Flow Pattern, Mixing, Gas Hold-Up and Mass ... - ACS Publications

Article pubs.acs.org/IECR

Flow Pattern, Mixing, Gas Hold-Up and Mass Transfer Coefficient of Triple-Impeller Configurations in Stirred Tank Bioreactors Minghui Xie, Jianye Xia,* Zhen Zhou, Ju Chu, Yingping Zhuang, and Siliang Zhang State Key Laboratory of Bioreactor Engineering, East China University of Science and Technology, Shanghai 200237, China ABSTRACT: Global and local gas−liquid characteristics (gas hold-up, volumetric mass transfer coefficient), and flow field, mixing time of the liquid phase are investigated for various triple-impeller configurations. Four types of impellers (Rushton turbine (RT), hollow blade turbine (HBT), wide-blade hydrofoil impeller pumping down (WHd), and pumping up (WHu)) were used to form four combinations (3RT, HBT+2WHd, HBT+2WHu, and 3WHu). The results show that the axial impellers combination (3WHu) provides more effective homogenization performance than the impellers with combined radial and axial flow (HBT+2WHu, HBT+2WHd), while the radial impellers combination (3RT) is the worst. When the gas superficial velocity is 1.625 mm s−1, 3WHu produces a 53% higher mass transfer coefficient than HBT+2WHd, HBT+2WHu, and 3RT lie between them. When the gas superficial velocity reaches up to 8.124 mm s−1, however, all of the tested configurations give almost similar mass transfer coefficients under equivalent power input. For 3RT, the highest hold-up is in the bottom impeller discharge stream and near the wall for the middle and top impellers. For the HBT+2WHd combination, there was no large variations of gas hold up in the bulk except the region around the bottom impeller. For HBT+2WHu and 3WHu, high gas hold-up was observed between the two up pumping impellers, and moderately low gas hold-up above the top impeller. There are three zones of higher interfacial area for the HBT +2WHu and 3RT combination, two zones of higher interfacial area for 3WHu, and only one zone of higher interfacial area for HBT+2WHd combination. At low gas velocity, the flow pattern generated by each impeller combination results in different gas bubble trajectory and different bubble breakup and coalescence kinetics, which in turn influences both local and average gas holdup directly, and also affects the local specific interfacial area, that is, influences the mass transfer coefficient indirectly. At higher gas velocity, the power drop also contributes to change of the gas hold-up and mass transfer at the same specific power consumption. researchers14−16 studied modifications of the Rushton turbine in agitated tanks and suggested that a better performance may be expected after retrofitting of Rushton turbines with streamlined impellers. Some axial impellers or axial and radical impeller combinations are able to disperse gas in liquids and yield good mass transfer performances. Most of the investigations done are limited to only measurement of the overall gas hold-up and volumetric mass transfer coefficient. However, without local information on both gas hold up, bubble size and volumetric mass transfer coefficient, it is hard to fully understand what difference between different impeller combinations causes the discrepancy on the overall characteristics. In fermentation process, some researchers have studied multiple impeller combinations in the fermentation field. Jüsten et al.17 examined the influence of the agitation conditions on the growth, morphology, vacuolation, and productivity of Penicillium chrysogenum in 6 L fed-batch fermentations using a standard Rushton turbine, a four-bladed paddle, and a sixbladed pitched blade impeller. They showed that changes of morphology, specific growth rate, and specific penicillin production rate depended on impeller geometry at given power input. Joo-Hyung et al.18 investigated the effect of

1. INTRODUCTION Multiple-impeller gas−liquid reactors are widely used in industrial bioprocess. The multiple-impeller systems produce larger hold-ups and improved volumetric mass transfer coefficient as compared to the single-impeller system at equivalent power dissipation due to higher residence time of the bubbles.1−4 For shear sensitive microorganisms, multipleimpeller systems are favored due to lower impeller speed for equivalent power dissipation resulting in lower values of the maximum shear, whereas for the rational design of stirred reactors power consumption, mixing, gas hold-up, and volumetric mass transfer coefficient are very important aspects. Sometimes, the mixing performance is important as well as interfacial mass transfer intensity. For instance, the local concentration gradients of substrate or pH occurring in large scale fermentations as a consequence of imperfect mixing are harmful for bioprocesses.5 The mixing performance for various impeller configurations has been studied.6−10 The axial flow impellers have higher mixing efficiency while reducing energy consumption when they are compared with radial turbines. The flow field generated by impeller plays a considerable role on mixing. Rushton turbine impellers are usually used in the laboratory and industry, which give very good dispersion and hence good gas mass transfer rates11,12 but are also responsible for a significant drop in power draw on gassing, low bulk liquid circulation in vessel, uneven distribution of shear, and energy dissipation.4,13 Both these disadvantages can be tackled by choosing an appropriate impeller to be combined with it. Many © 2014 American Chemical Society

