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Hydrodynamics and local heat transfer measurements in a bubble column with suspension of yeast
Biochemical Engineering Journal 9 (2001) 155–163
Hydrodynamics and local heat transfer measurements in a bubble column with suspension of yeast A. Prakash, A. Margaritis∗ , H. Li, M.A. Bergougnou Department of Chemical and Biochemical Engineering, The University of Western Ontario, London, Ont., Canada N6A 5B9 Received 12 January 2001; accepted 6 June 2001
Abstract Hydrodynamics and heat transfer were investigated in a 0.28 m diameter slurry bubble column for air–water–yeast cells system. Yeast cells of about 8 m diameter were used and the effects of gas velocity and yeast concentrations were studied. Gas holdups exhibited a maximum value around a gas superficial velocity of 0.10 m/s when foam height was included. Without the foam layer, gas holdups increased with increasing gas velocity. Bubble population was measured by means of dynamic gas disengagement technique. Rise velocity of small bubbles decreased, while rise velocity of large-bubbles fraction increased with yeast concentration which was varied from 0–0.4% w/w. Local heat transfer coefficients were measured at different axial and radial locations inside the column. Heat transfer in the foam section was significantly lower than in the main slurry column. In the bulk and developing regions, the addition of yeast cells into water increased heat transfer in the center and decreased at the wall. This work is one of very few studies reported in the literature on the use of actual yeast cells in slurry bubble columns and opens up new opportunities on their potential use as bioreactors in the fermentation industries, wastewater biological treatment and other applications. © 2001 Elsevier Science B.V. All rights reserved. Keywords: Slurry bubble column; Yeast cells; Hydrodynamics; Bubbles fractions; Heat transfer
1. Introduction Three-phase bubble columns are commonly used as chemical reactors and separation devices. The advantages of these columns include high heat and mass transfer rates, isothermal conditions, plug-free operation and ability for on-line catalyst addition and withdrawal. Three-phase bubble column have potential for applications in biochemical processes, such as fermentation and biological wastewater treatment [1–6]. When a slurry bubble column is used for suspensions containing micron-size particles such as yeast cells, the operation of the column can be different from that of a conventional slurry bubble column. Indeed, significant differences can arise in hydrodynamics and heat transfer characteristics due to the presence of different biochemical compounds associated with yeast cells and their growth medium. Moreover, the yeast cells can alter the air bubbles behavior, which in turn can alter the hydrodynamics as well heat transfer in the column. Therefore, there is a need to understand the hydrodynamics and heat transfer characteristics in the design of slurry bubble columns as bioreactors.
Gas holdup is a critical hydrodynamic parameter and it is generally found to be affected by several parameters including gas velocity, particle size and concentration and presence of surface-active agents [7–11]. Additional information from the study of bubble populations allow more detailed hydrodynamic analysis [12–16]. Heat transfer in slurry bubble columns has been investigated by several researchers [8,11,17]. However, most of the literature studies have used inert solid particles for their investigations. For applications in biochemical processes, it is important to study effects of micron-size cells concentration on column hydrodynamics and heat transfer. For example, micron-size cells, can alter bubble–bubble interactions as they can get immobilized on the surface of gas bubbles [18]. This study investigates hydrodynamics and heat transfer behavior of a slurry bubble column containing suspensions of yeast cells of about 8 m in diameter. The yeast samples were obtained from a local brewery and measurements have been made for gas holdups, bubble populations and their rise velocities and local heat transfer coefficients. 2. Experimental
∗ Corresponding author. E-mail address: [email protected] (A. Margaritis).
Experiments were conducted in a Plexiglas column of 0.28 m inside diameter and 2.4 m height (see Fig. 1). The
1369-703X/01/$ – see front matter © 2001 Elsevier Science B.V. All rights reserved. PII: S 1 3 6 9 - 7 0 3 X ( 0 1 ) 0 0 1 3 7 - 1
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Nomenclature A Ap Cp dp DC h havg HD HS HT k N Pv P/H q r R T Ub Ub,L Vb∞ VC Vg Vs w Rep Pr z
(m2 )
heat transfer area parameter in Eq. (7) specific heat (J/kg ◦ C) probe diameter (m) column diameter (m) heat transfer coefficient (kW/m2 ◦ C) average heat transfer coefficient (kW/m2 ◦ C) height of expanded slurry region (m) static height of the suspension (m) total height of dispersion including foam bed (m) thermal conductivity (kW/m◦ C) number of data points used to obtain time-averaged heat transfer coefficient power input per unit mass (m2 /s3 ) pressure gradient along column height (Pa/m) heat flow rate (kW) radial distance (m) column radius (m) temperature (K) bubble rise velocity (m/s) rise velocity of large bubbles population (m/s) terminal rise velocity of bubbles (m/s) average liquid circulation velocity (m/s) superficial gas velocity (m/s) hindered settling velocity of spherical particle (m/s) weight fraction in slurry Reynolds number based on probe diameter (ρUb dp /µ) Prandtl number for suspension (Cp µ/k) axial distance from column bottom (m)
Greek letters α thermal diffusivity, ksl /ρ sl Cp,sl (m2 /s) δo laminar boundary layer thickness (m) δ thermal boundary layer thickness (m) ε phase holdup εg,D gas phase holdup in expanded slurry region εg,T total gas phase holdup including foam section φs volume fraction of solids in gas-free slurry phase µ viscosity (Pa s) ν kinematic viscosity (m2 /s) θc contact time (s) ρ density (kg/m3 )
Subscripts b g l s sl S
bulk gas liquid solids slurry surface
liquid or slurry phase height before fluidization was maintained at 1 m. A six-arm gas distributor introduced air into the column bottom. The gas distributor arms were 5 mm in diameter and 0.14 m long. Each arm had four holes of 1.5 mm diameter facing downward. An electric heater was located near the bottom to maintain a constant temperature in the column. Compressed air and tap water constituted the gas and the liquid phases, respectively. The solid phase consisted of yeast cells obtained from a local brewery company. Before the use of fresh yeast cells in the slurry bubble column, they were treated with copper sulfate at 21◦ C for 4–5 h, in order to stop their biological activity by deactivating their respiration enzymes. Before the start of the experiments, the suspension was washed 3–4 times with water to remove most of CuSO4 and alcohol from it. The washing was carried out in the column by adding twice the volume of
Fig. 1. Schematic diagram of the experimental setup.
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157
Fig. 3. Structural arrangement of slurry bubble column with foam layer.
