Advanced Powder Technology 25 (2014) 1699–1708

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Advanced Powder Technology

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Original Research Paper

Hydrodynamic studies on fluidization of Red mud: CFD simulation

http://dx.doi.org/10.1016/j.apt.2014.06.0170921-8831/� 2014 The Society of Powder Technology Japan. Published by Elsevier B.V. and The Society of Powder Technology Japan. All rights reserved.

⇑ Corresponding author. Tel.: +91 9438600372 (M); fax: +91 6612472926.E-mail address: pranatisahoo02@gmail.com (P. Sahoo).

Pranati Sahoo ⇑, Abanti SahooChemical Engg. Department, National Institute of Technology, Rourkela 769008, Odisha, India

a r t i c l e i n f o a b s t r a c t

Article history:Received 13 February 2014Received in revised form 12 June 2014Accepted 22 June 2014Available online 3 July 2014

Keywords:FluidizationFine Red mud particlesHydrodynamic studiesEulerian modelCFD analysis

Hydrodynamic studies are carried out for the fluidization process using fine i.e. Geldart-A particles.Effects of superficial velocity on bed pressure drop and bed expansion is studied in the present work.Commercial CFD software package, Fluent 13.0 is used for simulations. Red mud obtained as wastematerial from Aluminum industry having average particle size of 77 microns is used as the bed material.Eulerian–Eulerian model coupled with kinetic theory of granular flow is used for simulating unsteadygas–solid fluidization process. Momentum exchange coefficients are calculated using the Gidaspow dragfunctions. Standard k–e model has been used to describe the turbulent pattern. Bed pressure drop andbed expansion studies are simulated by CFD which are explained with the help of contour and vectorplots. CFD simulation results are compared with the experimental findings. The comparison shows thatCFD modeling is capable of predicting the hydrodynamic behaviors of gas–solid fluidized bed for fineparticles with reasonable accuracy.� 2014 The Society of Powder Technology Japan. Published by Elsevier B.V. and The Society of Powder

Technology Japan. All rights reserved.

1. Introduction

Huge amounts of Red mud, nearly 1.5 tons per 1 ton of Aluminaare produced in any Aluminum plant. With increasing demand forproduction of Aluminum, more and more Red muds are being pro-duced daily for which its disposal is now a matter of great concern.That is why researchers are coming forward to know more aboutthe Red mud so that it can be utilized for useful purpose. It isalready tested by many researchers that Red mud can be usedfor making bricks, flooring materials and extraction of some noblematerials etc. Researchers is still going on to explore more aboutRed mud for which it is planned to know the fluidization character-istics of Red mud before using for any application. Fluidizationcharacteristics of bed material and hydrodynamic studies of thefluidizer are very much inter-related. Hydrodynamic studies giveprior information about the flow behavior along with the changesoccurring on the bed materials.

Literature revels that superficial velocity plays important role influidization process [1]. To know the effect of superficial velocityon fine particles, the present work is planned to be studied bothexperimentally and with the help of CFD simulation. At high gasvelocity the movement of bed materials is more vigorous whichresults in bubbling fluidized bed. With further increase in gasvelocity turbulent fluidization results and entrainment of solids

becomes appreciable. With further increase in gas velocity fast flu-idization is resulted. Therefore it is essential to understand thehydrodynamics of fluidized bed quantitatively as well as qualita-tively by which one can design fluidization unit properly. Selectionof the correct operating parameters is very much essential for theappropriate fluidization regimes [1–3]. Gas–solid flows inside thecolumn are quite complex because of the coupling of the turbulentgas flows and fluctuation of particle motions dominated by inter-particle collisions. These complexities lead to considerable difficul-ties in designing, scaling up and optimizing the operation. Thefundamental problem encountered in modeling hydrodynamicsof a gas–solid fluidized bed is the motion of two phases wherethe interface is unknown and transient, and the interaction isunderstood only for a limited range of conditions [1,2]. Computa-tional fluid dynamics (CFD) offers a new approach to understandthe complex phenomena between the gas and the solid particlesin the bubbling fluidized beds. Two different CFD models namelyEulerian–Lagrangian and Eulerian–Eulerian model are applied formodeling gas–solid fluidized beds. The Eulerian–Eulerian modelis considered to be more appropriate for fluidized beds [4] forwhich this model is preferred over Eulerian–Lagrangian model.

