Epstein N., Grace J.R. Spouted and Spout-Fluid Beds: Fundamentals and Applications

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Introduction 7

0.2 0.3 0.4 0.5

SUPERFICIAL AIR VELOCITY, m/s

BUBBLING

STATIC

BED

SLUGGING

BED DEPTH, cm

Figure 1.6. Regime map for Ottawa sand (d

= 0.589 mm). D

= 152 mm, D

= 15.8 mm. After

Mathur and Gishler.

spouted bed annulus, of at least one-half the cylinder diameter. If the bed is much

shallower, the system differs hydrodynamically from true spouting, and any generally

formulated principles of spouted bed behavior would not be expected to apply. A mini-

mum spoutable bed depth has, however, not been precisely deﬁned or investigated, except

in the case of conical beds,

nor have any detailed studies been made about the max-

imum spouting velocity at which transition from coherent spouting to either bubbling

ﬂuidization or slugging occurs. For most practical purposes, however, there is usually

sufﬁcient latitude between the minimum and maximum spouting velocity that the ﬂuid

ﬂow can be amply increased above the minimum without transition to ﬂuidization.

1.4 Nonaxisymmetric geometries of spouted beds

The photographs of Figure 1.2 and, in most cases, illustrations such as those of Figure 1.4

are obtained from the transparent ﬂat face of semicircular conical or conical-cylindrical

vessels, with aligned semicircular gas inlets. Qualitatively, what is seen and what is

measured in such half-sectional columns is comparable to what can be detected by

various techniques (e.g., piezoelectricity, stop-ﬂow, laser-Doppler anemometry, optical

ﬁber probes) and what is measured in full columns. Quantitatively, however, although

some characteristics of spouted beds are well matched in the fully cylindrical and

semicylindrical (labeled “full” and “half,” respectively) columns, others show signiﬁcant

differences.

Experimental data required to construct a regime map such as Figure 1.5 are, by and

large, quite similar when obtained for either a half or a full column. Thus, measured

values of U

from half-columns have been shown to be equal to,

17,19

or at most only

10 percent greater than,

for full columns, whereas H

from half columns has been

8 Norman Epstein and John R. Grace

measured as approximately equal to,

and never more than 32 percent smaller than,

for full columns. Bed pressure drops and pressure proﬁles during spouting have

also been found to be quite similar f or the two-column geometries, as have pitot tube

measurements of radial and longitudinal gas velocity proﬁles,

except in the immediate

vicinity of the half-column ﬂat wall.

The situation is different for details of particle motion and spout geometry. All walls

tend to retard vertical particle motion r elative to what occurs a few particle diameters

from the wall (typically by about 20 percent or more for smooth walls and smooth

spherical particles), whether the motion is downward in the annulus at ﬂat or circular

walls

or upward in the spout at the ﬂat wall of a half column (measured by laser-

Doppler anemometry

or by ﬁber optic probes

). Consequently, cinephotographic

measurements at the transparent ﬂat face, which are biased toward particles very near

the face, tend to underestimate particle velocities in the spout, especially at or near the

axis of a full column. Par ticle velocities in the fountain are, however, well matched

between half and full columns.

As for the annulus, although the downward particle

velocity at the ﬂat wall has sometimes been repor ted as being equal to that at the

corresponding half-round wall, as well as to that at the round wall of a full column,

2,19,20

Rovero et al.,

by a carefully monitored stop-ﬂow technique using tracer particles,

showed that particle trajectories at the half-column ﬂat face are not representative of full

column behavior at the same plane, and that only the middle 60

◦

sector of a half-column

is representative of a full column as far as particle velocity is concerned.

Spout geometry is more contentious. Mikhailik,

based on a piezoelectric probe

for measuring particle velocity in the spout of a fully circular bed, coupled with

visual observation of spouting in a larger diameter transparent semicircular column,

concluded that the shape and size of the spout were the same in both types of column.