Received: Revised: Accepted: Published: 5941

March 14, 2013 August 20, 2013 March 3, 2014 March 3, 2014 dx.doi.org/10.1021/ie400831s | Ind. Eng. Chem. Res. 2014, 53, 5941−5953

Industrial & Engineering Chemistry Research

Article

impeller types on morphology and protein expression in a submerged culture of Aspergillus oryzae. They found that cells cultured by propeller agitation grew in the form of a pellet, which showed high levels of protein production for both the intracellular heterologous protein (β-glucuronidase) and the extracellular homologous protein (α-amylase), whereas cells cultured by turbine agitation grew with freely dispersed-hyphal and clump form. Li et al.19 employed radial-flow impellers and radial-axial flow impellers for lincomycin fermentation and showed the yield of lincomycin in a 60 m3 industrial fermentor equipped with radial-axial flow impellers was increased by 46% compared to that obtained in the radial-flow impellers fermentor. But in general, thoroughly comparison between different multiple impeller combinations was limited reported. We have found some interesting results using different impeller combinations in some fermentation experiments.19,20 Glucoamylase fermentations using Aspergillus niger in a 50 L fermenter were conducted with 3RT, HBT+2WHu, and 3WHu. During the fermentation process, the rotational speed of impeller was adjusted for each combination to keep the same OUR (oxygen uptake rate) profile during the whole process (144h). The initial rotational speed of three combinations were all 100 rpm, the rotational speed of 3RT increased from 100 to 360 rpm, the rotational speed of HBT+2WHu increased from 100 to 330 rpm, the rotational speed of 3WHu increased from 100 to 470 rpm in the fermentation process. Furthermore, the final glucoamylase activity was 214, 263, and 275 AGI/mL respectively corresponding to power consumptions of 1.345, 0.473, and 0.662 kW/m3 for each combination. Therefore, it can be seen that culture with HBT+2WHu combination showed the highest enzyme activity, while culture with 3RT combination showed the lowest enzyme activity under the same specific power consumption. It can be seen that the flow filed generated by the impeller system has great impact on the fermentation process. Whereas, little is known about multipleimpeller configuration’s effect on the glucoamylase fermentation, so it is necessary to understand mass transfer and mixing performance of these combinations. The objective of this work was to provide global characteristics (gas hold-up, volumetric mass transfer coefficient, and mixing time) and local gas−liquid flow characteristics (using conductivity probes) for four impeller types (Rushton turbines, hollow blade turbine, and wide-blade hydrofoil impellers pumping down or up) in triple-impeller configurations. The flow fields measured by particle image velocimetry are to study the effect of these flow regimes on the mixing time for the impeller combinations. General correlations of mixing time, gas hold-up and volumetric mass transfer coefficient for all impeller combinations are developed for low-viscous coalescent system.

Figure 1. Schematic diagram of the experimental setup.

hollow blade turbine in bottom stage with dual wide-blade hydrofoil impellers pumping down in upper stages and 3WHu represents triple wide-blade hydrofoil impeller pumping up. The experiments were carried out at room temperature (25 ± 1 °C) and atmospheric pressure. Plain tap water was used as the working fluid (coalescent system) and air was introduced through a ring sparger with 18 equally spaced orifices (1.5 mm in diameter). The sparger with diameter of 90 mm is located 32 mm below the bottom impeller. The superficial gas velocity tested in this work ranges from 1.625 to 16.25 mm s−1 (equivalent to 0.2−2 vvm). The experiments were performed in the impeller speed ranging from 100 to 550 rpm. At the low gas flow rate (0.2vvm), the complete dispersion speed is 130 rpm for 3RT, 150 rpm for HBT+2WHu, 170 rpm for HBT +2WHd, and 200 rpm for 3WHu. At the high gas flow rate (1.0vvm), the complete dispersion speed is 220 rpm for 3RT, 250 rpm for HBT+2WHu, 280 rpm for HBT+2WHd, and 300 rpm for 3WHu. 2.2. PIV Measurement of Flow Pattern. The 2D PIV system employed in this investigation is a commercial system (ILA, German), including a laser (Leamtech, Nd:YAG, 200 mJ, 15 Hz), a PIV camera (PCO2000, 2048 × 2048 pixels), a synchronizer and Vidpiv4.7 software. The vertical laser plane in all of the experiments was placed at angle of 5° behind the nearest baffle, due to distortions produced by the imperfect matching of refractive index between the working fluid and the polymethyl methacrylate baffle.21 An ellipsoid bottom cylindrical perspex stirred tank is used for PIV measurement. And its diameter is 280 mm, filling liquid level is 500 mm, the height of the ellipsoid base is 70 mm. The rotating speed used in the test is 350 rpm, corresponding to a Reynolds number of Re = 8.4 × 104, thus satisfying operation within the turbulent flow regime (Re > 4 × 103). Hollow glass particles with diameter of 10 μm are used as tracer particles (the density is 1.1−1.2 g cm−3). Vidpiv4.7 is used to interrogate the image with two-frame cross-correlation between 24 × 24 pixels windows and 50% overlap between interrogation windows, and then the vector resolution is about 2.82 mm. Delay between the couple laser pulses is optimized to ensure that the maximum in-plane displacements of the particles are less than one-quarter of the interrogation size.22 The vector data obtained are filtered according to velocity magnitude and signal-to-noise ratio. This step causes patently invalid data to be excluded. A global filter is applied to remove velocities greater than twice the value of the impeller tip speed,