In Eulerian–Eulerian model, all phases are considered to becontinuous and fully interpenetrating. The equations employedare the generalization of the Navier–Stokes equations. One set ofthe mass and momentum conservation equations are solved foreach phase, where the momentum equations are linked by aninterphase exchange term. By comparing Eulerian–Eulerian and

Nomenclature

d Diameter (m)e Volume fractionq Density of fluid (kg/m3)u Velocity (m/s)p Pressure (Pa)s Stress–strain tensor (Pa)g Acceleration due to gravity (m/s2)F Forcel Viscosity (Pa S)Kgs The fluid–solid and Solid–solid exchange coefficientI2D Second invariant of the deviatoric stress tensorRe Reynolds numberess Coefficient of restitutiong0,ss Radial distribution co-efficientCD Drag co-efficientCfr,gs Coefficient of friction of solid phase particles

Hs Solid phase granular temperature (m2/s2)ls Solid shear viscosity (Pa S)ls,col Collision viscosity (Pa S)ls,kin Kinetic viscosity (Pa S)ls,fr Frictional viscosity (Pa S)ks Bulk viscosity (Pa S)/ Angle of internal friction (�)KHs Diffusion co-efficient�Hs Collisional dissipation of energy (kg/s3 m)Ugs Energy exchange of solid phase (kg/s3 m)g Rate exponent~vs Phase-weighted velocity» Gradient

1700 P. Sahoo, A. Sahoo / Advanced Powder Technology 25 (2014) 1699–1708

Eulerian–Lagrangian approaches [5] it can be concluded thatEulerian–Eulerian approach is potentially faster than theEulerian–Lagrangian method which requires the formulation ofconstitutive equations. Gas–solid simulations of bubbling fluidizedbeds obtained by Eulerian–Eulerian model are verified experimen-tally with the existing correlations for bubble size or bubblevelocity [6,7]. A computational study for the flow behavior of alab-scale fluidized bed is also carried out by Chiesa et al. [8]. Theresults obtained from a ‘discrete particle method’ (DPM) are qual-itatively compared with that of a multi fluid computational fluiddynamic (CFD) model by them. Hydrodynamic behaviors of nonre-active gas–solid fluidized bed reactor are investigated by usingmulti fluid Eulerian model where the effects of particle size andsuperficial gas velocity are studied by researchers [9,10]. Hydrody-namics of a gas–solid tapered fluidized bed with CFD simulationsare reported [11]. Simulation for minimum fluidization velocity,bubbling velocity and slugging velocity are studied for four typesof Geldart particles by Labview method [12].

Thus it is aimed to carryout CFD simulation for hydrodynamicstudies of bubbling gas–solid fluidized bed using fine particles inthe present work. For CFD simulations, it is planned to develop atwo fluid Eulerian model coupled with the kinetic theory granularflow for 77 lm sized solid (i.e. Red mud) particles. The standardk–e turbulence model is thought to apply to simulate the gas–solidflows at different superficial gas velocities. The effects of modelingparameter i.e. inlet air velocity on hydrodynamics of bubbling gas–solid fluidized bed are planned to be studied both experimentallyand computationally. It is also planned to carry out comparisonsof the model predictions and experimental measurements on thetime-averaged bed pressure drop, bed expansion and qualitativegas–solid flow pattern have been carried out at different operatingconditions in order to arrive at a conclusion.

1.1. Governing equation

The governing equations for the gas–solid flow include the con-servation of mass and momentum [13]. CFD simulation parametersare given in Table 1(a–c) where model equations, mesh size, timestep, convergence criteria, discretization method, geometry andboundary conditions are listed.

