Effect of Bed Height, Bed Diameter and Particle

Effect of Bed Height, Bed Diameter and Particle Shape on Minimum Fluidization in a Gas-Solid Fluidized Bed Md Rashedul H Sarker,1 Md Mahamudur Rahman1, Norman Love2, and Ahsan Choudhuri3 Center for Space Exploration and Technology Research (cSETR) University of Texas at El Paso, El Paso, Texas 79968

In this paper parameters most likely to effect pressure drop and particle behavior in gassolid fluidized bed is investigated. These parameters included bed height, bed diameter, and particle shape. For the experiments two laboratory scale fluidized bed systems were used one having a 3.5cm inner diameter and a second having a 12.4cm inner diameter to test the effect of bed diameter. The effect of particle shape was also investigated by testing different shapes of particles ranging from 0.85 to 1.0mm. Also, the bed height varied from 2 to 7 cm. Overall, it is observed that minimum fluidization velocity decreases with decrease of bed diameter. It was also determined that the minimum fluidization velocity did not show a significant variation with bed height. Testing particle shape showed spherical particles requires more pressure drop to reach a minimum fluidization velocity compared to non-spherical particles. Digital imaging of the particle was also recorded to observe incipient minimum fluidization, bubble flow, bubble shape, and particle collisions in the fluidization regime.

Nomenclature Dp Deq dsd H Vm Vs ∆P ∈ µ ! ρf ρp

I

= = = = = = = = = = = =

Particle Diameter Equivalent Diameter Sauter-mean Diameter Bed Height Minimum Fluidization Velocity Superficial Gas Velocity Bed Pressure Drop Void Fraction Viscosity Sphericity Density of fluid (air) Density of Particle (Borosilicate Glass Beads)

I. Introduction

N nature coal is an abundant fossil fuel and is used widely throughout the world. Combustion of coal for power generation exposes it constituents such as sulfur, nitrogen, and carbon to oxidation processes leading to unwanted pollutant emissions. One of the biggest challenges of 21st century is to continue to supply the necessary energy requirements while continuing to reduce unwanted pollutant emissions. Coal gasification is a process which can reduce some emissions without sacrificing performance. One method of converting coal or other feedstock to a synthetic gas (CO-H2) is through the use of a fluidized bed. The process breaks the fuel down helping to remove impurities and ash. The U.S. Department of energy is now emphasizing development of advanced coal gasifiers with enhanced efficiency and expanding the reliability1. In last quarter century significant advancement in coal burning technology such as fluidized bed technology have been observed2. In many cases a popular choice for fluidization, particularly with coal particles, is a gas-solid type 1

Graduate Research Assistant, Student Member AIAA Research Assistant Professor, Member AIAA 3 Director cSETR and Professor, Senior Member AIAA

2

1 American Institute of Aeronautics and Astronautics 092407

fluidized bed. The important parameters to design a gas-solid fluidized bed are solid mixing and their flow pattern. However, involving multiple scales with interaction of phases, gas-solid flows in fluidized bed are difficult to access3. In order to characterize fluidized bed, minimum fluidization velocity is one of the most important parameters. There are many factors that influence the minimum fluidization which have been described in literature4, 5-11 . A summary of these studies is presented in the following paragraphs. Gunn et al. 4 investigated the effect of particle diameter and bed height on the minimum fluidization in a cylindrical bed. The authors found that there was no significant effect of bed height on minimum fluidization. Geldart et al.5 used a rectangular bed (2D) and cylindrical bed (3D) to examine the effect of bed dimension and bed height on minimum fluidization. They used six (6) different bed heights for both 2D and 3D beds and found minimum fluidization increased with increase of bed height for rectangular bed and remained same at different bed heights for a cylindrical bed. Hilal et al.6 studied the effect of bed diameter with 0.29 m diameter cylindrical bed and compared results to another with a diameter of 0.089 m. The authors found minimum fluidization velocity decreased with the increase of bed diameter. Ramos et al.7 used a rectangular bed (2D) with dimensions of 1×0.2×0.012m and range of particles from 160-700 µm at different bed heights and bed widths. They found minimum fluidization velocity increased with the increase of bed height, particle diameter and decreases with the increase of column width. Zhong et al8 used Geldart Type-D particles at different heights ranging from 300-500 mm in a spouted fluidized bed with dimensions of 300mm × 300mm and 2000mm height. They concluded the influence of static bed height on the minimum fluidization increases with the increase of bed size. Sau et al.9 determined bed pressure drop and minimum fluidization over a conical tapered fluidized bed for various column heights. They concluded that bed height had no significant effect on minimum fluidization for conical tapered fluidized bed. However, the minimum fluidization velocity increased with the increase of taper angle. Zhon et al.10 used 0.4 m×0.4 m rectangular bed using wood chips, mung, beans, millet, corn stalks and cotton stalks to study the effect of particle size, shape, and density on minimum fluidization. They found minimum fluidization velocity increased with increasing the length to diameter ratio of the bed. Rao et al.11 used two fluidization segregation units with 1.6 cm and 2.4 cm column diameters. Glass beads (100-600µm) with a density of 2500 kg/m3 and polystyrene beads (250-354µm) with a desnity of 1250 kg/m3 were used in their experiment. The authors showed that minimum fluidization velocity was influenced by bed diameter and bed height. Escudero et al.12 investigated the effect of bed height and material density on minimum fluidization velocity in a cylindrical bed with glass beads, ground corncob and ground walnut shells at different bed heights. They found for each type of particle the minimum fluidization velocity is insensitive to bed height. They also found that minimum fluidization is sensitive to particle density. It increases with the increase of particle density. Therefore, based on the aforementioned studies bed height, bed diameter, and particles shape are among the parameters which may significantly effect minimum fluidization in a fluidized bed. This paper presents the experimental results of the effect of these parameters in cylindrical bed for both spherical and non-spherical particles.

