Chemical Reaction Engineering - Industrial & Engineering Chemistry

Related Content: Modern Production Methods Based on 1,3-Butadiene and 1-Butene. Industrial & Engineering Chemistry. Passmann. 1970 62 (5), pp 48–51...
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Chemical Reaction New plants present challenge for industrial-process oriented reaction engineering

his review covers recent reaction engineering literature (July T 1968 to July 1969) ranging from direct industrial applications to theoretical papers which, in the reviewer's opinion, will have a significant impact in the near future on reactor design or operation. It is encouraging to see the increasing number of papers dealing directly with reaction kinetics or transport processes taking place in commercial reactors. Among those represented are catalytic cracking and hydrocracking of gas oils, the reforming of naphthas, the production of ammonia, and the water-gas shift reaction. Almost all of the published industrial experience relates to existing processes, and very little experience has been published on applying reaction kinetic principles to the design of new plants. While applications of reaction engineering are apparently paying their way in the improvement of existing plants, the biggest potential payout comes when these same principles can be applied to new plant design. The design of a new plant, of course, is much more risky than improving a n existing one, and apparently a credibility gap still exists between plant designers and practitioners of reaction engineering. An increasing number of papers dealing with applications of zeolite catalysts shows them rapidly becoming one of the most important classes of industrial catalysts. From their initial use as the work horse of catalytic cracking, they are finding applications in hydrocracking and reforming as well as petrochemical uses. Clearly much work remains to be done also in understanding the complex transport processes taking place in these highly structured catalysts. Even though fluid beds have been employed industrially for over 25 years, the study of their fluid dynamic as well as heat and mass transfer characteristics remains a very active area. Seventeen papers in this area alone were reviewed. While multiphase flow packed reactors find widespread use in industry, little attention has been devoted to them in terms of reaction kinetic studies. The present year was no exception; however, one of the first papers experimentally measuring mass transfer resistances in two-phase gas-liquid flow over catalyst undergoing reaction has been reported. More hazards have been exposed in the indiscriminate fitting of kinetic rate constants from experimental data. I n one rather frightening example, 21 separate rate equations were shown to fit the same data a t the 99% confidence level. Fortunately, while more hazards are being exposed, techniques are also being developed to improve our ability to abstract meaningful rate constants from data. The increased industrial output is encouraging; however, the 52


literature still is largely composed of theoretical techniques and simulations of hypothetical reactors. More industrial feedback on the usefiilness of the techniques and the validity of the simulations is badly needed. The rapid identification of key phenomena influencing 'reactor behavior is another area requiring more attention. I n most academic work, one seeks to isolate a particular phenomenon for study, but in describing an industrial reactor the crucial problem is cutting through the jungle of existing phenomena to those few that critically influence the performance. Behavior and Performance of Fixed Beds

A theory to account for voidage fluctuations near the wall of packed bed reactors has been proposed by Ridgway and Tarbuck ( 4 A ) . The oscillatory variation in voidage near the containing wall has been described by means of a semitheoretical expression. The agreement with experimental data is quite good and represents the only available method for design calculations. The extreme parametric sensitivity of highly exothermic reactor systems has been demonstrated by Carberry and White ( I A ) in a digital simulation of naphthalene oxidation in a packed bed of V205 catalyst. The yield of phthalic anhydride was shown to be highly sensitive to the value of the thermal radial peclet number while the yield was virtually insensitive to the mass radial peclet number. The computations revealed significant interparticle radial temperature gradients as well as interphase concentration and temperature differences and severe intraphase mass diffusional restrictions. With this highly exothermic system, significant gradients would persist even a t the low conversion levels used in laboratory differential rate experiments. The paper highlights the value of simulating chemical reactors, particularly with respect to determining parametric sensitivities and potential experimental disguises. In a n experimental tour de force, Schertz and Bischoff ( 5 A ) have made point measurements of concentration, temperature, and velocity at various radial and axial positions in 4-inch packed column. Velocity and temperature measurements were made simultaneously with a specially designed probe. The data confirm early work showing a sharp increase in velocity near the wall of the packed column. The velocity profiles us. radius were much steeper under non-isothermal conditions compared to isothermal conditions in the packed bed. Correlations are presented for the point values of velocity, effective thermal conductivity, and the radial mass dispersion coefficient. The thermal conductivity results were well correlated by the Yagi-Kunii relationship. Littman, Barile, and Pulsifer ( 3 A ) have filled in an existing gap

tngineering for gas-particle heat transfer data at low Reynolds numbers. Using a frequency response technique, they report experimental data on the gas-particle heat transfer coefficients in packed beds at Reynolds numbers between 2 and 100. The model accounts for axial dispersion of heat in the gas phase and axial conduction in the solid phase. An excellent review of the available literature on the structure of packed beds is presented by Haughey and Beveridge ( Z A ) . The review includes modes and properties of both regular and random packings for spherical and nonspherical particles. A total of 239 references are given. Diffusion and Reaction in Zeolite Catalysts

Eberly ( 3 B ) has studied some of ?he diffusion properties of zeolites by means of gas chromatographic techniques using argonne, krypton, and helium in various mordenites and faujasites, as well as amorphous silica-alumina catalyst. The internal transport was best described as an activated diffusional process with activation energies increasing with molecular weight and ranging from 2.5 to 15 kcal/mol. The effect of molecular weight is much greater than predicted by Knudsen diffusion and the inert gases gave diffusion coefficients several orders of magnitude higher than those measured for the C r C 4 hydrocarbons studied by Satterfield and Frabetti (8B). Clearly much has yet to be learned in the flow oflarge organic molecules in zeolite catalysts. In another paper, Eberly ( 4 B ) has studied the adsorption behavior of c 5 - C ~normal paraffins on erionites and 5A molecular sieves. The rate of adsorption of erionites is low and decreases with molecular weight of the paraffins, whereas no such drop-off is noted for the 5A sieve. While Ficks second law was used to describe the rate behavior, he cautioned that it cannot be expected to describe adsorptive diffusion completely since it does not account for the observed nonlinear aspects of adsorption. The adsorption of a binary gas mixture on 5A and 1OX molecular sieves has been reported by Danner and Wenzel ( 2 B ) . The gases employed were carbon monoxide-nitrogen, carbon monoxide-oxygen, and oxygen-nitrogen mixtures. These zeolites gave a surface selectivity which is independent of any sieving effect based on the size of the adsorbed molecules. Experimental adsorption phase diagrams are given for all the gas mixtures at -200'F and 1 atm. While they were not able to arrive at a specific adsorption mechanism, they did find that van der Waals' forces alone are not adequate to interpret the experimental data. Hydrocracking and diffusion in mordenite catalysts have been studied by Beecher, Voorhies, and Eberly ( I B ) . Hydrocracking reactions were carried on with n-decane and decalin with independent diffusional measurements made with inert gases. Hydrocracking experiments were carried out at 450 psig and from 450610'F. Acid leaching of the mordenite gave an aluminum-deficient structure which gave a fourfold increase in catalyst activity for the hydrocracking reaction. The diffusivities, as measured by the inert gases, indicated that the hydrocracking reaction should not be diffusion limited; however, the actual experimental data indicated otherwise. This apparent contradiction was explained by the hypothesis that the flow of hydrocarbons in the mordenite is governed by strong forces of adsorption. On cracking mixtures of n-decane and decalin there was virtually no hydrocracking of n-decane and this suggested that the decalin almost completely denied the n-decane access to the active sites. In screening a series of zeolite hydroisomerization catalysts, Voorhies and Bryant (IOB) found that palladium-impregnated