Fig. 2. Optical microscope photograph of dead yeast cells (Saccharomyces cerevisiae).
fresh tap water. This was followed by air agitation for about 15 min. The air flow was then stopped and yeast cells allowed to settle. The aqueous layer above the settled bed was then decanted. This process was repeated again until there was significant reduction in foaming during air agitation. Washing was stopped when the foam layer height was less than 50 mm for the selected gas velocity. This procedure introduced reasonable reproducibility in subsequent measurements. However, presence of foam layer did indicate small amounts of surface-active molecules still remaining in the suspension. It was decided to conduct the experiments with the foam layer formation and investigate column hydrodynamics and heat transfer for such conditions. Fig. 2 is an optical microscope picture that shows a typical population of dead yeast cells Saccharomyces cerevisiae. For a few experiments, 11 m glass beads were also used as solid phase. The column was operated in semi-batch mode at various superficial gas velocities and slurry concentrations. Slurry samples were withdrawn with the help of sampling probe located at about 0.8 m from bottom. The yeast cells concentration in the samples was measured by the dry weight method. During operation, a foam layer was formed at the top resulting in two sections as shown in Fig. 3. In the lower section, slurry was the continuous phase while in the
foam section at the top, gas was the continuous phase. Gas holdups were calculated with and without foam section, using Eqs. (1) and (2). Gas holdup with foam bed: εg,T =
HT − H S HT
(1)
Gas holdup without foam bed: εg,D =
HD − H S HD
(2)
Gas holdups were also obtained from the axial pressure difference measured by two pressure transducers (Omega PX541) mounted at 0.07 and 0.93 m from the column bottom and were calculated using Eq. (3). This equation has been derived by assuming that frictional pressure drop and inertial effects are negligible and particles concentration is constant along the column height [12]. εg = 1 −
1 P ρl + φs (ρs − ρl ) H
(3)
Gas holdup structure and bubble rise velocities were investigated by means of dynamic gas disengagement technique (DGD). After shutting-off the gas flow by closing air inlet in Fig. 1, the drop of instantaneous gas holdup was recorded. The DGD technique can distinguish between different bubble classes in a dispersion, if there are significant differences between their rise velocities [16]. Generally two
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classes of bubbles called large and small-bubbles fractions have been distinguished by this technique. Initially, there is a fast drop in the dispersion height indicating escape of fast rising large bubbles. This is followed by a slower drop in the dispersion height with the escape of slow rising small bubbles. The time-course of the instantaneous gas holdup can be used to evaluate the holdup fractions of small and large-bubbles fraction and their rise velocities [12]. Various assumptions and sources of errors associated with the DGD technique and their relative effects are reviewed in literature and discussed in details by Li and Prakash [12]. Based on this procedure, the measurements in this study provided relative fractions of the large and small bubbles populations and their rise velocities in the dispersion. Heat transfer was measured by means of heat transfer probes located at different axial and radial locations inside the column. The probes were designed with a micro-foil heat flow sensor mounted on the surface of a brass cylinder. Inside the cylinder, cartridge heater was installed as a heat source. The probes registered both heat flux and surface temperature. The construction details of the probe were described in a previous study [17]. Bed temperature was monitored by means of a digital thermometer. The heat transfer probes were located at axial locations of 0.07, 0.52 and 1.28 m from the column bottom. The signals of heat flux and probe surface temperature were collected simultaneously at 60 Hz for about 35 s. The microvolt signals from the heat flux sensor were amplified to millivolts before collection by a data acquisition system. The time-averaged
heat transfer coefficient at a given location was obtained by averaging the instantaneous heat transfer data: N
havg =
1 q/A N TS − T b
(4)
i=1
The number of samples (N) selected for averaging was 2100 to ensure a stable value of heat transfer coefficient.
3. Results and discussion 3.1. Hydrodynamics Variations of gas holdups with and without foam layer are shown in Fig. 4 for different yeast concentrations. High total gas holdups at gas velocities below 0.2 m/s can be attributed to a foam layer formation at the top of the dispersion. It has been known that foam is created when small amount of surface-active substances (e.g. alcohols) are added into a liquid [19,20]. The solutions used in this investigation had varying concentrations of surfactants (ethanol, protein, etc.). The surfactants concentration could be expected to be proportional to the concentration of yeast in the suspension. The increased gas holdup with increasing yeast concentration indicates a positive effect of surfactants’ concentration on foam formation (Fig. 4). The foam structure is known to be affected by superficial gas velocity [7,20]. It can be seen from Fig. 4, that the total gas holdup exhibits a maximum
Fig. 4. Gas holdups in air–water–yeast cells system measured by the bed height method.
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at superficial gas velocity around 0.10 m/s. Pino and Solari [7] observed a similar tendency in air–kerosene system containing surfactants and 5 m inert solids. The maximum gas holdup occurred at superficial gas velocity around 0.05 m/s. The location of maximum gas holdup could be a function of the nature of surfactants, solid particles, distributor type, etc. As the superficial gas velocity is increased beyond 0.10 m/s, fast rising large bubbles induce turbulent flow. A high degree of turbulence breaks the foam structure and decreases the gas holdup in this regime. Meanwhile, the difference in gas holdups with and without the foam section is significantly reduced. In this flow regime, the foam layer consisted of a short layer of small bubbles and it oscillated with the expanded slurry section. Fig. 5 shows that the gas holdups obtained from expanded slurry height method are generally higher than the values estimated by the differential pressure method. The difference is small at 0.05 m/s but increases with gas velocities up to about 0.2 m/s followed by a slight drop at higher gas velocity. The accumulation of small bubbles at the top of dispersion can add to gas holdup measured by the expanded bed height method. The measurements by pressure differential method excluded the contribution by this layer (the top pressure transducer was generally located below this layer). Increased turbulence intensity at high gas velocities could reduce this contribution by breaking the foam layer and by dispersing the small bubbles into the bulk phase.The hydrodynamics of the column were further investigated based on bubbles populations measurements. Two bubbles fractions,
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Fig. 6. Rise velocities of small bubbles population in different systems.
namely, large and small, could be detected by the DGD technique. Fig. 6 shows that small-bubbles rise velocity decreased with increasing superficial gas velocity. The decrease in rise velocity of small-bubbles fraction with increasing gas velocities is documented in literature [12]. It is also noted from Fig. 6 that the rise velocity of small bubbles decreased with increasing yeast concentrations. This decrease
Fig. 5. Comparison of gas holdup in air–water–yeast cells system measured by different methods (yeast cells concentration: 0.4 wt.%).