II. Experimental Setup and Procedure Two fluidized beds with cylindrical shape were employed to perform the hydrodynamic analysis. Figures 1 and 2 show the packed fluidized bed columns with the basic components used for present investigation. The fluidized bed presented in Fig 1 was made of plexiglass tube with an outside diameter of 3.8 cm and a wall thickness of 0.318cm. For fluidization, air was supplied to the test section by a high-pressure rotary screw type compressor. Air supply was controlled with a pressure regulator and stainless steel integral bonnet needle valve. The fluid delivery system was developed using polyvinyl chloride (PVC) pipe with an inside diameter of 1.34 cm. a digital mass flow meter was used to measure the volumetric fluid flow rate. A honeycomb shaped distributor, made of 2.54 cm hexagonal brass tubes, was used in order to ensure the uniform fluid flow distribution to the bed. A mesh catch with a nominal diameter of 0.425 mm was installed at the bottom of the bed to prevent test particles from falling back into the tubing and at the top to prevent particles from being ejected out of the column. A six field selectable differential pressure transducer was installed at the bottom of the bed to measure the pressure fluctuations across the bed. Data acquisition software was used for averaging and storing the pressure fluctuation measurements. Figure 2 shows the second experimental fluidized bed setup used in the current study. The bottom section of the packed bed column was made of plexi glass tube with 12.7 cm outer diameter and 0.318 cm wall thickness. A quartz 2 American Institute of Aeronautics and Astronautics 092407

tube with outside diameter of 12 cm and a wall thickness of 0.5 cm inserted into the plexi glass tube to attain better optical access for the particle image velocimetry (PIV) analysis. A 3730 KW pressure blower with 34 m3/min flow rate was used to supply air flow through 12.7 cm diameter of sheet metal pipe. A wafer style 12.7 cm butterfly valve was used to control the flow of air to the bed. To estimate the volumetric flow rate across the bed an insertion type high performance thermal mass flow meter with 200 ms was used which was calibrated by the manufacturing company to the flow range of 0 to 4000 SLPM. For uniform distribution of air to the bed a honeycomb shape distributor with 2.54 cm in length has been placed at 8 cm below the bed. To hold the particle in the bed a mesh made from brass screen with 0.053 mm nominal diameter was installed. The same type of mesh catch was also installed at the top of the bed. To measure the pressure drop across the bed a tygon tube is connected at the bottom of the bed with 53 micron mesh at the entry part to restrain the particle entering into the tube. A digital manometer capable of measuring differential pressure with 0-13790 Pa range and 7 Pa of resolution was connected to the tygon tube exit part to measure the pressure drop across the bed at different volumetric flow rate of air.

Fig 1: Experimental set up of cylindrical Fluidized Bed (3.5 cm inner diameter)

Fig 2: Experimental set up of cylindrical Fluidized Bed (12.4 cm inner diameter)

A. Particles Borosilicate glass beads with 2230 kg/m3 density are used to characterize the fluidized bed. 1 mm glass beads are used as spherical particles. To produce non- spherical particles, a hydraulic compressor was used to crush 6 mm spherical glass beads into smaller pieces. Crushed particles contained different sized particles ranging from a few microns to 1mm nominal diameter. To categorize these crushed particles sieve shaker was used with different sieve sizes ranging from 20-1180 µm. For the current experimental study we selected only 850 to 1000 µm range non spherical particles. This range of particles has been considered as equivalent 1 mm non spherical particles for comparison. B. Bed Pressure Drop To predict the pressure drop across the bed many empirical and semi-empirical models have been developed. Most widely used semi-empirical model that satisfactorily predicts the pressure drop across a packed bed seen in Eq (1)13: ∆P = 150

(!!∈)! ! ! !! !∈

HV! + 1.75

(!!∈)!! !! ∈!

