Perry s chemical engineers handbook 8e section 17gas solid operations and equipment

Copyright © 2008, 1997, 1984, 1973, 1963, 1950, 1941, 1934 by The McGraw-Hill Companies, Inc. All rights reserved. Manufactured in the United

States of America. Except as permitted under the United States Copyright Act of 1976, no part of this publication may be reproduced or distributed

in any form or by any means, or stored in a database or retrieval system, without the prior written permission of the publisher.

0-07-154224-8

The material in this eBook also appears in the print version of this title: 0-07-151140-7.

All trademarks are trademarks of their respective owners. Rather than put a trademark symbol after every occurrence of a trademarked name, we use

names in an editorial fashion only, and to the benefit of the trademark owner, with no intention of infringement of the trademark. Where such

designations appear in this book, they have been printed with initial caps.

McGraw-Hill eBooks are available at special quantity discounts to use as premiums and sales promotions, or for use in corporate training programs.

For more information, please contact George Hoare, Special Sales, at [email protected] or (212) 904-4069.

TERMS OF USE

This is a copyrighted work and The McGraw-Hill Companies, Inc. (“McGraw-Hill”) and its licensors reserve all rights in and to the work. Use of this

work is subject to these terms. Except as permitted under the Copyright Act of 1976 and the right to store and retrieve one copy of the work, you may

not decompile, disassemble, reverse engineer, reproduce, modify, create derivative works based upon, transmit, distribute, disseminate, sell, publish

or sublicense the work or any part of it without McGraw-Hill’s prior consent. You may use the work for your own noncommercial and personal use;

any other use of the work is strictly prohibited. Your right to use the work may be terminated if you fail to comply with these terms.

THE WORK IS PROVIDED “AS IS.” McGRAW-HILL AND ITS LICENSORS MAKE NO GUARANTEES OR WARRANTIES AS TO THE

ACCURACY, ADEQUACY OR COMPLETENESS OF OR RESULTS TO BE OBTAINED FROM USING THE WORK, INCLUDING ANY

INFORMATION THAT CAN BE ACCESSED THROUGH THE WORK VIA HYPERLINK OR OTHERWISE, AND EXPRESSLY DISCLAIM

ANY WARRANTY, EXPRESS OR IMPLIED, INCLUDING BUT NOT LIMITED TO IMPLIED WARRANTIES OF MERCHANTABILITY OR

FITNESS FOR A PARTICULAR PURPOSE. McGraw-Hill and its licensors do not warrant or guarantee that the functions contained in the work will

meet your requirements or that its operation will be uninterrupted or error free. Neither McGraw-Hill nor its licensors shall be liable to you or

anyone else for any inaccuracy, error or omission, regardless of cause, in the work or for any damages resulting therefrom. McGraw-Hill has no

responsibility for the content of any information accessed through the work. Under no circumstances shall McGraw-Hill and/or its licensors be liable

for any indirect, incidental, special, punitive, consequential or similar damages that result from the use of or inability to use the work, even if any of

them has been advised of the possibility of such damages. This limitation of liability shall apply to any claim or cause whatsoever whether such claim

or cause arises in contract, tort or otherwise.

DOI: 10.1036/0071511407

This page intentionally left blank

FLUIDIZED-BED SYSTEMS

Gas-Solid Systems. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-2

Types of Solids . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-2

Two-Phase Theory of Fluidization . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-2

Phase Diagram (Zenz and Othmer). . . . . . . . . . . . . . . . . . . . . . . . . . . 17-3

Phase Diagram (Grace) . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-3

Regime Diagram (Grace) . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-3

Solids Concentration versus Height. . . . . . . . . . . . . . . . . . . . . . . . . . . 17-5

Equipment Types . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-5

Minimum Fluidizing Velocity. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-5

Particulate Fluidization . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-6

Vibrofluidization . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-6

Design of Fluidized-Bed Systems . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-6

Fluidization Vessel . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-6

Scale-up. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-9

Heat Transfer . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-11

Temperature Control . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-12

Solids Mixing. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-12

Gas Mixing . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-12

Size Enlargement . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-12

Size Reduction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-12

Standpipes, Solids Feeders, and Solids Flow Control. . . . . . . . . . . . . 17-12

Solids Discharge . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-13

Dust Separation . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-14

Example 1: Length of Seal Leg . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-15