However, the longitudinally uniform cylindrical spout shape reported by Mikhailik is

at odds with the recorded observations of most other workers, and the criterion used for

size equality, given the inequality of vessel diameters and the absence of geometrical

similarity between the two systems, was the adherence of measured spout diameters in

both cases to the same empirical equation relating average spout diameter to operating

system parameters. A more convincing study by He et al.,

using optical ﬁber probes

for voidage and particle velocity in geometrically matched semicylindrical and fully

cylindrical cone-based columns, showed cross-sectional modiﬁcation of the full column

circular spout to an approximate semiellipse (with the larger axis perpendicular to

the ﬂat wall) in the half column. There was longitudinal distortion from monotonic

divergence of spout width in the upward direction, both for the full column and for

the half-column at 90

◦

to the ﬂat wall, to an S-shape along the ﬂat wall itself. Average

spout diameters in the full column were underestimated by about 35 percent if based on

measurements along the ﬂat face of the half-column.

The corners of half columns contribute most to distor tion of full-column spouted bed

behavior.

◦

,60

◦

, and 30

◦

sector columns, which produce ever narrower corners as

the sector angle decreases, although they give rise to spouted bed behavior of sorts,

do not in most important respects compare well to measurements obtained in full or

Introduction 9

even half columns, and show increasing instability as the sector angle decreases.

27,28

Unlike half-sectional beds, which retain their value as qualitative and partly quan-

titative visual simulators of full beds, sector beds are not recommended for these

purposes.

As in the case of ﬂuidized bed research, “two-dimensional” beds – that is, beds of

particles encased between parallel transparent vertical planes a few millimeters apart

with a ﬂuid inlet that spans this thickness – have also been used in spouted bed research

to give qualitative insight into three-dimensional behavior. In the case of spouted beds,

it has been claimed that such slot-rectangular devices can be scaled up readily – for

instance, for grain drying purposes.

This claim is discussed in Chapter 17.

1.5 Spouted beds in the gas–solid contacting spectrum

1.5.1 The spectrum

Gas–solid contacting systems may be broadly classiﬁed as (1) nonagitated, (2) mechan-

ically agitated, and (3) gas-agitated. Categories (1) and (3), unlike (2), have no moving

parts.

Fixed and moving packed beds, which belong to the ﬁrst category, are applicable to

processes that do not call for maintaining high rates of heat and mass transfer between

the gas and the coarse solids, and in which uniformity of conditions in different parts of

the bed is either not critical or not desirable. In a ﬁxed bed, solids cannot be continuously

added or withdrawn, and treatment of the gas is usually the main objective – for example,

to recover solvent vapors by gas adsorption, to dry a gas, or to carry out a gas-phase

reaction catalyzed by the particles, in which the catalyst has an extended life (limited

deactivation). A moving bed provides continuous ﬂow of solids through the reaction

zone. Its application therefore extends to chemical and physical treatment of solids in

such processes as roasting of ores, calcining of limestone, and drying and cooling pellets

and briquettes. In both ﬁxed and moving beds, the gas movement is close to plug ﬂow,a

feature that is almost always advantageous for chemical reactions, where axial dispersion

tends to reduce conversions and yields.

Limited agitation can be imparted to the solids by mechanical means, either by move-

ment of the vessel itself, as in rotary driers and kilns, or by internal impellers. In either

case, most of the material remains in a packed-bed condition, but the differential move-

ment of particles improves contacting effectiveness because fresh surface is continually

exposed to the gas. Also, the blending of solids by agitation levels out interparticle gradi-

ents in composition and temperature. Mechanical systems are used mainly for processes

involving solids treatment, s uch as drying, calcining, and cooling, but are unsuitable for

processes that require the gas to be treated uniformly.

In gas-agitated systems, such as ﬂuidized beds or solids transport systems, intense

agitation is imparted to each solid particle by the action of the gas stream. In ﬂuidization,

whether of the dense, noncirculating (externally), particulate, or bubbling bed variety at

10 Norman Epstein and John R. Grace

one extreme, or dilute, circulating, or fast-ﬂuidization at the other, the large surface area

of the ﬁne well-mixed solids gives rise to high particle-to-gas heat and mass transfer

rates; coefﬁcients of heat transfer to or from the vessel walls or submerged surfaces

(e.g., of heat exchangers) are considerably higher than those at the same superﬁcial gas

velocity in the absence of solids or in the presence of a ﬁxed bed of solids. Continu-

ous operation of ﬂuidized beds can be readily accomplished because solids are easily

added to, and withdrawn from, such beds. In ﬂuidization, the average residence time

of the particles in the bed can be controlled by adjusting the solids feed rate and the

holdup of the bed. Because of these basic features, ﬂuidized beds are the preferred

contacting method for many processes, including chemical reactions (both catalytic

and noncatalytic) involving the gas, as well as physical treatment (e.g., drying) of the

solids.