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in small-bubbles rise velocities can be attributed mainly to a proportional rise in the concentration of surface-active molecules in suspension (ethanol, proteins, etc.) with increasing yeast concentrations in the suspension and/or due to coalescence hindrance effects of yeast particles on bubbles surface. These two phenomena can have synergistic effects under appropriate conditions in the system. The hindrance effect of particles in a suspension has been observed only at low slurry concentrations [20]. These authors observed significantly higher gas holdups for low slurry concentration (1 wt.%) of fine particles (7 m) in air–electrolyte solution. The increase in gas holdup was attributed to coalescence hindering effect of the fine particles at the low slurry concentration. However, most literature studies have generally reported an opposite trend, i.e. increase in small-bubbles rise velocities (and decrease in gas holdup) with the addition of solids in a liquid [12,13]. The increase in rise velocity of small bubbles with the slurry concentration has been attributed to increase in apparent slurry viscosity [21]. For the low concentration suspensions of yeast used in this study, the change in apparent slurry viscosity is very small [18]. Therefore, this effect is not expected to play a significant role in affecting bubble size distribution in this study. The effects of gas velocity and yeast concentration on large-bubbles rise velocities are shown in Fig. 7. It can be observed that, the rise velocities of large-bubbles fraction increased with increasing superficial gas velocity or yeast concentration. Similar variations of large-bubbles rise velocities were also observed by Li and Prakash [12] in air–water–glass beads system. However, the large-bubbles rise velocities in the air–water–yeast system are found to be higher than those reported by Li and Prakash [12] for similar concentration of inert glass beads in suspension. This difference in behavior could be attributed to very dif-
Fig. 8. Gas holdups due to large and small-bubbles fractions in the dispersion.
ferent nature of the particles (and associated contaminants) used in the two studies. Walter and Blanch [22] investigated bubble break-up in fermentation broths of yeast and other bacteria. The maximum stable bubble size was observed to be significantly higher than in distilled water or dilute solutions of alcohols. This was attributed to different surface elastic behavior of the broths due to the presence of associated surfactants such as alcohols, proteins and salts. The bubble break-up mechanism could also be affected by the sweeping action of immobile yeast on bubble surface due to fast rising bubbles [22]. This action will tend to accumulate the surfactant molecules at the rear end of the bubbles, thus creating surface tension gradients on the bubble surface. Therefore, several additional factors come into play when investigating the behavior of suspensions with yeast cells. The gas holdup contributions due to small and large bubbles populations are presented in Fig. 8. It can be seen that contributions due to both large and small-bubbles fractions increase with gas velocity in both systems. However, with increase in yeast cells concentration, the contribution due to large bubbles decreases while that due to small bubbles increases. The decrease in gas holdup due to large-bubbles fraction can be attributed to increase in rise velocity of this fraction with increasing yeast concentration. Such changes in bubbles populations are expected to influence the processes of heat and mass transfer in the system. Measurements of heat transfer in different systems and different locations are discussed below. 3.2. Heat transfer
Fig. 7. Rise velocities of large bubbles population in different systems.
Fig. 9 shows heat transfer coefficients measured in the central and wall regions at the axial location of 1.28 m from
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Fig. 9. Local heat transfer coefficients obtained in bulk region of column with different systems (Z = 1.28 m).
the column bottom. At superficial gas velocities of 0.05 and 0.10 m/s, heat transfer probe was above the interface between the expanded slurry and foam sections. It can be seen that heat transfer coefficients for both superficial gas velocities are quite low, indicating low turbulence in the foam section. Also heat transfer coefficients between central and wall regions are close for these gas velocities (0.05 and 0.10 m/s). As the gas velocity increases further, the interface rises and heat transfer probe gets closer to the foam-slurry interface. At the gas velocity of 0.15 m/s, the interface fluctuated around the heat transfer probe. The heat transfer coefficients in this region are significantly higher than in the foam section indicating higher turbulence in the region. It can also be noted that differences between the wall and center are still low in the interface region but increase significantly as the probe moves into the bulk region (below the interface) at higher gas velocities (≥0.2 m/s). This indicates different mixing patterns in the interface or exit region compared to the bulk region. Based on above observations, it may be reasonable to assume radial uniformity in the foam section and interface region but not in the bulk region. Fig. 9 also compares heat transfer coefficients obtained in air–water–yeast cells system with those obtained in air–water and air–water–glass beads (11 m) systems. It can be observed that heat transfer coefficients in air–water and air–water–glass bead system are very close. In the air–water–yeast system, however, the heat transfer coefficients are higher in the central region but lower in the wall region. These observations can be related to changes in bubble size distribution observed earlier in the air–water–yeast cells system. The higher heat transfer coefficients in the central region can be attributed to increases in rise velocities of large bubbles in the air–water–yeast cells system. These
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large bubbles pass through the central region of the column and thus contribute to enhanced heat transfer rate in the region [23]. In the wall region, heat transfer is mainly controlled by the circulating liquid or slurry back-flow velocity and the influence of changes in gas holdup structures [23]. For example, the small bubbles which tend to accumulate in the wall region could influence the turbulence in the region. The heat transfer coefficients obtained at the wall and center of the bulk region were analyzed based on consecutive film and surface-renewal theory of Wasan and Ahluwalia [24]. This model assumes that a thin liquid film of thickness δ exists surrounding the heating surface, through which the heat transfer takes place by conduction. The outer surface of the film is continuously renewed with fluid elements induced by the bubble wake. During the contact, the heat is transferred by the elements through unsteady-state conduction. The temperature of fluid element sweeping the outer surface of the film is assumed to be uniform and equal to the bulk temperature. Thus, the heat transfer phenomenon is a sequential process of diffusion followed by convection. The time average heat transfer coefficient from heating surface to the bed can be expressed by the physical properties of suspension, the film thickness (δ) and the contact time between the liquid elements and the film (θ c ) as √ 2ksl ksl δ αθc αθc 1 − erf h= √ + exp −1 αθc δ δ2 π αθc (5) √ the term ( αθc /δ) accounts for the contribution of film resistance to heat transfer. The contact time (θ c ) at the center can be estimated using large bubble rise velocity and probe diameter. dp θc = (6) Ub,L The thickness of laminar viscous sub-layer depends on the geography of the surface of heat transfer source. For the cylindrical probe used in this study, the thickness of laminar viscous sub-layer at different angular locations with respect to the point of incidence can be expressed as [25]: δo =
