Fluidized Bed Dryers — Recent Advances

Advanced Powder Technology 19 (2008) 403418www.brill.nl/apt

Review paper

Fluidized Bed Dryers Recent Advances

Wan Ramli Wan Daud

Department of Chemical and Process Engineering, Faculty of Engineering,University Kebangsaan Malaysia, 43600 UKM Bangi, Malaysia

Received 18 January 2008; accepted 7 April 2008

AbstractAlthough industrial fluidized bed dryers have been used successfully for the drying of wet solid particles formany years, the development of industrial fluidized bed dryers for any particular application is fraught withdifficulties such as scaling-up, poor fluidization and non-uniform product quality. Scaling-up is the majorproblem and there are very few good, reliable theoretical models that can replace the expensive laboratorywork and pilot-plant trials. This problem is mainly due to the different behavior of bubbles and mixingregimes in fluidized bed dryers of different size. Simple transformation of laboratory batch drying data tocontinuous back-mixed dryers using the residence time distribution of the solids is insufficient to account forthe complex flow and heat and mass transfer phenomena occurring in the bed. Although time scaling usingtemperature driving forces and solids mass flux for the same change in moisture content in the batch andcontinuous dryers has been successful in predicting moisture content profiles in the continuous dryer at theconstant rate period, it does not take into account solid mixing. Two-phase DavidsonHarrison models havebeen used in modeling of the continuous back-mixed dryer with various degrees of success. On the otherhand, the three-phase KuniiLevenspiel model is seldom used in modeling fluidized bed dryers because it istoo complex to handle. A combination of multi-phase models and residence time distribution could improvepredicting power for back-mixed dryers because this combination takes into account both the bubbles andsolid mixing phenomena. Incremental models were widely used to model continuous plug flow fluidizedbed dryers, but the cross-flow of drying medium has not been sufficiently modeled except by the author.In some incremental models, axial dispersion is modeled using the Peclet number, Pe. A combination ofan incremental model with an axial dispersion and cross-flow model of drying medium would improvepredicting power. Poor fluidization of Geldart group C particles could be improved by the assistance ofexternal means such as vibration, agitation, rotation and centrifugation. Both vibrated and agitated fluidizedbed dryers have been successfully used in industry, but rotating or centrifugal fluidized bed dryers are stillnot available for industrial use. Koninklijke Brill NV, Leiden and Society of Powder Technology, Japan, 2008

KeywordsFluidized bed, modeling, scaling up, vibrating fluidized bed, centrifugal fluidized bed, agitated fluidizedbed, cross flow

* E-mail: [email protected]

Koninklijke Brill NV, Leiden and Society of Powder Technology, Japan, 2008 DOI:10.1163/156855208X336675

404 W. R. W. Daud / Advanced Powder Technology 19 (2008) 403418

NomenclatureA bed area (m2)

E(t) RTD curve () bed porosity ()F dry basis solid flow rate (kg/h)f falling rate drying factor (C)G dry basis gas flow rate (kg/h)h bed height (m)L total length of dryer (m)

mB mass of bed (kg)NV local solids dying rate (kg/h/m2)

NTU number of transfer units ()Nu nusselt number ()Pe Peclet number ()Pr Prandtl number ()

Rep particle Reynolds number ()TG local gas temperature (C)TGI initial gas temperature (C)Ti inlet gas tenmperature (C)To exhaust gas temperature (C)TS local solids temperature (C)

Twb wet bulb temperature (C)t time (s) residence time (s)X dry basis solids moisture content (kg/kg)Xi inlet solids dry basis moisture content (kg/kg)Xo exit solids dry basis moisture content (kg/kg)Y dry basis gas humidity (kg/kg)Yo exhaust gas dry basis humidity (kg/kg)Yi inlet gas dry basis humidity (kg/kg)z distance along bed (m)Z scaling factor ()

1. IntroductionFluidized bed technology has been used in industrial dryers for the drying of wetsolid particles for many years. Fluidized bed dryers have successfully been used fordrying of products such as coal [1], maize [2], paddy [3], coconut [4], biosynthesisproducts [5], chillies [6, 7], nylon [8], bakers yeast [9], black tea [10] and bleachingagents (sodium perborate) [11]. Industrial drying operations require a high rate ofheat and mass transfer and a high rate of solid transport to or from the dryer. Themain advantages of fluidized bed technology in drying application are large contactsurface area between solids and gas, high thermal inertia of solids, good degree of

W. R. W. Daud / Advanced Powder Technology 19 (2008) 403418 405

solids mixing, and rapid transfer of heat and moisture between solids and gas thatshortens drying time considerably without damaging heat sensitive materials. Inaddition, fluidized solid particles can be easily transported into and out of the dryerby gravity (much like a liquid) and transported elsewhere by pneumatic conveyingwith less mechanical equipment.

