Gas-Solid-Liquid Mixing Systems

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Process Engineering Guide: GBHE-PEG-MIX-706

Gas-Solid-Liquid Mixing Systems Information contained in this publication or as otherwise supplied to Users is believed to be accurate and correct at time of going to press, and is given in good faith, but it is for the User to satisfy itself of the suitability of the information for its own particular purpose. GBHE gives no warranty as to the fitness of this information for any particular purpose and any implied warranty or condition (statutory or otherwise) is excluded except to the extent that exclusion is prevented by law. GBHE accepts no liability resulting from reliance on this information. Freedom under Patent, Copyright and Designs cannot be assumed.



Process Engineering Guide: Gas-Solid-Liquid Mixing Systems

CONTENTS SECTION 0 INTRODUCTION/PURPOSE 3 1 SCOPE 3 2 FIELD OF APPLICATION 3 3 DEFINITIONS 3 4 SELECTION OF EQUIPMENT 3 5 THREE-PHASE MASS TRANSFER WITH CHEMICAL

REACTION 3 6 STIRRED VESSEL DESIGN 4

6.1 Agitator Design 4 6.2 Design for Solids Suspension 4 6.3 Vessel Design 5 6.4 Gas-Liquid Mass Transfer Coefficient and Surface Area 5



7 THREE-PHASE FLUIDIZED BEDS 6 7.1 Gas and Liquid Hold-Up 7 7.2 Calculation Procedure 8 7.3 Bubble Size 9 7.4 Mass Transfer 9 7.5 Heat Transfer 11 7.6 Elutriation 12 8 SLURRY REACTORS 12 8.1 Gas Rate 12 8.2 Mass Transfer 12 9 NOMENCLATURE 12 10 BIBLIOGRAPHY 14



FIGURES 1 MASS TRANSFER COEFFICIENT vs % v/v SOLIDS

CONCENTRATION 6 2 THREE-PHASE FLUIDISED BED 7 3 PLOT OF EXPERIMENTAL DRIFT FLUX DATA vs GAS

HOLD-UP System : Air-Water Particle sizes : 1 - 6 mm 8 4 RANGES OF BUBBLE SIZES IN WATER-FLUIDIZED

BEDS OF GLASS BEADS 0.5 m FROM THE DISTRIBUTOR 9

5 DEPENDENCE OF SHERWOOD NUMBER ON PECLET

NUMBER 10 6 VOLUMETRIC ABSORPTION COEFFICIENTS IN GAS-

LIQUID FLUIDIZED BEDS OF 1 mm PARTICLES 11 DOCUMENTS REFERRED TO IN THIS PROCESS ENGINEERING GUIDE 16



0 INTRODUCTION/PURPOSE This Guide is one in a series of Mixing Guides and has been produced for GBH Enterprises. 1 SCOPE This Guide covers the selection and design of equipment suitable for the mixing of three phase (gas-solid-liquid) systems. It does not deal with systems where the solid phase is stationary and not itself mixed; thus packed towers and fixed beds are excluded. The main devices covered are agitated vessels, slurry reactors (also known as particle - bubble columns) and fluidized beds. 2 FIELD OF APPLICATION This Guide applies to Process Engineers in GBH Enterprises worldwide. 3 DEFINITIONS No specific definitions apply to this Guide. With the exception of terms used as proper nouns or titles, those terms with initial capital letters which appear in this document and are not defined above are defined in the Glossary of Engineering Terms. 4 SELECTION OF EQUIPMENT The choice of equipment depends primarily on the nature of the system: particle size, shape and density are important parameters as is the requirement to keep solids in suspension in the absence of gas and/or liquid flows. An agitator may also be required in a system which could be run as a slurry reactor if there is a demand to keep the solids suspended when the gas feed is shut off. In fluidized beds the particles are kept in suspension by the liquid flow and this permits particle sizes in a range from approximately 0.5 to 6 mm. Both the gas and liquid phases are back-mixed to some extent in an agitated vessel. In slurry reactors and fluidized beds there is little axial dispersion of the gas although there is generally back-mixing of the liquid. A slurry reactor can be run in either co-current or counter-current flow.