A p dp 1/2
(7)
2 Rep
here dp is diameter of the cylindrical probe, parameter Ap depends on angular location on probe. The heat flux sensor was located at an angle of 90◦ in front of liquid flow. For the probe arrangement used in this study, the value of parameter Ap was selected to be 2.5 [25]. Rep is Reynolds number based on probe diameter and rise velocity of large bubbles (ρ sl Ub,L dp /µsl ) passing through central region of column. The film thickness of thermal conduction δ is equivalent to the thickness of the diffusion sublayer and is related to the laminar viscous sublayer, δ o as δ=
δo Pr1/3
(8)
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Table 1 Comparison of experimental local heat transfer coefficients with calculated values by the model of Wasan and Ahluwalia [18] Gas velocity (m/s)
Heat transfer coefficients (kW/m2 ◦ C) Air–water system
Air–water–yeast system
Wall region
0.20 0.30
Central region
Wall region
Central region
Experimental
Calculated
Experimental
Calculated
Experimental
Calculated
Experimental
Calculated
5.70 5.83
5.46 5.95
7.07 7.36
6.45 6.80
5.15 5.45
5.21 5.77
7.49 7.80
7.37 7.95
Thus, the film thickness can be calculated by combining Eqs. (7) and (8), expressed as: δ=
1.25dp
(9)
1/2
Rep Pr1/3
For the wall region, the above procedure needs to be modified. This region is dominated mainly by turbulent eddies generated by motion of bubbles and their interactions in the wall region. The contact time θ c in the wall region can be estimated by applying Kolmogoroff’s concept of isotropic turbulence [11] 1/2 1/2 νsl µsl θc = (10) = Pv Vg ρsl g To calculate the Reynolds number for estimation of boundary layer thickness, Li and Prakash [23] proposed use of average bubble rise velocity (Vg /ε g ). The calculated and experimental values of heat transfer coefficients are presented in Table 1. It can be seen that predicted values are generally within 10% for all experimental data. A few measurements were also made at the axial location of 0.52 m from column bottom in both air–water and air–water–yeast systems. Since the ratio of axial distance to column diameter is less than 2, this axial position can be considered to lie in the developing flow region. In this region, bubbles rising from the distributor region would break-up and coalesce to reach stable bubble size in the bulk region above it. Fig. 10 shows that at low gas velocities (≤0.10 m/s) there are no significant or systematic differences between the air–water and the air–water–yeast cells systems. At these gas velocities, the flow regime is between dispersed bubble and churn turbulent regime. However, at higher gas velocities (≥0.15 m/s), the heat transfer coefficients in the air–water–yeast cells system are lower in the wall region
Fig. 10. Heat transfer coefficients in air–water and air–water–yeast cells systems at axial location of 0.52 m from column bottom (yeast cells concentration: 0.4 wt.%).
but higher in the central region compared to the air–water system. This again indicates an increase in the size of large-bubbles fraction and a decrease in size of smaller bubble fraction due to the addition of yeast into water. Fig. 11 compares the effects of adding yeast cells into water on heat transfer coefficients in the bulk and developing regions at the column center. It can be observed that the difference between the two regions decreases due to addition of yeast cells. As shown in Table 2, the difference is decreased by about one-third. It could be concluded from data in Figs. 10
Table 2 Comparison of heat transfer coefficients at the center of column in bulk and developing regions in expanded slurry section Gas velocity (m/s)
Heat transfer coefficient (kW/m2 ◦ C) Air–water system
0.2 0.3
Air–water–yeast system
Bulk region
Developing region
Difference (%)
Bulk region
Developing region
Difference (%)
7.07 7.36
5.85 6.39
21 15
7.49 7.80
6.52 7.15
15 9
A. Prakash et al. / Biochemical Engineering Journal 9 (2001) 155–163
Fig. 11. Comparison of heat transfer coefficients in column center at different axial locations in air–water and air–water–yeast cells systems (yeast cells concentration: 0.4 wt.%).
and 11 that the process of stable bubble formation begins in the developing region and continues into the bulk region. 4. Concluding remarks The addition of yeast cells (Saccharomyces cerevisiae) in water altered the hydrodynamic and heat transfer behavior of the suspension. The gas holdup due to small-bubbles fraction increased and their rise velocities decreased while gas holdup due to large-bubbles fraction decreased and their rise velocities increased with increasing concentration of yeast. The layer of a stable foam bed observed at low gas velocities (≤0.1 m/s) in air–water–yeast system disappeared at higher gas velocities. The changes in local heat transfer coefficients measured at the wall and center in bulk region could be related to changes in bubble size distribution due to the addition of yeast. Heat transfer in foam bed was significantly lower than in expanded slurry bed region. Acknowledgements This research project was supported by the Natural Science and Engineering Research Council of Canada (NSERC) through individual research grants awarded to Dr. A. Prakash, Dr. A. Margaritis and Dr. M.A. Bergougnou. References [1] N. Ellis, A. Margaritis, C.L. Briens, M.A. Bergougnou, Fluidization characteristics of biobone particles used for biocatalysts, AIChE J. 42 (1996) 87.
