Reactor With Mass Transfer

As indicated in our earlier discussions about kinetics in Chap. 2, chemical reactors sometimes have mass-transfer limitations as well as chemical reaction-rate limitations. Mass transfer can become limiting when components must be moved 

Reactant A is fed as a gas through a distributor into the bottom of the liquid-filled reactor. A chemical reaction occurs between A and B in the liquid phase to form a liquid product C. Reactant A must dissolve into the liquid before it can react. 

If this rate of mass transfer of the gas A to the liquid is slow, the concentration of A in the liquid will be low since it is used up by the reaction as fast as it arrives. Thus the reactor is mass-transfer limited. 

If the rate of mass transfer of the gas to the liquid is fast, the reactant A concentration will build up to some value as dictated by the steadystate reaction conditions and the equilibrium solubility of A in the liquid. The reactor is chemical-rate limited. 

Notice that in the mass-transfer-limited region increasing or reducing the concentration of reactant B will make httle difference in the reaction rate (or the reactor productivity) because the concentration of A in the liquid is so small. Likewise, increasing the reactor temperature will not give an exponential increase in reaction rate. The reaction rate may actually decrease with increasing temperature because of a decrease in the equihbrium solubihty of A at the gas-liquid interface. 

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The system discussed is easily extended to a three-phase reactor with mass transfer and reaction in series, for example a gas is absorbed in a liquid in which nonporous particles are suspended. Reaction occurs at the surface of the particles. Examples are the hydrogenation of organic liquids with a solid catalyst and the alkylation of a liquid re-... 

CHEMICAL REACTORS WITH MASS TRANSFER LIMITATIONS... 

Reactor with mole changes and variable density, Chapter 8, p. 130. Chemical reactors with mass transfer limitations, Chapter 8, p. 131. Continuous stirred-tank reactors. Chapter 8, pp. 135, 136. 

Chapter 4 eoncerns differential applications, which take place with respect to both time and position and which are normally formulated as partial differential equations. Applications include diffusion and conduction, tubular chemical reactors, differential mass transfer and shell and tube heat exchange. It is shown that such problems can be solved with relative ease, by utilising a finite-differencing solution technique in the simulation approach. 

We have frnished our discussions of the fundamentals of catalytic reactions, catalytic reactors, and mass transfer effects. While we noted that most catalytic reactions can be made to exhibit complicated kinetics, we have confined our considerations to the almost trivial reaction r" = k"CA- We did this because the algebra was messy enough with the simplest kinetics. 

These equations are only vahd if there is no density change in the reactor, because otherwise W 7 v and Cj is not an appropriate variable. With mass transfer to and from a phase, one expects the volumetric flow rates in and out of the phase to not be identical. For... 

The agitated cell reactor consists of two chambers, one for the liquid phase and another for the gas-phase, which can both be independently mixed by two mixers. In this reactor the mass-transfer area can be varied independently of the gas flow rate by installing various porous plates with a defined number of holes, i. e. a defined contact area gas-liquid, between the two chambers. The value of kL can then be determined from the measurement of kLa. 

Using this approach of a selectivity term SPFR Sunder and Hempel (1996) successfully modeled the oxidation of small concentrations of Tri- and Perchloroethylene (c(M)a = 300-1300 pg D) by ozone and hydrogen peroxide in a synthetic ground water (pH = 7.5-8.5 c(Sj) = 1-3 mmol C03 L"1). In this study an innovative reaction system was used the oxidation was performed in a tube reactor and mass transfer of gaseous ozone to pure water was realized in a separate contactor being located in front of the tube reactor. By this way a homogeneous system was achieved. Since the two model compounds react very slowly with molecular ozone (kD < 0.1 L mol-1 s "1), nearly the complete oxidation was due to the action of hydroxyl radicals, which were produced from the two oxidants (03/H202). With... 

An exothermal reaction is to be performed in the semi-batch mode at 80 °C in a 16 m3 water cooled stainless steel reactor with heat transfer coefficient U = 300 Wm"2 K . The reaction is known to be a bimolecular reaction of second order and follows the scheme A + B —> P. The industrial process intends to initially charge 15 000 kg of A into the reactor, which is heated to 80 °C. Then 3000 kg of B are fed at constant rate during 2 hours. This represents a stoichiometric excess of 10%.The reaction was performed under these conditions in a reaction calorimeter. The maximum heat release rate of 30Wkg 1 was reached after 45 minutes, then the measured power depleted to reach asymptotically zero after 8 hours. The reaction is exothermal with an energy of 250 kj kg-1 of final reaction mass. The specific heat capacity is 1.7kJ kg 1 K 1. After 1.8 hours the conversion is 62% and 65% at end of the feed time. The thermal stability of the final reaction mass imposes a maximum allowed temperature of 125 °C The boiling point of the reaction mass (MTT) is 180 °C, its freezing point is 50 °C. 

