5, where an increase in yield and reactant conversions with respect to residence time is observed at different volumetric flow rates through the reactor using 72 wt % sulfuric acid catalyst. The increase in yield with velocity also shows that there are mass

transfer limitations at such low velocities. The leveling off of the yield and reactant conversions at a residence time of about 120 s suggests that the reaction time for the cycloaddition reaction is 120 s using 72 wt % sulfuric acid catalyst at those operating conditions. Obtaining large residence times in microreactors usually requires the use of long microreactors even when low volumetric flow rates through the microreactor are used. However increasing the length of the packed microreactor also directly increases the pressure drop along the microreactor, which increases the power requirement for the microreactor as well as possibly introducing other operational problems during multichannel operation. The increase in CP yield when the catalyst concentration was increased from 83 to 90 wt %, in spite of the fact that the starting reactants were fully converted at a catalyst concentration of 83 wt %, could be attributed to the deoligomerization of the unsaturated DAMS: 2,4-diphenyl-4-methyl-1-pentene and 2,4diphenyl-4-methyl-2-pentene at high catalyst concentrations, and the subsequent cycloaddition of the deoligomerized DAMS with IA or DIA to form CP. Sharma et al.38 in their study of the cycloaddition reaction between IA and AMS in a batch reactor using Amberlyst-15 catalyst noticed a reduction in the amount of DAMS as the reaction progressed and showed that unsaturated DAMS could deoligomerize to AMS and then react with IA to form CP. Similar maxima in DAMS amounts during the experimental reaction time in a batch reactor were also noticed in the dimerization of AMS,41 and the alkylation of phenol using AMS.45 However, like the cycloaddition reaction, the very low residence times obtained in the microreactor compared to the batch or semibatch reactor means that there is insufficient reaction time for the deoligomerization of unsaturated DAMS, rather the rate of deoligomerization of the unsaturated DAMS is likely increased by increased catalyst concentration. This increase in the rate of deoligomerization of the unsaturated DAMS with an increase in catalyst concentration explains the increase in the yield of CP from 18% at 83 wt % catalyst to 34% at 90 wt % catalyst concentration even after the reactant conversions were nearly 100% at a catalyst concentration of 83 wt % as is observed in Figure 3. Since sulfuric acid is relatively cheap because of its extensive industrial use, a high concentration of the catalyst would not significantly affect the cost of production of CP. 3.2.3. Dependence of Conversion, Yield, and ARR on Reactants Inlet Molar Ratio. Cobb6 suggested that the molar ratio of the reactants did not have to be one to obtain optimum product yield, hence this effect was studied in the microreactor. The results from Figure 6 show that while an equimolar amount of the reactants seems to have given the highest ARR, the yield amounts for the different reactant inlet molar ratios are similar within experimental margins. Since the ARR is of greater industrial importance than the yield, a molar ratio of one was used for the rest of the experiments. It was observed that an excess in either of the reactants led to an increase in the yield of the respective reactant dimer. This ability to preferentially produce reactant dimers is especially relevant when there is a need for the production of the reactant dimers as intermediates in the production of other compounds.40,41 Also since IA is more expensive than AMS, the fact that an excess amount of IA in the inlet feed does not improve the ARR favors the economics of the cycloaddition reaction in the microreactor. The higher average reaction rates of the reactant dimerization reactions compared to the cycloaddition reaction also mean that the reactants readily dimerize at all IA/AMS molar ratios, leading to the high reactant conversions noticed at all IA/AMS ratios.


Figure 6. Dependence of yield, ARR, and conversion on reactants inlet molar ratio (reaction conditions: total flow rate, 1 mL/min; 90 wt % acid; temp, 65 °C; reactor length, 14 cm).

Figure 7. Dependence of yield, ARR, and conversion on temperature (reaction conditions: total flow rate, 3 mL/min; 90 wt % acid; reactor length, 6 cm).