hydrogen mordenite gave the highest activity. Varying the superficial velocity over a fourfold range, they found no mass transfer limitation between the fluid and the particles. By crushing the catalyst over a tenfold range, they were able to show experimentally that intraparticle diffusion played no significant role. Diffusion within the mordenite crystallite cells, however, could still be significant. A mathematical model is given based on a dual-site mechanism and Langmuir adsorption isotherms. The cracking activities of synthetic faujasite, mordenites, and L zeolite have been studied by Hopkins ( 5 B )for n-hexane, n-heptane, and ethylbenzene. The catalytic cracking of gas oil fractions over rare earth exchanged faujasite zeolite catalysts have been reported by Thomas and Barmby ( S B ) . Based on experimental observations, their proposal is that the gas oil initially cracks on the outside of the catalyst since the gas oil molecules are too large to enter the cavities. Gasoline molecules produced are postulated as now entering the cavities where secondary hydrogen transfer reactions take place. By contrast, amorphous silica-alumina catalyst provide no such separation of sites for primary and secondary reactions. The hypothesis is largely based on conventional calculation of diffusion rates as opposed to direct experimental measurement of diffusion in zeolites. Using a novel microreactor technique, Nace ( 6 B )has studied the product distribution and instantaneous rates of cracking for hydrocarbon reactants over zeolite and silica-alumina catalysts. Using a pulse technique, he obtained reaction rate data at only 0.45 sec catalyst on-stream time. The R E H X zeolite catalyst decreased in activity by a factor of 10 during the first 10 seconds on-stream. A greatly reduced rate of aging was observed u p to 120 seconds. Both silica-alumina catalysts and R E H X catalyst gave the same value of n in the tCn decay law. I n a companion paper, Nace ( 7 B ) has studied the effect of cracking paraffins, olefins, alkyl aromatics, and cycloparaffins over both X-type zeolites as well as amorphous silica-alumina catalysts. Olefins cracked much faster than the corresponding paraffins. Carbonium ion mechanisms are suggested for the various cracking steps. Industrial Process Kinetics

In one of the few kinetic studies from academe aimed a t a complex industrial process, Qader and Hill (9C) studied the hydrocracking of gas oils over a nickel tungsten sulfide catalyst on silicaalumina. They show the conversion of the gas oil to gasoline to be first order with respect to the fractions remaining unconverted. Desulfurization and dehydrogenation reactions were also shown to be first order with activation energies of about 17 kcal/mol. The cracking reaction to gasoline exhibited an activation energy of 21 kcal/mol. They find that hydrogenation reactions d o not apparently control the overall reaction rate and, surprisingly, total pressure had little effect on the overall reaction rate. Conversion level had a strong effect on gasoline composition. I n conflict with the above paper, Carr and Stahfeld (IC)report the overall: conversion reaction rate to be first order in hydrogen but zero order in the fraction of gas oil remaining unconverted. An overall activation energy of 26 kcal/g-mol was observed. By including the energy balance and accounting for quench gas and vaporization, an overall model was constructed. By employing fitted transfer functions, the steady state model was used to develop a conversion control scheme which was implemented successfully on a commercial hydrocracker. The regeneration of coke from catalysts in fixed beds has been VOL. 6 2

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studied theoretically by Ozawa (7C) where there are no mass transfer limitations. H e shows that when regeneration takes place completely in the intrinsic burning controlled region, minimums may exist in the coke profile with distance in the bed. This minimum in the coke profile results from the dynamic interaction of the temperature and oxygen concentration profiles. A semianalytical solution is provided which allows rapid simulation of the quasi-steady-state regime. A mathematical model of the radiant section of a steam methane reformer is given by Hyman (5C). The model includes descriptions of the heat transfer from the tube wall to the flowing gas as well as for the chemical change within a typical tube. The steam methane reforming reaction is represented by four reversible reaction steps. Using literature references for the rate expressions, the necessary heat and mass balances were constructed and the rate constants were based on a standard nickel catalyst. T h e model was compared with four sets of commercial data and gave reasonably good .predictions of methane conversion and outlet temperatures. The observed conversions and temperatures were, unfortunately, over relatively narrow range. An improved method of operating gasoline reforming units has been presented by Pfefferle et al. (8C) based on a kinetic model of the system. I t is based on five primary reaction steps and is apparently similar to a n earlier published model. No details of any modifications to the model were given. T h e principal differences in the new method compared to conventional reforming are the low temperature, low recycle, and higher space velocity in the first reactors with the higher recycle rate in the final reactors. I n addition, incentives were found for minimizing catalyst loading in the first reactors. Details of the improved yields are presented. New pilot plant kinetic rate data have been presented on the widely used carbon monoxide shift conversion reaction by Ting and Wan (72C). A generalized design equation is developed based on a reaction rate expression. A sample calculation is given for determining the optimum catalyst volume. Data on the effects of sulfur and pressure are also included. The kinetics of the thermal cracking of high molecular weight paraffins have been studied by digital simulation techniques by Woinsky (74C). H e considers a series of 28 separate reaction steps and obtains approximate frequency vectors and activation energies from a wide range of literature data. Using 1.3 order kinetics he was able to obtain a reasonable fit to the available experimental data. The author rightfully urges caution in the model’s use because of the uncertainty of some of the rate constants introduced. T h e model is claimed to be valid in the region of 750-900’K and from 0.5-70 atmospheres. Long-term cyclic processes many times bring unpleasant surprises to the reactor designer. The subtle interplay between diffusional restrictions and catalyst attrition in the catalytic cracking process has been presented by Weisz (73C). Part of the burning of carbon from catalyst particles in the therrnofor catalytic cracking process (TCC) takes place in the strongly diffusion limited shell progressive mode. The presence of the unburned carbon cores causes strains to be set u p in the catalyst particles, which on repeated exposure to both the reactor and regenerator conditions, resulted in catalyst breakage. The introduction of a less diffusionally restricted catalyst made it possible to significantly reduce the amount of catalyst breakage by eliminating the unburned carbon cores. Dyson and Simon (3C) present a new kinetic rate expression for ammonia synthesis on typical industrial catalysts which includes a diffusional term. The rate expression proposed is valid from 150 to 300 atm and is claimed to be suitable for design, optimization, and control of modern ammonia plants. The rate expression is based on the earlier work of Temkin et al. A polynomial expression in terms of nitrogen conversion and temperature is given which adequately accounts for diffusion in the 6- to 10-millimeter particle range used commercially. A method for calculating the heat of catalytic cracking has been proposed by Orochko and Chernakova (6C) based on assuming that the charge and cracked products can be represented by a series of pure hydrocarbons. Using the known heats of formation of these pure hydrocarbons, it was possible to obtain an estimate of the heat of cracking. Results compared favorably with some experimental data. Unfortunately, no details were given of the type of experimental procedures employed. No references were given to the extensive work in western literature on heats of cracking. 54


Some useful results on the particle size distribution and attrition of cracking catalysts have been given by Gwyn (4C). H e shows that the rate of attrition of the catalyst is a function only of its initial. diameter and time. The decrease in attrition rate as the catalyst ages can be expressed solely as a function of its age. The proposed attrition model compares well with available commercial catalyst results. The paper is based solely on silicaalumina catalyst; and, unfortunately, the author gives no clue if the same analysis holds for the ubiquitous zeolite catalyst. Simulation and optimization of an ethylbenzene dehydrogenation reactor have been reported by Sheel and Crawe ( 7 7C). An adiabatic reactor model is developed for a six-step mechanism and compared with a small amount of data from a commercial reactor. An optimization was conducted based on six process variables using Rosenbrock’s multivariable search technique. Published kinetic data for the commercial water-gas shift catalyst have been interpreted by Ruthven (7OC) in terms of pore diffusional effects. By accounting for the catalyst diffusivity, intrinsic rate constants were calculated which were similar for all catalysts employed. The calculated intrinsic activities also gave activation energies of 22-27 kcal/g-mol which are similar to experimental data on small catalyst particles. Data from a commercial plant have been used by Chartrand and Crowe (2C) in optimizing the catalytic oxidation of sulfur dioxide. A model was proposed for the multibed adiabatic reactor and the optimization of a number of process variables was achieved. Apparently, the plant operators have done their job well since the optimization resulted in a n improvement in SO2 conversion of from 97.59 to 97.84%. Mass TFanSfer in Gas-Solid gluidized Beds