HV!!

(1)

Where !! is the diameter of the particle for a sphere. In this equation the pressure drop is subject to the particle size and bed porosity but in case of non-spherical particles the application of this equation differs slightly. For nonspherical particles this equation utilizes equivalent diameter Deq = φ x Dsd, where Dsd is the Sauter-mean diameter. Li et al.13 found their experimental data are well correlated if equivalent diameter is chosen for non spherical particles and is presented in Eq (2): ∆P = 150

(!!∈)! ! ! !! !" ∈

HV! + 1.75

(!!∈)!! !!" ∈!

HV!!


(2)

C. Minimum Fluidization Velocity Minimum fluidization can be obtained from Eq (1,2) by replacing the pressure drop (∆P) across the bed with the weight of the bed per unit area for a particular bed height (H). Buoyant force of the displaced fluid is also taken into consideration14. ∆P = g 1 − ϵ ρ! − ρ! H Using Eq (2) the following relationship is obtained: ∆P =

!"#(!!∈)! !!! ! !! !∈

HV! + 1.75

(3) !! !! (!!∈) !! ∈!

! HV!

(4)

D. Void Fraction: Packing characteristics are required to understand the design and operation of a packed bed. During our experiment the bed is densely packed. This is categorized by the voidage fraction which ranges from 0.37 to 0.39 for the bed used for the experiments15. For non-spherical particles, very little theoretical and experimental work has been performed. It is suggested by Brown15 that the packing void fraction depends on particles sphericity and based on experiment it has shown the voidage can relate to the sphericity. E. Sphericity It was necessary to determine the sphericity for the non-spherical particles used in the present experiment (850-1000 µm). Krumbein15 expressed the sphericity formula in Eq (5):

!=

!

!"#$%& !" !!! !"#$%&'(

!"#$%&' !"#$%&%' !" !!! !"#$%&'(

≈ !"#$%&' !"#$%&%' !"#!$%&!#"'() !" !!! !"!!"! !"!"#$ !" !!! !"#!$%&!#"'() !"!!"!

(5)

Nominal diameter of the particle for this study is considered as the average diameter of sieve size used and the largest diameter circumscribed by the sphere shown in Fig 3. An example of measuring this diameter is shown in Fig 3 using a microscopic video camera.

Fig 3: Longest diameter tracking of non spherical particle


III. Results & Discussion Pressure drop and minimum fluidization velocity across the bed at increasing superficial gas velocity was measured for both experimental setups. 1 mm spherical glass beads and non-spherical glass beads with a nominal diameter of 1 mm were used. The experimental data are presented in the following sections. A. Effect of Bed Height: Minimum fluidization velocity of 1 mm spherical particles in the 12.4 cm bed is plotted in Fig 4. For these tests bed height ranged from 2-7 cm. During experimental testing it was observed that the deviation in analytical and experimental maximum bed pressure drop decreased with the increase of bed height and did not significantly vary from theoretical calculations after a 5.00 cm bed height.

Fig 4: Minimum Fluidization Velocity at Different Bed Height (1mm Spherical Particles), 12.4 cm Bed Diameter. Examining the effects of bed heights for non-spherical particles with the same nominal diameter were also done. Figure 5 shows the results for the non-spherical particles. For the non-spherical particles bed heights between 2 and 5.5 cm were tested. Non-spherical particles showed a larger deviation from analytical pressure drop calculations. Figure 5 shows the plot of experimental minimum fluidization velocity for different bed heights for non-spherical particles.

Fig 5: Minimum Fluidization Velocity at Different Bed Height (1mm non-spherical Particles), 12.4 cm Bed Diameter. 5 American Institute of Aeronautics and Astronautics 092407

For both spherical and non-spherical particles the minimum fluidization occurs at approximately the predicted value resulting from Eq. (1, 2). Deviation from the theoretical curve may be effected by assuming constant values for sphericity, particle diameter, void fraction, or equivalent diameter. B. Effect of Particle Shape To identify the effect of particle shape on packed bed, the experimental data of 1 mm non-spherical particles were compared with the 1 mm spherical particles. For these tests the bed height was maintained same for both and particles size. Some significant differences are observed in the experimental data because of changes is particle shapes and drag acting on these surfaces. Figure 6 shows a plot of spherical and non-spherical particles at a 5.0 cm bed height.

