Fluidizing with hot air is an attractive means for drying many moist powders and granular products. The technique has been used industrially since 1948, and today it enjoys widespread popularity for drying crushed minerals, sand, polymers, fertilizers, pharmaceuticals, crystalline materials, and many other industrial products.
The main reasons for this popularity are as follows:
(a) Efficient gas solids contacting leads to compact units and relatively low capital cost combined with relatively high thermal efficiency.
(b) The handling of the particles is quite gentle compared to some other types of dryer. This is important with fragile crystals.
(c) The lack of moving parts, other than feeding and discharge mechanisms, keeps reliability high and maintenance costs low.
The main limit on the applicability of fluid bed dryers is that the material being dried must be fluidizable. Some potential feedstocks are too wet to fluidize satisfactorily. This is usually due to an excessive amount of surface moisture on the particles, causing them to agglomerate. This problem may be overcome by flashing off the surface moisture in a pneumatic conveying dryer preceding the fluid bed dryer. Another limitation is encountered if the product has a very wide size distribution, so that at an air velocity high enough to fluidize the large particles there is an unacceptable loss of small particles from the bed. To some extent this limitation can be overcome nowadays by the use of a vibrating fluid bed (see below).
The potential user of fluid bed drying has a wide variety of equipment to choose from. The main categories are described briefly below. Within each category there are many variations offered by equipment vendors.
10.2.1 Batch dryers
The first distinction to be made is between batch and continuous fluid dryers.
Batch dryers are normally used when the production scale is small and several different products have to be made on the same production line. Figure 10.1
Dried product
-Fuel Air
10.2.2 Continuous 'well-mixed' dryers
The first continuous fluid bed dryer was the 'well-mixed' type, illustrated in Fig. 10.2, which was introduced in the United States in 1948. It is usually of circular cross-section, and takes its name from the fact that the particle residence time distribution approaches the perfect mixing law:
E(t)
=
~exp (- t~ )In this expression, E(t)dt is the fraction· of particles with residence times betweent
+
dt, andtRis the mean residence time. Because of the near-perfect mixing the bed has a nearly uniform composition and temperature equal to the composition and temperature of the outlet product stream. Hence, the moist feed falls into a bed of almost dry particles. For this reason this type of continuous fluid bed dryer can handle wetter feedstocks than can other types to be described later.The main disadvantage of the 'well-mixed' type of fluid bed dryer is the wide particle residence time distribution, leading to a wide range of moisture content in the product particles. The average moisture content
?f the product may be acceptable, but 40 per cent. of the particles stay In the dryer for less than half the average residence time and 10 per cent.
for less than one-tenth of the average (see Chapter 5, Section 5.6). Hence, some particles will emerge quite wet. For some products, particularly many polymers, this is unacceptable. Furthermore, the wide distribution makes it difficult to achieve a very low average product moisture content.
Despite this drawback, the 'well-mixed' bed is still the most popular type of continuous fluid bed dryer in North America. In Western Europe it has been superseded in many applications by the 'plug flow' bed or by the vibrated fluid bed.
shows a typical batch fluid bed dryer. The wet feed is loaded into the cabinet and clamped to the filter sock module. The cabinet doors are then closed and the blower started. An adjustable damper controls the degree of air recirculation. The circulating air may be heated by a steam tube battery or by gas firing. Batch fluid bed dryers are particularly popular in the pharma-ceutical and dyestuffs industries.
Adjustable weir
In addition to reducing the spread of product particle moisture contents, a plug flow bed will normally require a smaller bed volume than a well-mixed bed to achieve the same average product moisture content. However, the simple plug flow bed has some disadvantages. Firstly, the moist feed falls into an area where the particles are still comparatively wet. Consequently, there may be fluidization difficulties at the feed end with some feeds which could be handled quite satisfactorily in a well-mixed bed. Secondly, the hot air passing through the bed towards the discharge end does comparatively little drying and therefore does not give up much of its heat. Hence, the thermal efficiency of a simple plug flow bed is lower than that of a well-mixed bed. Thirdly, the temperature of the particles rises as they flow along the bed and towards the discharge end it approaches the inlet air temperature. With a heat-sensitive product this limits the inlet air temperature which can be used, thereby reducing the thermal efficiency still further. Variations on the simple plug flow bed have been developed to overcome these difficulties. The most important of these variations are the vibrated fluid bed and the multi-stage bed.