I n an important contribution to understanding gas flow in fluidized beds, Kunii and Levenspiel ( 6 D ) have proposed a “bubbling bed” model. Their model envisions uniformly sized bubbles surrounded by gas clouds and followed by wakes rising through an emulsion of downward moving solids. There is continuous transport of gas between emulsion, cloud, wake, and bubble regions. Surprisingly, the model contains only one adjustable parameter-the effective bubble size. They show the model is capable of predicting with good accuracy the exchange coefficients developed in previous experimental work. The authors indicate the model is most satisfactory in beds where the height is no greater than the diameter. I n a related study, Kunii and Levenspiel (70) have compared a number of experimental studies with their bubbling bed model of fluidized beds. Their model successfully predicts a wide range of experimental data in fluidized beds with only one adjustable parameter-the effective bubble size. Further comfort for the theory is provided by the bubble sizes being physically quite reasonable. Comparisons are made with gas-solid mass transfer experiments, gas-solid heat transfer experiments, as well as reactions taking place in fluidized beds. The model did a particularly excellent job in predicting the results of the catalytic decomposition of ozone. In this case, experimental data gave conversions that were poorer than back-mixed flow, which the model successfully predicted. Using photographic techniques, Carlos and Richardson (20, 30) have measured individual particle velocities in liquid fluidized beds. The bed of uniformly sized glass spheres was fluidized by dimethyl phthalate which has a refractive index close to that of glass. Using similar techniques, they also measured axial mixing coefficients in the same experimental fluidized beds. The mixing coefficients measured were about ten times greater than those predicted by previous investigators. The movement of particles in agglomerates was thought to be the reason for this difference. Yoshida, Kunii, and Levenspiel (720) present new experimental data on the axial dispersion of gas in bubbling fluidized beds. The experimental measurements were performed in a 20-cm I D fluidized bed with both adsorptive and nonadsorptive gases. The actual dispersion is strongly sensitive to the extent of adsorption of gas on the solids. Furthermore, the equilibrium between the adsorbing gas and the solid was not achieved under the conditions of fluidization investigated. This was attribuied to the very short contact time between the emulsion, solids, and the gas passing through as fast rising bubbles. An experimental study of nucleate boiling heat transfer from stainless steel balls fluidized in a water system has been reported by Young and Holman (730). A radio frequency induction heater was used to heat the balls while they were in a fluidized

state. An unusually early on-set of nucleate boiling was observed in the fluidized system and was attributed to cavitation. Correlations are presented which represent the experimental data. Heat transfer rates from a heated surface to a vibrating fluidized bed have been presented by Bukareva, Chlenov, and Mikhailov (7D). A continuous increase in the heat transfer coefficient was noted with the increase of the vibration frequency with n o maximum indicated. Experimental data and techniques are given in the article. Heat transfer between fluidized beds and bundles of horizontal tubes have been studied by Gel’perin, Ainshtein, and Korotyanskaya ( 5 D ) . Experimental data on the effects of vertical and horizontal spacing as well as velocity are presented. An empirical correlation is given which relates the important design parameters to the Nusselt number. A review of experimental studies on the influence of turbulence on particulate heat and mass transfer is given by Clamen and Gauvin ( 4 0 ) . Seventy-three literature references are given along with analysis of various correlations. Interphase mass transfer coefficients in bubbling fluidized beds have been measured experimentally by Oigenblik et a l . ( B D ) . Tracer experiments with helium and COZ were conducted in fluidized beds of silica gel with diameters of 116 and 416 mm. Based on experimental results, interphase transfer coefficients were calculated for a variety of commercial reactors including oxidation of methylene, chlorination of methane, oxidation of S O ~ S O S , decomposition of nitric oxide, hydrogenation of ethylene, and decomposition of ozone. T h e interphase transfer coefficient values varied from 0.9 to 0.14 sec-1 for these various reactors. Yet another model has been proposed for gas transfer between bubble and continuous phase in the gas-solid fluidized bed by Toei et a l . ( 7 7 0 ) . They present experimental results derived from high-speed photography of NO2 gas in fine glass beads. T h e behavior of the bubbles was generally similar to that previously observed by Rowe, Partridge, and Lyall (QD).Gas transfer coefficients were based on experimental results using COZas the tracer gas. The proposed model was claimed to fit experimental results satisfactorily. Bed to wall radiation heat transfer measurements were made by Szekely and Fisher (7OD) in a gas-solid fluidized bed. A model was proposed which allows for transient heat transfer from the source to the particles during their brief exposure to the wall. The heat transfer rates predicted by the model were in good agreement with experimental results. Hydrodynamics of Fluidized Beds

T h e theory of contact time distributions in gas-fluidized beds has been discussed by Nauman and Collinge ( 4 E ) . They show that the contact time distribution is the exact heterogeneous counterpart of the ordinary residence time distribution. With a known contact time distribution, conversion of a first order reaction system can be predicted exactly. Bounds on the conversion can be calculated for other reaction orders from the mixing extremes of complete segregation and maximum mixedness. I n a companion paper ( 5 E ) , they present a new method for measuring contact time distributions employing tracers which are weakly adsorbed on the solid surface. Experimental results are reported using a silica-alumina cracking catalyst and nitrous oxide tracer gas in a six-inch fluidized bed. Vakhrushev and Basov (6E)present theoretical expressions describing heterogeneous fluidization. Tables are provided for estimating bubble velocities and bed heights under both turbulent and laminar conditions. A small amount of experimental data are presented which seem to verify some of the theory. T h e behavior of freely bubbling fluidized beds has been studied by Grace and Harrison (3E). They find that the velocity of bubbles in fluidized beds is greater when the bubble is in a swarm as opposed to the bubble in isolation. They attribute this phenomena to the acceleration caused by overtaking bubbles during the coalescence process. Available theories tend to overestimate the flow in the bubble phase and underestimate the through flow of gas relative to bubbles. The discrepancy can be accounted

AUTHOR Vern W . Weekman, J r . , is supervisor of the Systems Research Group of the Applied Research and Development Division at M o b 2 Oil Corp. Laboratory, Paulsboro, N . J . 08066.

for by considering the effects of bubble shape and high local concentrations of bubbles. Similar results on bubble velocities in freely bubbling beds have been presented by Godard and Richardson ( I E ) . They show experimental data which is in agreement with the work of Grace and Harrison ( 3 E ) . Godard and Richardson (ZE) have used the Ergun equation for pressure drop through fixed beds to express the ratio of the free falling velocity of a n individual particle to the minimum fluidizing velocity as a function of the Galileo nuniber. T h e Ergun equation proved quite consistent with the experimental results. The analysis holds for particulately fluidized systems. Wect of Catalyst Decay on Reactor Performance