Other types of three-phase contactors include the 'irrigated fluidized bed' and the three-phase spouted bed, investigated by Vukovic et al 1973, 1974 [Ref. 2]. 5 THREE-PHASE MASS TRANSFER WITH CHEMICAL REACTION The problem of gas absorption into a slurry with reaction has been treated theoretically (Uchida and Wen, 1977 [Ref. 3]) by looking at reaction in the liquid phase preceded by dissolution of the gas and solid and with the added simplification that only fast, irreversible reactions and no gas-side resistance to mass transfer were considered. Six different cases could be identified depending on the concentration profiles of the dissolved gas and solid. One particular case has been verified experimentally (Sada et al, 1977 [Ref. 4]) by performing experiments with carbon dioxide and sulfur dioxide in a calcium hydroxide slurry. 6 STIRRED VESSEL DESIGN In spite of the wide application of three-phase stirred vessels, there is little information available on their performance. For reasonably dilute suspensions (< 10% solids) it is probably better to design for a gas-liquid system with a low agitator clearance to favor solids suspension, i.e. by the use of a disc turbine with Z = T/4 (see GBHE-PEG-MIX-705 Figure 11). If full suspension of solids is important then the design should be checked to ensure that solids shall be suspended. This can be done by determining the power requirement to just suspend the solids, PJS, for the un-gassed conditions (see GBHE-PEG-MIX-703) and using an actual P/V greater than PJS. For floating solids, consult GBHE-PEG-MIX-703. There is little relevant company experience in the design of agitators for three-phase systems, although there is substantial experience in the operation of some large (up to 100 m3) three-phase reactors (e.g. KA, Terephthalic Acid, Synprolam and Aniline). These reactors are generally part of licensed processes and they incorporate the contractor's know-how. However, they look quite different from each other and no general conclusion can be drawn from their design.



The two basic requirements are: (a) adequate dispersion of gas to achieve the required gas-liquid mass

transfer, and (b) complete suspension of solid to prevent accumulation, although

homogeneous dispersion of solid may not be required. In general the presence of a gas phase increases the speed and power required to achieve both solid suspension and gas dispersion. 6.1 Agitator Design Chapman et al, 1983 [Ref. 4], however, show that a pitched blade turbine, pumping upwards, appears to be the best way to achieve simultaneous gas dispersion and solid suspension because its specific power input is almost independent of the gas flow and is therefore very stable during operation. The chosen agitator shall provide sufficient power, Pg, to perform the necessary gas-liquid mass transfer duty (see 6.4 and GBHE-PEG-MIX-705) which may favor the disc turbine with its higher Power number (Po u = 5.5 compared with 1.2 for the pitched blade turbine). Downward pumping impellers should not be used when both solids and gas are present because they tend to develop unstable behavior which results in solid sedimentation. 6.2 Design for Solids Suspension The power required to suspend solids in a liquid, PJS, has to be maintained when a gas is added, and if on increasing gas flow, the gassed power, Pg becomes smaller than PJS then the solids will settle. The agitator speed, N, and hence the un-gassed and gassed power, Pu and Pg should be increased to a value such that Pg > PJS to achieve solids suspension. The agitator speed required for complete suspension of particles without gas, NJS, has been determined by Zwietering, 1958 [Ref. 5] for various agitator geometries (see GBHE-PEG-MIX-703).



In general more power (higher agitator speed) is required to re-suspend the solid once the gas has been introduced. However, there is not enough information from large scale vessels to predict scale-up conditions. Experimental work using small vessels of exactly the same configuration and the real fluids and particles (or models with the same density, viscosity and particle size and shape) is recommended. 6.3 Vessel Design The aspect ratio (H/T) is very important for three-phase systems because of the tendency of the gas to disengage and the solids to sink. A single turbine is recommended for a 1:1 aspect ratio. However, for an aspect ratio greater than 1.5:1 a pitched blade turbine located near the bottom of the tank to guarantee complete suspension of the solids and a disc turbine higher up to redistribute the gas and lift the particles further are required. Sometimes the gas is fed in below the upper impeller so that it does not affect the solids suspension characteristics of the bottom impeller. However, the bottom part of the vessel may then contain only a very small amount of gas and be under-utilized for mass transfer. 6.4 Gas-Liquid Mass Transfer Coefficient and Surface Area If the reaction rate in a three-phase stirred system is limited by the rate of dissolution of the gas due to liquid film resistance, then the effect of the solid particles on the liquid film mass transfer coefficient, KL, and the specific surface area, a, shall be taken into account.