3.2.4. Dependence of Conversion, Yield, and ARR on Temperature. The effect of temperature was also investigated to determine the optimum operating temperature in the microreactor. An increase in temperature generally increases the average reaction rate of reactions, and this effect along with a corresponding increase in product yield is shown in Figure 7. AMS conversion was very high at all the temperatures used, showing the high selectivity to the formation of AMS dimers. However the conversion of IA decreases slightly with an increase in temperature. This is most likely due to the vaporization of some IA molecules at high temperatures because of its low boiling point (38 °C). The vaporized molecules were unable to take part in the liquid phase reaction, and thus they condensed in their unreacted form at ambient temperature outside the constant temperature water bath, just before sample collection. A way to counteract this would be to operate the microreactor system at a pressure higher than the vapor pressure of the reactant mixture, but this would introduce more operational costs that would make the microreactor system less economical. 3.2.5. Dependence of Conversion, Yield, and ARR on Velocity. Convective mass transfer in the continuous packedbed microreactor depends on the flow velocity because convection increases with an increase in flow velocity and vice versa. When the reaction is controlled by convective mass transfer, an increase in velocity should lead to an increase in the average reaction rate; however, there is no increase in average reaction rate if the reaction is in the kinetically controlled regime.

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Figure 8. Dependence of yield, ARR, and conversion on velocity (reaction conditions: residence time, 5.43 s; 90 wt % acid; 35 °C).

Figure 9. Dependence of conversion and product yield on experimental reaction time with 70 wt % catalyst.

However, merely increasing the velocity of the feed (reactants and catalyst) through the reactor may lead to results that are not conclusive because the residence time of the feed (reactants and catalyst) in the microreactor changes with velocity, thus the two effects are superimposed on each other. An increase in velocity for the same reactor length may not only increase the reaction rate because of the increased mass transfer effect, but it might also increase the reaction rate because of the increase in the average concentration of the reactants in the reactor since there is a lower conversion of reactants at higher velocities. Therefore, to separate these two effects, the residence time for all the experiments aimed at investigating the effect of velocity on reaction rate was kept the same by varying the reactor length. Figure 8 shows the effect of velocity at a constant residence time of 5.43 s with a range that covers the velocities used in this work. Since the average reaction rate is independent of the velocity at the operating conditions used, it can be assumed that those experiments were performed in the kinetically controlled regime and free of any convective mass transfer effects. 3.3. Experimentation in the Semibatch Reactor. 3.3.1. Dependence of Conversion and Product Yield on Experimental Reaction Time. Figure 9 shows the product yields and reactant conversions for experiments with a total reaction time of 2, 4, and 6 h, using 70 wt % acid catalyst at 65 °C with a stirrer speed of 1500 rpm. Each experiment was initially run in semibatch mode and then allowed to stay in batch mode (without any inlet feed) for the last hour. The flow rates

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Figure 10. Dependence of conversion and product yield on reactor temperature with 90 wt % catalyst.

of the reactants into the semibatch reactor were selected such that the same total volume of reactants was fed into the reactor for each experiment. From the data, the final product yields were similar for all the experimental reaction times. However during the experiment with a total reaction time of 2 h, it was noticed that while the reactants were being fed into the semibatch reactor, the reactor temperature fluctuated much more than the system accuracy of (2 °C during the first hour of the experiment. It is likely that the amount of heat released by the exothermic reaction was not quickly dissipated by the semibatch reactor cooling jacket. This fluctuation in reactor temperature was not noticed in the experiments with total reaction times of 4 and 6 h. All the other semibatch experiments in this work were performed with a total reaction time of 4 h since the smaller reactant feed rates with this reaction time allowed for more efficient heat dissipation from the semibatch reactor. 3.3.2. Dependence of Conversion and Product Yield on Reactor Temperature. Figure 10 shows the dependence of conversion and product yield on temperature using 90 wt % acid catalyst with a stirrer speed of 1500 rpm. From the data it appears that the cycloaddition reaction was not controlled by kinetics at the selected experimental conditions because the product yield would be expected to increase sharply with an increase in temperature if the reaction was kinetically controlled. Ende et al.,10 in a preliminary investigation of the temperature effects on the interfacial area between hydrocarbons and sulfuric acid catalyst, suggested that temperature affects the interfacial area when they noticed that interfacial area was increased as the reactor temperature was increased from 15 to 19.5 °C. A similar increase in interfacial area between the aqueous catalyst and reactants during the cycloaddition reaction with temperature could also contribute to the increase in product yields shown in Figure 10. The high reactant conversions is due to the high selectivity to the formation of product isomers and DAMS, which is most likely also a function of the high catalyst concentration used, as is explained in the following section. The lower conversion of IA at higher temperatures, also noticed in the microreactor, is most likely due to the vaporization of small portions of IA because the system was run at autogenous pressure. Thus the vaporized IA molecules are unable to take part in the liquid phase reaction. This effect was also noticed in the microreactor, but is probably more prominent in the semibatch reactor because of the larger void space above the reactants in the semibatch reactor, compared to the smaller void space in the microreactor.

