Gas bubble diameter

The gas hold-up in the liquid is of interest for the correct dimensioning of the height of the gas space of a stirred tank, or a bubble column. This quantity is more important, however, in connection with the average gas bubble diameter db> since it enables the interfacial area per unit volume a to be determined db = (o/ /a-, see equation (4.51). 

Tab. 4.3 Dependence of sorption number Y upon terminal gas bubble diameter d.

Specify the arrangement to be used and estimate the average gas-bubble diameter, gas holdup, specific interfacial area, and mass-transfer coefficient to be expected. 

Microencapsulation is the coating of small solid particles, liquid droplets, or gas bubbles with a thin film of coating or shell material. In this article, the term microcapsule is used to describe particles with diameters between 1 and 1000 p.m. Particles smaller than 1 p.m are called nanoparticles particles greater than 1000 p.m can be called microgranules or macrocapsules. 

These design fundamentals result in the requirement that space velocity, effective space—time, fraction of bubble gas exchanged with the emulsion gas, bubble residence time, bed expansion relative to settled bed height, and length-to-diameter ratio be held constant. Effective space—time, the product of bubble residence time and fraction of bubble gas exchanged, accounts for the reduction in gas residence time because of the rapid ascent of bubbles, and thereby for the lower conversions compared with a fixed bed with equal gas flow rates and catalyst weights. 

Cf, C y, and Cq are the concentrations of the substance in question (which may be a colligend or a surfactant) in the feed stream, bottoms stream, and foamate (collapsed foam) respectively. G, F, and Q are the volumetric flow rates of gas, feed, and foamate respectively, is the surface excess in equilibrium with C y. S is the surface-to-volume ratio for a bubble. For a spherical bubble, S = 6/d, where d is the bubble diameter. For variation in bubble sizes, d should be taken as YLnid fLnidj, where n is the number of bubbles with diameter dj in a representative region of foam. 

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. 

This form is partieularly appropriate when the gas is of low solubility in the liquid and "liquid film resistanee" eontrols the rate of transfer. More eomplex forms whieh use an overall mass transfer eoeffieient whieh ineludes the effeets of gas film resistanee must be used otherwise. Also, if ehemieal reaetions are involved, they are not rate limiting. The approaeh given here, however, illustrates the required ealeulation steps. The nature of the mixing or agitation primarily affeets the interfaeial area per unit volume, a. The liquid phase mass transfer eoeffieient, kL, is primarily a funetion of the physieal properties of the fluid. The interfaeial area is determined by the size of the gas bubbles formed and how long they remain in the mixing vessel. The size of the bubbles is normally expressed in terms of their Sauter mean diameter, dj, whieh is defined below. How long the bubbles remain is expressed in terms of gas hold-up, H, the fraetion of the total fluid volume (gas plus liquid) whieh is oeeupied by gas bubbles. 

The flow pattern of gas within the emulsion phase surrounding a bubble depends on whether the bubble velocity Ug is less than or greater than minimum fluidization velocity U . For Ubflow lines. For Ub> U , the much different case of Figure 4(B) results. Here a gas element which leaves the bubble eap rises much more slowly than the bubble, and as the bubble passes, it remms to the base of the bubble. Thus, a cloud of captive gas surrounds a bubble as it rises. The ratio of eloud diameter to bubble diameter may be written... 

In a trayed absorber the amine falls from one tray to the one below in the same manner as the liquid in a condensate stabilizer (Chapter 6, Figure 6-4). It flows across the tray and over a weir before flowing into the next downcomer. The gas bubbles up through the liquid and creates a froth that must be separated from the gas before it reaches the underside of the next tray. For preliminary design, a tray spacing of 24 in. and a minimum diameter capable of separating 150 to 200 micron droplets (using the equations developed in Volume 1 for gas capacity of a vertical separator) can be assumed. The size of packed towers must be obtained from manufacturer s published literature. 

In bubble-column slurry reactors, momentum is transferred to the liquid phase by the movement of gas bubbles. The liquid medium is stationary in most cases. Finely divided solids with particle diameters of the order of 0.01 mm are used. The operation is usually carried out in columns with high height-to-diameter ratios. The operation may be employed for batchwise conversion of a liquid reactant, or for continuous reaction between gaseous reactants. 

Houghton et al. (HI3) have reported data on the size, number, and size-distribution of bubbles. Distinction is made between bubble beds, in which bubble diameter and gas holdup tend to become constant as the gas velocity is increased (these observations being in agreement with those of other workers previously referred to), and foam beds, in which bubble diameter increases and bubble number per unit volume decreases for increasing gas velocity. Pore characteristics of the gas distributor affect the properties of foam beds, but not of bubble beds. Whether a bubble bed or a foam bed is formed depends on the properties of the liquid, in particular on the stability of bubbles at the liquid surface, foam beds being more likely to form in solutions than in pure liquids. 

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