Optimum temperature policies for reactors undergoing catalyst deactivation are presented by Szepe and Levenspiel (3F). For simple reaction and deactivation kinetics, they provide an analytical solution to the optimal temperature policy in batch reactors. Results are applicable for arbitrary orders of reaction and deactivation. Catalyst decay has been shown by Weekman (4F) to have a strong effect on the selectivity of catalytic cracking reactions in fixed beds. If the product from a decaying fixed bed is time averaged, the selectivity behavior will be greatly different from that observed in a steady state reactor with the same catalyst and charge stock. Experimental verification is given. I n evaluating catalyst i n decaying fixed beds, a more selective catalyst may actually appear inferior because of the disguise imposed by time averaging. Performance of fixed bed catalytic reactors with poison in the feed has been treated by Wheeler and Robell (5F). A theory is presented which predicts poisoning profiles in the bed as a function of time. I n addition, it is possible to calculate the conversion from the reactor a t various times during the poisoning process. Experimental data are presented for the oxidation of carbon monoxide over a platinum-palladium catalyst supported on alumina. Hydrogen sulfide was present in the feed gas as the poison. I t was comforting that the theory was capable of fairly accurately predicting the conversion over the entire period of poisoning. I n a n excellent blend of theory and experiment, Murakami et a l . (2F) have studied the effect of interparticle diffusion on catalyst fouling. Fouling from both parallel and series reactions were studied theoretically and effectiveness factors were determined. I n the experimental study the disproportionation of toluene over alumina boria catalyst is used as the example parallel reaction scheme while the dehydrogenation of primary alcohols over a n alkaline alumina catalyst was used for the series reaction. Under strong diffusion limitations the parallel disproportionation of toluene gave a shallow carbon layer on the outside of the particle as predicted by the theory. For the series reaction involving the dehydrogenation of n-butyl alcohol under low diffusion conditions, the carbon formed in the center of the pellet where the concentration of the coke producing species was highest. At a higher temperature (higher diffusion conditions) the coke was observed experimentally to be highest on the outer portions of the catalyst. This is in apparent contradiction with the previous theoretical results for series reactions by Massmuni and Smith. The experimental results confirmed the significant effect of series fouling on the yield of the intermediate as predicted by the theory. Using a time on-stream catalyst decay function, Wojciechowski, Juusola, and Downie ( 6 F ) suggest a theoretical classification of catalyst based on decay. The classifications are based on both the behavior of the average conversion with space time and on the behavior of the instantaneous reaction rate as a function of the fraction of surface poisoned. It is difficult in industrial reactors to obtain good estimates of the rate of catalyst decay. Gavalas and Seinfeld (7F)present a theoretical method which offers hope of solving this problem through the use of on-line sequential estimation of kinetic parameters for fixed bed reactors undergoing slow catalyst decay. T h e method, which is computationally simple, can be employed with on-line digital computers. Some numerical examples are given. Determination of Rate Constants from Data

I n a rather frightening example Lumpkin, Smith, and Douglas (6G) have shown how 21 separate rate models could be fit to the VOL. 6 2

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same set of isothermal data at the 99% confidence level. All was not lost, however, when they showed that experiments performed in an adiabatic reactor would more clearly distinguish between the rival models. They further showed that the optimal temperature profiles resulting from the different models could be the most definitive means of discriminating between models. Since many industrial reactors are run adiabatically, the indistinguishable isothermal rate expressions could lead to unpleasant surprises on scale-up. The proposed methods of choosing models based on adiabatic or optimal temperature profiles offer hope of avoiding these pitfalls. I n estimating rate constants from experimental data, Draper, Kanemasu, and Mezaki (4G) have shown that the time elimination procedure can lead to errors in estimating individual rate constants. Elimination of time in the rate equations allows one to estimate the ratios of rate constants. However, in their examples they show that the individual rate constants may be substantially in error. By introducing an extra time parameter, they show by example how the estimates of the rate constants can be substantially improved. This extra time parameter could represent a time lag before complete mixing is achieved in a batch stirred tank reactor. T h e question of whether to use differential or integral rate data in the determination of rate constants is tackled squarely by Barnard and Mitchell (2G). They argue strongly in favor of integral reactor data because of their accuracy at higher conversions. They also show that a particular rate expression is required to satisfy more rigid and accurate conditions when using integral reactor data as opposed to differential data. Ironically they point out that sophisticated experimental design techniques often result in conditions where it is impossible to measure the rate of reaction. An example is given for experiments dealing with the oxidation of benzene over platinum catalyst. I n a subsequent paper (3G), they apply their improved method for selecting Hougen-Watson rate equations to four reaction systems involving the oxidation of various c 6 - C ~hydrocarbons on noble metal catalysts. The least squares treatment of kinetic data to obtain apparent activation energies has been studied by Bagg ( I G ) . H e finds that while the frequency factor may be a function of temperature, such a n effect on the curvature of the Arrhenius curve will have little effect within 5 kcal/mol of the true value. In comparing activation energies between two different reactions, he finds that statistically only differences greater than 3-4 kcal will be significant. Hunter, Hill, and Henson (5G) have proposed a method for designing experiments for the estimation of constants in a mechanistic reactor model. A useful result is the ability to design experiments for estimating certain of the more crucial rate constants in the model. Many important industrial reactions may have only one or two critical reaction steps which strongly influence the process economics. The proposed method enables the experimenter to focus on these values.

Experimental Techniques

Tajbl ( 5 H ) has studied the kinetics of ethane and propane ruthenium-alumina cathydrogenolysis for a commercial 0.5 alyst. H e employed the Carberry reactor which mounts the catalyst in a spinning basket and gives a close approach to a perfectly mixed reaction while minimizing mass transfer resistances. The rate expression for propane hydrogenolysis is first order in propane and three halves order in hydrogen partial pressure. T h e kinetics of ethane hydrogenolysis were quite similar. Using the same reactor, Tajbl ( 6 H ) has studied the kinetics of ethane hydrogenolysis using a commercial nickel on kieselguhr catalyst. Experimental data are presented along with a kinetic rate expression which describes the data. For hydrogenation catalysts Sinfelt ( 4 H ) has proposed a useful experimental reactor for kinetic studies. Kinetic data are obtained during short reaction periods followed by periods of hydrogen flow. The hydrogen period serves to restore the catalyst t o its initial condition and provides a reproducible catalyst activity. A unique experimental technique to differentiate between uniform catalyst poisoning and pore mouth poisoning has been proposed by Balder and Petersen ( 7 H ) for first order reactions. A single catalyst pellet is mounted in a reactor such that both the overall reaction rate and the concentration a t the cenrer line of the particle can be measured. Experimental data from the hydrogenolysis of cyclopropane on platinum-alumina catalyst indicated 56


that the poisoning mechanism was between the limiting cases of uniform poisoning and pore mouth poisoning. Using an interesting experimental method, Basov et a!. ( 2 H ) have experimentally determined density profiles in fluidized beds. A radioactive cesium-137 isotope was placed in the center of the fluidized bed and three scintillation counters mounted 120’ apart could be moved simultaneously over the height of the bed to give complete density profiles. Sharp drops in density were found from 50 to 100 millimeters above the inlet grid. Data are also included for a variety of different grid designs. Lakshmanan and Rouleau ( 3 H ) have studied the kinetics of cupric oxide catalyzed oxidation of propylene in a novel reactor. A conventional stirred tank reactor was employed ; however, the walls were catalytic. This was achieved by a thin film of catalyst on a stainless steel tube which fitted snugly inside the reactor. Acrolein was the desired product. A Hougen-Watson type rate law was proposed. Using both chromatographic and radioactive isotope techniques, Zimin et al. (7H)have studied the kinetics and mechanisms of the catalytic oxidative dehydrogenation of butenes to form divinyl. Experiments were carried out over a Bi-Mo catalyst. The reaction was described by a series of nine steps, and the rate constants were fit from the chromatographic data using a gradient -technique. The rate constants gave a good fit to the experimental data. Using I4C labeled reactants, they independently measured the rate constants by following the radioactive counts of the various constituents. The resulting rate constants from the isotope method agreed very well with those from the chromatographic technique. Reaction Engineering Design Techniques