Figure 11. (a) Dependence of conversion and product yield on stirrer speed with 70 wt % catalyst; (b) dependence of conversion and product yield on stirrer speed with 90 wt % catalyst.

3.3.3. Dependence of Conversion and Product Yield on Stirrer Speed and Catalyst Concentration. Increasing the rate of mass transfer in a batch or semibatch reactor can be simply achieved by increasing the stirrer speed of the reactor. This increased level of agitation also increases the interfacial area between the two liquid phases,10 which should also lead to increased product yields. Figure 11 panels a and b show the dependence of conversion and product yield on stirrer speed at 65 °C using 70 and 90 wt % sulfuric acid catalyst. The graphs show clearly increased product yield with an increase in stirrer speed at these catalyst concentrations, which was due to an increased dispersion of the acid catalyst within the reactants caused by a decrease in the acid droplet sizes. However the continued increase in product yield with stirrer speed, even up to 1500 rpm, shows that the reaction was still controlled by mass transfer. So even though the upper limit of the magnetic drive stirrer is 1500 rpm, it is likely that the product yields will continue to increase at higher agitation rates. Such mass transfer dependent reactions are common in the industry and are excellent candidates for operation and study in the microreactor.16 The lower product yields observed at 90 wt % acid catalyst compared to the higher product yields observed at 70 wt % acid catalyst in the semibatch reactor were most likely caused by the increased viscosity of the sulfuric acid catalyst with increased concentration. As the viscosity of the catalyst increased, the energy requirements for obtaining dispersions also increased.10 So, since the viscosity of the 90 wt % catalyst (7.5 cP at 65 °C)46 is nearly twice the viscosity of the 70 wt % catalyst (4 cP at 65 °C),46 better dispersions are obtained at the


