Velocity bubble flow

Moderate viscosity (> 100 cP) will tend to keep the flow regime bubbly at higher superficial velocities. Bubbly flow at high superficial velocity approximates to homogeneous flow. At very high viscosity, completely different flow regimes may occur (see 10.2.3). 

CARPT experiments in bubble columns demonstrate that while liquid circulation is always present in the column, stationary multiple recirculation cells with height equal to column diameter are not. Secondary cells may exist in the entry region at lower gas velocities (bubble flow regime). The entry length is then limited to 1 to 1.5 column diameter. Similar recirculation profiles were reported for solids in fluidized beds (Moslemian etal., 1989). 

For group B and D particles, nearly all the excess gas velocity (U — U,nj) flows as bubbles tnrough the bed. The flow of bubbles controls particle mixing, attrition, and elutriation. Therefore, ehitriation and attrition rates are proportional to excess gas velocity. Readers should refer to Sec. 17 for important information and correlations on Gel-dart s powder classification, minimum fluidization velocity, bubble growth and bed expansion, and elutriation. 

Cocurrent three-phase fluidization is commonly referred to as gas-liquid fluidization. Bubble flow, whether coeurrent or countereurrent, is eonveniently subdivided into two modes mainly liquid-supported solids, in which the liquid exeeeds the minimum liquid-fluidization veloeity, and bubble-supported solids, in whieh the liquid is below its minimum fluidization velocity or even stationary and serves mainly to transmit to the solids the momentum and potential energy of the gas bubbles, thus suspending the solids. 

Zabor et al. (Zl) have described studies of the catalytic hydration of propylene under such conditions (temperature 279°C, pressure 3675 psig) that both liquid and vapor phases are present in the packed catalyst bed. Conversions are reported for cocurrent upflow and cocurrent downflow, it being assumed in that paper that the former mode corresponds to bubble flow and the latter to trickle-flow conditions. Trickle flow resulted in the higher conversions, and conversion was influenced by changes in bed height (for unchanged space velocity), in contrast to the case for bubble-flow operation. The differences are assumed to be effects of mass transfer or liquid distribution. 

Some data on gas holdup are also reported by Stemerding (SI6). Hoogendoorn and Lips (H10) have reported gas-holdup data for counter-current bubble flow in the experimental system described in Section V,A,4. Gas holdup was not influenced by changes of liquid flow rate, but increased with nominal gas velocity in the range from 0.03 ft/sec to 0.3 ft/sec. The results are somewhat lower than those obtained by Weber, the difference being explained as due to the difference in gas distributor. Weber used a porous plate and Hoogendoorn and Lips a set of parallel nozzles. 

A theory of two-phase bubble flow has been developed by Nicklin (N1), who shows that the motion of bubbles arises partly from buoyancy and partly from the nominal velocity caused by the entry of the two phases into the tube. Theoretical and semiempirical studies of bubble flow have also been presented by Azbel (A2) and by Azizyan and Smirnov (A3), and further experimental data on holdup have been recently reported by Yoshida and Akita (Yl), by Braulick et al. (B9) and by Towell et al. (T3). 

Gal-Or and Resnick (Gl) have developed a simplified theoretical model for the calculation of mass-transfer rates for a sparingly soluble gas in an agtitated gas-liquid contactor. The model is based on the average gas residencetime, and its use requires, among other things, knowledge of bubble diameter. In a related study (G2) a photographic technique for the determination of bubble flow patterns and of the relative velocity between bubbles and liquid is described. 

The flow patterns of agitated liquid have been studied extensively (Al, B11, F6, K5, M6, N2, R12, V5), usually by photographic methods. Apparently no work has been reported on bubble-flow patterns and relative velocities in agitated gas-liquid dispersions. Some simple pictures have been presented that only show the same details that may be seen with the unaided eye (Bll, F6, Y4). 

Dispersed bubbly flow (DB) is usually characterized by the presence of discrete gas bubbles in the continuous liquid phase. As indicated in Fig. 5.2, for the channel of db = 2.886 mm, dispersed bubbles appeared at a low gas superficial velocity but a very high liquid superficial velocity. It is known that for large circular mbes dispersed bubbles usually take a sphere-like shape. For the triangular channel of dh = 2.886 mm, however, it is observed from Fig. 5.2 that the discrete bubbles in the liquid phase were of irregular shapes. The deformation of the gas bubbles was caused by rather high liquid velocities in the channel. 

Figure 3.19 Configuration of bubble layer as affected by flow rate at high subcooling (Freon-118) (a) low-velocity boiling flow (b) high-velocity boiling flow. (From Tong et al., 1966b. Copyright 5 1966 by American Society of Mechanical Engineers, New York. Reprinted with permission.)...

The velocity profile effect—For a convex flow velocity profile, for instance, in a steady bubbly flow, the centerline velocity is higher than the average velocity. With bubbles usually concentrating at the center, they attain a higher velocity than the liquid. 

For a single-phase turbulent flow the ratio of the maximum to the average flow velocity is approximately 1.2, and the value of Co may also be close to 1.2 for a bubbly flow. Zuber and Findlay (1965) pointed out that, as the mixture velocity increases, the value of the exponent increases and flatter profiles result. 

Measurement of the velocity of a large particle. The investigation of the turbulence characteristics in the liquid phase of a bubbly flow has generated detailed studies on the use of thermal anemometry and optical anemometry in gas-liquid two-phase flows. These techniques have been proved to be accurate and reliable for the measurement of the instantaneous liquid velocity in bubble flow. However, the velocity of the gas bubbles—or, more precisely, the speed of displacement of the gas-liquid interfaces—is still an active research area. Three techniques that have been proposed to achieve such measurement were reviewed by Delhaye (1986), as discussed in the following paragraphs. 

A steady homogeneous model is often used for bubbly flow. As mentioned previously, the two phases are assumed to have the same velocity and a homogeneous mixture to possess average properties. The basic equations for a steady one-dimensional flow are as follows. 

In a vertical upward gas-liquid flow, a continuous swarm of bubbles flows upward with the liquid stream due to a buoyancy effect, and the gas slips past the liquid with a relative velocity U0 (rise velocity) ... 

Bubbly flow in horizontal pipes. High-velocity flow in horizontal pipes presents a minimum effect of the gravitational field and reduces one potential pa-... 

The effect of local enthalpy at CHF is due primarily to the wall voidage, which impairs the critical flux, and secondarily to the bulk voidage, which affects the flow pattern. The coupled effects of local subcooling and flow velocity in a subcooled bubbly flow were first reported by Griff el and Bonilla (1965), neglecting the pressure effect ... 

Geldart (1970) showed a substantial distinction between bubble sizes in two dimensional and three dimensional beds. He used 128 pm river sand in a 30.8 cm round bed and a 68 1.27 cm rectangular cross section bed. The results, shown in Fig. 12, show that the bubbles in the three dimensional bed are larger. There were differences in the visible bubble flow rate at the same superficial velocity. Geldart ascribes the differences in bubble diameter to differences invisible bubble flow rate as well as to out-of-line coalescence in the three dimensional bed. 

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