Instrumentation . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-15

Uses of Fluidized Beds. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-16

Chemical Reactions . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-16

Physical Contacting. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-20

GAS-SOLIDS SEPARATIONS

Nomenclature . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-21

Purpose of Dust Collection . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-24

Properties of Particle Dispersoids . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-24

Particle Measurements. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-24

Atmospheric-Pollution Measurements . . . . . . . . . . . . . . . . . . . . . . . . 17-24

Process-Gas Sampling . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-24

Particle-Size Analysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-24

Mechanisms of Dust Collection. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-26

Performance of Dust Collectors . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-27

Dust-Collector Design . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-27

Dust-Collection Equipment. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-28

Gravity Settling Chambers . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-28

Impingement Separators . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-28

Cyclone Separators . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-28

Mechanical Centrifugal Separators . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-36

Particulate Scrubbers . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-36

Dry Scrubbing. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-43

Fabric Filters . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-46

Granular-Bed Filters. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-51

Air Filters . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-52

Electrical Precipitators . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17-55

17-1

Section 17

Gas-Solid Operations and Equipment

Mel Pell, Ph.D. President, ESD Consulting Services; Fellow, American Institute of Chemical

Engineers; Registered Professional Engineer (Delaware) (Section Editor, Fluidized-Bed Systems)

James B. Dunson, M.S. Principal Division Consultant (retired), E. I. duPont de Nemours

& Co.; Member, American Institute of Chemical Engineers; Registered Professional Engineer

(Delaware) (Gas-Solids Separations)

Ted M. Knowlton, Ph.D. Technical Director, Particulate Solid Research, Inc.; Member,

American Institute of Chemical Engineers (Fluidized-Bed Systems)

Copyright © 2008, 1997, 1984, 1973, 1963, 1950, 1941, 1934 by The McGraw-Hill Companies, Inc. Click here for terms of use.

FLUIDIZED-BED SYSTEMS

Geldart categorized solids into four different groups (groups A, B, C,

and D) that exhibited different properties when fluidized with a gas.

He classified the four groups in his famous plot, shown in Fig. 17-1.

This plot defines the four groups as a function of average particle size

dsv, µm, and density difference s − f, g/cm3

, where s = particle den￾sity, f = fluid density, and dsv = surface volume diameter of the parti￾cles. Generally dsv is the preferred average particle size for fluid-bed

applications, because it is based on the surface area of the particle.

The drag force used to generate the pressure drop used to fluidize the

bed is proportional to the surface area of the particles. Another widely

used average particle is the median particle size dp,50.

When the gas velocity through a bed of group A, B, C, or D particles

increases, the pressure drop through the bed also increases. The pres￾sure drop increases until it equals the weight of the bed divided by the

cross-sectional area of the column. The gas velocity at which this

occurs is called the minimum fluidizing velocity Umf. After minimum

fluidization is achieved, increases in velocity for a bed of group A (gen￾erally in the particle size range between 30 and 100 µm) particles will

result in a uniform expansion of the particles without bubbling until at

some higher gas velocity the gas bubbles form at a velocity called the

minimum bubbling velocity Umb. For Geldart group B (between 100

and about 1000 µm) and group D (1000 µm and larger) particles, bub￾bles start to form immediately after Umf is achieved, so that Umf and Umb

are essentially equal for these two Geldart groups. Group C (generally

smaller than 30 µm) particles are termed cohesive particles and clump

together in particle agglomerates because of interparticle forces (gen￾erally van der Waals forces). When gas is passed through beds of cohe￾sive solids, the gas tends to channel or “rathole” through the bed.

Instead of fluidizing the particles, the gas opens channels that extend

from the gas distributor to the surface of the bed. At higher gas veloci￾ties where the shear forces are great enough to overcome the interpar￾ticle forces, or with mechanical agitation or vibration, cohesive

particles will fluidize but with larger clumps or clusters of particles

formed in the bed.