Dilute solids transport (e.g., pneumatic conveying) may be more suitable than ﬂu-

idization for some gas–solid contacting operations. In such systems the contact time

between a given particle and a gas is very short – no more than a few seconds – because

of very high gas velocities. Intense turb ulence f acilitates high coefﬁcients of heat and

mass transfer, but the extents of heat and mass transfer to or from the particles is limited

by the small residence time of the par ticles in the reaction zone. Dilute solids transport

systems are especially suitable f or processes that are surface-rate controlled rather than

internal-diffusion controlled with respect to the solid particles. Examples of established

applications are combustion of pulverized coal, ﬂash roasting or smelting of metallic

sulﬁdes, and ﬂash drying of s ensitive materials that can tolerate exposure to heat for

only a few seconds.

Because the annular region of a spouted bed, which contains most of the solids,

is virtually a moving packed bed, with counterﬂow of gas, spouted beds have some-

times been categorized as moving-bed systems.

However, the dilute solids trans-

port that characterizes the spout region is crucial to overall spouted-bed performance

and, inasmuch as bed solids recirculate many times, therefore becoming well mixed,

the operation comes closer to conventional ﬂuidization in its attributes. Although the

early statement that spouting “appears to achieve the same purpose for coarse par-

ticles as ﬂuidization does for ﬁne materials”

remains true for certain applications,

such as solids drying, the differences in the two operations are important for oth-

ers. For example, the systematic cyclic movement of particles in a spouted bed, in

contrast to more random motion in ﬂuidization, is a feature of critical value for gran-

ulation and particle coating processes. In the spectrum of contacting systems, there-

fore, spouted beds occupy a position that overlaps moving and ﬂuidized beds to some

extent, but at the same time they have a place of their own by virtue of cer tain unique

characteristics.

1.5.2 Spouting versus ﬂuidization

In what has become the well-entrenched C-A-B-D classiﬁcation scheme for distin-

guishing between various types of ﬂuidization by air at ambient conditions, Geldart

Introduction 11

proposed the criterion (for air at room temperature and atmospheric pressure as the

motive ﬂuid)

(ρ

− ρ)d

> 10

−3

kg/m

as characterizing particles that “can form stable spouted beds.” This criterion was

rederived by Molerus

(ρ

− ρ)gd

> 15.3Pa

for particles of diameter d

and density ρ

, ﬂuid of density ρ, and gravitational accel-

eration g. Although both these criteria, which neglect the important role of ﬂuid inlet

diameter on spoutability, have been shown experimentally to be strictly invalid,

they

have the virtue of emphasizing that there is a minimum particle diameter below which

spouting becomes impractical. This diameter is about 0.5 to 1.0 mm – that is, within the

ﬂuidization transition range of what Squires

labeled a relatively ﬂexible “ﬂuid bed”

of ﬁne particles and a relatively inﬂexible “teeter bed” (a term arbitrarily “borrowed

from ore dressing technology where it means to ﬂuidize particulately with water”) of

coarse particles, requiring a wide particle size distribution to avoid channeling. It is

possible to operate a miniature spouted bed with considerably ﬁner particles

using a

very small gas inlet. Indeed, perforated plate distributors for ﬂuidized beds often give

rise to formation of small spouts above each hole, with the s pouts breaking into bubbles

further up the bed.

If, however, a single small hole is used to spout ﬁne particles,

the allowable holdup and capacity of the bed would also be small, and any attempt at

scaleup by enlarging the inlet hole would lead to nonhomogeneous ﬂuidization rather

than spouting.

37,38

The total frictional pressure drop across a fully spouting bed is always lower by at least

20 percent

than that required to support the buoyed weight of the bed – in other words,

than the frictional pressure drop for particulate ﬂuidization. In this respect, a spouted

bed is qualitatively similar to a channeling ﬂuidized bed.

Channeling, however, is an

undesirable feature of a ﬂuidized bed that involves passage of gas through part of the

bed without inducing much movement of, or contact with, the surrounding particles.