same agitator speeds using the 70 wt % catalyst compared to the 90 wt % catalyst because of the relatively reduced energy input. Better dispersions using the 70 wt % catalyst means that the interfacial area available for reaction is increased, and hence the increased yields at the same agitator speeds using 70 wt % compared to 90 wt % catalyst. This dependence of yield on the catalyst viscosity was also confirmed by the reduced reactant conversions noticed using the more viscous 90 wt % catalyst in the semibatch reactor, where the cycloaddition reaction was limited by mass transfer. This contrasts with the case of the microreactor, where the more efficient mass transfer in the microreactor overcame any mass transfer resistance caused by an increase in catalyst viscosity, leading to high product yields at 90 wt % catalyst concentration. The high reactant conversions were a result of the reactant dimerization and product isomerization reactions that occurred even at low stirrer speeds. The high selectivity to reactant dimerization seemed less for the dimerization of IA than the dimerization of AMS, which proceeded readily even at the high viscosity of the 90 wt % acid catalyst, as was shown by the high AMS conversions in Figure 11b. It is therefore likely that the average reaction rate of AMS dimerization was higher than the average reaction rate of IA dimerization. This trend was also noticed when the reaction was carried out in the microreactor. A higher selectivity for DIA formation was noticed using 70 wt % acid catalyst, compared to the 90 wt % catalyst concentrations, which showed higher selectivities for DAMS and CP isomer formation. This selectivity is important because the major byproducts, DIA and DAMS, are used by different industries, and so with different catalyst concentrations the reaction can be tailored to increase the yield of the desired byproduct. 3.4. Comparison of Performance of the Microreactor to that of the Semibatch Reactor. 3.4.1. Comparison of ARR. Losey et al.,47 showed that the gas-liquid mass transfer rates in a microfabricated packed bed reactor were increased by more than a 100-fold when compared to the values reported for a traditional multiphase packed-bed reactor. Tadepalli et al.17 also showed that the microreactor mass transfer parameter values were 2 orders of magnitude higher than those obtained in the semibatch reactor for the catalytic hydrogenation of o-nitroanisole. With such high mass transfer rates, the use of microreactors translates to improved overall reaction rates and reactor performance for multiphase reactions limited by mass transfer effects, as is the case with the cycloaddition reaction in this work. A comparison of the average reaction rates obtained in the semibatch reactor and the microreactor at a similar reaction temperature and catalyst concentration is shown in Figure 12a. The comparison is made between the productivity using a microreactor for the cycloaddition reaction with a total reactant flow rate of 1.5 mL/min and catalyst flow rate of 1.5 mL/min; and a semibatch reactor, both at a reaction temperature of 35 °C. Figure 12a shows that the ARR values obtained in the microreactor were greater than those in the semibatch reactor by 3 orders of magnitude. Since a microreactor is a continuous reactor, increasing the residence time while maintaining a constant flow velocity through the microreactor requires that the reactor length be increased to increase the residence time, as is explained in section 3.2.2. This increase in the microreactor length translates to an increase in the amount of the catalyst in the microreactor, which is the reason why the ARR decreases with an increase in the microreactor residence time in Figure 12a. This is unlike the case with the semibatch reactor, where the amount of catalyst is constant throughout the reaction time.

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Figure 12. (a) Comparison of ARR between the microreactor and semibatch reactor using 90 wt % acid at 35 °C; (b) comparison of maximum yield values between the microreactor and semibatch reactor using 90 wt % acid at 35 °C; (c) comparison of maximum yield values between the microreactor and semibatch reactor using 70 wt % acid at 35 °C; (d) comparison of reactant conversions between the microreactor and semibatch reactor using 90 wt % acid at 35 °C.

