Stirred liquid phase

In the case of a laminar flow, the flow velocity is zero at the plane electrode surface, then continuously increasing within a given layer (it is called -> Prandtl boundary layer) and eventually reaches the value characteristic to the stirred liquid phase. 

In a typical synthesis, a mixture of C13H12O2 bisphenols was prepared in 80% yield by slow addition of a solution of trioxane (0.036 mole) in benzene (over 1.75 hour) to a stirred, liquid phase suspension of phenol (0.64 mole) and HY zeolite (5 gm) at 182°. The ratio of the 2,2, 2,4, and 4,4 isomers was 1.3 1.8 1.0. This technique, which afforded very high instantaneous ratios of phenol to aldehyde, prevented rapid catalyst aging. Generally, high yields were observed for carbonyl reactants with no a-hydrogens, since competitive intracrystalline aldol condensation reactions were eliminated. 

Rorvik et al. studied unsupported Nafion for isobutane/1-butene alkylation in a stirred liquid phase batch reactor [12]. The production of trimethylpentanes (the most desirable alkylate product) was shown to cease within 30 minutes of operation. More recently, silica-supported Nafion was used to catalyze the same reaction [13]. Once again, rapid deactivation with respect to trimethylpentane formation was observed. It was hypothesized that the strongest acid sites—the most active for alkylation—are also the first to be poisoned. 

The most simple, generic model of a separation process is a single equilibrium flash unit. Applying the usual assumptions (i.e. negligible vapor holdup, equilibrium, perfect stirred liquid phase, no chemical reactions) and assuming a constant molar holdup in the vessel, the model equations for the general dynamic case relating the distribution functions F (, ) and F (, t) of the vapor and liquid phase read as follows ... 

The use of an unnecessarily hot utility or heating medium should be avoided. This may have been a major factor that led to the runaway reaction at Seveso in Italy in 1976, which released toxic material over a wide area. The reactor was liquid phase and operated in a stirred tank (Fig. 9.3). It was left containing an uncompleted batch at around 160 C, well below the temperature at which a runaway reaction could start. The temperature required for a runaway reaction was around 230 C. ... 

By contrast, if the reactor is continuous well-mixed, then the reactor is isothermal. This behavior is typical of stirred tanks used for liquid-phase reactions or fluidized-bed reactors used for gas-phase reactions. The mixing causes the temperature in the reactor to be effectively uniform. 

There have been many modifications of this idealized model to account for variables such as the freezing rate and the degree of mix-ingin the liquid phase. For example, Burton et al. [J. Chem. Phy.s., 21, 1987 (1953)] reasoned that the solid rejects solute faster than it can diffuse into the bulk liquid. They proposed that the effect of the freezing rate and stirring could be explained hy the diffusion of solute through a stagnant film next to the solid interface. Their theoiy resulted in an expression for an effective distribution coefficient k f which could be used in Eq. (22-2) instead of k. 

In order to allow integration of countercurrent relations like Eq. (23-93), point values of the mass-transfer coefficients and eqiiilibrium data are needed, over ranges of partial pressure and liquid-phase compositions. The same data are needed for the design of stirred tank performance. Then the conditions vary with time instead of position. Because of limited solubihty, gas/liquid reactions in stirred tanks usually are operated in semibatch fashion, with the liquid phase charged at once, then the gas phase introduced gradually over a period of time. CSTR operation rarely is feasible with such systems. 

Two complementai y reviews of this subject are by Shah et al. AIChE Journal, 28, 353-379 [1982]) and Deckwer (in de Lasa, ed.. Chemical Reactor Design andTechnology, Martinus Nijhoff, 1985, pp. 411-461). Useful comments are made by Doraiswamy and Sharma (Heterogeneous Reactions, Wiley, 1984). Charpentier (in Gianetto and Silveston, eds.. Multiphase Chemical Reactors, Hemisphere, 1986, pp. 104—151) emphasizes parameters of trickle bed and stirred tank reactors. Recommendations based on the literature are made for several design parameters namely, bubble diameter and velocity of rise, gas holdup, interfacial area, mass-transfer coefficients k a and /cl but not /cg, axial liquid-phase dispersion coefficient, and heat-transfer coefficient to the wall. The effect of vessel diameter on these parameters is insignificant when D > 0.15 m (0.49 ft), except for the dispersion coefficient. Application of these correlations is to (1) chlorination of toluene in the presence of FeCl,3 catalyst, (2) absorption of SO9 in aqueous potassium carbonate with arsenite catalyst, and (3) reaction of butene with sulfuric acid to butanol. 

In many important cases of reactions involving gas, hquid, and solid phases, the solid phase is a porous catalyst. It may be in a fixed bed or it may be suspended in the fluid mixture. In general, the reaction occurs either in the liquid phase or at the liquid/solid interface. In fixed-bed reactors the particles have diameters of about 3 mm (0.12 in) and occupy about 50 percent of the vessel volume. Diameters of suspended particles are hmited to O.I to 0.2 mm (0.004 to 0.008 in) minimum by requirements of filterability and occupy I to 10 percent of the volume in stirred vessels. 