The governing equations of solid and gas phases based on theEulerian–Eulerian model are used for CFD simulation. The volumefractions of the phases sum to one i.e.

eg þ es ¼ 1 ð1Þ

The continuity equation for gas and solid phases in the absenceof interphase mass transfer are expressed as

@

@tðegqgÞ þ rðegqgugÞ ¼ 0 ð2Þ

@

@tðesqsÞ þ rðesqsusÞ ¼ 0 ð3Þ

The conservation of momentum for the gas and solid phases aredescribed by

@

@tðqgegugÞ þ rðqgegugugÞ ¼ �egrpþrsg þ qgegg þ Fi;g ð4Þ

@

@tðqsesusÞ þ rðqsesususÞ ¼ �esrp�rps þrss þ qsesg þ Fi;s ð5Þ

Where the terms Fi;g and Fi;s of the above momentum equationsrepresent the interphase momentum exchange for gas phase andsolid phase respectively. Thus the gas–solid interphase drag forceis expressed as

FD;gs ¼ Kgsðug � usÞ ð6Þ

Here g is gravitational constant and (�esrpþ Kgsðug � usÞÞ is aninteraction force (drag and buoyancy forces) representing themomentum transfer between gas and solid phases.

The terms sg and ss are the stress–strain tensors for gas andsolid phases respectively. They are expressed as follows.

sg ¼ eglgðrug þruTg Þ þ eg kg �

23lg

� �rugI ð7Þ

ss ¼ esls rus þruTs

� �þ es ks �

23ls

� �rusI ð8Þ

Here I is unity tensor (dimensionless).In the present study Gidaspow model has been chosen for sim-

ulation which is a combination of Wen and Yu model and theErgun equation. The fluid solid exchange coefficient, Kgs isexpressed in the following form.

1.1.2. Gidaspow drag model

When eg > 0:8; Kgs ¼34

CDesegqg jug � usj

dpe�2:65

g ð9Þ

Table 1CFD simulation parameters with base case settings.

Description Base case setting Change compared to base case settings

Model equationsKinetic viscosity Syamlal and O’Brien [15] Fixed valueGranular bulk viscosity Lun et al. [17] Fixed valueFrictional viscosity Schaeffer [16] Fixed valueAngle of internal friction 30� Fixed valueGranular conductivity Syamlal and O’Brien [15] Fixed valueDrag law Gidaspow [2] Fixed valueCoefficient of restitution for particle–particle collisions 0.90 Fixed value

Mesh size, time step, convergence criteria and discretization methodMesh resolution 0.002 m � 0.002 m grids Fixed valueConvergence criteria 10�3 Relative/fixed valueMaximum iterations 30 Fixed valueDiscretization method First order upwind Fixed valueTime step 0.001 s Fixed value

Geometry, boundary, initial and operating conditionsBed width 12 cm Fixed valueBed length 70 cm Fixed valueInitial bed height 10 cm Fixed minimum valueInitial solids packing (minimum voidage) 0.80 Fixed valueOutlet boundary condition Pressure outlet Fully developed flowWall boundary condition No slip condition –Gravitational acceleration 9.81 m/s2 Fixed valueOperating pressure 1.013 � 105 Pa Fixed valueSuperficial gas velocity 0.01 m/s 0.008 m/s,0.012 m/s, 0.014 m/s,0.016 m/s, 0.018 m/sInlet boundary condition Uniform velocity inlet Fixed value

P. Sahoo, A. Sahoo / Advanced Powder Technology 25 (2014) 1699–1708 1701

When eg � 0:8 Kgs ¼ 150esð1� egÞlg

egd2p

þ 1:75qgesjug � usj

dp

ð10Þ

where CD ¼24

egRep½1þ 0:15ðegRepÞ0:687� ð11Þ

The particle Reynolds number is defined as follows

Rep ¼qgdpjug � usj

lgð12Þ

1.2. Constitutive equations

Constitutive equations are required to close the governing rela-tions. The Constitutive equations are expressed as follows.