10.2.3 Continuous 'plug flow' dryers
These are beds of shallow depth (typically 0.1 m) in which the particles flow along a channel whose length/width ratio is much greater than unity (typically in the range 4:1 to 30:1). The objective is a more or less close approach to plug flow of the particles, so that they all emerge with approximately the same moisture content. For length/width ratios up to about 10 the particle flow channel may be straight, but for higher ratios a reversing path or spiral path is more practicable as shown in Fig. 10.3. Straight channel designs are often provided with baffles normal to the direction of particle flow in an attempt to improve the approach to plug flow.
Exhaust -air to
cyclone
I
Driedproduct
10.2.4 Vibrated fluid bed dryers
This is simply a straight channel plug flow bed with a vibrating distributor, as shown in Fig. 10.4. Alternatively, it may be viewed as a vibrating fluid bed
Flexible Z:=couplings
Plan view Feed--BL....
~G_
Product(bl
Feedl' ~
Plan view _
_______ ~I
ProductF = feed P= product
conveyor which uses hot air as the fluidizing medium. Compared to the simple plug flow bed, it has the advantage that any agglomerates arising at the feed will be kept moving by the vibrations of the distributor, hopefully until they have dried sufficiently to break up.
Furthermore, feeds with a wide size distribution can be processed successfully in this type of bed. The air velocity can be set low enough to avoid excessive elutration of the smaller particles while the largest particles are kept moving by the vibration of the distributor.
Figure 10.3 Continuous 'plug flow' fluid bed dryer: (a) straight path; (b) reversing path; (c) spiral path.
Finally, these beds are often used to dry feeds consisting entirely of large particles with a minimum fluidization velocity of the order of 1 m/s. With a static distributor an air velocity considerably in excess of the minimum would be needed to ensure adequate fluidization and conveying, but this would be far more air than is required to satisfy mass and heat balance considerations.
If the distributor is vibrated the air velocity can safely be kept in the vicinity of the minimum fluidization velocity, with consequent savings in capital and operating costs.
Present technology imposes two limitations on vibrated fluid bed dryers.
Firstly, the materials of construction available for vibration mountings limit air inlet temperature to about 400°C. Secondly, if the bed is to vibrate at its natural frequency, which is the simplest and most economical way of operating, bed lengths are limited to about 8 m. Therefore, vibrated fluid bed dryers are limited to intermediate temperatures and intermediate through-puts. Finally, they have a higher maintenance requirement than static fluid beds.
(b) To combine the ability of a 'well-mixed' bed to handle wet feeds with the ability of a 'plug flow' bed to achieve a comparatively uniform product moisture content, through having a 'well-mixed' section followed by a 'plug flow' section (see Fig. 1O.5b)
Purposes (a) and (b) can be combined in a three-stage unit.
In case (b) the temperature of the air supplied to the 'well-mixed' stage can be higher than that supplied to the 'plug flow' stage because the higher evaporation rate in the first stage will help to keep the bed temperature down.
A higher air velocity may also be used in the first stage to help in rapid dispersion of the wet feed.
1:::-
Feed~tt'---==-.~-===;...~~t~~? =...---1~ld~
Exhaust -air
10.2.6 Fluid bed dryers with internal heating
When a finely divided, heat-sensitive powder is being dried in a fluidized bed there are limitations on both the velocity and the temperature of the inlet air.
Consequently, the rate of heat input from the air per unit distributor area may be quite low, and if the air is the only heat source the distributor area required to perform the drying duty may be large. It can be reduced by supplying part of the heat through steam tubes or heated baffles immersed in the bed. For example, if half the heat is supplied in this way the distributor area needed is halved. This can give substantial savings in capital and operating costs.
