Mass transfer operating regimes

Yeong et al. [100,101] used a microstructured film reactor for the hydrogenation of nitrobenzene to give aniline in ethanol at a temperature of 60 °C, a H2 partial pressure of 0.1-0.4 MPa, and residence times of 9-17 s. Palladium catalysts were deposited as films or particles on a microstructured plate. Confocal microscopy was used to measure the liquid film thickness, which increased from 67 to 92 pm as flow rates were increased from 0.5 to 1.0 cm3 min-1. The value of kha characteristic of this system was estimated to be 3-8 s 1 at an interfacial surface area (per reactor volume) of 9000-15000 m2 m 3. Conversion was found to be affected by both liquid flow rate and H2 partial pressure, and the reactor operated between the kinetic and mass transfer-controlled regimes. 

The mass transfer coefficient increases only slightly with temperature, so above a certain temperature the reaction becomes mass transfer controlled. Further increases in temperature give almost no change in conversion. The transition to mass transfer control occurs at a lower temperature for very reactive species, such as H2 and CO, than for hydrocarbons, but the kinetics of oxidation are often not known. The design temperature and flow rate are based on lab tests or experience with similar materials. The reactor is usually operated in the mass transfer control regime, where the conversion depends on the rate of mass transfer and the gas flow rate. 

Regime B 0.3 < Ha< 1.0 and k a sJc B — In this case, the gas-liquid mass transfer rate decides the rate of absorption of the gas. As an example in a stirred tank reactor with relatively low value of k a, the inequality given by Equation 2.7 is likely to hold, and the reactor will operate in a mass transfer-controlled regime. The solved reactor design problem in Section 7A.10 supports this observation. 

The inclusion in the model of the three peculiar values 0,1/2, and 1 means that the three fundamental regimes of mass transfer can be deduced from the generalized mass transfer operator as shown in Table 10.3. 

Over 25 years ago the coking factor of the radiant coil was empirically correlated to operating conditions (48). It has been assumed that the mass transfer of coke precursors from the bulk of the gas to the walls was controlling the rate of deposition (39). Kinetic models (24,49,50) were developed based on the chemical reaction at the wall as a controlling step. Bench-scale data (51—53) appear to indicate that a chemical reaction controls. However, flow regimes of bench-scale reactors are so different from the commercial furnaces that scale-up of bench-scale results caimot be confidently appHed to commercial furnaces. For example. Figure 3 shows the coke deposited on a controlled cylindrical specimen in a continuous stirred tank reactor (CSTR) and the rate of coke deposition. The deposition rate decreases with time and attains a pseudo steady value. Though this is achieved in a matter of rninutes in bench-scale reactors, it takes a few days in a commercial furnace. 

In the design of optimal catalytic gas-Hquid reactors, hydrodynamics deserves special attention. Different flow regimes have been observed in co- and countercurrent operation. Segmented flow (often referred to as Taylor flow) with the gas bubbles having a diameter close to the tube diameter appeared to be the most advantageous as far as mass transfer and residence time distribution (RTD) is concerned. Many reviews on three-phase monolithic processes have been pubhshed [37-40]. 

At high gas velocities in the bed, the stable bubbles break down into unstable voids that continuously disintegrate and reform. This type of bed is said to be operating in the turbulent fluidized-bed regime, and is characterized by higher heat- and mass-transfer rates than in the bubbling bed. As the gas velocity is increased further, the... 

Here, issues in relation to the trickle flow regime—isothermal operation and plug flow for the gas phase—will be dealt with. Also, it is assumed that the flowing liquid completely covers the outer surface particles (/w = 1 or aLS = au) so that the reaction can take place solely by the mass transfer of the reactant through the liquid-particle interface. Generally, the assumption of isothermal conditions and complete liquid coverage in trickle-bed processes is fully justified with the exception of very low liquid rates. Capillary forces normally draw the liquid into the pores of the particles. Therefore, the use of liquid-phase diffusivities is adequate in the evaluation of intraparticle mass transfer effects (effectiveness factors) (Smith, 1981). 

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