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[2] N.A. Mensour, A. Margaritis, C.L. Briens, H. Pilkington, I. Russell, New developments in the Brewing Industry using immobilized yeast cell bioreactor systems, J. Inst. Brew. 103 (1997) 363. [3] A. Margarirts, D.W. teBokkel, D. Karamanev, Bubble rise velocities and drag coefficients in non-Newtonian polysaccharide solutions, Biotechnol. Bioeng. 64 (1999) 257. [4] H. Pilkington, A. Margaritis, N. Mensour, J. Sobezak, I. Hancock, I. Russell, Kappa-Carrageenan gel immbolization of larger Brewing yeast, J. Inst. Brew. 105 (1999) 398. [5] H. Heggart, A. Margaritis, R. Stewart, I. Russell, T. Dowhanick, Factors affecting yeast viability and vitality characteristics, MBAA Tech. Quart. 36 (1999) 383. [6] A. Prakash, A. Margaritis, R.C. Saunders, S. Vijayan, High concentrations ammonia removal by the cyanobacterium plectonema boryanum in a photobioreactor system, Can. J. Chem. Eng. 77 (1999) 99. [7] L.Z. Pino, R.B. Solari, Effect of operating conditions on gas holdup in slurry bubble columns with a foaming liquid, Chem. Eng. Comm. 117 (1992) 367. [8] S.C. Saxena, R. Vadivel, A.C. Saxena, Gas holdup and heat transfer from immersed surfaces in two- and three-phase systems in bubble columns, Chem. Eng. Commun. 85 (1989) 63. [9] E. Sada, H. Kumazawa, C. Lee, T. Iguchi, Gas holdup and mass-transfer characteristics in a three-phase bubble column, Ind. Eng. Chem. Process Des. Dev. 25 (1986) 472. [10] S. Kara, B.G. Kelkar, Y.T. Shah, N.L. Carr, Hydrodynamics and axial mixing in a three-phase bubble column, Ind. Eng. Chem. Process Des. Dev. 21 (1982) 584. [11] W.D. Deckwer, Y. Luoisi, A. Zaidi, M. Ralek, Hydrodynamic properties of the fischer–tropsch slurry process, Ind. Eng. Chem. Process Des. Dev. 19 (1980) 699. [12] H. Li, A. Prakash, Influence of slurry concentrations on bubble population and their rise velocities in a three-phase slurry bubble column, Powder Technol. 113 (2000a) 158. [13] R. Krishna, J.W.A. de Swart, J. Ellenberger, G.B. Martina, C. Maretto, Gas holdup in slurry bubble column: effect of column diameter and slurry concentrations, AIChE J. 43 (1997) 311. [14] C.L. Hyndman, F. Larachi, Christophe guy, understanding gas-phase hydrodynamics in bubble columns: a convective model based on kinetic theory, Chem. Eng. Sci. 52 (1997) 63. [15] J.G. Daly, S.A. Patel, D.B. Bukur, Measurement of gas holdups and sauter mean bubbles diameters in bubble column reactors by dynamic gas disengagement method, Chem. Eng. Sci. 47 (1992) 3647. [16] A. Schumpe, G. Grund, The gas disengagement technique for studying gas holdup structure in bubble column, Can. J. Chem. Eng. 64 (1986) 891. [17] H. Li, A. Prakash, Heat transfer and hydrodynamics in a three-phase slurry bubble column, Ind. Eng. Chem. Res. 36 (1997) 4688. [18] T. Oolman, H.W. Blanch, Non-Newtonian fermentation systems, CRC Crit. Rev. Biotechnol. 4 (1986) 133. [19] L.Z. Pino, M.M. Yepez, A.E. Saez, An experimental study of gas holdup in two-phase bubble columns with foaming liquids, Chem. Eng. Comm. 89 (1990) 155. [20] Y.T. Shah, S. Joseph, D.N. Smith, J.A. Ruether, On the behavior of the gas phase in a bubble column with ethanol–water mixtures, Ind. Eng. Chem. Process Des. Dev. 24 (1985) 1140. [21] L.S. Fan, Gas–Liquid–Solid Fluidization Engineering, Butterworth, Boston, 1989. [22] J.F. Walter, H.W. Blanch, Bubble break-up in gas–liquid bioreactors: break-up in turbulent flows, Chem. Eng. J. 32 (1986) B7. [23] H. Li, A. Prakash, Survey of heat transfer mechanism in different regions of a slurry bubble column, Can. J. Chem. Eng., 2000b, submitted for publication. [24] D.T. Wasan, M.S. Ahluwalia, Consecutive film and surface-renewal mechanism for heat and mass transfer from wall, Chem. Eng. Sci. 24 (1969) 1535. [25] H. Schlichting, Boundary Layer Theory, MaGraw-Hill, New York, 1960.
Hydrodynamics and local heat transfer measurements in a bubble column with suspension of yeast A. Prakash, A. Margaritis∗ , H. Li, M.A. Bergougnou Department of Chemical and Biochemical Engineering, The University of Western Ontario, London, Ont., Canada N6A 5B9 Received 12 January 2001; accepted 6 June 2001
Abstract Hydrodynamics and heat transfer were investigated in a 0.28 m diameter slurry bubble column for air–water–yeast cells system. Yeast cells of about 8 m diameter were used and the effects of gas velocity and yeast concentrations were studied. Gas holdups exhibited a maximum value around a gas superficial velocity of 0.10 m/s when foam height was included. Without the foam layer, gas holdups increased with increasing gas velocity. Bubble population was measured by means of dynamic gas disengagement technique. Rise velocity of small bubbles decreased, while rise velocity of large-bubbles fraction increased with yeast concentration which was varied from 0–0.4% w/w. Local heat transfer coefficients were measured at different axial and radial locations inside the column. Heat transfer in the foam section was significantly lower than in the main slurry column. In the bulk and developing regions, the addition of yeast cells into water increased heat transfer in the center and decreased at the wall. This work is one of very few studies reported in the literature on the use of actual yeast cells in slurry bubble columns and opens up new opportunities on their potential use as bioreactors in the fermentation industries, wastewater biological treatment and other applications. © 2001 Elsevier Science B.V. All rights reserved. Keywords: Slurry bubble column; Yeast cells; Hydrodynamics; Bubbles fractions; Heat transfer
1. Introduction Three-phase bubble columns are commonly used as chemical reactors and separation devices. The advantages of these columns include high heat and mass transfer rates, isothermal conditions, plug-free operation and ability for on-line catalyst addition and withdrawal. Three-phase bubble column have potential for applications in biochemical processes, such as fermentation and biological wastewater treatment [1–6]. When a slurry bubble column is used for suspensions containing micron-size particles such as yeast cells, the operation of the column can be different from that of a conventional slurry bubble column. Indeed, significant differences can arise in hydrodynamics and heat transfer characteristics due to the presence of different biochemical compounds associated with yeast cells and their growth medium. Moreover, the yeast cells can alter the air bubbles behavior, which in turn can alter the hydrodynamics as well heat transfer in the column. Therefore, there is a need to understand the hydrodynamics and heat transfer characteristics in the design of slurry bubble columns as bioreactors.
Gas holdup is a critical hydrodynamic parameter and it is generally found to be affected by several parameters including gas velocity, particle size and concentration and presence of surface-active agents [7–11]. Additional information from the study of bubble populations allow more detailed hydrodynamic analysis [12–16]. Heat transfer in slurry bubble columns has been investigated by several researchers [8,11,17]. However, most of the literature studies have used inert solid particles for their investigations. For applications in biochemical processes, it is important to study effects of micron-size cells concentration on column hydrodynamics and heat transfer. For example, micron-size cells, can alter bubble–bubble interactions as they can get immobilized on the surface of gas bubbles [18]. This study investigates hydrodynamics and heat transfer behavior of a slurry bubble column containing suspensions of yeast cells of about 8 m in diameter. The yeast samples were obtained from a local brewery and measurements have been made for gas holdups, bubble populations and their rise velocities and local heat transfer coefficients. 2. Experimental
∗ Corresponding author. E-mail address: [email protected] (A. Margaritis).