Heat transfer coefficients from immersed heating surfaces to the bed particles increase with decreasing particle size (see Chapter 9). The converse is true for gas-to-particle heat transfer, when account is taken of the enforced reduction in air velocity with decreasing particle size. Hence, the use of internal heating surfaces becomes more attractive the smaller the particle size.
When internal heating surfaces are used, the bed depth is usually determined by the need to keep the heat transfer surfaces immersed, rather than by consideration of drying kinetics.
10.2.5 Multi-stage beds
Multi-stage beds are often used for the following purposes:
(a) To accomplish drying and cooling in the same vessel, which is usually a straight channel plug flow unit with a divided chamber below the distributor (see Fig. lO.Sa.)
~
Product 10.2.7 Fluid bed granulation
Fluidized beds can also be used for making dry powder from a feed which is a slurry or solution. The feed is sprayed on to the bed, usually with a pneumatic atomizer to give a very fine atomization. The particles in the bed are continually growing, by one or both of two mechanisms. A drop may strike a particle and form a thin layer of liquid on the particle surface, whereupon the layer immediately dries and the particle grows by one layer. This mechanism gives a hard, dense particle structured like an onion. Alternatively, the particle may strike another particle before the layer has dried, in which case the liquid may act as a binder and hold the particles together. This gives a
IWell-mixed I Plug flow lsectian : section
I I
Figure 10.5 Multi-stage fluid bed dryers: (a) dryer plus cooler;
(b) plug flow fluid bed following well-mixed fluid bed (plan view).
product conslstmg of agglomerates of finer particles. Which mechanism predominates depends very much on the materials and operating conditions, and must be ascertained by experiment.
In batch granulation the bed initially contains fine seed particles. The required amount of feed is then added at an appropriate rate. After feed addition is complete, there may be a further period during which the grown particles are thoroughly dried, followed finally by a cooling period when the fluidizing medium is changed from hot to cold air. Batch fluid bed granulation is very popular in the pharmaceutical industry.
In continuous fluid bed granulation, a stream of particles is continually withdrawn from the bed and classified into fines, size, and oversize. The fines and crushed oversize are returned to the bed for further growth, while product of the required size is taken off. Figure 10.6 shows the principal options.
Top spraying
Wet ~R feed~
Submerged feed
10.3.1 Batch drying curves
If a batch of moist powder is dried in a fluidized bed with periodic withdrawal of samples for determination of moisture content, the resulting curve of powder moisture content X (defined as the weight of H20 divided by the weight of dry solid) versus time t will probably look something like the curve of Fig. 1O.7(a). This is conventionalIy divided into two portions, the first
Separated solids~
ar dust Direct~OR
to fluid bed ~
Solids treatment
'-- (size reduction and/ Product
Surface or classification
---""""solids ..,...
OR removal
(w --- _
df'--~I
-Bottom~
x-x.
Figure 10.7 Batch drying curves.
The main potential problem with fluid bed granulation is an operational one. If the ratio of liquid to solid in the bed becomes too high the particles will rapidly form large agglomerates the size of golf balls and fluidization will be lost. The plant must then be shut down and the bed dug out. The operator gets only a few minutes warning of the onset of this 'wet quenching' process.
called the constant rate period and the second called the falling rate period.
The moisture content at the transition between the two periods is called the critical moisture content Xcr' If drying is continued for long enough, X will approach the equilibrium moisture content Xe. For a given material this is a function of relative humidity and temperature. At any point on the curve the amount of removable moisture remaining, X - Xc, is called the free moisture content. It should be noted that in reality Xcr rarely appears as a sharply
defined point on the experimental drying rate curve. There is usually some curvature at the transition from constant rate to falling rate. Xcr is probably best defined as the point at which the forward extrapolation of the constant rate line intersects the backwards extrapolation of the falling rate curve.
The rate of drying - dXldt can be determined at any point by differentiating the X versus t curve. A graph of - dXldt versus the free moisture content X - Xe is an alternative way of representing the drying characteristics of a material (see Fig. 1O.7b).
