Penberthy Houdaille

*Data by courtesy of Norton Company, Worcester, Mass. To convert inches to centimeters, multiply by 2.54; to convert feet per minute to meters per second, multiply by 0.0051.

*Data by courtesy of Norton Company, Worcester, Mass. To convert inches to centimeters, multiply by 2.54; to convert feet per minute to meters per second, multiply by 0.0051.

FIG. 14-97 Installation and dimensions of a tube stirrer: h/d = 1; H/D ~ 1; D/8 = 10; A = 1.5 dw, D/dN = 10; d/dN = 3; d/dri = 7.5; d/dra = 6. [Zlokarnik, Ullman's Encyclopedia of Industrial Chemistry, Sec. 25, VCH, Wein-heim, Germany, 1988.]

guidelines are given as to what equipment might be feasible and what equipment might prove most economical.

For producing foam for foam-separation processes, perforated-plate or porous-plate spargers are normally used. Mechanical agitators are often not effective in the light foams needed in foam fractionation. Dissolved-air flotation, based on the release of a pressurized flow in which oxygen was dissolved, has been shown to be effective for particulate removal when sparged air failed because the bubbles formed upon precipitation are smaller—down to 80 |m—than bubbles possible with sparging, typically 1000 |m [Grieves and Ettelt, AIChE J., 13,1167 (1967)]. Mechanically agitated surface aerators such as the Wemco-Fagergren flotation unit (Fig. 14-102) are used extensively for ore flotation.

To produce foam in batch processes, mechanical agitators are used almost exclusively. The gas can either be introduced through the free surface by the entraining action of the impeller or alternatively sparged beneath the impeller. In such batch operation, the liquid level gradually rises as the foam is generated; thus, squatly impellers such as turbines are rapidly covered with foam and must almost always be sparged from below. Tall impellers such as wire whips (Fig. 14-103) are especially well suited to entrain gas from the vapor space. Inter-meshing wire whips are standard kitchen utensils for producing foamed meringues, consisting of air, vegetable oil, and egg whites. For a new application, generally some experimentation with different impellers is necessary in order to get the desired fine final bubble size without getting frothing over initially. For producing foams continually, an aspirating venturi nozzle and restrictions in pipes such as baffles and metal gauzes are generally most economical.

For gas absorption, the equipment possibilities are generally packed columns; plate distillation towers, possibly with mechanical agitation on every plate; deep-bed contactors (bubble columns or sparged lagoons); and mechanically agitated vessels or lagoons. Packed towers and plate distillation columns are discussed elsewhere. Generally these devices are used when a relatively large number of stages (more than two or three) is required to achieve the desired result practically.

The volumetric mass-transfer coefficients and heights of transfer units for bubble columns and packed towers have been compared for absorption of CO2 into water by Houghton et al. [Chem. Eng. Sci., 7, 26 (1957)]. The bubble column will tolerate much higher vapor velocities, and in the overlapping region (superficial gas velocities of 0.9 to 1.8 cm/s), the bubble column has about three times higher masstransfer coefficient and about 3 times greater height of transfer unit. The liquid in a bubble column is, for practical purposes, quite well mixed; thus, chemical reactions and component separations requiring significant plug flow of the liquid cannot be effected with bubble columns. Bubble columns and agitated vessels are the ideal equipment for processes in which the fraction of gas absorbed need not be great, possibly the gas can be recycled, and the liquid can or should be well mixed. The gas phase in bubble columns is not nearly so well back-mixed as the liquid, and often plug flow of the gas is a logical assumption, but in agitated vessels the gas phase is also well mixed.

The choice of a bubble column or an agitated vessel depends primarily on the solubility of the gas in the liquid, the corrosiveness of the liquid (often a gas compressor can be made of inexpensive material, whereas a mechanical agitator may have to be made of exotic, expensive materials), and the rate of chemical reaction as compared with the mass-transfer rate. Bubble columns and agitated vessels are seldom used for gas absorption except in chemical reactors. As a general rule, if the overall reaction rate is five times greater than the mass-transfer rate in a simple bubble column, a mechanical agitator will be most economical unless the mechanical agitator would have to be made from considerably more expensive material than the gas compressor.

In bubble columns and simply sparged lagoons, selecting the sparger is a very important consideration. In the turbulent regime (superficial gas velocity greater than 4.6 to 6 cm/s), inexpensive

Surface Aerator
FIG. 14-98 The Cyclox surface aerator. (Cleveland Mixer Co.)
Fagergren Flotation Machine
FIG. 14-99 Propeller-type surface aerator. (Ashbrook-Simon-Hartley Corp.)