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The reason for this huge increase in ARR in the microreactor compared to the semibatch reactor is the mass transfer enhancement that is obtained via process intensification in the microreactor.17 Halder and Lawal15 also reported that much higher ARR values were obtained in the microreactor compared to the batch reactor in the liquid phase nitration of toluene. 3.4.2. Comparison of Yields. A comparison of the CP yield at the same conditions mentioned in section 3.4.1 is shown in Figure 12b, while Figure 12c shows a comparison using a catalyst concentration of 70 wt % with a total flow rate of 0.5 mL/min through the microreactor. These comparisons show that the optimal catalyst concentration in the semibatch reactor was observed to be 70 wt % rather than the 90 wt % observed in the microreactor. In the semibatch reactor, increasing the viscosity of either or both phases hinders both the dispersing and coalescing steps that occur during reactor operation. Such hindrances directly affect the interfacial area between the aqueous and organic phases, thereby leading to lower interfacial areas at high viscosities and vice versa.10 The higher viscosity of the 90 wt % catalyst compared to the 70 wt % acid catalyst also increased the energy requirement for obtaining dispersions in the semibatch reactor, thereby leading to a reduction in the mass transfer between the two phases when additional energy is not proportionally added to the semibatch reactor. Apart from the effect of viscosity, the relatively high density differences between the two phases created a driving force for the coalescing of the two phases and hence a reduction in interfacial area,10 which led to low product yields for the 90 wt % acid catalyst in the semibatch reactor. However the microreactor was more easily able to overcome such mass transfer limitations caused by the viscous nature of the catalyst because of the enhanced mass transfer obtained in the microreactor, and this led to optimal product yields using 90 wt % acid catalyst in the microreactor. Shah et al.38 observed that AMS had a much greater tendency toward dimerization rather than cycloaddition at the initial stages of the reaction between IA and AMS using Amberlyst 15 as a catalyst in the batch reactor, and that the selectivity toward the formation of DAMS peaked during the course of the reaction in the batch mode. They showed that the adsorption of AMS was greater than the adsorption of DAMS. Thus at the early stages of the reaction, when the concentration of AMS was high, AMS occupied the active site and inhibited the adsorption of DAMS. As the concentration of unreacted AMS reduced, the inhibition effect reduced thereby allowing the adsorption of DAMS to the active sites and the deoligomerization of the DAMS to produce AMS which could react with IA to give CP. However, the very small residence times obtained in the microreactor because of its small linear dimensions means that the dimerization of AMS, rather than the cycloaddition reaction, becomes the favored reaction, leading to the low CP product yields obtained using the 70 wt % acid catalyst in the microreactor. The low conversion because of insufficient residence time in the microreactor is also observed in the increasing yields with an increase in residence time shown in Figure 12c. Increasing the catalyst concentration in the microreactor directly increases the amount of catalyst available for the reaction, thereby reducing the residence time requirement, which is shown by the increased product yields at high catalyst concentrations. However, at optimized reaction conditions, the highest yield obtained in the microreactor (62%) was greater than the highest yield obtained in the semibatch reactor (41%) at the reaction conditions used in the work. Also the highest yields in the microreactor were obtained immediately the microreactor at-

tained steady state, which occurred within a matter of seconds. However a survey of all the semibatch reactor experiments shows that the yields in the semibatch reactor continued increasing during the experiment until the highest yields were obtained well toward the end of the total experimental run time of 4 h. This quick attainment of steady-state in the microreactor for this particular process is one of the major advantages of continuous processing, apart from the reduction in the total run time and number of startup and shutdown times, and reduced equipment size. While there is a possibility of increased yields in the semibatch reactor using even higher stirrer speeds than used in this work, it should be noted that such high stirrer speeds are difficult to replicate in conventional sized (100-10000 gallons) semibatch reactors; apart from the fact that the mass transfer efficiency of semibatch reactors generally decreases with an increase in reactor size. The mass transfer efficiency of a typical conventional size semibatch reactor will expectedly be much lower than that of the 22 mL compact semibatch reactor used in this work. 3.4.3. Comparison of Reactant Conversions. A comparison of the reactant conversions in the microreactor and semibatch reactor, at the conditions stated in section 3.4.1, is shown in Figure 12d. From the figure, the reactant conversions are greater than 95% in both reactors. Since the catalyst used in both reactors is the same, it is likely that the high selectivity toward the dimerization of the starting reactants is similar in both reactors since there is only an enhancement of heat and mass transfer in the microreactor, and not a change in the chemistry of the dimerization reactions.27 However, a difference in the reactant conversions obtained in the two reactors is noticed at lower catalyst concentrations such as 70 wt %, where the low residence time available in the microreactor is insufficient for the complete conversion of the reactants as was explained in the previous section. Increasing the length of the microreactor should theoretically lead to complete reactant conversions; however, huge pressure drops during reactor operation make this impracticable. 4. Conclusions The use of high stirrer speeds in the semibatch reactor is energy inefficient and also difficult to sustain for conventionalsized industrial semibatch reactors, thus making it difficult to obtain good mass transfer even after improvements such as improved impeller design, efficient impeller positioning, and the addition of baffles are implemented. The exothermic cycloaddition reaction between AMS and IA to obtain CP, involving the two immiscible phases, aqueous catalyst and organic reactants phases, was successfully carried out in the microreactor. The microreactor offers enhanced mass transfer for this mass transfer limited reaction compared to the mass transfer obtained in the semibatch reactor even at very high stirrer speeds. This enhancement of mass transfer in microreactors is evident from the ARR values obtained in the microreactor for the cycloaddition reaction, which were three to four times orders of magnitude higher than those obtained in the semibatch reactor because of the process intensification enabled by microreaction technology. An optimum catalyst concentration of 70 wt % was observed in the semibatch reactor rather than the 90 wt % observed in the microreactor because of the difference in the typical residence time obtained in the two reactors, and the density and viscosity differences between the 70 and 90 wt % catalysts. Sufficient residence time in the semibatch reactor allowed for complete reactant conversions leading to increased CP yields at a catalyst concentration of 70