1.2.1. Solid shear stressesThe solid shear stresses contain shear and bulk viscosities

arising from particle momentum exchange due to translation andcollision.

ls ¼ ls;col þ ls;kin þ ls;fr ð13Þ

Where collision viscosity [14] is given as

ls;col ¼45esqsdsgo;ssð1þ essÞ

Hs

p

� �1=2es

ð14Þ

Kinetic viscosity [15] is expressed as

ls;kin ¼esdsqs

ffiffiffiffiffiffiffiffiffiffiHspp

6ð3� essÞ1þ 2

5ð1þ essÞð3ess � 1Þesgo;ss

� �ð15Þ

And Frictional viscosity [16] is defined as

ls�fr ¼ps sin /

2ffiffiffiffiffiffiI2Dp ð16Þ

1.2.2. Bulk viscosityThe bulk viscosity accounts for the resistance of the granular

particle to compression and expansion [17].

ks ¼43esqsdsgo;ssð1þ essÞðHspÞ

12 ð17Þ

1.2.3. Solid pressureFor granular flow in the compressible regime where the solid

volume fraction is less than its maximum allowable value, a solidpressure is calculated independently. The solid pressure [17] iscomposed of a kinetic term and a secondary term due to particlecollisions. It is expressed as

ps ¼ esqsHs þ 2qsð1þ essÞe2s g0;ssHs ð18Þ

1.2.4. Radial distribution functionThe radial distribution function go is a correction factor [18]

that modifies the probability of collision of solid granular particles.It is expressed as for a one solid phase

go ¼ 1� es

es;max

� �1=3" #�1

ð19Þ

1.2.5. Granular temperatureThe granular temperature for solids phase is proportional to the

kinetic energy of random motion of particles. The transport equa-tion derived from kinetic theory takes the following form.

32

@

@tqsesHsð Þ þ r � ðqses~v sHsÞ

� �¼ ð�psI þ ssÞ

: r �~v s þr � ðKHsr �HsÞ �!Hs þUgs ð20Þ

where KHs is the diffusion coefficient for granular energy [15]. It isexpressed as

KHs ¼15dsqses

ffiffiffiffiffiffiffiffiffiffiHspp

4ð41�33gÞ 1þ125

g2ð4g�3Þesg0;ssþ16

15pð41�33gÞgesg0;ss

� �ð21Þ

Where; g ¼ 12ð1þ essÞ

The collision dissipation of energy [17] is denoted as YHs and isdefined as

1702 P. Sahoo, A. Sahoo / Advanced Powder Technology 25 (2014) 1699–1708

YHs ¼12ð1� e2

ssÞg0;ss

dsffiffiffiffipp :qse

2s H

3=2s ð22Þ

The transfer of the kinetic energy is denoted as

Ugs ¼ �3KgsHs ð23Þ

2. Experimentation

The fluidization characteristics of fine particles are studied in afluidizer, a cylindrical column made of perspex material, 12 cminside diameter and 70 cm high. A filter cloth with pores of approx-imately 40 lm is tightly attached at the bottom of column whichacts as the distributor. The calming section is packed with spheri-cal glass beads of size 5 mm for uniform distribution of gas. Thecolumn is also covered with filter cloth at the top to prevent theentrainment of the particles. The average particle size of bed mate-rials is determined by sieve analysis. The column is filled withknown amount of Red mud up to certain height. Air is suppliedfrom the bottom of the column through the distributor at theambient conditions. Air flow is measured by a Rotameter. A U-tubemanometer is connected to the fluidizer for measuring pressuredrop across the bed. Carbon tetra chloride (CCl4) is used as themanometric fluid. The pressure drop and the expanded bed heights(maximum and minimum heights within which the bed fluctuates)are noted against different air flow rates. Details about the equip-ment/instruments used in the experimentation are shown inTable 2.

3. Details of CFD simulations

The CFD software package, FLUENT 13 is used to simulate thegas - solid fluidization process. Red mud particles of 77 micronsin size are used as the bed material to study the hydrodynamicbehavior of the gas–solid fluidized bed. The results of simulationsare compared with the experimental data to check the effective-ness of the model.