Experiments were conducted in a Plexiglas column of 0.28 m inside diameter and 2.4 m height (see Fig. 1). The
1369-703X/01/$ – see front matter © 2001 Elsevier Science B.V. All rights reserved. PII: S 1 3 6 9 - 7 0 3 X ( 0 1 ) 0 0 1 3 7 - 1
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Nomenclature A Ap Cp dp DC h havg HD HS HT k N Pv P/H q r R T Ub Ub,L Vb∞ VC Vg Vs w Rep Pr z
(m2 )
heat transfer area parameter in Eq. (7) specific heat (J/kg ◦ C) probe diameter (m) column diameter (m) heat transfer coefficient (kW/m2 ◦ C) average heat transfer coefficient (kW/m2 ◦ C) height of expanded slurry region (m) static height of the suspension (m) total height of dispersion including foam bed (m) thermal conductivity (kW/m◦ C) number of data points used to obtain time-averaged heat transfer coefficient power input per unit mass (m2 /s3 ) pressure gradient along column height (Pa/m) heat flow rate (kW) radial distance (m) column radius (m) temperature (K) bubble rise velocity (m/s) rise velocity of large bubbles population (m/s) terminal rise velocity of bubbles (m/s) average liquid circulation velocity (m/s) superficial gas velocity (m/s) hindered settling velocity of spherical particle (m/s) weight fraction in slurry Reynolds number based on probe diameter (ρUb dp /µ) Prandtl number for suspension (Cp µ/k) axial distance from column bottom (m)
Greek letters α thermal diffusivity, ksl /ρ sl Cp,sl (m2 /s) δo laminar boundary layer thickness (m) δ thermal boundary layer thickness (m) ε phase holdup εg,D gas phase holdup in expanded slurry region εg,T total gas phase holdup including foam section φs volume fraction of solids in gas-free slurry phase µ viscosity (Pa s) ν kinematic viscosity (m2 /s) θc contact time (s) ρ density (kg/m3 )
Subscripts b g l s sl S
bulk gas liquid solids slurry surface
liquid or slurry phase height before fluidization was maintained at 1 m. A six-arm gas distributor introduced air into the column bottom. The gas distributor arms were 5 mm in diameter and 0.14 m long. Each arm had four holes of 1.5 mm diameter facing downward. An electric heater was located near the bottom to maintain a constant temperature in the column. Compressed air and tap water constituted the gas and the liquid phases, respectively. The solid phase consisted of yeast cells obtained from a local brewery company. Before the use of fresh yeast cells in the slurry bubble column, they were treated with copper sulfate at 21◦ C for 4–5 h, in order to stop their biological activity by deactivating their respiration enzymes. Before the start of the experiments, the suspension was washed 3–4 times with water to remove most of CuSO4 and alcohol from it. The washing was carried out in the column by adding twice the volume of
Fig. 1. Schematic diagram of the experimental setup.
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Fig. 3. Structural arrangement of slurry bubble column with foam layer.
Fig. 2. Optical microscope photograph of dead yeast cells (Saccharomyces cerevisiae).
fresh tap water. This was followed by air agitation for about 15 min. The air flow was then stopped and yeast cells allowed to settle. The aqueous layer above the settled bed was then decanted. This process was repeated again until there was significant reduction in foaming during air agitation. Washing was stopped when the foam layer height was less than 50 mm for the selected gas velocity. This procedure introduced reasonable reproducibility in subsequent measurements. However, presence of foam layer did indicate small amounts of surface-active molecules still remaining in the suspension. It was decided to conduct the experiments with the foam layer formation and investigate column hydrodynamics and heat transfer for such conditions. Fig. 2 is an optical microscope picture that shows a typical population of dead yeast cells Saccharomyces cerevisiae. For a few experiments, 11 m glass beads were also used as solid phase. The column was operated in semi-batch mode at various superficial gas velocities and slurry concentrations. Slurry samples were withdrawn with the help of sampling probe located at about 0.8 m from bottom. The yeast cells concentration in the samples was measured by the dry weight method. During operation, a foam layer was formed at the top resulting in two sections as shown in Fig. 3. In the lower section, slurry was the continuous phase while in the
foam section at the top, gas was the continuous phase. Gas holdups were calculated with and without foam section, using Eqs. (1) and (2). Gas holdup with foam bed: εg,T =
HT − H S HT
(1)
Gas holdup without foam bed: εg,D =
HD − H S HD
(2)
Gas holdups were also obtained from the axial pressure difference measured by two pressure transducers (Omega PX541) mounted at 0.07 and 0.93 m from the column bottom and were calculated using Eq. (3). This equation has been derived by assuming that frictional pressure drop and inertial effects are negligible and particles concentration is constant along the column height [12]. εg = 1 −
1 P ρl + φs (ρs − ρl ) H
(3)
Gas holdup structure and bubble rise velocities were investigated by means of dynamic gas disengagement technique (DGD). After shutting-off the gas flow by closing air inlet in Fig. 1, the drop of instantaneous gas holdup was recorded. The DGD technique can distinguish between different bubble classes in a dispersion, if there are significant differences between their rise velocities [16]. Generally two
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classes of bubbles called large and small-bubbles fractions have been distinguished by this technique. Initially, there is a fast drop in the dispersion height indicating escape of fast rising large bubbles. This is followed by a slower drop in the dispersion height with the escape of slow rising small bubbles. The time-course of the instantaneous gas holdup can be used to evaluate the holdup fractions of small and large-bubbles fraction and their rise velocities [12]. Various assumptions and sources of errors associated with the DGD technique and their relative effects are reviewed in literature and discussed in details by Li and Prakash [12]. Based on this procedure, the measurements in this study provided relative fractions of the large and small bubbles populations and their rise velocities in the dispersion. Heat transfer was measured by means of heat transfer probes located at different axial and radial locations inside the column. The probes were designed with a micro-foil heat flow sensor mounted on the surface of a brass cylinder. Inside the cylinder, cartridge heater was installed as a heat source. The probes registered both heat flux and surface temperature. The construction details of the probe were described in a previous study [17]. Bed temperature was monitored by means of a digital thermometer. The heat transfer probes were located at axial locations of 0.07, 0.52 and 1.28 m from the column bottom. The signals of heat flux and probe surface temperature were collected simultaneously at 60 Hz for about 35 s. The microvolt signals from the heat flux sensor were amplified to millivolts before collection by a data acquisition system. The time-averaged
heat transfer coefficient at a given location was obtained by averaging the instantaneous heat transfer data: N
havg =
1 q/A N TS − T b
(4)
i=1
The number of samples (N) selected for averaging was 2100 to ensure a stable value of heat transfer coefficient.
3. Results and discussion 3.1. Hydrodynamics Variations of gas holdups with and without foam layer are shown in Fig. 4 for different yeast concentrations. High total gas holdups at gas velocities below 0.2 m/s can be attributed to a foam layer formation at the top of the dispersion. It has been known that foam is created when small amount of surface-active substances (e.g. alcohols) are added into a liquid [19,20]. The solutions used in this investigation had varying concentrations of surfactants (ethanol, protein, etc.). The surfactants concentration could be expected to be proportional to the concentration of yeast in the suspension. The increased gas holdup with increasing yeast concentration indicates a positive effect of surfactants’ concentration on foam formation (Fig. 4). The foam structure is known to be affected by superficial gas velocity [7,20]. It can be seen from Fig. 4, that the total gas holdup exhibits a maximum
Fig. 4. Gas holdups in air–water–yeast cells system measured by the bed height method.