The effect of physical processes on reactor performance is more complex than for two-phase systems because both gas-liquid and liquid-solid interphase transport effects may be coupled with the intrinsic rate. The most common types of three-phase reactors are the slurry and trickle-bed reactors. These have found wide applications in the petroleum industry. A slurry reactor is a multi-phase flow reactor in which the reactant gas is bubbled through a solution containing solid catalyst particles. The reactor may operate continuously as a steady flow system with respect to both gas and liquid phases. Alternatively, a fixed charge of liquid is initially added to the stirred vessel, and the gas is continuously added such that the reactor is batch with respect to the liquid phase. This method is used in some hydrogenation reactions such as hydrogenation of oils in a slurry of nickel catalyst particles. Figure 4-15 shows a slurry-type reactor used for polymerization of ethylene in a sluiTy of solid catalyst particles in a solvent of cyclohexane. 

Homogeneous reactions are those in which the reactants, products, and any catalysts used form one continuous phase (gaseous or liquid). Homogeneous gas phase reactors are almost always operated continuously, whereas liquid phase reactors may be batch or continuous. Tubular (pipeline) reactors arc normally used for homogeneous gas phase reactions (e.g., in the thermal cracking of petroleum of dichloroethane lo vinyl chloride). Both tubular and stirred tank reactors are used for homogeneous liquid phase reactions. 

A mixture consisting of 22.7 g potassium o-bromobenzoate, 16.6 g 2,6-dichloro-3-methvlani-line, 12 ml N-ethylmorpholine, 60 ml diethylene glycol dimethyl ether, and 1.0 g anhydrous cupric bromide is heated in a nitrogen atmosphere at 145 C to 155°C for 2 hours. The reaction mixture is diluted with 60 ml diethylene glycol dimethyl ether and acidified with 25 ml concentrated hydrochloric acid. The acidic mixture is diluted with 100 ml of water and the liquid phase decanted from the insoluble oil. The insoluble oil is stirred with methanol and the crystalline N-(2,6-dichloro-3-methylphenyl)anthranilic acid which separates is collected and washed with methanol. The product, after recrystallization from acetone-water mixture melts at 248 C to 250°C. 

To 19 8 of well-agitated distilled water plus 18 g of ditertiary-butyl-ppinene oxide that was about half racemic, half d-form. The temperature was maintained at 30°C to 50°C, first with ice bath cooling and then with tap water cooling. The addition of the pinene oxide required 1 h hours. After the addition was complete and the exothermic reaction was about over, the mixture was stirred for 1 h hours at about 30°C, and then centrifuged to separate the crude sobrerol from the liquid phase consisting of oil and water. 

In the aqueous biphasic hydroformylation reaction, the site of the reaction has been much discussed (and contested) and is dependent on reaction conditions (temperature, partial pressure of gas, stirring, use of additives) and reaction partners (type of alkene) [35, 36]. It has been suggested that the positive effects of cosolvents indicate that the bulk of the aqueous liquid phase is the reaction site. By contrast, the addition of surfactants or other surface- or micelle-active compounds accelerates the reaction, which apparently indicates that the reaction occurs at the interfacial layer. 

It is not unusual for the full chemical potential of a reaction to be diminished by slower transport processes (i.e., to be transport limited). In fast liquid phase enzyme reactions, mechanical stirring rates can have a strong influence on the observed kinetics that may be limited by the rate of contacting of the reactants and enzymes. Most heterogeneous catalytic reactions take... 

In the second class, the particles are suspended in the liquid phase. Momentum may be transferred to the particles in different ways, and it is possible to distinguish between bubble-column slurry reactors (in which particles are suspended by bubble movement), stirred-slurry reactors (in which particles are suspended by bubble movement and mechanical stirring), and gas-liquid fluidized reactors (in which particles are suspended by bubble movement and cocurrent liquid flow). 

In stirred-slurry reactors, momentum is transferred to the liquid phase by mechanical stirring as well as by the movement of gas bubbles. Small particles are used in most cases, and the operation is usually carried out in tank reactors with low height-to-diameter ratios. The operation is in widespread use for processes involving liquid reactants, either batchwise or continuous— for example, for the batchwise hydrogenation of fats as referred to in Section II. 

The typical bioreactor is a two-phase stirred tank. It is a three-phase stirred tank if the cells are counted as a separate phase, but they are usually lumped with the aqueous phase that contains the microbes, dissolved nutrients, and soluble products. The gas phase supplies oxygen and removes by-product CO2. The most common operating mode is batch with respect to biomass, batch or fed-batch with respect to nutrients, and fed-batch with respect to oxygen. Reactor aeration is discussed in Chapter 11. This present section concentrates on reaction models for the liquid phase. 

A mass balance for an arbitrary liquid-phase component in the stirred tank reactor is thus written as follows dci... 

There is an additional point to be made about this type of processing. Many gas-phase processes are carried out in a continuous-flow manner on the macro scale, as industrial or laboratory-scale processes. Hence already the conventional processes resemble the flow sheets of micro-reactor processing, i.e. there is similarity between macro and micro processing. This is a fimdamental difference from most liquid-phase reactions that are performed typically batch-wise, e.g. using stirred glass vessels in the laboratory or stirred steel tanks in industrial pilot or production plants. 

The catalysts most frequently used are based on noble metals (mainly palladium and platinum) on various supports, or on nickel catalysts (mainly Raney type). Hydrogenations are generally performed in the liquid phase, under relatively mild conditions of temperature and pressure (1—40 bar). Most processes are performed batch-wise using powder catalysts in stirred tank or loop-type reactors with sizes up to 10 m . 

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