Fig 6: Bed Pressure Drop Vs Superficial Gas Velocity for Spherical and Non Spherical Particles, 12.4 cm Bed Diameter. From Fig 6 bed pressure drop and minimum fluidization velocity decreased for non-spherical particles. This may be due to non-spherical particles uneven distribution in the fluidized bed allowing for more space and higher voidage fractions to exist. This was also observed when particle bed weights were measured, spherical particles with the same bed heights contained higher mass then the non spherical particles. Therefore, the spherical bed requires more pressure drop and minimum fluidization velocity to fluidized the bed. C. Effect of Bed Diameter To identify the effect of bed diameter on packed bed spherical particles were tested at two different bed heights. Figure 7 shows a plot comparing the 12.5cm and 3.5 cm bed diameters. From Fig 7 it can be seen that by decreasing the bed diameter pressure drop increased and minimum fluidization velocity decreased.


Fig 7: Bed Pressure Drop Vs Gas Velocity for Two Different Bed Diameter at 5 cm Bed Height. Some authors have shown different results to the findings presented in Fig. 7. In previous studies minimum fluidization velocity increased with a decreasing bed diameter. Compared to other studies’, however, results in the present experimental investigation consisted of a large particle sizes, significant variations in sphericity, and larger differences in bed diameters. Further tests over a wider range of these bed characteristics may further reveal possible empirical correlations for fluidized bed systems. D. High Speed Flow Visualization Fluidization can be classified in several regimes such as fixed bed, particulate fluidization, bubbling fluidization, slugging fluidization, and turbulent fluidization. Particles are inactive in the fixed bed regime. Particles disperse uniformly with no identifiable bubble in particulate fluidization. In bubbling fluidization bubbles rise from bottom of the bed to the surface of the bed. Slugging fluidization occurs as the superficial velocity is increased and in the turbulent regime bed surfaces are diffused and are difficult to observe15. A high-speed digital image optical technique has been used for better observation of bubbling in the 12.4cm bed. A high –speed camera (5 kHz) captured different flow regimes in the fluidized bed for both spherical and nonspherical particles. Some images are presented for the various fluidized bed regimes in the following sections. 1) Incipient Fluidization This regime occurs at the point of minimum fluidization. At this point upward fluid drag force acting on the particle is equal to the weight of particles. Figures 8 and 9 show the progression from a fixed bed to an incipient fluidization state (left to right). In the case of spherical particles bed expansion is observed at the beginning of minimum fluidization. Bed expansion is described by a significant observable increase in bed height as the superficial velocity is increased. This occurs due to the more densely packed spherical particles compared to the non-spherical particles. Channeling is also observed for both types of particles with larger channeling occurring for the non-spherical type. Non-spherical particles irregular shapes reduce the void fraction and create larger channels along the walls which fluid flows. Near the top points of this regime bubbles begin to form at the bottom of the bed and rise to the surface of bed. Minimum fluidization appears at this point. From plotted pressure drop and superficial velocity smooth minimum fluidization curves are observed for spherical particles since channeling is reduced.


Fig 8: Non Spherical Particles

Fig 9: Spherical Particles 2) Bubbling Fluidization This regime occurs at higher gas flow rates beyond the minimum fluidization15. Figures 10 and 11 show periodic steps from point of minimum fluidization to bubbling fluidization. From both types of particles it is observed that bubbles occur at the bottom of bed and rise. It is observed that bubbles coalescence in vertical and horizontal directions. Pressure reduces towards the top of bed resulting in tailing bubbles merging with the leading bubbles. For horizontal coalescence the bubbles merge with the adjacent bubbles.

Fig 10: Non Spherical Particles

Fig 11: Spherical Particles Figures 12 and 13 present the binary images of bubbles for particles. These images were processed to obtain a better imaging of the bubble formation in the bed. The black region represents the gas phase and the white region represents the emulsion phase. Small bubbles are also observed adjacent to the large bubble formations. In the present experiment irregular bubble shapes occurred for both particles. 8 American Institute of Aeronautics and Astronautics 092407

Fig 12: Binary Image of Bubble ( Non Spherical Particles)

Fig 13: Binary Image of Bubble (Spherical)

IV. Conclusion The minimum fluidization velocity was determined for both spherical and non-spherical particles in a cylindrical packed fluidized bed. The bed height, particle shape, and bed diameter were varied and the effect on pressure drop across the bed determined. It was determined that the minimum fluidization velocity did not show a significant variation with bed height. Testing particle shape showed spherical particles requires more pressure drop to reach a minimum fluidization velocity compared to non-spherical particles. Bed diameter also influenced the minimum fluidization velocity decreasing with the decrease of bed diameter. High speed imaging technique used for flow visualization in incipient minimum fluidization and bubbling fluidization regimes for both shape of particles. Larger channeling observed before minimum fluidization for non-spherical particles and in bubbling fluidization regimes revealed that bubbles shapes are irregular for both shape of particles.

Acknowledgments This research is supported by the U.S. Department of Energy, under award DE-FE0003742 (Project Manager Robie Lewis). However, any opinions, findings, conclusions, or recommendations expressed herein are those of the authors and do not necessarily reflect the views of the Department of Energy.

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