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at superficial gas velocity around 0.10 m/s. Pino and Solari [7] observed a similar tendency in air–kerosene system containing surfactants and 5 m inert solids. The maximum gas holdup occurred at superficial gas velocity around 0.05 m/s. The location of maximum gas holdup could be a function of the nature of surfactants, solid particles, distributor type, etc. As the superficial gas velocity is increased beyond 0.10 m/s, fast rising large bubbles induce turbulent flow. A high degree of turbulence breaks the foam structure and decreases the gas holdup in this regime. Meanwhile, the difference in gas holdups with and without the foam section is significantly reduced. In this flow regime, the foam layer consisted of a short layer of small bubbles and it oscillated with the expanded slurry section. Fig. 5 shows that the gas holdups obtained from expanded slurry height method are generally higher than the values estimated by the differential pressure method. The difference is small at 0.05 m/s but increases with gas velocities up to about 0.2 m/s followed by a slight drop at higher gas velocity. The accumulation of small bubbles at the top of dispersion can add to gas holdup measured by the expanded bed height method. The measurements by pressure differential method excluded the contribution by this layer (the top pressure transducer was generally located below this layer). Increased turbulence intensity at high gas velocities could reduce this contribution by breaking the foam layer and by dispersing the small bubbles into the bulk phase.The hydrodynamics of the column were further investigated based on bubbles populations measurements. Two bubbles fractions,
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Fig. 6. Rise velocities of small bubbles population in different systems.
namely, large and small, could be detected by the DGD technique. Fig. 6 shows that small-bubbles rise velocity decreased with increasing superficial gas velocity. The decrease in rise velocity of small-bubbles fraction with increasing gas velocities is documented in literature [12]. It is also noted from Fig. 6 that the rise velocity of small bubbles decreased with increasing yeast concentrations. This decrease
Fig. 5. Comparison of gas holdup in air–water–yeast cells system measured by different methods (yeast cells concentration: 0.4 wt.%).
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in small-bubbles rise velocities can be attributed mainly to a proportional rise in the concentration of surface-active molecules in suspension (ethanol, proteins, etc.) with increasing yeast concentrations in the suspension and/or due to coalescence hindrance effects of yeast particles on bubbles surface. These two phenomena can have synergistic effects under appropriate conditions in the system. The hindrance effect of particles in a suspension has been observed only at low slurry concentrations [20]. These authors observed significantly higher gas holdups for low slurry concentration (1 wt.%) of fine particles (7 m) in air–electrolyte solution. The increase in gas holdup was attributed to coalescence hindering effect of the fine particles at the low slurry concentration. However, most literature studies have generally reported an opposite trend, i.e. increase in small-bubbles rise velocities (and decrease in gas holdup) with the addition of solids in a liquid [12,13]. The increase in rise velocity of small bubbles with the slurry concentration has been attributed to increase in apparent slurry viscosity [21]. For the low concentration suspensions of yeast used in this study, the change in apparent slurry viscosity is very small [18]. Therefore, this effect is not expected to play a significant role in affecting bubble size distribution in this study. The effects of gas velocity and yeast concentration on large-bubbles rise velocities are shown in Fig. 7. It can be observed that, the rise velocities of large-bubbles fraction increased with increasing superficial gas velocity or yeast concentration. Similar variations of large-bubbles rise velocities were also observed by Li and Prakash [12] in air–water–glass beads system. However, the large-bubbles rise velocities in the air–water–yeast system are found to be higher than those reported by Li and Prakash [12] for similar concentration of inert glass beads in suspension. This difference in behavior could be attributed to very dif-
Fig. 8. Gas holdups due to large and small-bubbles fractions in the dispersion.
ferent nature of the particles (and associated contaminants) used in the two studies. Walter and Blanch [22] investigated bubble break-up in fermentation broths of yeast and other bacteria. The maximum stable bubble size was observed to be significantly higher than in distilled water or dilute solutions of alcohols. This was attributed to different surface elastic behavior of the broths due to the presence of associated surfactants such as alcohols, proteins and salts. The bubble break-up mechanism could also be affected by the sweeping action of immobile yeast on bubble surface due to fast rising bubbles [22]. This action will tend to accumulate the surfactant molecules at the rear end of the bubbles, thus creating surface tension gradients on the bubble surface. Therefore, several additional factors come into play when investigating the behavior of suspensions with yeast cells. The gas holdup contributions due to small and large bubbles populations are presented in Fig. 8. It can be seen that contributions due to both large and small-bubbles fractions increase with gas velocity in both systems. However, with increase in yeast cells concentration, the contribution due to large bubbles decreases while that due to small bubbles increases. The decrease in gas holdup due to large-bubbles fraction can be attributed to increase in rise velocity of this fraction with increasing yeast concentration. Such changes in bubbles populations are expected to influence the processes of heat and mass transfer in the system. Measurements of heat transfer in different systems and different locations are discussed below. 3.2. Heat transfer
Fig. 7. Rise velocities of large bubbles population in different systems.
Fig. 9 shows heat transfer coefficients measured in the central and wall regions at the axial location of 1.28 m from
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Fig. 9. Local heat transfer coefficients obtained in bulk region of column with different systems (Z = 1.28 m).