However, fluidized bed technology also has several serious undesirable charac-teristics that lower fluidization quality and hence fluidized bed performance whichare detrimental to dryer operation. The fluidized bed dryer for drying specific ma-terial could only be scaled-up using empirical pilot-plant data and not by usingmathematical models. Hence, the performance of the fluidized bed drier could notbe reliably predicted a priori before pilot-plant trials are undertaken. This is due tothe lack of reliable mathematical models for fluidized bed dryers. For some solidparticles that belong to Geldart group C and D, fluidization quality is poor and prod-uct quality is not uniform. For more fragile solid particles, attrition or pulverizationproduces fine particles that can easily be entrained in the gas, causing product lostthrough the dust control system. Fine particles also agglomerate and may seri-ously change product size distribution, causing lower drying rates and non-uniformproduct quality. A wider particle size distribution produced by attrition may alsolower fluidization quality and hence lower product quality. The vigorous mixing ofsolid particles especially in back-mixed fluidized bed dryers also induces a widerresidence time distribution (RTD) of the solids going out of the dryer leading tonon-uniform product quality. The operational life of the fluidized bed dryer and as-sociated pipes and vessels are shorter than that of other types of dryers because oferosion of the dryer, pipe and vessel walls by violent particlewall collisions. Inaddition, the higher pressure drop across the bed requires a more powerful com-pressor, and hence higher capital and operational costs.

2. Scaling-up and Modeling of Fluidized Bed Dryers

By far the most serious problem is scaling-up of the fluidized bed dryer becausebubble size remains essentially the same, but flow patterns of solid particles andgas may differ in larger fluidized bed dryers [12]. In addition solids mixing in smallequipment at low gas velocities is due to upwards solids motion in the wake ofrising bubbles, whereas in larger equipment, more vigorous mixing is due to large-scale toroidal circulation patterns, up in the center and down at the wall [12]. SeeFig. 1.

The dispersion or diffusion coefficient as a measure of mixing increases pro-portionally with bed diameter [12]. These limitations require that fluidized beddrying of any material must be tested both at the laboratory and pilot-plant scaleso that proper scale-up procedure is used to evaluate the performance of the indus-trial dryer. It is therefore difficult to design the full-scale fluidized bed dryer fromfirst principles using laboratory data only and predict its performance for any par-ticular application [12]. These limitations render the development of fluidized bed


Figure 1. Different solid and bubble flow patterns in small (left) and large (right) fluidized beds.

dryers for any particular application expensive. Hence, the main impetus for devel-oping a successful fluidized bed dryer model is the savings in development costswhen the model could successfully be used to evaluate the full-scale dryer usingminimum laboratory data without recourse to expensive pilot-plant trials.

An enduring difficulty is the problem of estimating heat and mass transfer coeffi-cients with sufficient accuracy. Most heat transfer correlations are based on data forturbulent flow around stationary spheres, whereas the motion of bubbles and wakesas well as the circulating flow of solids in larger fluidized beds creates a very differ-ent gas flow patterns around the solid particles. An old but effective correlation thatcan estimate the heat transfer coefficient more accurately than others is the Ranzand Marshall correlation [13, 25]:

Nu = 2 + 0.6Re0.5p Pr0.33. (1)The other serious difficulty is to model the solid and gas mixing in the fluidized

bed, and the interaction between the solid and the gas phase. There are two maintypes of continuous fluidized bed dryers based on RTD: the back-mixed continuousfluidized bed dryer with a broad residence time (Fig. 2) and the plug flow continuousfluidized bed dryer with a relatively narrower RTD (Fig. 3). Mixing regimes inactual industrial fluidized bed dryers may lie between the two.

Both back-mixed and plug flow fluidized bed dryers have been used in the dry-ing of particulate solids for many years. However, modeling and simulation of theperformance of any particular design of the dryer has always been problematic dueto the inadequate overall empirical models used, which are often too inflexible andtoo specific to the particular design, and the inadequate theoretical models that en-counter difficulty in modeling the cross-flow of solids and gas. Black box empiricalmodels have been used for drying of specific products, but the models are not gen-eral enough to be used for scaling-up of other fluidized bed dryers drying differentmaterials.


Figure 2. Back-mixed fluidized bed dryer.

Figure 3. Plug flow fluidized bed dryer.

2.1. Back-Mixed Continuous Fluidized Bed Models

There are three main types of models of back-mixed continuous fluidized bed dry-ers:

(i) Simple empirical model.(ii) Transformation of batch drying data using RTD:

(a) Solids mixing model.(b) Time scaling of driving forces model.