For a group of industrially important second order reactions, Russell and Buzzelli ( 1 3 5 ) have analyzed the effect of reactor variables on the product distributions. Among the classes of reactions for which the analysis applies are the addition of alkene oxides to proton donors and halogenation or nitration of hydrocarbons. The general effects of temperature, backmixing, and recycle on the resulting product distributions are discussed. As part of a refresher series, Lederman (105)has published a review of available flow sheet simulation programs. The article reviews some of the general problems involved in developing flow sheet simulators as well as discusses existing programs. A graphical method has been presented by Koo and Ziegler (8J,9J) who simplify the design and process calculations for single-stage and multistage reactors. Dimensionless graphs relate input/output variables for chemical conversion from zero, half, first, and second order reactions. I n addition, the curves account for complete microscopic, complete macroscopic, segregated laminar, and plug flow mixing regimes. A simple correction factor has been developed to compute the effect of backmixing on conversion in reactors by Corrigan and Dean ( 4 J ) . The approximation is valid for a useful range of conditions and should be a good rule-of-thumb in designing or interpreting reactor data. The problem of design of commercial plants based on pilot plant data has been given in a series of three papers by Kitzen, Wall, and Cijfer (75);Kern ( 6 J ) ; and Small ( 7 4 J ) . A variety of design experiences was discussed ranging from gas oil pyrolysis to blast furnaces. A common theme was the tranquilizing effect on a commercial designer of smooth orderly pilot plant operation. T o sell the commercial process, smooth pilot plant operation, unfortunately, is stressed but ironically the very smoothness of the operation drastically reduces the useful information obtained, Commercial plants typically operate with large heat and mass transfer gradients which encourage catalyst decay, damage, scaling, and plugging all of which may be missed with smooth pilot plant operation. Deliberately looking for trouble in the pilot plant was advocated to provide more useful design information. A somewhat negative note was struck by Bunn et al. (25)for the use of mathematical models or reaction kinetic principles in the design of a fluid catalytic cracking unit. I n developing a new design for fluid catalytic cracking units, they claim that the “deductive reasoning and the inductive process of discovering explanations from . . . sets of facts are at least as important, , . as the construction of mathematical models and the manipulation of reaction kinetics which are currently so much in vogue.’’ While they showed they were able to improve the design of the unit without the use of reaction engineering principles, a combination

of practical process knowledge coupled with reaction engineering principles might have lead to a n even more effective design. Using both pulse and sinusoidal response techniques, Chung and Wen ( 3 5 )have studied the axial dispersion of liquids flowing through fixed and fluid beds. They found fluid velocity to have a strong effect on the dispersion coefficient with a n increase in velocity causing an increase in the dispersion coefficient. Using both their own data as well as a n extensive amount of literature data, they present a generalized correlation of axial dispersion coefficients for both fixed and fluid beds. While the data show a fair amount of scatter, they do represent the general trend over more than six orders of magnitude in Reynolds number. The effect of imperfect mixing on the stability of auto-refrigerated reactors has been studied by Luyben (77J). The analyses show that hot spots, runaways, and large temperature gradients are possible unless high mixing rates are achieved. The theoretical simulation also showed that severe control problems can be anticipated in operating auto-refrigerated reactor systems. A number of recommended design equations for fluidized beds in both dense fluidized beds and perforated grid tray beds are given by Mukhlenov ( 7 2 J ) . They account for pressure drop, void fraction, and fluidization velocities based on the author’s experience as well as published results. A theoretical study of the effect of catalyst dilution on the degree of conversion in fixed bed catalytic reactors has been given by Van Den Bleek, Van Der Wiele, and Van Den Berg (75J). Dilution of active catalysts is commonly practiced to promote isothermal conditions. A stochastic model was developed which demonstrated that different distributions of catalysts and inert particles at different dilutions can have an effect on the conversion. A dilution criterion is given to determine the maximum allowable dilution and the minimum amount of catalyst. Reaction between two immiscible fluids in counter-current flow in a series of stirred tanks has been considered by Ahluwalia and Levenspiel (7J). An explicit solution for conversion has been found for second order reactions through a series of equally sized backmixed reactors. Useful design charts are presented. A rigorous analysis of a perfectly mixed photochemical reactor has been presented by Jacob and Dranoff ( 5 J ) . A detailed but practical method is presented for the design and radial scale-up of batch and continuous flow reactors based on polychromatic light intensity profiles, light absorption, and quantum efficiency. A method for determining the quantum efficiency dependency has also been given.

Multiphase Reactors

One of the first direct studies of mass transfer limitations in trickle beds undergoing reaction has been presented by Satterfield, Pelossof, and Sherwood ( 3 K ) . Using a single string of porous palladium on alumina particles, they studied the liquid phase hydrogenation of a-methyl styrene to cumene. The experimental results were compared to kinetic studies using powdered catalyst in well stirred reactors. They were able to show that the overall reaction rate was halved by the presence of mass transfer limitations in the outside liquid film. Under these conditions, the effectiveness factor of the particle itself was extremely low. Criteria are presented for estimating whether mass transfer through the liquid film will be significant in industrial trickle flow reactors. Buchanan ( I K ) has studied the pressure drop and liquid hold-up for counter-current flow of gas and liquid in packed columns. He presents extensive experimental data for water, water and sucrose, and diesel oil as the liquid phase, with air as the gas phase. Raschig rings were used as the packing and empirical correlations are presented for hold-up and pressure drops. Axial dispersion measurements have been made in two-phase flow in pipes by Gomezplata and Brown ( 2 K ) . Using two thermal conductivity probes and sodium chloride pulses, they report axial diffusion coefficients varying from 0.2 to 0.3 sq ft/sec with void fractions varying from 10% to 30% and average liquid velocities varying from 6 to 3 ft/sec. Residence time distributions by means of tracer injection have been measured for the liquid phase in two-phase trickle flow in packed beds by Van Swaaij, Charpentier, and Villermaux ( 4 K ) . They found that two different mechanisms caused a spread in the residence time : namely, an eddy diffusion process and a mass exchange with stagnant regions. O n a plot of Bodenstein number us. Reynolds number, they find their experimental data for trickle flow fall within the general region found in the literature for single

phase flow. A model is presented to account for the observed experimental dispersions. Useful Theoretical Developments

In a series of two papers, Wei and Kuo (3L) have made a significant contribution to the understanding of reaction systems involving a large number of separate reacting species. Because of this large number of constituents, it has always been necessary to lump many of them together into pseudo compounds. Wei and Kuo show under what conditions species may be lumped together so as to preserve the important characttristics of the reaction system. They treat exactly lumpable systems in which the lumped classes can be exactly described by complex first order reaction schemes. I n a second category they treat approximately lumpable systems and describe how much error is involved in the lumped system compared to the more complex unlumped system. This analysis offers hope of a more rational design of reactors involving large classes of reacting compounds. How to study nonlinear tubular reactors without solving the equations has been presented by Amundson and Luss ( 7 L ) . I n a general treatment of the tubular steady state reactor, they enumerate important properties which can be deduced from a nonlinear differential equation describing the behavior, Where solutions can be verified only by more numerical calculation, the method provides bench marks for checking the solutions. Optimization of reactors by delaying the addition of reactants has been studied by Jackson and Senior (2L). They show that in certain cases it is possible to selectively increase the yield of the desired product by delaying the addition of one or more reactants. A general criterion is derived to determine cases in which the delay is advantageous. Noncatalytic Reactors