In spouting, on the other hand, agitation of the entire bed is achieved by the gas jet,

and intimate contact between the particles and the gas occurs both in the jet and in the

surrounding dense-phase annulus, the latter being fed gas by crossﬂow from the jet.

In a typical spouted bed, well over half the jet gas has already permeated the annulus

halfway up the bed. In addition, channeling in ﬂuidized beds is favored by very ﬁne

particles,

11,40,41

whereas spouting is normally a coarse particle phenomenon. Thus the

similarity between spouting and channeling, stressed by several authors,

11,12,42

can be

misleading.

The shape of the longitudinal proﬁle in a spouted bed also distinctly differs from that

of a ﬂuidized bed when a cylindrical column, or even a conical-cylindrical column with

no bed support, is used in both cases. Thus, for a spouted bed, the longitudinal pressure

gradient varies with height, rising to a maximum at the top of the bed. For a ﬂuidized bed,

on the other hand, the pressure gr adient is nearly constant over the cylindrical portion of

12 Norman Epstein and John R. Grace

the column, despite the fact that for gas ﬂuidization, as for spouting, the usual pattern

of solids movement is up at the center and down near the wall.

Apart from achieving favorable gas–solid contact, spouting, more than ﬂuidization, is

a means of agitating coarse particles with a gas and of causing solid-solid contacts, the

interaction between gas and solid particles being incidental. These additional features

of a spouted bed are particularly applicable to mechanical operations such as blending,

comminution, and dehusking of solids. The absence of a ﬂuid distributor and its asso-

ciated problems (e.g., plugging by solids) is another signiﬁcant aspect of spouting that

distinguishes it from ﬂuidization.

Table 1.1, from Cui and Grace,

summarizes key differences between gas-ﬂuidized

and gas-spouted beds.

1.5.3 Spout-ﬂuid beds

A hybrid reactor that shares some characteristics of both spouting and ﬂuidization,

and which is especially useful for coarse, sticky, or agglomerating solids, is commonly

referred to as a spout-ﬂuid bed. Such a reactor involves a substantial ﬂuid ﬂow through

a single central inlet oriﬁce, as in a spouted bed, and an auxiliary ﬂuid ﬂow through a

series of holes in a surrounding distributor, as in a ﬂuidized bed. The distributor can be

incorporated into either a ﬂat or a conical base of the spout-ﬂuid column, whereas the

ﬂuid can be a gas,

a liquid,

or a gas–liquid mixture.

Often, in spout-ﬂuid beds, and less frequently in spouted beds, a centrally located

draft tube is inserted at a small distance above the inlet oriﬁce to decrease crossﬂow

of the spout ﬂuid into the annular region, thus increasing the maximum spoutable bed

depth. The draft tube also reduces the intermixing of annular and spout solids, and

thus reduces the residence time spread of the solids for solids-continuous processes.

A limiting design of a draft-tube spout-ﬂuid bed is the “internally circulating ﬂuidized

bed,”

in which the draft tube becomes a riser by extending it downward all the way

to the ﬂuid inlet oriﬁce, and in which the annular solids are received as spillover from

the top of the riser and then returned to the spout via smaller oriﬁces located in the

riser wall near its base. By vir tue of the net positive pressure difference between the

concentrated solids annular region and the dilute solids riser, the riser ﬂuid is prevented

from escaping into the annulus, whereas a small fraction of the auxiliary ﬂuid ﬂows

through the riser oriﬁces, enhancing the radial inﬂux and hence the circulation of the

solids.

1.6 Layout of chapter topics

The three chapters that follow this one, as well as Chapter 8, deal with various aspects

of the ﬂuid-particle mechanics of CSBs. Chapter 5 is restricted to conical spouted beds,

but covers a broad range of subjects from hydrodynamics to applications. Chapters 6

and 7 again focus mainly on ﬂuid-particle mechanics, but this time respectively for

spout-ﬂuid beds without, and for both spouted and spout-ﬂuid beds with, draft tubes.

Introduction 13

Table 1.1. Signiﬁcant differences between gas-spouted beds and gas-ﬂuidized beds.