(iii) Transformation of single-particle heat and mass transfer using two- or three-phase fluidized bed models:

(a) One-dimensional two-phase model with and without solids mixingmodel.

(b) Three-dimensional two-phase continuum model.(c) One-dimensional three-phase model.


The simple empirical models can only be used for specific designs of fluidizedbed dryers that dry specific products [3].

The back-mixed continuous fluidized bed dryer has a relatively small aspect ra-tio (length to width ratio) and the mixing of solids is similar to that in a continuousstirred tank reactor. The RTD of the solids particle is therefore quite broad, leadingto non-uniform final solids moisture content. A very popular approach is to directlytransform batch data for use in the continuous model by integrating the batch flu-idized bed drying curve with the RTD of the solids to calculate the average solidsmoisture content at the outlet of the continuous dryer [10, 1421]:

Xo = 1

0

X(t)E(t)dt. (2)

This approach presupposes that the conditions in the original batch drying aresimilar to those in the continuous back-mixed fluidized bed dryer. This assumptionis, however, erroneous because the flow patterns of solids and gas in the batch andcontinuous dryers are very different. This error often leads to overestimation ofdryer performance.

Other researchers who develop similar models rectify this error by assumingthe temperature is uniform throughout the bed and scaling the temperature drivingforces with the solid flow rates of batch and continuous drying for the same changein moisture content [2225]:

Fast drying materials Z = 21

= (mB/A)2G1(TGI Twb)1(mB/A)1G2(TGI Twb)2 , (3)

Slow drying materials Z = 21

= (TGI Twb)1(TGI Twb)2 . (4)

This method was further extended by including the effect of mass transfer in theform of number of transfer units for the batch and continuous dryers in the scalingfactor [25]:

Z = 21

= (mB/A)2G1(TGI Twb)1(1 efNTUz)1

(mB/A)1G2(TGI Twb)2(1 efNTUz)2 . (5)The RTD of solids and the heat and mass transfer between the gas and solid

phases are combined in a latest extension of this model [2729].It has been suggested that the model for the solids phase should be based on

heat and mass transfer on a single particle that could then be scaled-up to a batchor continuous back-mixed fluidized bed dryer using the two- or three-phase modelwithout having to transform batch drying data for continuous application [3032].Another approach for non-bubbling fluidized bed dryers is to model the gas flowthrough the bed using Darcys Law and the interparticle heat and mass transferbetween stationary particles [33]. However, since industrial fluidized bed dryers usebubbling fluidized beds, there is limited use of this model. The steady-state, one-dimensional two-phase models (suspension of particles and bubbles) of Harrison


Figure 4. Schematic of a two-phase model of a fluidized bed.

and Davidson [3436] were widely used successfully [3740]. A more elaborateone-dimensional two-phase model that also incorporates the RTD of solids was alsodeveloped [29, 41]. The three-phase model (cloud, emulsion and bubbles) of Kuniiand Levenspiel [42, 43] was never used successfully for fluid bed dryers because ofits complexity. See Fig. 4.

The unsteady- and steady-state three-dimensional two-phase continuum modelwas developed [44, 45] by remodeling the solids phase as a solids-only contin-uum and applying fluid-like differential mass and momentum balance on the solidscontinuum and solving the resulting set of equations along with the differential mo-mentum balance of the gas phase continuum by using computational fluid dynamictechniques of the constant volume method. The model has been successfully ap-plied to fluidized bed reactors in the petrochemical industry. However, although thecomputed solids temperature was in close agreement with experimental values, thesolids moisture content deviated widely from experimental data.

Both the two- and three-phase models estimate the performance of back-mixedcontinuous dryers better than both the RTD-batch dryer transformation model andthe time scaling of batch to continuous dryer model. The one-dimensional two-phase model with RTD can estimate the dryer performance even better. Howeverthe three-dimensional continuum two-phase model was promising, but the largedeviations between the computed and experimental solids moisture content weredisappointing. The three-phase model is not often used in fluidized bed dryer mod-els because of the many parameters required by the model.2.2. Plug Flow Fluidized Bed Dryer ModelsThere are two main types of models for plug flow fluidized bed dryers:

(i) Simple black box model.


(ii) Transformation of batch drying data:(a) Discrete solids mixing model.(b) Incremental plug flow model.(c) Incremental plug flow model with dispersion using the Peclet num-

ber, Pe.(d) Incremental plug flow model with gas cross-flow.