Extensive experimental data from the high temperature pyrolysis isobutene (900-1O5O0C) is presented by Schugerl and Happel ( 6 M ) . Experimental equipment allowed the use of high temperature steam from the hydrogen-oxygen burner. All reactions were shown to be first order and ratios of reaction rate constants are given. Selectivity of a n acetylene production was as high as 35% a t 50y0 conversion of isobutene. Simple first or second order reaction kinetic models were not sufficient to describe the pyrolysis of propane in tubular flow reactors as reported by Crynes and Albright ( 3 M ) . They show that both first and second order rate constants vary significantly with propane conversion. Treatment of the walls of the stainless steel reactors with oxygen-hydrogen bromide in steam promotes secondary reactions. Hydrogen sulfide prevents the secondary reactions by forming a protective metal sulfide film. A test of the validity of the steady state approximation in the pyrolysis of n-butane has been given by Blakemore and Corcoran ( 7 M ) . The equations for a detailed mechanistic model were evaluated on the computer utilizing the best available published experimental data. The resulting solutions were then compared (favorably) with the solution resulting from the simplified steady state approximation. A thorough review of mathematical models for noncatalytic heterogeneous solid-fluid reactions has been given by Wen ( 9 M ) . The article covers a variety of mathematical techniques to treat reactant solids reacting with fluids or gases. Various versions of the unreacted shrinking core model are presented in terms of effectiveness factors. A simplified version of the oil-shale gasification process is given as a n example along with all of the physical constants required by the model. A noncatalytic oxidation of methane with pressures up to 6600 atmospheres has been reported by Hardwicke, Lott, and Sliepcevich ( 4 M ) . Pressure increased the yield of methanol and other oxygenated organic liquid products. The rate of oxygenation was negative second order indicating an inhibiting effect of oxygen. The temperature range studied was 250-365°C and a n activation energy of 42.9 kcal/mol was found for the overall reaction. The kinetics and mechanism of the thermal reaction of ethylene has been studied at temperatures from 703-854OC at atmospheric pressure by Kunugi et al. ( 5 M ) . Extensive chromatographic data are presented for a large range of secondary products produced during the reaction. The rate of ethylene disappearance was expressed as a three halves order reaction with a n activation energy of 49 kcal. From their present work plus extensive literature VOL. 6 2

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sources, the authors have compiled a table of 60 elementary reaction steps occurring in the process with associated activation energies and frequency factors. That bane of chemical engineers, the “plug,” has finally received attention in the form of a kinetic study. Taylor and Wallace (7M) have studied the kinetics of deposit formation due to trace levels of sulfur compound from 200-450’F in the presence of oxygen. The presence of sulfur compounds at the 1000-ppm level increased deposit rates as much as 20 times. Activation energies and rate constants for deposit formation are given. The kinetics and mechanism of the thermal decomposition of n-butanes have been studied by Vedeneev and Kudryavtseva ( 8 M ) . Experimental data are presented over a temperature range from 479 to 547OC and pressure from 50 to 500 mm of mercury. A mechanism in which methyl radicals play an important role is presented along with reaction rate constants and activation energies for the important steps in the mechanism. Reactor dynamics and molecular weight distributions for continuous tubular polymerization reactors have been presented by Cintron-Cordero, Mostello, and Biesenberger (2M). A theoretical study was made of the effect of a residence time distribution on molecular weight dispersion in continuous nonchain and chain polymerization reactions. The effect of nonuniform thermal history was also considered. Kinetic-Catalytic Studies

Rao, Kumar, and Kuloor (5N)present experimental data for the dehydrogenation of butyl alcohol over a copper chromia catalyst supported on a pumice base. Simple first order kinetics were sufficient to describe the rate of reaction and no diffusion limitation was noted. Thomas, Carretto, and Nobe ( 8 N ) have studied the antipollution reaction of carbon monoxide to carbon dioxide over a copper oxide catalyst. They correlate their experimental data with a rate expression which is a function of the ratio of carbon monoxide over carbon dioxide to the three-tenths power. An activation energy of 24 kcal is also reported. The presence of carbon dioxide strongly inhibits the carbon monoxide oxidation. The oxidation of o-xylene on a vanadium pentoxide catalyst has been studied by Herten and Froment ( 3 N ) . Extensive experimental data are given and a power function rate law dependent on the partial pressure of o-xylene and oxygen is proposed. The o-xylene order changed significantly with temperature ; however, the oxygen order remains constant at approximately one third. Product distribution data are also given at 341 ‘C. Ritchie and Nixon ( 6 N ) have studied the dehydrogenation of monocyclic naphthenes over platinum catalyst with the interesting goal of cooling high speed aircraft operating above mach 3. Experimental data are presented for platinum on alumina catalyst at 10 atm (without added hydrogen) at temperatures of 8401200’F. The rate constants and activation energies are given for a number of monocyclic naphthenes including cyclohexane, ethylcyclohexane, dimethylcyclohexane, and di- and tri-ethylcyclohexane. Data for the cis and trans forms of these compounds are also given. An extensive review of oxidative dehydrogenation of hydrocarbons has been presented by Skarchenko ( 7 N ) . The article references both western and Soviet literature as well as U. S . and foreign patents. A limited amount of kinetic data is given. Caretto and Nobe ( I N ) found good agreement between calculated and experimental diffusivity results in the catalytic oxidation of ethylene over a copper oxide-alumina catalyst. In an earlier study involving Cg hydrocarbons, they had found that surface diffusivity played a significant role, while the present study minimizes this effect. Their theoretical predictions match the experimental results well over a range of 200-400OC. The hydrogenation of 1-butene and subsequent isomerization steps have been studied by Ragaini and Somenzi ( 4 N ) over a supported palladium catalyst. Experimental results are given along with proposed mechanism and rate expression. An extensive review of kinetic rate expressions for the catalytic oxidation of sulfur dioxide is given by Weychert and Urbanek ( 9 N ) . Thirty-four rate expressions are given for a variety of platinum and vanadium catalysts. The rate expressions, which are obtained from both western literature as well as eastern Europe and Russia, are discussed. It is quite startling to see the variety of greatly different kinetic rate expressions which describe the same reaction. 6iR


An organic catalyst consisting of pyrolized polyacrylonitrile has been used by Cutlip and Peters ( 2 N ) for dehydration of tertiary butyl alcohol. Possible reaction mechanisms as well as rate laws are given. A typical inorganic alumina catalyst was approximately 20 times as active as the organic catalyst. Mass Transfer in Porous Catalysts

Most diffusivity measurements are made without reaction; yet most diffusivity measurements are used to predict the results of catalytic reaction in catalysts. Sterrett and Brown (9P)have investigated the foundations of this wide-spread practice and the results are not particularly comforting. Using as a reaction the ortho- para shift of hydrogen over ferric oxide catalyst, they have studied the diffusivities measured during reaction and compared them with those predicted by some of the standard nonreacting methods. The experimentally determined diffusivities measured during reaction agreed with each other extremely well. A number of theoretical diffusion models, however, predicted diffusivities which were on the average 40% below the expeqimentally determined values. The authors suggest that some form of surface transport may account for the differences. I n a theoretical study, Hlavacek and Marek ( 2 ) report the effects of simultaneous heat and mass transfer on nonisothermal zero order reactions within a porous catalyst particle. Approximate analytic solutions were given for slabs and cylindrical shapes with a semianalytic method being given for spheres. Interestingly, the authors claim as many as five steady states are possible with spherical catalyst particles. I n a subsequent paper, Hlavacek, Marek, and Kubicek ( 3 P ) present criteria for the existence of multiple solutions within nonisothermal catalyst particles. A generalized treatment for reaction within a porous support containing a homogeneous liquid phase catalyst has been given by Rony (7P). Theoretical relations determining diffusion and reaction kinetic characteristics in such a catalyst are derived. Results are similar to those of Thiele for heterogeneous catalysis without a liquid phase. The ever popular shell progressive mechanism has been employed by Beveridge and Goldie (7P)to study the effectiveness factors and instability in noncatalytic gas-solid reactions. Instabilities caused by both the particle geometry and thermal effects are presented. Using a chromatographic technique, Schneider and Smith (8P) were able to study surface diffusion coefficients for ethane, propane, and n-butane on silica gel. At 5OoC, surface migration accounted for 73% of the total transport for propane in silica gel. The activation energy for the surface diffusion of n-butane was 4.4 kcal. A compilation of surface diffusivity measurements of ethane, propane, and n-butane by a variety of authors is presented in the paper. Luss (5P)has developed a useful modification of the TinklerMetzner approximation of effectiveness factors for exothermic catalytic reactions. The approximation can be used to obtain upper and lower bounds on the exact value of the effectiveness factor. I n addition, the method can be used for a variety of catalyst shapes. A check on interparticle diffusion theory has been presented by Jiracek, Horak, and Pasek ( 4 P ) by hydrogenation of benzene over a nickel alumina catalyst. They were able to measure interparticle temperature by means of an imbedded thermister. Predictions by the theory of Weisz and Hicks (70P)compared satisfactorily with the experimental data. The temperature rise in catalyst particles controlled by a shell progressive mechanism may be much higher than previously anticipated. Luss and Amundson (6P)have shown theoretically that very high temperatures may wcur during diffusion controlled gas-solid reactions operating in the shell progressive regime. Previous investigators have neglected the heat capacity of the unreacted core which has resulted in too low a n estimate of the potential temperature rise. The transient analysis reveals that the temperature rise may be severe enough to cause sintering damage to the catalyst particles. BIBLIOGRAPHY

