Fluidization particle diameter

Most industrial fluidized beds using fine powder are found to be in the category of turbulent beds typical of these is fluidized catalytic cracking. On the other hand, most noncatalytic reactions and physical operations use coarse particles and are operated as bubbling beds. Squires (S14) has pointed out that the state of fluidization varies significantly with fluidized particle diameter, and has identified the following two categories ... 

Stirred tank paddles power input suspend solids, 0.2 to 1.6 kW/m UD = 0.7 to 1.05/1. Baffle, four 90° baffle width = 0.08 x tank diameter off-the-wall distance = 0.015 x tank diameter. Minimum level of liquid = 0.15 x tank diameter for impeller tank diameter 0.28 1 and minimum level = 0.25 x tank diameter for impeller tank diameter = 0.4 1. Use a foot bearing plus a single, main axial hydrofoil impeller diameter = 0.28 x tank diameter located 0.2 x tank diameter from the bottom plus a pitched blade impeller diameter = 0.19 x tank diameter located 0.5 x tank diameter from the bottom. Liquid fluidized bed in general, particle diameter 0.5 to 5 mm with density and diameter of the particle dependent on the application. The superficial liquid velocity to fluidize the bed depends on both the diameter and the density difference between the liquid and the particle. Usually, the operation is particulate fluidization. Particle diameter 0.2 to 1 mm reactors superficial liquid velocity 2 to 200 mm/s. Fluidized adsorption bed expands 20 to 30% superficial liquid velocity for usual carbon adsorbent = 8 to 14 mm/s. Fluidized ion exchange bed expands 50 to 200% superficial liquid velocity for usual ion exchange resin = 40 mm/s. Backwash operations fixed-bed adsorption superficial liquid velocity = 8 to 14 mm/s fixed-bed ion exchange superficial backwash velocity = 3 mm/s. 

Particle Size. The soHds in a fluidized bed are never identical in size and foUow a particle size distribution. An average particle diameter, is generally used for design. It is necessary to give relatively more emphasis to the low end of the particle size distribution (fines), which is done by using the surface mean diameter, to calculate an average particle size ... 

This reaction is carried out in tall fluidized beds of high L/dt ratio. Pressures up to 200 kPa are used at temperatures around 300°C. The copper catalyst is deposited onto the surface of the silicon metal particles. The product is a vapor-phase material and the particulate silicon is gradually consumed. As the particle diameter decreases the minimum fluidization velocity decreases also. While the linear velocity decreases, the mass velocity of the fluid increases with conversion. Therefore, the leftover small particles with the copper catalyst and some debris leave the reactor at the top exit. 

Particle diameter is a primary variable important to many chemical engineering calculations, including settling, slurry flow, fluidized beds, packed reactors, and packed distillation towers. Unfortunately, this dimension is usually difficult or impossible to measure, because the particles are small or irregular. Consequently, chemical engineers have become familiar with the notion of equivalent diameter of a partiele, which is the diameter of a sphere that has a volume equal to that of the particle. 

As can be seen, two factors are particularly critical (a) the density of the particle, since heavier particles are more difficult to fluidize, and (b) particle size, since the necessary gas velocity varies as the square of the particle diameter. The design of the reactor is also important since gas velocity at the top must be less than the terminal velocity of the particles, otherwise they would be blown out of the bed.P l... 

The value for is conservatively interpreted as the particle diameter. This is a perfectly feasible size for use in a laboratory reactor. Due to pressure-drop limitations, it is too small for a full-scale packed bed. However, even smaller catalyst particles, dp 50 yum, are used in fluidized-bed reactors. For such small particles we can assume rj=l, even for the 3-nm pore diameters found in some cracking catalysts. 

The catalyst was prepared by impregnating porous alumina particles with a solution of nickel and lanthanum nitrates. The metal loading was 20 w1% for nickel and 10 wt% for lanthanum oxide. The catalyst particles were A group particles [8], whereas they were not classified as the AA oup [9]. The average particle diameter was 120 pm, and the bed density was 1.09 kg m . The minimum fluidization velocity was 9.6 mm s. The settled bed height was around 400 mm. The superficial gas velocity was 40-60 mm s. The reaction rate was controlled by changing the reaction temperature. 

In general, u,/ is a function of the particle diameter and gas properties, as well as 0 and emf. Once the fluidizing gas and the length of scale of the model is chosen, the proper particle diameter is that which gives the value of umf needed in Eq. (82). 