Two dimensional (2D) computational geometry of the bed isgenerated by using commercial software GAMBIT as shown inFig. 1(a). Quadrilateral element structure (height to width ratioof 1) is used for meshing the geometry (Fig. 1(b)). In this study,total of 21,154 numbers of cells with each cell size of0.002 m � 0.002 m and 21,570 numbers of nodes are employedfor simulating the fluidized bed. The time step is chosen as 0.001of 1000 steps. The convergence criteria for all the numerical simu-lations are based on monitoring of the mass flow residual and theresidual value is observed to be converging in the range of 1.0 e�03

as shown in Fig. 1(c). The simulation is carried out using differentflow quantities till the system reaches quasi-steady state.

Standard k–e dispersed Eulerian granular multiphase modelwith standard wall functions are used for modeling the transitionnature of bubbling fluidized bed. The success of Eulerian–Eulerian

Table 2Details of instruments used in experimentation.

Sl. No. Instruments/equipments Capacity Company/model

1 Multistage aircompressor

25 kgf/cm2 ELGI equipmentlimited

2 Air Rotameter (made upfiber)

0–10 lpm APEX

3 U tube manometer 0–100 mm –(made up Glass) (CCl4 manometric

fluid)4 Fluidizer 12 cm ID,70 cm High –

(made up Perspex)

approach depends on the proper description of all possible intraand interphase interactions, such as gas–solid interactions, colli-sion and frictional interactions between particles, and interactionsbetween wall and particles. Air is taken as continuous phase whileRed mud solid particles are taken as dispersed phase which aretreated as continua, interpenetrating and interacting with eachother and everywhere in the computational domain.

The assumptions made for CFD simulations are isothermal, non-reactive, unsteady state gas–solid system, no lift force, no masstransfer between gas and solid phase. Constant pressure gradientand constant density of each phase are also assumed in the presentwork. In this work, bubbling fluidization is observed with bedmaterial where viscosity is considered to be negligible.

Initially particle velocity is set at zero (i.e. in minimum fluidiza-tion) and the inlet gas velocity at the bottom of fluidized bed isassumed to be uniform along the axial direction. The pressure isnot specified at the inlet because of the assumption of incompress-ible gas phase i.e. of relatively low pressure drop system. At theoutlet, only pressure boundary condition is specified and no slipboundary conditions are assumed. The phase coupled SIMPLEmethod [19] is chosen for pressure–velocity coupling and firstorder upwind scheme is used for discretization of volume fractionequation whereas second order upwind scheme is used for discret-ization of momentum, turbulent kinetic energy and turbulent dis-sipation rate.

4. Results and discussion

Studying the effect of inlet velocity is very much essential as itplays an important role in fluidization process. In this work, fluid-ization velocities higher than the minimum fluidization velocityand terminal velocity are used for experimentation which impliesthe occurrence of bubbling fluidization. The bubbling behavior isobserved for Geldart-A particles in between the minimum fluidiza-tion and terminal velocity of gas in a gas–solid fluidization processas reported by Kunii and Levenspiel [1].

4.1. Effects of inlet gas velocity on bed dynamics

Experimental analysis is performed to achieve the steady bedpressure drop and expansion ratio at different superficial gasvelocities varying from 0.008 m/s to 0.018 m/s. The hydrodynamicbehaviors of fluidized bed are analyzed by monitoring the contourplots for volume fraction of bed materials, static pressure and fluidvelocity etc. The bed pressure drop and expansion ratios are alsomeasured experimentally for different superficial gas velocities.These outputs are then compared with the CFD simulated results.

Fig. 2 compares the contours of solid volume fraction distribu-tion against gas velocity at different times of simulation. It isobserved that bubbles grow in size as time increases at any partic-ular gas velocity. With increase on time, the volume fraction ofsolid material is observed to decrease indicating the bed expan-sion. It is also seen that the average bed heights increase withincreasing time at any particular inlet gas velocity.