Behavior and Performance of Fixed Beds (1A) Carberry, J. J., and White, D., TND. ENC. CHEM., 61 (7), 27-35 (1968). (2A) Haughe D. P., and Beveridge, G. S. G., Can. J. Chem. Eng., 47, 130-140 (Apri! 19697:

(3A) Littman, H., Barile, R. G., and Pulsifer, A. H., IND. ENC. CHEM.,FUNDAM., 7,554-561 (1968). (4A) Ridgway,K., andTarbuck,K. J., Chem. Eng.Sci., 23,1147-1155 (1968). (5A) Schertz, W. W., andBischoff, K.B., A.1.Ch.E. J.,15,597-604 (1969).

Diffusion and Reaction in Zeolite Catalysts (1B) Beecher R Voorhies A., and Eberly, P., IND.ENC. CHEM.,PROD. RES. DEVELOP., 2d3-209 (196h). (2B) Danner, R. P., and Wenzel, L.A., A.1.Ch.E. J.,15,5151520 (1969). (3B) Eberly, P. E., Jr., IND. ENC.CHEM.,FUNDAM., 8,25-35 (1969). (4B) Eberly, P. E., Jr., IND.ENC.CHEM.,PROD.RES.DEVELOP., 8,140-144 (1969). (5B) Hopkins, D. P., J. Catol., 12,325-334 (1968). (6B) Nace, D. M., IND.END.CHEM.,PROD.RES.DEVELOP., 8,24-31 (1969). (7B) Nace, D. M., ibid., pp 31-38. (8B) Satterfield, C.N., and Frabetti, A. J., A.1.Ch.E. J.,13,731 (1967). (9B) Thomas, C. L., and Barmby, D. S., J. Catol., 12,341-346 (1968). (10B) Voorhies,A., andBryant,P.A., A.Z.Ch.E.J., 14,852-856 (1968).


Industrial Process Kinetics (1C) Carr, N. L., and Stahfeld, D. L., Pre rint A.1.Ch.E. Mtg., Los Angcles (Decemberl, 1968); Summary, OilGasJ., I t 0 (September 15, 1969). (2C) Chartrand, G., and Crowe, C. M., Can. J. Chem. Eng., 47, 296-301 (June 1969). 7, 605-610 (3C) Dyson, D. C., and Simon, J. M., IND.END.CHEM.,FUNDAM., (1968). (4C) Gwyn, J. E., A.l.Ch.E. J . , 15,35-38 (1969). (5C) Hyman, M. H., Hydrocarbon Proc., 47 (7), 131-137 (July 1968). (6C) Orochko, D. I., and Chernakova, G. N., Int. Chem. Eng., 9,219-221 (1969). (7C) Ozawa, Y.,INn. ENC.CHEM.,PROCESS DES.DEVELOP., 8,378-383 (1969). (8C) Pfefferle, W. C., Dalson, M. H., Nevison, J. A., Kopf, F. W., and Decker, W. H.,HydrocarbonProc.,48(5), 111-115 (May1969). (9C) Qader, S. A., and Hill, G. R., IND.ENC. CHEM.,PROCESS DES. DEVELOP., 8, 98-105 (1969). (1OC) Ruthven, D. M., Can. J.Cham. Eng., 47,327-331 (June 1969). (11C) Sheel, J. G. P., and Crowe, C.M.,ibid.,pp 183-187 (April 1969). (12C) Ting,A. P., and Wan, S. W., Chem. Eng., 76,185-191 (May 19,1969). (13’2) Weisz, P.B., IND.ENG. CHEM.,FUNDAM., 8,325-329 (1969). (14C) Woinsky, S. G., IND.ENG.CHEM., PROCESS DES.DEVELOP., 7,529-538 (1968).

Mass and Heat Transfer in Gas-Solid Fluidized Beds (1D) Bukareva, M. F., Chlenov, V. A., and Mikhailov, N. V., Intern. Chem. Eng., 9,119-121 (1969). (2D) Carlos, C. R., and Richardson, J. F., Chem. Eng.Sci., 23,813-824 (1968). (3D) Carlos, C. R., and Richardson, J. F., ibid., pp825-831. (4D) Clamen, A., and Gauvin, W. €I., Can. J. Chcm. Enp., 46, 223-228 (August 1968). (5D) Gel’perin, N. I., Ainshtein, V. G., and Korotyanskaya, L. A., int. Chem. Eng., 9,137-142 (1969). (6D) Kunii, D., and Levenspiel, O., IND.END.CHEM., FUNDAM., 7,446-452 (1968). (7D) Kunii, D., and Levenspiel, O., IND.ENC. CHEM.,PROCESS DES. DEVELOP., 7, 481-492 (1968). (ED) Oigenblik, A. A., Genin, L. S., Goikhman, I. D., and Filippova, L. A., I n f . Chem. Ens., 9,201-204 (1969). (9D) Rowe, R. N., Partridge, B. A., and Lyall, E., Chem. Eng. Sci., 19,973 (1964). (10D) Szekely, J., and Fisher, R. J., ibid., 24,833-849 (1969). (11D) Toei, R., Matsuno, R., Miyagawa, H., Nishitani, K., and Komagawa, Y., Znt. Chem. Eng., 9,358-364 (1969). (12D) Yoshida, K., Kunii, D., and Levenspiel, O., INn. ENC. CHEM.,FUNDAM., 8,402-406 (1969). (13D) Young, F. M., and Holman, J. P., ibid., 7, 561-567 (1968).

Hydrodynamics of Fluidized Beds (1E) (2E) (3E) (4E) (5E) (6E)

Godard, K., and Richardson, J. F., Chem. Eng.Sci., 24,363-367 (1969). Godard, K., and Richardson, J. F., ibid., pp 663-670. Grace, J. R., and Harrison D., ibid., pp 497-508. Nauman, E. B., and Collinge, C. N., ibid., 23, 1309-1316 (1968). Nauman, E. B., and Collinge, C. M., ibid., pp 1317-1326. Vakhrushev, 1.A., and Basov, V.A., Znl. Chem. Enz., 9,83-87 (1969).

Effect of Catalyst Decay on Reactor Performance (1F) Gavalas, G. R., and Seinfeld, J. H., Cham. Eng.Sci., 24,625-636 (1969). (2F) Murakami, Y., Koba ashi, T Hattori, T., and Masuda, M., IND. END. cHEM., FUNDAM., 7,599-685 (1968)’: (3F) Szepe, S., and Levenspiel, O., Chem. Eng.Sci., 23,881-894 (1968). (4F) Weekman, V. W., Jr., IND.END.CHEM.,PROCESS DES.DEVELOP., 8, 385-391 (1969). (5F) Wheeler, A., and Robell, A. J., J . Catal., 13,299-305 (1969). (6F) Wojciechowski, B. W., Juusola, J. A., and Downie, J., Can. J. Chem. En$., 47,338-340 (June 1969).