FIGURE 1 MASS TRANSFER COEFFICIENT vs % v/v SOLIDS CONCENTRATION

Generalizations from limited experimental data is unwise; however for solid particles in the size range 50 - 250 µm the volumetric mass transfer coefficient, KLa, is apparently hardly affected by the presence of solids as long as the solids concentration does not exceed 20% (v/v), see Figure 1 (Joosten et al, 1977 [Ref. 6]). 7 THREE-PHASE FLUIDIZED BEDS The term 'three-phase fluidized bed' is applied to a system in which gas is passed continuously into the bottom of a liquid fluidized bed. The gas may be introduced with the liquid or via a separate sparger as shown in Figure 2. In such a device the liquid is well-mixed while the gas is almost in plug flow with only a small amount of axial diffusion. The particle sizes used so far have been in the range 0.5 - 6 mm diameter.



If it is necessary to use smaller particles, then a slurry or stirred vessel is preferable. A major advantage of the fluidized bed is its high thermal conductivity, reducing the likelihood of hot spots in exothermic reactions. FIGURE 2 THREE-PHASE FLUIDIZED BED



7.1 Gas and Liquid Hold-Up Regimes of behavior are identified from the drift flux. The drift flux of gas, W, can be defined as the volumetric flux of gas relative to a surface moving at the average velocity (see Wallis 1962 [Ref. 7]). It is then possible to derive a relationship between drift flux and the superficial velocities, see Darton, 1975 [Ref.8].

The drift flux calculated from some experiments by Michelsen and Ostergaard, 1970 [Ref. 9] is plotted as a function of gas voidage in Figure 3. The point of transition from one regime to the other depends on the physical properties of the gas and liquid, the particle size and the fluid flow rates. These are the factors which principally determine the onset of coalescence. FIGURE 3 PLOT OF EXPERIMENTAL DRIFT FLUX DATA vs GAS HOLD-

UP System Air-Water Particle sizes : 1 - 6 mm



Superficial Gas Velocities : 0.042 - 0.2 m/s 7.2 Calculation Procedure For a bed operating in the uniform bubbling regime, the following data are required: vSG, vSL, v tf and n, where n is the fluidization expansion index, which varies with particle properties between 2 and 5 (Richardson, 1954 [Ref. 10]). If no measured value of n is available, n = 4.5 should be used. The calculation procedure is as follows:



7.3 Bubble Size General quantification of bubble size is currently impossible. Lee et al, 1973 [Ref. 11] show that bubbles are smaller in beds of large particles, see Figure 4 (a) and 4 (b). FIGURE 4 RANGES OF BUBBLE SIZES IN WATER-FLUIDISED BEDS OF

GLASS BEADS 0.5!m FROM THE DISTRIBUTOR (Lee et al, 1973 [Ref. 11

These data have been obtained for air bubbles in water-fluidized beds. As the mechanism of bubble break-up depends on the interfacial tension and surface activity effects, prediction of bubble sizes for other systems is uncertain.



7.4 Mass Transfer 7.4.1 Gas Film Coefficient A numerical solution of the diffusion equation, allowing for recirculation in the bubble (see Zaritzky & Calvelo, 1979 [Ref.12]) gives the Sherwood number (ShGL = kG db/DG) as a function of the Peclet number (P e = db v b/DG) as shown in Figure 5. Note: DG is the diffusivity in the gas phase. FIGURE 5 DEPENDENCE OF SHERWOOD NUMBER ON PECLET

NUMBER

At low bubble Reynolds numbers (< 300) the circulation inside the bubble can be disregarded and the result

7.4.2 Liquid-Particle Mass Transfer The transfer rate to the particle surface may be estimated from:



Selection of a suitable velocity for use in this equation is difficult, the terminal velocity could be used, selecting the minimum value from those calculated from each of the Re ranges. An alternative correlation, based on data from liquid fluidized beds, see Beek, 1971 [Ref. 13], may be easier to use. This is given in equation (8).