As a rough approximation, the constant rate period may be regarded as corresponding to the removal of surface moisture from the particles, while the falling rate period corresponds to the removal of internal moisture. Since most materials will not fluidize satisfactorily if there is substantial surface moisture, the constant rate period in fluidized bed drying may be so short as to be unobservable, except under very mild drying conditions. On the other hand, a non-porous material such as sand may show virtually no falling rate period.
for example, Keey, 1978). Hence, if either Kp or hgp is known, Nc can be calculated for any drying conditions.
There is a plethora of data on gas-to-particle heat transfer coefficients in fluidized beds of small particles of diameter less about 1 mm. This subject is covered in Chapter 9. One of the most frequently used correlations is that proposed by Kothari and discussed in Kunii and Levenspiel (1969):
Nu =0.03Re~3 (10.4)
This suggests that the heat transfer coefficient, and hence the drying rate in the constant rate period, should be proportional to the gas velocity Uraised to the power 1.3. This was confirmed by Mostafa (1977) in experiments on the drying of small particles of silica gel, molecular sieve, and vermiculite in a fluidized bed.
Many fluid bed dryers, particularly those with vibrating distributors, operate on particles of diameter greater than 1 mm. Many less data are available for this size range, and what there are suggest a much lower dependence of drying rate on gas velocity. No generally applicable correla-tion is yet available. Zabeschek (1977) found that with a fluidized bed of aluminium silicate particles the particle-to-gas mass transfer coefficient was roughly proportional to If'.5 when the particle diameter was 2.76 mm, and was almost independent of U when the particle diameter was 4.30 mm.
Subramanian, Martin, and Schliinder (1977) also concluded, from an analysis of other investigators' results, that transfer rates are effectively independent of gas velocity when the particle diameter is much greater than 1 mm.
Equations (10.2), (10.2), and (10.4) have a number of other practical uses.
They enable one to predict the effects of bulk gas temperature, bulk gas humidity, and small particle diameter on the drying rate in the constant rate period.
10.3.2 Constant rate period
In the constant rate period the surface of the particle is wet enough for the layer of air adjacent to the surface to be saturated. Hence, the drying rate is determined by the rate at which vaporized moisture can be transported across the boundary layer surrounding a particle. During this period the tempera-ture of the particle surface remains constant at the wet bulb temperatempera-ture Twb of the air. If Pwb is the partial pressure of vapour at the wet bulb temperature, P is the partial pressure in the bulk air stream, and Kp is the mass transfer coefficient based on partial pressure, the rate of moisture removal Nc per unit particle surface area (called the drying flux) in the constant rate period is given by:
Nc
=
Kp(Pwb - p) (10.2)Since the particle temperature does not rise during this period, all the heat transferred across the boundary layer from gas to particle must be used for evaporation. Hence, an alternative formulation for Eq. (10.2) is:
hgp
NC=T(T- Twb)
10.3.3 Falling rate period
In the falling rate period the rate of moisture migration to the surface of a particle is insufficient to keep the layer of air adjacent to the particle surface saturated. Hence, the drying rate is no longer determined solely by conditions in the boundary layer. It also depends on the pore structure of the material and on the mechanism of moisture migration. There may, in fact, be several simultaneous mechanisms. These include capillary action, vapour diffusion, diffusion along internal surfaces, and, in the case of cellular materials, diffusion across cell walls. The balance between these mechanisms may change as drying proceeds. For example, capillary motion may predominate during the early part of the falling rate period when the pores are relatively full, while vapour diffusion may dominate towards the end when only small pockets of moisture remain in the solid structure. In general, the falling rate where hgp is the gas-to-particle heat transfer coefficient,A is the latent heat
of vaporization and Tis the bulk gas temperature.
For the air-water system the wet bulb temperature of the air is almost equal to the adiabatic saturation temperature, which is readily obtained from psychrometric charts. For organic solvents in air or an inert gas, TWbcan be calculated if the psychrometric ratio of the solvent-gas system is known (see,
curve cannot be predicted a priori and must be determined by experiment.
However, having determined the falling rate curve experimentally at one set of operating conditions, it is often possible to predict approximately how it will change when the operating conditions are changed. This is useful in that it
However, having determined the falling rate curve experimentally at one set of operating conditions, it is often possible to predict approximately how it will change when the operating conditions are changed. This is useful in that it