Using Gidaspow drag model, the contours of solid volume frac-tion distribution are obtained for six different gas velocities in therange of 0.008–0.018 m/s (Fig. 3) simulated for a time period morethan 10 s. The results show an increase in bed expansion withincreasing inlet gas velocity. Initially small bubbles are formed atthe distributor which causes the movement of the particles. Withthe increase in gas velocity, bubble size increases as more gasesare processed within the bubble. These bubbles gradually coalesceand convert to slug within the bed. At higher velocity, bubblesgrow larger and consequently the bed expands significantly. It isobserved that at the superficial gas velocity of 0.014 m/s, the bed

Fig. 1. Fluidized bed details.

Fig. 2. Solid volume fractions at different inlet gas velocities from 1 to 10 s simulation time.

P. Sahoo, A. Sahoo / Advanced Powder Technology 25 (2014) 1699–1708 1703

surface is highly fluctuating. These fluctuations may be consideredas the indication of transition from stable fluidization to turbulentfluidization, which agrees with the experimental observations. It isalso observed that when the bed expands a large portion of bedmaterials is pushed toward the wall region as a result a high valueof solid volume fraction is seen near the walls implying the forma-tion of wall slug.

Fig. 4 shows the comparison among the vector plots of solidphase velocity and gas phase velocity against inlet gas

velocities predicted by the CFD different simulation times. Inall cases, the internal circulation of particles is observed tooccur while gas is not found to be distributed evenly. Thecore-annulus structure shows that the solid and gas velocitiesin the core region are much higher than those in the annulusregion while solid velocity and gas velocity near the wall aregreatly decreased. This may be due to the back mixing andinternal circulation which is also observed with the simulationof gas–solid flow.

Fig. 3. Solid volume fraction for different gas inlet velocities.

Fig. 4. Velocity vector plots of solids phase and gas phase.

1704 P. Sahoo, A. Sahoo / Advanced Powder Technology 25 (2014) 1699–1708

Fig. 5(a) shows variation in axial velocity for solid phasesagainst radial positions. Profiles of axial velocity of particles arefound to be smoother in higher gas velocities because of more tran-sition in flow patterns. This transient flow pattern results asmoother velocity profile in the bed after reaching a quasi-steady-state condition. With the increasing inlet gas velocity themaximum axial velocity of particles decreases to certain limit inthe central region but it increases near the wall region. The reasonfor this may be as follows. Bed materials move upward with themovement of rising bubble. When bubble breaks particles are dis-persed in the radial direction. As a result axial velocity decreases

but in the wall region particles movement increases. Thus adecreased value of axial velocity causes the particles near the wallto show particle slip on the wall.

Fig. 5(b) shows the computed gas velocity distribution versusradial distance for different inlet gas velocities. The simulatedresults show that in the central region of the bed, the velocity ofgas is maximum which decreases from the core toward the wallof the bed. With the increase of inlet gas velocity, the energy ofthe high pressure gas is quickly converted to the kinetic energyof particles. With the increase of inlet gas velocity within thebed, the radial distribution of gas velocity becomes more uniform.

Fig. 5. Radial profiles for different gas inlet velocities.

P. Sahoo, A. Sahoo / Advanced Powder Technology 25 (2014) 1699–1708 1705

Decreased velocities for solid particles in gas phase appear near thewalls with increasing inlet gas velocities. This happens due tosevere back mixing in the axial direction of the bed.

Fig. 5(c) and (d) shows the radial profiles for predicted turbu-lent kinetic energy of the gas phase and predicted granular temper-ature distribution of the solid phase respectively. Both the outputparameters are observed to increase significantly with the increasein inlet gas velocity. The lower velocity is observed to give a lowgranular temperature where as a high fluctuating velocity per unitof mass is observed at higher inlet gas velocity. The particle fluctu-ating energy per unit of mass is found to increase from the central

Fig. 6a. Contour of bed pressure drop against air velocity.

region of the bed toward the wall of the bed. At the wall, the gran-ular temperature is found to decrease because of the wall effects.