wt %. The less viscous 70 wt % catalyst also allowed for increased mass transfer between the organic and catalyst phase. However in the microreactor, the much smaller residence times led to incomplete reactant conversions and a higher selectivity to reactant dimerization because there was insufficient time for the deoligomerization or cracking of unsaturated DAMS. Increasing the catalyst concentration to 90 wt % ensured for increased catalytic activity leading to complete reactant conversions, possible deoligomerization of DAMS, and thus increased CP yields. The increased density and viscosity of the 90 wt % was inconsequential to the operation of the microreactor because of its enhanced mass transfer. Many of the suggestions in the literature for improving the yields of the cycloaddition reaction in a semibatch reactor such as the addition of acid-soluble oils and surfactants as hydride transfer agents,10 and the use of other additives,5,6 can most likely also be successfully applied to the microreactor to obtain even greater product yields. Also with the use of the microreactor numbering-up approach rather than the scale-up approach, a lot of time and cost savings can be obtained in the transfer of any improvements in the cycloaddition reaction from the research to the production stage. These savings will be made by bypassing the very difficult reaction scale-up problems that usually arise in the scale-up of such complex liquid-liquid reactions because of the complexities of drop formation and coalescence that change with scale.11 Acknowledgment The financial and technical support from International Flavors and Fragrances Inc., and Graduate Tuition support from the Stevens Institute of Technology are gratefully acknowledged. The authors thank all the members of NJCMCS at Stevens who provided valuable help at different stages of this project. Literature Cited (1) Rowe, D. J. Chemistry and Technology of FlaVors and Fragrances; Blackwell: Oxford, U.K., 2005; p 144-155. (2) Heeringa, L. G.; Beets, M. G. J. Tricyclic isochromans and processes for making same. U.S. Patent 3,360,530, December 26, 1967. (3) Ferber, G. J.; Goddard, P. J. Production of indanes. U.K. Patent 4,440,966, April 3, 1984. (4) Barbier, H. Process for preparation of butylcymene. U.S. Patent 1,951,123, March 13, 1934. (5) Cobb, R. L. Acid catalyzed reactions of monovinyl aromatic compounds. U.S. Patent 4,596,896, June 24, 1986. (6) Cobb, R. L. Acid-sulfolane catalyzed production of cyclic alkylated compounds. U.S. Patent 5,034,562, June 23, 1991. (7) Henkes, E.; Halbritter, K.; Striepe, W.; Pock, R. Preparation of indanes. Ger. Patent 4,740,646, April 26 1988. (8) Wiegers, W. J.; Sprecker, M. A.; Vock, M. H.; Schmitt, F. L. Flavoring with indane alkanols and tricyclic isochromans. U.S. Patent 4,315,951, February 16, 1982. (9) Ghosh, B.; Chaudhuri, B. Dimerization of R-methylstyrene (AMS): Kinetic study of the liquid-liquid process. AIChE J. 2006, 52 (5), 1847– 1854. (10) Ende, D. J.; Eckert, R. E.; Albright, L. F. Interfacial area of dispersions of sulfuric acid and hydrocarbons. Ind. Eng. Chem. Res. 1995, 34 (12), 4343–4350. (11) Paul, E. L.; Atiemo-Obeng, V. A.; Kresta, S. M. Handbook of Industrial Mixing: Science and Practice; Wiley-Interscience: Hoboken, NJ, 2004; p 381, 730. (12) Besser, R. S.; Ouyang, X.; Surangalikar, H. Hydrocarbon hydrogenation and dehydrogenation reactions in microfabricated catalytic reactors. Chem. Eng. Sci. 2003, 58 (1), 19–26. (13) Hessel, V.; Knobloch, C.; Lowe, H. Review on patents in microreactor and microprocess engineering. Recent Pat. Chem. Eng. 2008, 1, 1–16. (14) Okafor, O. C.; Tadepalli, S.; Tampy, G.; Lawal, A. Cycloaddition of isoamylene and R-methylstyrene in a microreactor using Filtrol-24

