Fluidized catalyst beds velocity

It is interesting to see how far modern technology for fluidized catalyst beds has served to achieve good fluidization. Our criterion of good fluidization is a gas-fluidized bed of a low viscosity liquid (such as water), where the low-viscosity liquid would set the lower limit to the fluidity of the emulsion. Such a gas-fluidized liquid bed is the well-known bubble column, which has been studied extensively. Our objective is to understand the behavior in the recirculation flow regime, since the superficial gas velocity of practical interest is usually more than 30 cm/sec for fluidized catalyst beds and for these conditions intense recirculation of the emulsion has been observed (note Fig. 2 and Section II,D,3). 

The mean bubble size in a fluidized bed has been discussed in Section II,B. As discussed, for a fluidized catalyst bed of good fluidity may be taken as approximately 5.0 cm [cf. Figs. 10 and 11, and Eq. (2-11)] for Uc, > 10 cm/sec. With Eq. (3-33), this (I32 gives = 49.5 cm/sec, which is shown in Fig. 34 as a dashed line. It is interesting that the mean slip velocity is essentially the same as for a bubble column, when Uq > 20 cm/sec. As noted in Section II,B, and Mg are very sensitive to change in particle size, size distribution, shape, and density. 

Most data available from past studies are summarized in Figs. 40 and 41 for the longitudinal dispersion coefficient of the emulsion phase in fluidized catalyst beds of good fluidity. Such coefficients were first measured in the pioneer work of Bart (1950, B20) for a wide range of gas velocity (Uq = 7.5-90 cm/sec). His values with a 3.2-cm-diam column are approximately one-half those by Morooka et al. (M40) in the higher flow rate region. The basis for their correlation follows (M27). 

When the velocity profile of the emulsion phase is similar to that of the liquid phase in a bubble column, Eq. (4-11) will apply to the fluidized catalyst bed. This similarity seems to be well justified as mentioned in Sections III,A,4-5, although there is no direct calculation of the turbulent kinematic viscosity from the measurement of velocity profile in the fluidized catalyst bed. 

The turbulent kinematic viscosity vt of the fluidized catalyst bed has been determined, as Eq. (3-3 la), from the use of axial dispersion coefficient This is a natural consequence of the analogy between the bubble column and the fluidized catalyst bed of good fluidity (such as in fluidized catalytic cracking). The mean gas holdup (Fig. 36) and the mean bubble velocity along the bed axis (Fig. 37) are reasonably well predicted by applying Eq. (3-3 la) for the fluidized cracking catalyst bed. 

III,D,5). In this chapter, the equation is further examined in relation to bed performance, since the turbulence properties of the bed result from interaction between bubbles and the continuous phase. As shown in Fig. 34, the mean slip velocity of bubbles in a fluidized catalyst bed of good fluidity is essentially the same as that for a bubble column when Uq > 20 cm/sec. A criterion under which bubble size approaches a dynamic equilibrium is obviously needed for predicting or evaluating the performance of scaled-up beds. 

Taking Uq - instead of U, Eq. (5-3) is equally applicable to the fluidized bed, where is the superficial velocity at incipient fluidization and is usually negligible in comparison with for the fluidized catalyst bed. 

The mean bubble size that concerns us here is on the order of 5 cm, so that the Eotvos number Eo (equal to dlgpi/a) is well over 40 for usual bubble-column liquids. The bubbles are of spherical-cap type under this condition, which is essentially equivalent to a Weber-number criterion We (equal to dy piul/a) > 20, since Ug = Vgrfb/2 (H4, H5). The bubbles in a fluidized catalyst bed satisfy the above criterion, since a- 0. Consequently, surface tension has relatively little effect, and instead the splitting is closely related to disturbances induced by the bulk turbulence, the intensity and the scale of which are mainly governed by the fluidity of the continuous phase and the operating gas velocity. 

A reactor model is developed to include reaction taking place in the dilute phase, and to be reasonably consistent with the known flow properties of fluidized catalyst beds operated under relatively high gas velocity. According to this model, reaction proceeds successively in the dense phase and in the dilute phase. 

In packed bed reactors the solid catalyst is held stationary by plates at the top and bottom of the bed. In contrast, in fluidized bed reactors, the catalyst bed is relatively loosely packed, and there is no plate at the top. Rapid fluid flow from the bottom raises the bed and ensures good mixing, leading to insignificant temperature or concentration gradients. However, due to high fluid velocity some catalyst carryover is common. 

Catalytic butane dehydrogenation can be successfully carried out in a laboratory scale fluidized bed reactor operating at 310 °C and at atmospheric pressure. The catalytic particles have diameter 310 pm and density 2060 kg/m. Such a reactor is 150 mm in diameter and has a fixed 500 mm long catalytic bed. When the catalyst bed is fluidized with butane blown at a velocity of 0.1 m/s, it becomes 750 mm thick. 

However, high gas velocity, may lead to problems in bed operation. It increases the entrainment loss of fluidizing catalyst particles. It also may give rise to excessive reaction in the particle-disengaging space, which sometimes will lead to reactor instability or to decreased selectivity (as discussed in later sections). Attrition or erosion of the reactor is more likely at higher gas velocity. Thus, there is an optimum gas velocity for fluid beds, which for most catalysts is in the range of 20-80 cm/sec, usually 40-60 cm/sec. 

As stated in Section II,D,3, the pioneer work of Lewis et al. (Lll) showed intense recirculation of the emulsion phase in an air-fluidized MS-catalyst bed. The time-averaged velocity profile in the recirculation-flow regime is still lacking, although Yamazaki (Y3a) has revealed the profile at a low gas velocity ingeniously by utilizing thermal response. The radial position < >, where the mean velocity is zero relative to the bed wall, measured by Morooka (M43) for a FCC bed of 7.9-cm diameter, is shown in Fig. 29. The deviation from prediction by Eq. (3-20) is about 0.3 cm and is comparable with the resolution power of the strain-gauge probe. 

The averaged volume fraction Cb. calculated by Eq. (3-25), is shown in Fig. 36, for bubble columns and also for FCC-catalyst beds (M40). The mean slip velocity of bubbles is again taken as ii , = 49.5 cm/sec. Also, Vi is calculated by Eq. (3-31a) for curve FQF and by Eq. (3-31) for curve FQP, while curve EE is an empirical fit of the data. As in the case of a bubble column, curve FQP matches better with curve EE, although FQP is consistently higher. Curve EE tends to decrease the slope for Uq s 7-8 cm/sec, perhaps due to the decrease in Ms the scatter of data makes the behavior unclear. This is explained by the difference in the region of Uc < 20 cm/sec. The bubble column shows higher 6b values than those for the fluidized bed, which is due to the bubble column s lower Us (cf. Fig. 34). 

Minimum fluidization velocities for spherical particles in air are shown in Figure 9.2. Equation (9.2) applies for particles up to about 300 microns in size, which includes most fluidized catalysts. For fluidized-bed combustion or metallurgical processes, the particles are much larger, and Eq. (9.1) must be used. For very large sizes, the laminar-flow term in Eq. (9.1) becomes unimportant, and varies with the square root of dp ... 

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