Two-Phase Theory of Fluidization The two-phase theory of

fluidization assumes that all gas in excess of the minimum bubbling

velocity passes through the bed as bubbles [Toomey and Johnstone,

Chem. Eng. Prog. 48: 220 (1952)]. In this view of the fluidized bed,

the gas flowing through the emulsion phase in the bed is at the mini￾mum bubbling velocity, while the gas flow above Umb is in the bubble

phase. This view of the bed is an approximation, but it is a helpful way

Consider a bed of particles in a column that is supported by a distribu￾tor plate with small holes in it. If gas is passed through the plate so that

the gas is evenly distributed across the column, the drag force on the

particles produced by the gas flowing through the particles increases as

the gas flow through the bed is increased. When the gas flow through

the bed causes the drag forces on the particles to equal the weight of the

particles in the bed, the particles are fully supported and the bed is said

to be fluidized. Further increases in gas flow through the bed cause

bubbles to form in the bed, much as in a fluid, and early researchers

noted that this resembled a fluid and called this a fluidized state.

When fluidized, the particles are suspended in the gas, and the flu￾idized mass (called a fluidized bed) has many properties of a liquid.

Like a liquid, the fluidized particles seek their own level and assume

the shape of the containing vessel. Large, heavy objects sink when

added to the bed, and light particles float.

Fluidized beds are used successfully in many processes, both cat￾alytic and noncatalytic. Among the catalytic processes are fluid cat￾alytic cracking and reforming, oxidation of naphthalene to phthalic

anhydride, the production of polyethylene and ammoxidation of

propylene to acrylonitrile. Some of the noncatalytic uses of fluidized

beds are in the roasting of sulfide ores, coking of petroleum residues,

calcination of ores, combustion of coal, incineration of sewage sludge,

and drying and classification.

Although it is possible to fluidize particles as small as about 1 µm

and as large as 4 cm, the range of the average size of solid particles

which are more commonly fluidized is about 30 µm to over 2 cm. Par￾ticle size affects the operation of a fluidized bed more than particle

density or particle shape. Particles with an average particle size of

about 40 to 150 µm fluidize smoothly because bubble sizes are rela￾tively small in this size range. Larger particles (150 µm and larger)

produce larger bubbles when fluidized. The larger bubbles result in a

less homogeneous fluidized bed, which can manifest itself in large

pressure fluctuations. If the bubble size in a bed approaches approxi￾mately one-half to two-thirds the diameter of the bed, the bed will

slug. A slugging bed is characterized by large pressure fluctuations

that can result in instability and severe vibrations in the system. Small

particles (smaller than 30 µm in diameter) have large interparticle

forces (generally van der Waals forces) that cause the particles to stick

together, as flour particles do. These type of solids fluidize poorly

because of the agglomerations caused by the cohesion. At velocities

that would normally fluidize larger particles, channels, or spouts, form

in the bed of these small particles, resulting in severe gas bypassing.

To fluidize these small particles, it is generally necessary to operate at

very high gas velocities so that the shear forces are larger than the

cohesive forces of the particles. Adding finer-sized particles to a

coarse bed, or coarser-sized particles to a bed of cohesive material

(i.e., increasing the particle size range of a material), usually results in

better (smoother) fluidization.

Gas velocities in fluidized beds generally range from 0.1 to 3 m/s

(0.33 to 9.9 ft/s). The gas velocities referred to in fluidized beds are

superficial gas velocities—the volumetric flow through the bed

divided by the bed area. More detailed discussions of fluidized beds

can be found in Kunii and Levenspiel, Fluidization Engineering, 2d

ed., Butterworth Heinemann, Boston, 1991; Pell, Gas Fluidization,

Elsevier, New York, 1990; Geldart (ed.), Gas Fluidization Technology,

Wiley, New York, 1986; Yang (ed.), Handbook of Fluidization and

Fluid Particle Systems, Marcel Dekker, New York, 2003; and papers

published in periodicals, transcripts of symposia, and the American

Institute of Chemical Engineers symposium series.

GAS-SOLID SYSTEMS

Researchers in the fluidization field have long recognized that parti￾cles of different size behave differently in fluidized beds, and several

have tried to define these differences. Some of these characterizations

are described below.

Types of Solids Perhaps the most widely used categorization

of particles is that of Geldart [Powder Technol. 7: 285–292 (1973)].

FIG. 17-1 Powder-classification diagram for fluidization by air (ambient con￾ditions). [From Geldart, Powder Technol., 7, 285–292 (1973).]

17-2

of understanding what happens as the gas velocity is increased

through a fluidized bed. As the gas velocity is increased above Umb,

more and larger bubbles are formed in the bed. As more bubbles are

produced in the bed, the bed expands and the bed density decreases.