the column bottom. At superficial gas velocities of 0.05 and 0.10 m/s, heat transfer probe was above the interface between the expanded slurry and foam sections. It can be seen that heat transfer coefficients for both superficial gas velocities are quite low, indicating low turbulence in the foam section. Also heat transfer coefficients between central and wall regions are close for these gas velocities (0.05 and 0.10 m/s). As the gas velocity increases further, the interface rises and heat transfer probe gets closer to the foam-slurry interface. At the gas velocity of 0.15 m/s, the interface fluctuated around the heat transfer probe. The heat transfer coefficients in this region are significantly higher than in the foam section indicating higher turbulence in the region. It can also be noted that differences between the wall and center are still low in the interface region but increase significantly as the probe moves into the bulk region (below the interface) at higher gas velocities (≥0.2 m/s). This indicates different mixing patterns in the interface or exit region compared to the bulk region. Based on above observations, it may be reasonable to assume radial uniformity in the foam section and interface region but not in the bulk region. Fig. 9 also compares heat transfer coefficients obtained in air–water–yeast cells system with those obtained in air–water and air–water–glass beads (11 m) systems. It can be observed that heat transfer coefficients in air–water and air–water–glass bead system are very close. In the air–water–yeast system, however, the heat transfer coefficients are higher in the central region but lower in the wall region. These observations can be related to changes in bubble size distribution observed earlier in the air–water–yeast cells system. The higher heat transfer coefficients in the central region can be attributed to increases in rise velocities of large bubbles in the air–water–yeast cells system. These
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large bubbles pass through the central region of the column and thus contribute to enhanced heat transfer rate in the region [23]. In the wall region, heat transfer is mainly controlled by the circulating liquid or slurry back-flow velocity and the influence of changes in gas holdup structures [23]. For example, the small bubbles which tend to accumulate in the wall region could influence the turbulence in the region. The heat transfer coefficients obtained at the wall and center of the bulk region were analyzed based on consecutive film and surface-renewal theory of Wasan and Ahluwalia [24]. This model assumes that a thin liquid film of thickness δ exists surrounding the heating surface, through which the heat transfer takes place by conduction. The outer surface of the film is continuously renewed with fluid elements induced by the bubble wake. During the contact, the heat is transferred by the elements through unsteady-state conduction. The temperature of fluid element sweeping the outer surface of the film is assumed to be uniform and equal to the bulk temperature. Thus, the heat transfer phenomenon is a sequential process of diffusion followed by convection. The time average heat transfer coefficient from heating surface to the bed can be expressed by the physical properties of suspension, the film thickness (δ) and the contact time between the liquid elements and the film (θ c ) as √ 2ksl ksl δ αθc αθc 1 − erf h= √ + exp −1 αθc δ δ2 π αθc (5) √ the term ( αθc /δ) accounts for the contribution of film resistance to heat transfer. The contact time (θ c ) at the center can be estimated using large bubble rise velocity and probe diameter. dp θc = (6) Ub,L The thickness of laminar viscous sub-layer depends on the geography of the surface of heat transfer source. For the cylindrical probe used in this study, the thickness of laminar viscous sub-layer at different angular locations with respect to the point of incidence can be expressed as [25]: δo =
A p dp 1/2
(7)
2 Rep
here dp is diameter of the cylindrical probe, parameter Ap depends on angular location on probe. The heat flux sensor was located at an angle of 90◦ in front of liquid flow. For the probe arrangement used in this study, the value of parameter Ap was selected to be 2.5 [25]. Rep is Reynolds number based on probe diameter and rise velocity of large bubbles (ρ sl Ub,L dp /µsl ) passing through central region of column. The film thickness of thermal conduction δ is equivalent to the thickness of the diffusion sublayer and is related to the laminar viscous sublayer, δ o as δ=
δo Pr1/3
(8)
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Table 1 Comparison of experimental local heat transfer coefficients with calculated values by the model of Wasan and Ahluwalia [18] Gas velocity (m/s)
Heat transfer coefficients (kW/m2 ◦ C) Air–water system
Air–water–yeast system
Wall region
0.20 0.30
Central region
Wall region
Central region
Experimental
Calculated
Experimental
Calculated
Experimental
Calculated
Experimental
Calculated
5.70 5.83
5.46 5.95
7.07 7.36
6.45 6.80
5.15 5.45
5.21 5.77
7.49 7.80
7.37 7.95
Thus, the film thickness can be calculated by combining Eqs. (7) and (8), expressed as: δ=
1.25dp
(9)
1/2
Rep Pr1/3
For the wall region, the above procedure needs to be modified. This region is dominated mainly by turbulent eddies generated by motion of bubbles and their interactions in the wall region. The contact time θ c in the wall region can be estimated by applying Kolmogoroff’s concept of isotropic turbulence [11] 1/2 1/2 νsl µsl θc = (10) = Pv Vg ρsl g To calculate the Reynolds number for estimation of boundary layer thickness, Li and Prakash [23] proposed use of average bubble rise velocity (Vg /ε g ). The calculated and experimental values of heat transfer coefficients are presented in Table 1. It can be seen that predicted values are generally within 10% for all experimental data. A few measurements were also made at the axial location of 0.52 m from column bottom in both air–water and air–water–yeast systems. Since the ratio of axial distance to column diameter is less than 2, this axial position can be considered to lie in the developing flow region. In this region, bubbles rising from the distributor region would break-up and coalesce to reach stable bubble size in the bulk region above it. Fig. 10 shows that at low gas velocities (≤0.10 m/s) there are no significant or systematic differences between the air–water and the air–water–yeast cells systems. At these gas velocities, the flow regime is between dispersed bubble and churn turbulent regime. However, at higher gas velocities (≥0.15 m/s), the heat transfer coefficients in the air–water–yeast cells system are lower in the wall region
Fig. 10. Heat transfer coefficients in air–water and air–water–yeast cells systems at axial location of 0.52 m from column bottom (yeast cells concentration: 0.4 wt.%).
but higher in the central region compared to the air–water system. This again indicates an increase in the size of large-bubbles fraction and a decrease in size of smaller bubble fraction due to the addition of yeast into water. Fig. 11 compares the effects of adding yeast cells into water on heat transfer coefficients in the bulk and developing regions at the column center. It can be observed that the difference between the two regions decreases due to addition of yeast cells. As shown in Table 2, the difference is decreased by about one-third. It could be concluded from data in Figs. 10
Table 2 Comparison of heat transfer coefficients at the center of column in bulk and developing regions in expanded slurry section Gas velocity (m/s)
Heat transfer coefficient (kW/m2 ◦ C) Air–water system
0.2 0.3
Air–water–yeast system
Bulk region
Developing region
Difference (%)
Bulk region
Developing region
Difference (%)
7.07 7.36
5.85 6.39
21 15
7.49 7.80
6.52 7.15
15 9
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Fig. 11. Comparison of heat transfer coefficients in column center at different axial locations in air–water and air–water–yeast cells systems (yeast cells concentration: 0.4 wt.%).
and 11 that the process of stable bubble formation begins in the developing region and continues into the bulk region. 4. Concluding remarks The addition of yeast cells (Saccharomyces cerevisiae) in water altered the hydrodynamic and heat transfer behavior of the suspension. The gas holdup due to small-bubbles fraction increased and their rise velocities decreased while gas holdup due to large-bubbles fraction decreased and their rise velocities increased with increasing concentration of yeast. The layer of a stable foam bed observed at low gas velocities (≤0.1 m/s) in air–water–yeast system disappeared at higher gas velocities. The changes in local heat transfer coefficients measured at the wall and center in bulk region could be related to changes in bubble size distribution due to the addition of yeast. Heat transfer in foam bed was significantly lower than in expanded slurry bed region. Acknowledgements This research project was supported by the Natural Science and Engineering Research Council of Canada (NSERC) through individual research grants awarded to Dr. A. Prakash, Dr. A. Margaritis and Dr. M.A. Bergougnou. References [1] N. Ellis, A. Margaritis, C.L. Briens, M.A. Bergougnou, Fluidization characteristics of biobone particles used for biocatalysts, AIChE J. 42 (1996) 87.
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