Fig. 5(e) shows the simulated time-averaged volume fraction ofsolid particles for static bed height of 10 cm at six different inletgas velocities. It is observed that the volume fraction of solid par-ticles increases toward the walls. It is also seen that at the highervelocity the volume fraction of particles increases more near thewalls. At lower velocity of particles the volume fraction remainsnearly same as before. As the inlet gas velocity increases, theparticles tend to accumulate more at the walls than in the centralportion of the fluidized column.

Fig. 6b. Variation in static pressure along the bed height.

Fig. 7a. Variation of bed pressure drop against time.

Fig. 7b. Comparison plot of bed pressure drop.

1706 P. Sahoo, A. Sahoo / Advanced Powder Technology 25 (2014) 1699–1708

4.1.1. Pressure dropVariation in bed pressure drop is mainly due to the gas–solid

interactions during the fluidization process. Fig. 6a shows the con-tours of pressure drop against the gas velocity and static pressurevariations along the bed height with simulation time step of 10 s.The bed pressure drop for a fluidized bed varies from maximumvalue at the bottom of the bed to minimum value at the top ofthe bed. It is evident from this figure that the pressure is maximumat the inlet which decreases gradually and becomes zero at the top.

Fig. 8. Comparison of volume frac

It is also found from Fig. 6b that the higher velocity yields thehigher static pressure as pressure drop is directly proportional tosuperficial air velocity. The pressure drop is found to be minimumin the bed height zone above 0.25 m which is the free board regionfor this system.

Fig. 7 shows the variations of the bed pressure drop againsttime for different velocities. It is observed that as time increases,the bed pressure drop fluctuates and increases significantly. Thehigher superficial gas velocity gives the lower pressure drop thanthe lower superficial gas velocity (Fig. 7a) for more drag force beingexerted on particles. Initially the bed pressure drop is found toincrease linearly with superficial velocity (Fig. 7b) indicatingpacked bed behavior. The pressure drop becomes constant whenall the materials start fluidizing. This is justified from orifice the-ory. It is also found that simulated results are in good agreementwith experimental results with a deviation of 4.09% approximately.

4.1.2. Bed expansionThe time-averaged voidage profiles for the gas–solid fluidized

bed are shown in Fig. 8(a) for six different velocities varying from0.008 m/s to 0.018 m/s. It is found that initially, the bed heightincreases with bubble formation. As a result gas volume fractionincreases. After some time expanded bed height remains constantat steady state of fluidization. In the beginning of the simulation,waves of voidage are created which travel through the bed. Subse-quently bubbles coalesce to form large bubbles as the simulationprogresses. It is also observed from this figure that for higher gasvelocities the gas volume fraction is larger indicating more bedexpansion. It is further observed that there are fluctuations in gasvolume fractions. This may be due to the frequent formation andbreakage of bubbles with increase in gas velocity within the bed.

In the bottom region of the column, concentration of solid par-ticles is larger than that in the upper part. Therefore, the maximumgas volume fraction/voidage is found to occur in the top part of thecolumn. The voidage then increases sharply to 1 at the top of thecolumn which corresponds to the region with no solid particlespresent. Thus the expanded bed represents a clear interfacebetween the fluidized regions and the free board regions. Gas vol-ume fraction approaches the saturation condition when it is equalto 1. The maximum expanded bed heights for different velocity aregiven in Table 3.

Solid volume fraction against bed heights for six differentsuperficial gas velocities in the range of 0.008–0.018 m/s areshown in Fig. 8(b). It is seen that at higher superficial gas velocity,the distribution of solid volume fraction decreases in the bed. Atlower superficial gas velocity, the bed shows higher solid volume

tions for different velocities.

Table 3Comparisons of static bed height and expanded bed height.