As is shown in Figure 2, in the two-phase model the fluid bed reactor is assumed to be divided into two phases with mass transfer across the phase boundary. The mass transfer between the two phases and the subsequent reaction in the suspension phase are described in analogy to gas/liquid reactors, i.e. as an absorption of the reactants from the bubble phase with pseudo-homogeneous reaction in the suspension phase. Mass transfer from the bubble surface into the bulk of the suspension phase is described by the film theory with 6 being the thickness of the film. D is the diffusion coefficient of the gas and a denotes the mass transfer coefficient based on unit of transfer area between the two phases. 6 is given by 6 = D/a. 

Whereas the laboratory fluidized bed is generally operated with no internals, plant equipment often must contain bundles of heat-exchanger tubes. Screens, baffles, or similar internals are frequently used to redisperse the bubble gas in industrial reactors. The mass-transfer area is thus increased relative to the fluidized bed without internals the extra area can be utilized to partially offset the conversion-reducing effects of bed diameter and gas distributor [122]. 

Hofmann, H., Reaction Engineering Problems in Slurry Reactor in Mass Transfer with Chemical Reaction in Multiphase Systems, Vol. II (E. Alper, ed.) NATO ASI Series, Martinus Nijhoff Publishers, The Hague, 1983. 

This has the effect of removing the mass transfer limitation, and the result should be the same as shown in Figure 8.3, and it is the same. The code is returned to its proper form, and the program mn ratel mass calculates the output (see Figure 8.9). With mass transfer resistance included, the outlet concentration of B is 0.61. When there was no mass transfer limitation, the outlet concentration of B was 0.85. Thus, the reactor is not able to produce as much product, and a bigger reactor is required. 

As long as these unit operations are not available, process development departments will have to develop their own unit operations to close the gap. Figure 2.1-7 shows an example of an extraction unit developed and used by the central process development department of Merck KGaA for extractions and hydrolysis on a laboratory scale. In this special case an organic solution was hydrolyzed with a weak acid. Compared with batch reactors the mass transfer is very good and there is no problem with settling the phases. 

Mass transfer involves establishing a transfer between the elementary regions of the reactor and between individual phases (interfacial mass transfer coefficients gas phase mass transfer, liquid phase mass transfer, mass transfer with reaction, liquid-solid mass transfer), as well as other elementary phenomena and processes connected with mass transfer gas phase phenomena and processes (gas hold-up, bubble size, interfacial area and bubble coalescence/redispersion), volumetric mass transfer and power consumption during mass transfer (2). 

The overall description (model) of a reactor is obtained through process synthesis by combining models of reactor hydrodynamics, mass transfer and heat exchange with an appropriate cell (subcellular) or population model ( 1).Description of a population should take into consideration possible dispersed or aggregated (the distinct morphological appearances of a culture pellets, mycelium, flocks, growth on reactor wall in the form of microbial film) forms of population. Biomass support particles are gaining appreciable importance in aerobic (40) as well as in anaerobic processes. 

The effect of incorporating tin into titanium sUicahte-1 (TS-1) on the kinetics of phenol hydroxylation to dihydroxybenzenes with aqueous hydrogen peroxide has been investigated (197). The pathways are illustrated in Figure 1.9. The hydroxylation reaction was modeled using the results obtained with a batch reactor, whereby mass transfer Hmitations were carefully excluded. The analysis of the kinetics indicated that under the same reaction conditions, titanium-tin sUicalite-l (Ti—Sn-S-1) gave a h her phenol conversion rate than TS-1. This difference was attributed to the presence of active tin sites. The incorporation of tin influences the initiation of... 

Simulations show that, in order to achieve the same conversion degree of a fluidized bed membrane reactor without mass transfer limitations, the membrane area installed in the reactor needs to be increased 2.4 times with respect to the case without limitations as reported in the figure. Figure 10.10 shows that a decrease of 10 times in the mass transfer limitations is enough to reach the limit conversion required. 

Interfacial areas per unit volume in falling-film microreactors have been reported to be as high as 25,000 m m, as compared with the values of 1-200 m m typical in bubble columns. This effect is particularly important in gas-liquid reactions because the rate of mass transfer from the gas to the liquid limits the reaction rate. For the hydrogenation of cyclohexene to cyclohexane in this type of reactor, the mass transfer rate constant Kid) was found to be in the range 3-7 s which is two orders of magnitude higher than that for conventional reactors. 

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