Most of the existing models of plug flow fluidized bed dryers concentrate onlyon the solid mixing model and neglect the cross-flow pattern of the drying medium.The non-ideal flow of solids in plug flow fluidized bed dryers has been modeled asseveral continuous back-mixed fluidized bed dryers connected in several differentways [10, 23, 4648] that use models originally validated for back-mixed batch andcontinuous fluidized bed dryers [10, 2022] similar to the modeling of non-idealflow reactors. The problem with this type of model is the difficulty of estimatinga sufficient number of completely mixed dryers for any particular dryer a priori. Itcould only be determined by analysis of pilot-plant data. The position in the dryeris transformed from the residence time using similar arguments to the incrementalmodels. The essential feature of the cross-flow of drying medium and solids in plugflow fluidized bed dryer was not modeled at all.

Many incremental models of plug flow fluidized bed cross-flow dryers had tointroduce a time increment in the model that is later substituted with the lengthincrement by using the bed velocity in order to track the changes in moisture contentand temperature of solids and air as functions of distance along the bed [15, 16, 26,49, 50]. The use of the time increment in a steady-state model is confusing. Thecross-flow dryer was modeled as a series of incremental step-wise co-current dryers[50]. See Figs 5 and 6.

The incremental model was modified by modeling the cross-flow of the dryingmedium and removing the need to use the time increment and bed velocity forcomputing the distance along the dryer [51, 52]. To account for non-ideal flow,an axial dispersion term was introduced in the moisture balance of the solids andcharacterized the dispersion with the Peclet number, Pe [15].

A similar plug flow fluidized bed dryer model was developed that included a mo-mentum balance of the particles [53] based on the back-mixed batch and continuousfluidized bed dryer model [29]. The essential feature of the cross-flow of drying

Figure 5. Moisture and energy balance in cross-flow dryers as stepwise co-current flow [50].


Figure 6. Moisture and energy balance in cross-flow dryers [51, 52].

medium and solids in plug flow fluidized bed dryers was not modeled. If the in-cremental model with axial dispersion is combined with the cross-flow model ofdrying medium, the predicting power of the model will be enhanced.

3. Improvement of Fluidization Quality of Difficult to Fluidize Particles3.1. Mechanically Assisted Fluidization

It is well known that fluidization quality of group A and B particles in the Gel-dart classification is excellent. However, the fluidization quality of group C (fineand ultra fine particles) as well as group D (large particles) is comparatively poor.In addition to natural group C particles, vigorous solids mixing, energetic bubbleeruption at the bed surface, and violent solidsolid and solidwall collisions leadto the formation of many fine group C particles in the system through attritionor pulverization of particles. This causes unsatisfactory contact efficiency betweensolid and fluid phases, and a drop in the performance of the fluidized bed dryer.These effects could be reduced and the performance of the fluidized bed dryer im-proved by installation of inner vertical baffles that break up larger bubbles beforethey break through the surface [55]. Fluidization quality of fine and ultra-fine par-ticles (group C particles) is poor because of the strong cohesive forces betweenfine particles. As a result, channeling in the bed of particles and agglomeration offines frequently occur. In addition, entrainment of fine particles is severe becauseof the slower terminal falling velocity of fine particles. However, the fluidizationquality of fine particles can be enhanced by the assistance of external means suchas vibration, agitation, rotation and centrifugation [56].3.2. Vibrated Fluidized Bed

Vibrated fluidized bed dryers were first studied in 1969 as an alternative to the con-ventional fluidized bed dryer [57]. Vibrated fluidized beds are often associated withspray drying systems to produce agglomerated instant powder or act as a powder


conditioning plant for special coating operations. Vibration and the upward and for-ward flow of air in a fluidized bed enables the drying particles to fluidize smoothly[39]. This is especially useful for processing group C particles [54]. Vibration alsoassists fluidization by creating large tensile stresses during transient periods of thecyclic pressure pulses [58, 59] and inducing higher particle accelerations [60] thatbreak up cohesive assemblies of group C particles into smaller agglomerates. Inaddition, the gas velocities for minimum and complete fluidization are relativelylower compared to the conventional fluidized bed dryer [61]. Attrition due to vigor-ous particleparticle and particlewall interactions is thus minimized appreciably.Hence, the application of fluidized beds can be further extended to fragile, abra-sive and heat-sensitive materials [56]. As the operating gas velocity is lower than ina conventional fluidized bed, the problem of fine particles entrainment is avoided.