This correlation predicts that the maximum vertical jet penetration into a fluidized bed varies with gas density to the 0.67 power, and decreases with increasing fluidizing gas velocity and increasing particle diameter. 

The second approach assigns thermal resistance to a gaseous boundary layer at the heat transfer surface. The enhancement of heat transfer found in fluidized beds is then attributed to the scouring action of solid particles on the gas film, decreasing the effective film thickness. The early works of Leva et al. (1949), Dow and Jacob (1951), and Levenspiel and Walton (1954) utilized this approach. Models following this approach generally attempt to correlate a heat transfer Nusselt number in terms of the fluid Prandtl number and a modified Reynolds number with either the particle diameter or the tube diameter as the characteristic length scale. Examples are ... 

As noted earlier, increasing gas velocity for any given fluidized bed beyond the terminal velocity of bed particles leads to upward entrainment of particles out of the bed. To maintain solid concentration in the fluidized bed, an equal flux of solid particles must be injected at the bottom of the bed as makeup. Operation in this regime, with balanced injection of particles into the bed and entrainment of particles out of the bed, may be termed fast fluidization, FFB. Figure 10 presents an approximate map of this fast fluidization regime, in terms of a dimensionless gas velocity and dimensionless particle diameter. 

Figure 18. Parametric effects of solid flux and particle diameter on heat transfer in fast fluidized beds. (From Furchi et al, 1988).

Another parametric effect is the apparent dependence of the heat transfer coefficient on the physical size of the heat transfer surface. Figure 24, from Burki et al. (1993), graphically illustrates this parametric effect by showing that the effective heat transfer coefficient can vary by several hundred percent with different vertical lengths of the heat transfer surface, for circulating fluidized beds of approximately the same particle diameter and suspension density. This size effect significantly contributed to confusion in the technical community since experimental measurements by inves-... 

The median particle diameter, dm, seems to provide a better indication of fluidization performance, as described by Geldart... 

Gas-particle flows in fluidized beds and riser reactors are inherently unstable and they manifest inhomogeneous structures over a wide range of length and time scales. There is a substantial body of literature where researchers have sought to capture these fluctuations through numerical simulation of microscopic TFM equations, and it is now clear that TFMs for such flows do reveal unstable modes whose length scale is as small as ten particle diameters (e.g., see Agrawal et al., 2001 Andrews et al., 2005). 

Catalysis by solids depends on the amount of surface exposed to the fluid. Large specific surface is obtained with small particles, but primarily with highly porous structures. For instance, to achieve 1 m2/cc the diameter of a sphere must be reduced to 6(10-4) cm, but porous catalysts may have several hundred m2/cc. Practical limitations exist to the smallness of particles that can be used, such as pressure drop and entrainment. In fixed or moving beds, particle diameters are several millimeters, in fluidized beds they may be less than 0.1 mm. 

For non-spherical particles, values of sphericity lie in the range 0 < < 1. Thus, the effective particle diameter for fluidization purposes is the product of the surface-volume mean diameter and the sphericity (Kunii and Levenspiel, 1991). The sphericity of regular-shaped particles can be deduced by geometry whilst the sphericity of irregular-shaped... 

This equation allows the prediction of minimum fluidizing velocity from a knowledge of the mean particle diameter, the particle density, the density of fluidizing medium and fhe viscosity of fluidizing medium (SI units). Couderc (1985) quotes data which show that the inaccuracy of Leva s equation increases significantly outside the range 2 < Re < 30. 

Figure 3.5 Variation in gas-particle heat transfer coefficient with particle diameter (A = fluidized bed at maximum air velocity B = fluidized bed at Fr = 120 C = flat surface velocity as for curve B). Adapted from Persson, ASHRAE Journal, June 1967. American Society of Heating, Refrigerating and Air-Conditioning...

According to Ruthven (1984), the linear velocity should not be higher than a certain maximum value, so that extended friction between the packing material is avoided in both down-and upflow operations. This maximum velocity is 0.8 times the minimum fluidization velocity for upflow operation and 1.8 times the same velocity for downflow operation. The minimum fluidization velocity can be estimated using the equations presented in Section 3.8.2. In Figure 6.1, the maximum linear velocityversus the particle diameter is presented. 

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