For all Geldart groups (A, B, C, and D), as the gas velocity is

increased, the fluidized-bed density is decreased and the turbulence

of the bed is increased. In smaller-diameter beds, especially with

group B and D powders, slugging will occur as the bubbles increase

in size to greater than one-half to two-thirds of the bed diameter.

Bubbles grow by vertical and lateral merging and increase in size as

the gas velocity is increased [Whitehead, in Davidson and Harrison

(eds.), Fluidization, Academic, London and New York, 1971]. As the

gas velocity is increased further, the stable bubbles break down into

unstable voids. When unstable voids characterize the gas phase in flu￾idized beds, the bed is not in the bubbling regime anymore, but is said

to be in the turbulent regime. The turbulent regime is characterized

by higher heat- and mass-transfer rates than bubbling fluidized beds,

and the pressure fluctuations in the bed are reduced relative to bub￾bling beds. As the gas velocity is increased above the turbulent flu￾idized regime, the turbulent bed gradually changes into the pneumatic

conveying regime.

Phase Diagram (Zenz and Othmer) As shown in Fig. 17-2,

Zenz and Othmer, (Fluidization and Fluid Particle Systems, Reinhold,

New York, 1960) developed a gas-solid phase diagram for systems in

which gas flows upward, as a function of pressure drop per unit length

versus gas velocity with solids mass flux as a parameter. Line OAB in

Fig. 17-2 is the pressure drop versus gas velocity curve for a packed

bed, and line BD is the curve for a fluidized bed with no net solids

flow through it. Zenz indicated that there was an instability between

points D and H because with no solids flow, all the particles will be

entrained from the bed. However, if solids are added to replace those

entrained, system IJ (generally known as the pneumatic conveying

region) prevails. The area DHIJ will be discussed in greater detail

later.

Phase Diagram (Grace) Grace [Can. J. Chem. Eng., 64: 353–363

(1986); Fig. 17-3] has correlated the various types of gas-solid systems

in which the gas is flowing vertically upward in a status graph using the

parameters of the Archimedes number Ar for the particle size and a

nondimensional velocity U* for the gas effects. By means of this plot,

the fluidization regime for various operating systems can be approxi￾mated. This plot is a good guide to estimate the fluidization regime

for various particle sizes and operating conditions. However, it should

not be substituted for more exact methods of determining the actual

fluidization operating regime.

Regime Diagram (Grace) Grace [Can. J. Chem. Eng., 64,

353–363 (1986)] approximated the appearance of the different

regimes of fluidization in the schematic drawing of Fig. 17-4. This

drawing shows the fluidization regimes that occur as superficial gas

velocity is increased from the low-velocity packed bed regime to the

pneumatic conveying transport regime. As the gas velocity is increased

from the moving packed bed regime, the velocity increases to a value

Umf such that the drag forces on the particles equal the weight of the

bed particles, and the bed is fluidized. If the particles are group A par￾ticles, then a “bubbleless” particulate fluidization regime is formed. At

a higher gas velocity Umb, bubbles start to form in the bed. For Geldart

group B and D particles, the particulate fluidization regime does not

form, but the bed passes directly from a packed bed to a bubbling flu￾idized bed. As the gas velocity is increased above Umb, the bubbles in

the bed grow in size. In small laboratory beds, if the bubble size grows

to a value equal to approximately one-half to two-thirds the diameter

FLUIDIZED-BED SYSTEMS 17-3

FIG. 17-2 Schematic phase diagram in the region of upward gas flow. W = mass flow solids, lb/(h  ft2); ε = frac￾tion voids; ρp = particle density, lb/ft3

; ρf = fluid density, lb/ft3

; CD = drag coefficient; Re = modified Reynolds

number. (Zenz and Othmer, Fluidization and Fluid Particle Systems, Reinhold, New York, 1960.)

Key:

OAB = packed bed IJ = cocurrent flow AC = packed bed FH = dilute phase

BD = fluidized bed = (dilute phase) = (restrained at top) MN = countercurrent flow

DH = slugging bed ST = countercurrent flow OEG = fluid only = (dilute phase)

= (dense phase) = (no solids) VW = cocurrent flow

= (dense phase)