For a polydisperse system, low gas velocity fluidizes the fine particles gentlywhile the vibration keeps the coarse particles in a mobile state. Heat and masstransfer occur more effectively and efficiently. A vibrating fluidized bed dryer istherefore useful for drying materials that are fine, poorly fluidizable, have a broadparticle distribution, have a highly irregular particle shape or require relatively lowfluidization velocities to prevent attrition [56]. Vibrating fluidized bed dryers aregenerally of the plug flow type with shallow beds. This gives a much lower productresidence time per unit bed area than non-vibrating beds [62]. Vibration increasesparticle diffusivity, constant drying rate [6264] and falling drying rate [6567].Drying rate also increases with increasing amplitude and gas superficial velocity[59, 64]. However, the understanding of transport processes in the vibrated flu-idized bed dryer is still little known even 20 years after the first effort was started[68]. A LangragianEuler model of the vibrated fluidized bed dryer was developedusing the discrete element model for the particles and the NavierStokes equationfor the gas [59], but the model was validated using published experimental bed ex-pansion data [69] only, which is not adequate. More work has to be done in orderto extend our understanding of the phenomena. See Fig. 7.

Figure 7. Vibrating fluidized bed.


Figure 8. Agitated fluidized bed.

3.3. Agitated Fluidized BedAnother way to improve the fluidization quality of fine particles is to use agitatingdevices within the fluidized bed dryer. By agitating the bed of particles, a homo-geneous fluidizing bed of fine particles is formed without any channeling or bigbubbles [56]. In addition, deeper bed depth is possible if the bed is agitated whileits fluidization quality is maintained [56]. Mechanical agitation improves fluidiza-tion by reducing agglomeration and channeling, as well as increasing void fraction[70]. See Fig. 8.

The drying rate of agitated fluidized bed dryers increases with the rate of ag-itation, reaches a maximum and then decreases as the rate of agitation is furtherincreased [71]. This is because at low agitation speed, agitation and fluidizationincrease interparticle and gasparticle contact, while at higher agitation speed, thegreater centrifugal force hurls more particles towards the wall away from the mainair flow at the center. The greater interparticle and gasparticle contact at loweragitation speeds increases both heat and mass transfer, and the drying rate. On theother hand, the lower interparticle and gasparticle contacts at agitation rates higherthan that at the maximum drying rate, decreases both heat and mass transfer, andthe drying rate. An earlier work that found drying rate decreases with increasingagitation rate operated the agitated fluidized bed beyond the maximum drying ratepoint [72].3.4. Centrifugal and Rotating Fluidized BedThe centrifugal or rotating fluidized bed balances the centrifugal force generatedby chamber rotation with the particle drag force caused by the radial fluidizationgas [73]. The fluidization gas velocity can be easily varied by varying the rota-


Figure 9. Centrifugal/rotating fluidized bed.

tional speed of the fluidization chamber [74]. The centrifugal force is generatedby rapidly rotating the fluidization chamber to form an annular particle bed thatis radially fluidized by gas injected through the porous or sintered outer wall [75]and removed via a central chimney. The annular fluidized bed is therefore morecompact than conventional fluidized beds. The centrifugal force can be increasedby several fold of gravity to increase fluidization and gassolid slip velocities, toimprove inter-phase mass and heat transfer through good contact efficiency, and toprevent agglomeration and entrainment of particles [7678]. Recently, models ofrotating fluidized beds were proposed for the bubble size [79] as well as for thewhole rotating fluidized bed [80]. A novel rotating fluidized bed formed by inject-ing gas tangentially in a static chamber was recently invented [81]. See Fig. 9.

4. Conclusions

The main problem of industrial fluidized bed dryers is scaling up because there arevery few good, reliable theoretical models that can replace expensive laboratorywork and pilot-plant trials. The models must take into account both the behaviorof bubbles and mixing regimes in fluidized beds of different size. Simple trans-formation of laboratory batch drying data to continuous back-mixed dryers usingRTD of the solids is insufficient. Time scaling using temperature driving forces andsolids mass flux for the same change in moisture content is successful only in cer-tain cases. Multi-phase models such as the two-phase DavidsonHarrison modeland three-phase KuniiLevenspiel model have been used in modeling continuousback-mixed dryers but combinations of these models and RTD could improve thepredicting power. Most incremental models of continuous plug flow fluidized beddryers with and without axial dispersion neglect the cross-flow of drying medium.A combination of an incremental model with an axial dispersion and cross-flowmodel of drying medium would improve the models predicting power. Poor flu-idization of Geldart group C particles could be improved by the assistance ofexternal means such as vibration, agitation, rotation and centrifugation in novel


fluidized bed dryers. Models of these novel fluidized bed dryers are not well devel-oped because the theoretical understanding of them is still far from complete. Bothvibrated and agitated fluidized bed dryers have been successfully used in industry,but rotating or centrifugal fluidized bed dryers are still not available for industrialuse.

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