Superficial airvelocity (m/s)

Uo/Umf

Static bedheight (m)

Expanded bedheight (m)

Percentageincrease (%)

0.008 0.8 0.1 0.171 710.010 1 0.1 0.185 850.012 1.2 0.1 0.2 1000.014 1.4 0.1 0.214 1140.016 1.6 0.1 0.228 1280.018 1.8 0.1 0.242 142

P. Sahoo, A. Sahoo / Advanced Powder Technology 25 (2014) 1699–1708 1707

fraction because the solid particles are accumulated in the lowerportion of the bed. When the superficial gas velocity increases,the solid volume fraction distributions fluctuate more in the axialdirection. Gradually solid volume fraction decreases and thenreduces to zero at the bottom of the column. With increasingsuperficial gas velocity, the solid volume fraction generallyincreases with height in the bed. Finally solid volume fractionincreases to 1 at the top of the bed beyond which there is no solidparticles indicating it to be free board region.

Fig. 9 shows the time-averaged solid volume fraction as a func-tion of bed height for different inlet gas velocities for differenttimes of simulation. Initially the solid volume fraction fluctuatesand then it decreases as fluidization starts. It can also be seen athigher superficial gas velocity the distribution of solid volume

Fig. 9. Distributions of volume fra

fractions decreases more in bed than at the lower superficial gasvelocity. Then solid volume fraction decreases sharply to zero atthe bottom of the bed which is known as saturation condition. Itis also observed from Fig. 6 that with simulation time step of10 s the solid volume fraction is approaching saturation conditionfor all velocities (Table 3).

Fig. 10a shows the bed expansion against time at six differentinlet gas velocities varying from 0.008 to 0.018 m/s. The resultsshow an increase in bed height with increasing inlet gas velocity(Table 3). Fig. 10b shows the plot of variation in bed height againstdifferent superficial air velocities. The bed height is observed toincrease linearly with inlet gas velocity indicating that the bedexpands with increased velocity till steady state is attained.Fig. 10c shows the comparison of variation in bed expansion ratioagainst superficial/inlet air velocities for experimental and simula-tion results. The bed expansion ratio is found to increase linearlywith inlet gas velocity. It is also found that simulated results arein good agreement with experimental results with a deviation of9.76% approximately.

5. Conclusion

Experimental observation revealed that pressure drop and bedexpansion ratio increase with superficial velocity. As the superficialvelocity is about two times the minimum fluidization velocity, it is

ction with variation of time.

Fig. 10a. Bed expansion against time.

Fig. 10b. Expanded bed height versus inlet air velocity.

Fig. 10c. Comparison plot of bed expansion ratio.

1708 P. Sahoo, A. Sahoo / Advanced Powder Technology 25 (2014) 1699–1708

ensured that the bubbling fluidization has occurred in the presentwork. CFD simulation was also carried out for bubbling fluidiza-tion. Effect of velocity on bed dynamics such as pressure dropand bed expansion ratio were analyzed both experimentally and

computationally. It was observed that the modeling predictionsagree reasonably well with experimental pressure drop, bedexpansion ratio and gas–solid flow patterns. Pressure drops ingas–solid flow predicted by the simulations were found to be inrelatively close agreement with experimental measurements atall superficial gas velocities higher than the minimum fluidizationvelocity ensuring for the same hydrodynamics of bubbling fluidiza-tion both computationally and experimentally. Simulation resultsalso indicated that small bubbles are produced at the bottom ofthe bed. These bubbles coalesce and grow as they move upwardsforming larger bubbles with the increased air velocity. By selectingproper velocity the amount of fines entrained from the bed can bedecreased. Further experimental and modeling efforts are requiredwith respect to other parameters such as time and particle size forvalidation of CFD models for fluidized bed. As velocity of fluidizingmedium is one of the important parameters to determine the qual-ity of fluidization, this study on hydrodynamics of fluidized bedusing fine particles can be considered as a strong base for develop-ment of a pilot plant unit.

Acknowledgments

This work was supported by National Institute of Technology,Rourkela for which the authors would like to thank Director ofNational Institute of Technology, Rourkela, for necessary fundingand support without which this work would not have beenpossible.

References

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