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catalyst: Microreactor performance study and comparison with semi-batch reactor performance. Int. J. Chem. React. Eng. 2010, 8, A71. (15) Halder, R.; Lawal, A.; Damavarapu, R. Nitration of toluene in a microreactor. Catal. Today 2007, 125 (1-2), 74–80. (16) Tadepalli, S.; Lawal, A. The catalytic hydrogenation of aromatic nitro ketone in a microreactor: Reactor performance and kinetic studies. Int. J. Chem. React. Eng. 2008, 6, 1–37. (17) Tadepalli, S.; Qian, D.; Lawal, A. Comparison of performance of microreactor and semi-batch reactor for catalytic hydrogenation of onitroanisole. Catal. Today 2007, 125 (1-2), 64–73. (18) Voloshin, Y.; Halder, R.; Lawal, A. Kinetics of hydrogen peroxide synthesis by direct combination of H2 and O2 in a microreactor. Catal. Today 2007, 125 (1-2), 40–47. (19) Roberge, D. M.; Ducry, L.; Bieler, N.; Cretton, P.; Zimmermann, B. Microreactor technology: A revolution for the fine chemical and pharmaceutical industries? Chem. Eng. Technol. 2005, 28 (3), 318–323. (20) Burns, J. R.; Ramshaw, C. Development of a microreactor for chemical production. Chem. Eng. Res. Des. 1999, 77 (3), 206–211. (21) Chan, T. Y.; Priestman, G. H.; MacInnes, J. M.; Allen, R. W. K. Development of a micro-channel contactor-separator for immiscible liquids. Chem. Eng. Res. Des. 2008, 86 (1A), 65–74. (22) Hessel, V.; Hardt, S.; Lowe, H. Chemical Micro Process EngineeringsFundamentals, Modeling and Reactions; Wiley-VCH: Weinheim, FRG, 2004; p 379-567. (23) Robins, I.; Shaw, J.; Miller, J.; Turner, C.; Harper, M. In Solute transfer by liquidsliquid exchange without mixing in micro-contactor deVice, 1st Int. Conf. on Microreaction Technology, 1998; 1998; pp3546. (24) Schwesinger, N.; Frank, T.; Wurmus, H. A modular microfluid system with an integrated micromixer. J. Micromech. Microeng. 1996, 6 (1), 99–102. (25) Seong, G. H.; Crooks, R. M. Efficient mixing and reactions within microfluidic channels using microbead-supported catalysts. J. Am. Chem. Soc. 2002, 124 (45), 13360–13361. (26) Adeosun, J. T.; Lawal, A. Mass transfer enhancement in microchannel reactors by reorientation of fluid interfaces and stretching. Sens. Actuators, B 2005, 110 (1), 101–111. (27) Ehrfeld, W.; Hessel, V.; Lowe, H. Microreactors, 1st ed.; WileyVCH: Weinheim, FRG, 2001; p 1-12. (28) Harper, M. J. Method and apparatus for diffusive transfer between immiscible liquids. U.S. Patent 6,306,305, October 23, 2001. (29) Burns, J. R.; Ramshaw, C. A microreactor for the nitration of benzene and toluene. Chem. Eng. Commun. 2002, 189 (12), 1611–1628. (30) Cherlo, S. K. R.; Kariveti, S.; Pushpavanam, S. Experimental and numerical investigations of two-phase (liquid-liquid) flow behaviour in rectangular microchannels. Ind. Eng. Chem. Res. 2010, 49, 893–899. (31) Zhao, Y.; Guangwen, C.; Quan, Y. Liquid-liquid two-phase flow patterns in a rectangular microchannel. AIChE J. 2006, 52, 4052–4060. (32) Kashid, M. N.; Agar, D. W. Hydrodynamics of liquid-liquid slug flow capillary microreactor: Flow regimes, slug size and pressure drop. Chem. Eng. J. 2007, 131 (1-3), 1–13. (33) Stroock, A. D.; Dertinger, S. K. W.; Ajdari, A.; Mezic, I.; Stone, H. A.; Whitesides, G. M. Chaotic mixer for microchannels. Science 2002, 295, 647–651. (34) Mengeaud, V.; Josserand, J.; Girault, H. H. Mixing processes in a zigzag microchannel: finite element simulations and optical study. Anal. Chem. 2002, 74 (16), 4279–4286. (35) Yamaguchi, Y.; Ogino, K.; Yamashita, K.; Maeda, H. Rapid micromixing based on multilayer laminar flows. J. Chem. Eng. Jpn. 2004, 37 (10), 1265–1270. (36) Grob, R. L. Modern Practice of Gas Chromatography, 3rd ed.; Wiley-Interscience: New York, 1995; p 438-439. (37) Felder, R. M.; Rousseau, R. W. Elementary Principles of Chemical Processes, 3rd ed.; John Wiley & Sons: Hoboken, NJ, 2005; p 135. (38) Shah, N. F.; Bhagawat, M. S.; Sharma, M. M. Cross-dimerization of R-methylstyrene with isoamylene and aldol condensation of cyclohexanone using a cation-exchange resin and acid-treated clay catalysts. React. Polym. 1994, 22 (1), 19–34. (39) Cruz, V. J.; Izquierdo, J. F.; Cunill, F.; Tejero, J.; Iborra, M.; Fite, C. Acid ion-exchange resins catalysts for the liquid-phase dimerization/ etherification of isoamylenes in methanol or ethanol presence. React. Funct. Polym. 2005, 65 (1-2), 149–160. (40) Shah, N. F.; Sharma, M. M. Dimerization of isoamylene: Ion exchange resin and acid-treated clay as catalysts. React. Polym. 1993, 19 (3), 181–190. (41) Chaudhuri, B.; Sharma, M. M. Some novel aspects of the dimerization of R-methylstyrene with acidic ion-exchange resins, clays, and other acidic materials as catalysts. Ind. Eng. Chem. Res. 1989, 28 (12), 1757–1763.