Since only a very rough estimate of mass transfer rates can be obtained, they should always be confirmed by experimental data on the system of interest. 7.4.3 Gas-Liquid Mass Transfer An approximate value of the liquid film coefficient for transfer from the liquid-gas interface can be obtained by using the correlations recommended for bubble columns (see GBHE-PEG-MIX-705). Work by Ostergaard et al, 1972 [Ref. 14] using oxygen, water and ballotini shows the variation of KLa with distance from the distributor for two different particle sizes in a bubble column; see Figure 6.



FIGURE 6 VOLUMETRIC ABSORPTION COEFFICIENTS IN GAS-LIQUID FLUIDISED BEDS OF 1 mm PARTICLES



7.5 Heat Transfer Experimental studies have been confined to measurements of heat transfer from a surface to a water-air-glass beads system. Work to date is adequately reviewed by Baker et al, 1978 [Ref.15] who also present their own data. Heat transfer coefficients increase rapidly when gas is injected into a liquid-fluidized bed but the rate of increase tails off at high gas rates. The heat transfer coefficients also increase with increasing particle size and liquid velocity. For the water air- glass system all the data could be adequately correlated by the expression:

This gives some idea of the effect of those parameters which have been investigated. 7.6 Elutriation A height of liquid is required above the bed to reduce particle entrainment; particles will not be carried far out of the bed if the upward velocity of liquid in the bubble wake is less than the terminal velocity of the particles. If elutriation is likely to be a problem, a wire mesh baffle in the freeboard can be used to reduce it by splitting bubbles and encouraging wake shedding. Although a mesh can significantly reduce the amount of elutriation it does not affect the maximum height above the bed that particles reach. A comprehensive review of three-phase fluidized beds has been published by Wild et al, 1984 [Ref. 16]. Their main conclusions are that further research is required using organic systems as well as columns heights and diameters close to industrial scale.



8 SLURRY REACTORS Slurry reactors have much in common with bubble columns containing no solid particles. Considerably more data are available on bubble column design (see GBHE-PEG-MIX-705) and that Guide should be consulted in conjunction with this Clause. 8.1 Gas Rate The solid particles in a slurry reactor are suspended by momentum transfer from the gas phase via the liquid medium and thus the relationship between solids hold-up and the gas flowrate is of importance. The range of operating gas velocities decreases with increasing solids concentration; the lower limit increases because of an increasing tendency towards sedimentation and the upper limit decreases because of bubble coalescence and a corresponding increase in bubble size. The gas rate also affects the solids distribution in the reactor: an increase in gas velocity leads to a more uniform solids distribution. Quantitative information is sparse: Kölbel et al, 1964 (Ref.17) observed large deviations from the average solids concentration when operating at Vsg = 0.035 m/s, using 5% by weight sand particles of 0.1 - 0.125 mm diameter. Kato, 1963 [Ref.18] and Roy et al, 1964 [Ref. 19] have made some measurements on the amount of solids that can be suspended at a given gas velocity. 8.2 Mass Transfer The usual practice is for a solid (e.g. a catalyst) to be in a finely divided form to produce a large surface area. The calculation of mass transfer coefficients for the transfer from the gas phase to the liquid is described in GBHE-PEG-MIX-705. Transfer from the liquid to the particle can be calculated from:

where Re is the particle Reynolds number:



DOCUMENTS REFERRED TO IN THIS PROCESS ENGINEERING GUIDE This Process Engineering Guide makes reference to the following documents: GBHE ENGINEERING GUIDES GBH Enterprises Glossary of Engineering Terms

(referred to in Clause 3) GBHE-PEG-MIX-703 Mixing of Solid-Liquid Systems

(referred to in Clause 6 and 6.2) GBHE-PEG-MIX-705 Mixing of Gas Liquid Systems

(referred to in Clause 6, 6.1, 7.4.3, 8 and 8.2).

Gas-Solid-Liquid Mixing Systems

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