Determination of Rate Constants from Data (1G) Bagg, J., J . Cold., 13,271 (1969). (2G) Barnard, J. A., and Mitchel1,D. S., ibid., 12,376-385 (1968). (3G) Barnard, J. A., and Mitchell, D. S., ibid., pp 386-397. (4G) Draper, N. R., Kanemasu, H., and Mezaki, R., IND.ENC. CHEM.,FUNDAM., 8,423-427 (1969). (5G) Hunter, W. G., Hill, W. J., and Henson, T. L., Can. J. Chcm. Eng., 47, 76-80 (February 1969). (6G) Lumpkin, R. E., Smith W. D., and Douglas, J. M., IND.ENC.CHEM.,FUNDAM., 8,407-411 (1969).

Experimental Techniques (1H) Balder, J. R., and Petersen, E. E.,Chcm. Eng.Sci.,23,1287-1291 (1968). (2H) Basov, V. A., Markhevka, V. I., Melik-Akhnazarov, T. Kh., and Orochko, D. I., Znl. Chem. Ens., 9,263-266 (1969). (3H) Lakshmanan, R., and Rouleau, D., Can. J. Chem. Eng., 47, 45-50 (February 1969). (4H) Sinfelt, J.H.,Chem.Eng.Sci., 23, 1181-1184 (1968). (5H) Tajbl, D. G., IND. ENC.CHEM.,PROCESS.DES.DEVELOP., 8,364-369 (1969). (6H) Tajbl, D. G., Can. J . Chcm. Eng., 47,154-159 (April 1969). (7H) Zimin, R. A., Gagarin, S . G., Berman, A. D., and Yanovskii, M. I., Kinetics Calol. (USSR)(Engl. Transl.), 9, 95-100 (1968).

Reaction Engineering Design Techniques (1J) Ahluwalia, M. S., and Levenspiel, O., Can. J.Chem. Eng., 46, 443-446 (December 1968). (25) Bunn, D. P., Gruenke, G. F., Jones, H. B., Luessenhop, D. C., and Youngblood, D. J., Chem. Ens. Progr., 65 (6), 88-93 (June 1969). (35) Chung, S. F., and Wen, C. Y . ,A.1.Ch.E. J.,14,857-866 (1968). (45) Corrigan, J. E., and Dean, M. J., Hydrocarbon Proc., 47 (7), 149-152 (July 1968). (5J) Jacob, S. M., and Dranoff, J. S., CEPSymp. Series 89, Vol. 64, 54-63 (1968). (6J) Kern, D. Q., Chem. Eng. Progr., 65 (7), 77-80 (July 1969). (75) Kitzen, M. R., Wall, F. M., and Cijfer, H. J., ibid., pp 71-76. (8J) Koo, L., and Ziegler, E. N., Chem. Eng.Sci., 24,217-222 (1969). (9J) Koo, L., and Ziegler, E. N., Chem. Ens., 76,91-95 (February 24,1969). (1OJ) Lederman, P. B., Chcm. Eng., 75,127-132 (December 2,1968). (11J) Luyben, W. L., A.Z.Ch.E.J., 14,880-885 (1968). (125) Mukhlenov, I. P., Znt. Chem. Eng., 9,37-42 (1969). (135) Russell, T. W. F., and Buzzelli, D. T., IND.ENG.CHEM.,PROCESS DES.DEVELOP., 8,2-9 (1969). (145) Small, W. M., Chcm. Eng. Progr. 65 (7),81-82 (July 1969). (155) Van Den Bleek, C. M., Van Der Wiele, K., and Van Den Berg, P. J., Chcm. Eng. Scr., 24,681-694 (1969).

Multiphase Reactors (1K) Buchanan, J. E., IND. END.CHEM.,FUNDAM., 8,502-51 1 (1969). (2K) Gomezplata, A.,andBrown,R. W., A.Z.Ch.E. J., 14,657-658 (1968). (3K) Satterfield, C. N., Pelossof, A. A., and Sherwood, T. K., ibid., 15, 226-234 (1969). . . (4K) Van Swaai’ W. P. M., Charpentier, J. C., and Villermaux, J., Chcm. Ens. Sci.,24, 1083-1&35 (1969).

Useful Theoretical Developments (1L) Amundson, N. R., and Luss, D., Can. J. Chem. Eng., 46, 424-433 (December 1968). (2L) Jackson, R., and Senior, M. G., Chem. Eng. Sci.,23,971-980 (1968). (3L) Wei, J., and Kuo, J. C. W., IND.ENC.CHEM.,FUNDAM., 8,114-133 (1969).

Noncatalytic Reactors (1M) Blakemore, J. E., and Corcoran, W. H., IND.ENC.CHEM.,PROCESS DES.DEVELOP., 8,206-209 (1969). (2M) Cintron-Cordero, R., Mostello, R. A., and Biesenberger, J. A., Can. J . Ckcm. Enp., 46,434-443 (December 1968). (3M) Crynes, B. L., and Albright, L. F., IND.END.CHEM.,PROCESS DES.DEVELOP., 8,25-31 (1969). (4M) Hardwicke, N. L., Lott, J. L., and Sliepcevich, C. M., IND.END.CHEM.,PROD. 8, 133-140 (1969). RES.DEVELOP., (5M) Kunugi, T., Sakai, T., Soma, K., and Sasahi, Y.,IND.ENC.CHEM.,FUNDAM., 8,374-383 (1969). (6M) Schugerl, K., and Happel, J., IND.ENC.CHEM.,PROCESS DES.DEVELOP., 8, 419-431 (1969). (7M) Taylor W F and Wallace, T. J., IND.ENC. CHEM.,PROD.RES.DEVELOP., 7,198-202’(1i683. (EM) Vedeneev V. I and Kudryavtseva, Y.I., Kinetics Cotal. (USSR) (Engl. Transl.), 9,81b-814 ii968). (9M) Wen, C. Y., IND.ENC.CHEM.,60 ( 9 ) , 34-54 (1968).

Kinetic-Catalytic Studies (1N) Caretto, L. S.,andNobe,K., A.I.Ch.E.J., 15, 18-24 (1969). (2N) Cutlip, M.B., and Peters, M. S.,CEPSymg.Scrier89, Vol. 64,l-11 (1968). (3N) Herten, J., and Froment, G. F., IND.ENO. CHEM.,PROCESS DES. DEVELOP. 7,516-526 (1968). (4N) Ragaini, V., and Somenzi, G., J. Cotal., 13,20-24 (1969). (5N) Rao, U. R., Kumar, R., and Kuloor, N. R., IND.ENO.CHEM.,PROCESS DES. DEVELOP., 8,9-16 (1969). (6N) Ritchie A. W and Nixon, A. C., IND.ENG.CHEM.,PROD.RES.DEVELOP., 7,209-21 5 ’(1968):’ (7N) Skarchenko, V. K.,Int. Chem. Ens., 9,1-23 (1969). (EN) Thomas, N. T., Carretto, L. S., and Nobe, K., IND. ENC.CHEM.,PROCESS DES. DEVELOP., 8,282-287 (1969). (9N) Weychert, S., and Urbanek, A., Int. Chem. Eng., 9,396-403 (1969).

Mass Transfer in Porous Catalysts (1P) Beveridge, G. S. G., and Goldie, P. J., Chcm. Eng.Sci., 23,913-929 (1968). (2P) Hlavacek, V., and Marek, M., ibid., pp 865-880. (3P) Hlavacek, V., Marek, M., andKubicek, M., ibid., pp 1083-1097. (4P) Jiracek, F., Horak, J., and Pasek, J., A.1.Ch.E. J.,15,400-404 (1969). (5P) Luss, D., ibid., 14, 966-969 (1968). (6P) Luss, D., and Amundson, N. R., ibid., 15, 194-199 (1969). (7P) Rony, P. R., Chem. Eng.Sci., 23,1021-1034 (1968). (8P) Schneider, P., and Smith, J. M., A.l.Ch.E. J.,14,886-895 (1968). (9P) Sterrett, J. S., and Brown, L. F., ibid., pp 696-702. (1OP) Weisz, P. B., and Hicks, J. S., Chem. Ens. Sci., 17,256 (1962).

VOL 6 2

NO. 5

MAY 1970