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(42) Cohen, N.; Benson, S. W. Estimation of heats of formation of organic compounds by additivity methods. Chem. ReV. 1993, 93 (7), 2419– 2438. (43) Hessel, V.; Lo¨we, H.; Scho¨nfeld, F. MicromixerssA review on passive and active mixing principles. Chem. Eng. Sci. 2005, 60 (8-9), 2479–2501. (44) Fogler, H. S. Elements of Chemical Reaction Engineering, 4th ed.; Prentice Hall: Upper Saddle River, NJ, 2006; p 37-69, 816, 839. (45) Chaudhuri, B.; Sharma, M. M. Alkylation of phenol with R-methylstyrene, propylene, butenes, isoamylene, 1-octene, and diisobutylene: Heterogeneous vs homogeneous catalysts. Ind. Eng. Chem. Res. 1991, 30 (1), 227–231.

(46) Fasullo, O. T. Sulfuric Acid: Use and Handling; McGraw-Hill Book Company: New York, 1965; p 305. (47) Losey, M. W.; Schmnidt, M. A.; Jensen, K. F. Microfabricated multiphase packed-bed reactors: Characterization of mass transfer and reactions. Ind. Eng. Chem. Res. 2001, 40 (12), 2555–2562.

ReceiVed for reView November 12, 2009 ReVised manuscript receiVed April 13, 2010 Accepted April 23, 2010 IE901794P

Microreactor Performance Studies of the Cycloaddition of Isoamylene

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