perforated-pipe spargers should be used. Often the holes must be placed on the pipe bottom in order to make the sparger free-draining during operation. In the quiescent regime, porous septa will often give considerably higher overall mass-transfer coefficients than perforated plates or pipes because of the formation of tiny bubbles that do not coalesce. Chain and coworkers (First International Symposium on Chemical Microbiology, World Health Organization, Monograph Ser. 10, Geneva, 1952) claimed that porous disks are about twice as effective as open-pipe and ring spargers for the air oxidation of sodium sulfite. Eckenfelder [Chem. Eng. Progr., 52(7), 290 (1956)] has compared the oxygen-transfer capabilities ofvarious devices on the basis of the operating power required to absorb a given quantity of O2. The installed cost of the various pieces of equipment probably would not vary sufficiently to warrant being including in an economic analysis. Surface mechanical aerators are not included in this comparison. Of the units compared, it appears that porous tubes give the most efficient power usage. Kalinske (Adv. Biol. Waste Treatment, 1963, p. 157) has compared submerged sparged aerators with mechanical surface aerators. He has summarized this comparison in Water Sewage Works, 33 (January 1968). He indicates that surface aerators are significantly more efficient than subsurface aeration, both for oxygen absorption and for gas-stripping operations.

Zlokarnik and Mann (paper at Mixing Conf., Rindge, New Hampshire, August 1975) have found the opposite of Kalinske, i.e., subsurface diffusers, subsurface sparged turbines, and surface aerators compare approximately as 4:2:1 respectively in terms of O2 transfer efficiency; however, Zlokarnik [Adv. Biochem. Eng., 11, 157 (1979)] later indicates that the scale-up correlation used earlier might be somewhat inaccurate. When all available information is considered, it appears that with near-optimum design any of the aeration systems (diffusers, submerged turbines, or surface impellers) should give a transfer efficiency of at least 2.25 kg O2/kWh. Thus, the final selection should probably be made primarily on the basis of operational reliability, maintenance, and capital costs.

Mass Transfer Mass transfer in plate and packed gas-liquid contactors has been covered earlier in this subsection. Attention here will be limited to deep-bed contactors (bubble columns and agitated vessels). Theory underlying mass transfer between phases is discussed in Sec. 5 of this handbook.

To design deep-bed contactors for mass-transfer operations, one must have, in general, predictive methods for the following design parameters:

• Flooding (for both columns and agitator impellers)

• Agitator power requirements

Aerator Ejector
FIG. 14-100 Aeration ejector. (Penberthy, a division of Houdaille Industries, Inc. )

FIG. 14-101 Impingement aerator.

• Gas-phase and liquid-phase mass-transfer coefficients

• Interfacial area

• Interface resistance

• Mean concentration driving force for mass transfer

In most cases, available methods are incomplete or unreliable, and some supporting experimental work is necessary, followed by scale-up. The methods given here should allow theoretical feasibility studies, help minimize experimentation, and permit a measure of optimization in final design.

Flooding of Agitator Impellers Impeller flooding correlations for six-blade disk (6BD) Rusthon turbines and six-blade disk Smith turbines (Chemineer designation: CD-6) are presented by Bakker, Myers, and Smith [Chem. Eng., 101, 98 (Dec. 1994)] and a review of impeller flooding. The Bakker et al. (loc. cit.) correlation is

where CFL = 30 for a 6BD impeller and CFL = 70 for a concave blade CD-6 impeller and NFr = Froude number = N2D/g; Q = gas flow rate at flooding, m3/s; N = impeller speed, rps; D = impeller diameter, m; and T = tank diameter, m. Note that the CD-6 impeller will handle 70/30 = 2.33 times the gas a 6BD will handle, without flooding, at the

Flotation Cells Wemco
FIG. 14-102 The Wemco-Fagergren flotation machine. [From www. tucottbus.de/BTU/Fak4/Aujbtech/pages/pbrr_N 7 Sep2-(elstat-flotat).pdf].
FIG. 14-103 Wire whip.

same N and D; this is the great advantage of the CD-6 along with lower power decrease as the gas flow rate increases.

Gassed Impeller Power Bakker et al. (op. cit.) have given a gassed power correlation for the 6BD and CD-6 impellers.

where Pg = gassed power, W; Pu = ungassed power, W; NA = Q/ND3; and the constants of Eq. (14-216a) are given in Table 14-25.

As mentioned previously, the CD-6 suffers much less power decrease with increased gassing compared to the 6BD. For example, at Na = 0.15, Pg/Pu = 0.7 for the CD-6 and 0.5 for the CD-6.

The ungassed power is calculated by

where the impeller power numbers Np are given Table 14-25.

Bakker et al. (op. cit.) and Sensel et al. (op. cit.) have given correlations for gas holdup in agitated vessels. The Bakker et al. correlation is e = Ce=(Pg/V)AvB

where Ce = 0.16, A = 0.33, B = 0.67; V = batch volume, m3; vsg = superficial gas velocity = p/[(n/4)T2]; T = tank diameter, m. Equation (14217) applies for both 6BD and CD-6.

Interfacial Area This consideration in agitated vessels has been reviewed and summarized by Tatterson (op. cit.). Predictive methods for interfacial area are not presented here because correlations are given for the overall volumetric mass transfer coefficient liquid phase controlling mass transfer.

TABLE 14-25 Numbers

Constants in Eq. (14-216) and Impeller Power

TABLE 14-25 Numbers

Constants in Eq. (14-216) and Impeller Power

Impeller type

a

b

c

d

Np

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