Bed diameter

This equation predicts that the height of a theoretical diffusion stage increases, ie, mass-transfer resistance increases, both with bed height and bed diameter. The diffusion resistance for Group B particles where the maximum stable bubble size and the bed height are critical parameters may also be calculated (21). 

Fig. 17. TDH above vigorously bubbling or turbulent fluidized beds as a function of bed diameter from 0.025 to 7.5 m (27).

A typical catalyst bed is very shallow (10 to 50 mm) (76,77). In some plants the catalyst is contained in numerous small parallel reactors in others, catalyst-bed diameters up to 1.7 and 2.0 m (77,80) and capacities of up to 135,000 t/yr per reactor are reported (78). The silver catalyst has a useful life of three to eight months and can be recovered. It is easily poisoned by traces of transition group metals and by sulfur. 

Increasing bed diameter increases spoutable depth. By employing a bed-orifice diameter ratio of 12 for air spouting, a 9-in-diameter bed was spouted at a depth of 65 in while a 12-in-diameter bed was spouted at 95 in. 

As indicated by Eq. (12-62) the superficial fluid velocity required for spouting increases with bed depth and orifice diameter and decreases as the bed diameter is increased. 

Glaser and Lichtenstein (G3) measured the liquid residence-time distribution for cocurrent downward flow of gas and liquid in columns of -in., 2-in., and 1-ft diameter packed with porous or nonporous -pg-in. or -in. cylindrical packings. The fluid media were an aqueous calcium chloride solution and air in one series of experiments and kerosene and hydrogen in another. Pulses of radioactive tracer (carbon-12, phosphorous-32, or rubi-dium-86) were injected outside the column, and the effluent concentration measured by Geiger counter. Axial dispersion was characterized by variability (defined as the standard deviation of residence time divided by the average residence time), and corrections for end effects were included in the analysis. The experiments indicate no effect of bed diameter upon variability. For a packed bed of porous particles, variability was found to consist of three components (1) Variability due to bulk flow through the bed... 

Wall-to-bed heat-transfer coefficients were also measured by Viswanathan et al. (V6). The bed diameter was 2 in. and the media used were air, water, and quartz particles of 0.649- and 0.928-mm mean diameter. All experiments were carried out with constant bed height, whereas the amount of solid particles as well as the gas and liquid flow rates were varied. The results are presented in that paper as plots of heat-transfer coefficient versus the ratio between mass flow rate of gas and mass flow rate of liquid. The heat-transfer coefficient increased sharply to a maximum value, which was reached for relatively low gas-liquid ratios, and further increase of the ratio led to a reduction of the heat-transfer coefficient. It was also observed that the maximum value of the heat-transfer coefficient depends on the amount of solid particles in the column. Thus, for 0.928-mm particles, the maximum value of the heat-transfer coefficient obtained in experiments with 750-gm solids was approximately 40% higher than those obtained in experiments with 250- and 1250-gm solids. 

Also, Schier s experiments revealed the necessity to consider the directional influence of the walls of the filter bed on the bed porosity e. This effect is very important especially for small sized beds that are usually used in laboratories for investigations. For beds made of spherically shaped collectors several correlations exist describing the e(y/dc) function where y denotes the distance from the wall in the radial direction. However, for relative large bed diameters DB/dc ranging from 5 to 25 it proved to be sufficient to use an averaged e in Eq. (3.2.5), as proposed by Jeschar [6],... 

Several pilot plants have been built to test periodic flow direction reversal. Pilot-scale reactors with bed diameters from 1.6 to 2.8 m were operated with flow reversal for several years. The units, described by Bunimovich et al. (1984,1990) and Matros and Bunimovich (1996), handled 600 to 3000 m3/h and operated with cycle periods of 15 to 20 min. Table VIII shows the performance of these plants for different feeds and potassium oxide promoted vanadia catalysts. The SVD catalyst was granular the IK-1-4 was in the form of 5 (i.d.) x 10-mm cylinders, while the SYS catalyst was... 

In this section, representative results are reviewed to provide a prospective of reactor modeling techniques which deal with bed size. There probably is additional unpublished proprietary material in this area. Early studies of fluidized reactors recognized the influence of bed diameter on conversion due to less efficient gas-solid contacting. Experimental studies were used to predict reactor performance. Frye et al. (1958) used... 

Figure 5. Mass transfer unit for ozone conversion for different bed diameters.

Bauer et al. (1981) measured the influence of bed diameter on the catalytic decomposition of ozone. Figure 6 shows the decrease of the conversion with bed diameter for Bauer s data. This figure also shows the influence of distributor design on conversion. In many small scale experiments, a porous plate is used which will give better performance than the distributors used in large shallow bed commercial designs. 

Figure 6. Conversion catalytic decomposition of ozone for different bed diameters and distributors. (From Werther, 1992.)...

Figure 7. Bed expansion as a function of bed diameter at a fluidization velocity of 0.20 m/sec. (From DeGroot, 1967.)...

Yerushalmi and Avidan (1985) suggest that the axial dispersion coefficient of solids in slugging and turbulent flow varies approximately linearly with the bed diameter, similar to Thiel and Potter (1978). The data are shown in Fig. 17 although May s results are probably in the bubbling fluidization regime rather than turbulent flow. 

Figure 17. The effect of bed diameter on solid mixing. (From Yerushalmi Avidan, 1985.)...

The thickness of the downflowing layers at the wall of the CFB is typically defined as the distance from the wall to the position of zero vertical solid flux. Measurements of the layer thickness were made on a 12 MW and 165 MW CFB boiler by Zhang, Johnsson and Leckner (1995). They found that the thickness increased for the larger bed. They related data from many different beds (Fig. 19), with the equivalent bed diameter, taken as the hydraulic diameter, using the following form... 

Figure 19. Empirical correlation and experimental data of thickness of downflowing layer at the wall of a CFB as a function of the equivalent bed diameter. (From Zhang etal., 1995.)...

Patience et al. (1992) developed a dimensionless correlation for the mean slip factor between the gas and solid by using solid suspension data from various small laboratory beds. The proposed correlation relates the slip to the Froude number based on the bed diameter. It remains to be seen if the correlation will hold at Froude numbers typical of large beds and if other dimensionless factors are important. 

Figure 20. Use of scale models with different bed diameters to simulate the influence of diameter on the hydrodynamics of a hot commercial reactor.

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