Aspect Fluidized Bed Spouted Bed

Mean particle size ∼0.03–3 mm; usually <1mm ∼0.6–6 mm; usually >1mm

Particle size distribution Usually broad Usually narrow

Pressure drop/height

within the bed

96–100% of that needed to support

the particles, i.e.,

∼(ρ

− ρ) g(1 − ε)

Less than ∼75% of that needed to support

the weight of the particles, i.e.,

<0.75 × (ρ

− ρ)g(1 − ε)

Pressure drop across

entry oriﬁces

Usually 30–50% of that across the

bed

As small as possible consistent with

satisfying the other constraints, e.g.,

oriﬁce dia. <25 mean particle dia.

Axial gradient of

pressure

Virtually independent of height in

the column

Varies with height

Temperature gradient Very uniform temperature over

entire ﬂuidized bed region

Signiﬁcant temperature gradients both

axially and radially

Column geometry Usually cylindrical columns Usually diverging conical base with or

without cylindrical portion above

Oriﬁce diameter No restriction Must not exceed 25 mean particle

diameters

Oriﬁce number density

and conﬁguration

Large number of oriﬁces (typically

>100 per square meter); many

geometries; no constriction

needed at the entrance

Usually a single oriﬁce. For multiple

spouting, smaller number density of

oriﬁces (typically no more than

∼10–50 per square meter); it is helpful

to have a constriction at the entrance

Oriﬁce orientation Most frequently horizontal,

downward facing, or oblique with

downward component, but may

also be upward-facing

Always upward-facing

Oriﬁce feed system Windbox (sometimes called plenum

chamber) feeds all oriﬁces

Separate gas supply and control of each

oriﬁce

Gas motion Less ordered; depends on ﬂow

regime and speciﬁc geometry

Outward from the spout into the dense

phase, except just above inlet

Particle motion Complex ﬂow regimes and particle

motion; region surrounding the

gas entry oriﬁces is usually

ﬂuidized with few

particle-particle contacts

Systematic circulation patterns, up the

spout and slowly downward in the

annulus. Annulus is in moving packed

bed ﬂow with substantial

particle–particle contacts

Particle segregation Very little segregation providing

that particles are well ﬂuidized

Signiﬁcant segregation according to both

size and density of particles

Attrition Normally little except in cyclone or

jet regions

Signiﬁcant in spout region and in fountain

above

Bed depth Broad range of depths, e.g., from

0.1 to 20 m dense bed depth

More limited range of depths, usually

between 0.2 and 2.0 m

Superﬁcial gas velocity Broad range, typically

(U − U

) = 0.2to10m/s

More limited range, typically 1.1 to

1.8U

Oriﬁce velocity variation Some variation with time across

each individual oriﬁce as bubbles

form and leave the oriﬁce or jet

Essentially steady, as ﬂow is controlled

to each hole from an upstream

valve

Internals Common in ﬂuidized bed processes Rare in spouted bed processes, except for

some use of draft tubes and heat

transfer coils

14 Norman Epstein and John R. Grace

Chapter 9 examines various aspects of heat transfer and, to a lesser extent, mass transfer.

Chapter 10 reports on a relatively new development known as a powder-particle spouted

bed, in which ﬁne particles that serve some useful function (e.g., as an adsorbent) are

continuously fed and removed from a spouted bed of coarse particles.

Chapters 11 through 14 describe the four principal physical operations in which

spouted or spout-ﬂuid beds have achieved industrial application – drying of particulate

solids, evaporative drying of liquids containing dissolved or suspended solids, granula-

tion, and particle-coating. The Wurster process (Chapter 14) is a tablet-coating technique

that has achieved much exposure in pharmaceutical journals; it involves a spout-ﬂuid

bed with a draft tube. Gasiﬁcation and combustion of coal and other fuels are dealt with

in Chapter 15, whereas catalytic chemical reactor modeling is treated in Chapter 19.

Mechanical spouting, in which the solids are elevated mechanically rather than hydro-

dynamically, is the subject of Chapter 18. Liquid and liquid–gas spouting are examined

in Chapter 20, whereas an important potential application of liquid spouting, its use as

an electrochemical reactor, is treated in Chapter 16. Scaleup of spouted beds, either by

application of dimensional similitude principles or by alternative designs, is the subject

of Chapter 17.

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Introduction 15

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16 Norman Epstein and John R. Grace

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