2. MATERIAL AND METHODS 2.1. Experimental Setup. Schematic diagram of the ellipsoid bottom stirred tank used in this work is shown in Figure 1. The tank’s diameter, T, is 280 mm, and filling liquid level, H, is 500 mm. Height of ellipsoid base is 35 mm. The tank has four vertical baffles (width B = 0.1T), which are equally mounted around the tank wall. Three impellers are mounted along the shaft with equal distance of 130 mm, and the bottom impeller’s clearance to the tank bottom is 112 mm. Impeller types tested in this work are shown in Figure 2. They are selected and combined to form four different impeller combinations. The abbreviation used for defining the impeller combinations is straightforward: e.g., HBT+2WHd means 5942

dx.doi.org/10.1021/ie400831s | Ind. Eng. Chem. Res. 2014, 53, 5941−5953


Article

Figure 2. Illustrations of impellers used in this study.

2Utip. A local filter is applied to remove vectors three times bigger in magnitude than the standard deviation of the magnitude of surrounding vectors in a 5 × 5 grid. The velocity vectors removed are replaced by interpolated values from the same 5 × 5 grid. Meanwhile, the valid vectors should be greater than 98% of the processed vectors after the filtering procedures. 2.3. Measurement of Mixing Time. Mixing time measurements were performed using conductivity measurements. NaCl solution as a tracer was rapidly injected below the liquid surface with the aid of a syringe. To eliminate the influence of high salt concentration on bubble size/overall gas volume fraction,23 less than 0.05 vol % of NaCl solution (8 mL, 200 g/dm3) was used for measurements of mixing time in this work. A conductivity electrode (MI915, 0.635 cm in diameter, Microelectrodes Inc. the United States) is mounted under the bottom impeller (Figure 1), and its signal is acquired via a data acquisition board (PCI9118 A/D converter card from ADLINK Technology Inc.). This method proposed by Gupta et al.24 was used to filter out the biased noise from the raw signals measured by the conductivity probe in this work. The mixing time t0.05 (time cost to reach 95% homogeneity) is used in this work to represent mixing intensity of different impeller configurations. 2.4. Measurement of Global Gas−Liquid Flow Characteristics. Power consumption was measured by a toque sensor mounted on the impeller shaft. Average gas hold-up was measured by visual inspection. The overall volume mass transfer coefficient, kLa, was measured by dynamic method.25,26 The oxygen dissolved in the liquid phase is eliminated by means of bubbling nitrogen. After the oxygen concentration is equal to zero, air is introduced into the stirred tank again, and variation of oxygen concentration with time is logged. Dissolved oxygen level is measured using oxygen probe (Hamilton OXYFERM VP 120), the response time (63%) of the probe is 5 s. We also employed the correction model proposed by Garcia-Ochoa and Gomez27 to correct kLa value. The probes located at 30 mm above each impeller and 90 mm from the wall. The measurements for kLa were repeated three times, which gave reproducibility of measurements within ±10%. The kLa was calculated as an arithmetic mean of individual stage values. In order to compare with literatures, mass transfer coefficients were normalized to 20 °C according to Jackson and Shen:28 (kLa)20 ° C = 1.022(20 − θ)(kLa)θ

2.5. Conductivity Probe Technique. The dual-tip conductivity probe technique is widely used in two-phase bubbly flow due to its capability for measuring the local gas− liquid flow characteristics.29−32 The measuring system consists of a conductivity probe, a measurement circuit, a digital highspeed acquisition board (PCI-9118HG), and software designed to process signals. Three platinum needles K2−0.26||(Kida electronics CO., LTD) of 0.15 mm in diameter, coated with insulating green varnish, were used to serve as the sensor probes as shown in Figure 1. The diameter of the tip of each platinum needle is about 20 μm. The vertical distance between needle tips of probes 1 and 2 is 0.50 mm. When a bubble hits the tip of each needle, it will switch off the corresponding circuit and a high voltage output signal will be produced. However, when a bubble leaves the needle, the corresponding circuit connects again and a low voltage signal will be recorded. The high-speed acquisition board can transfer these voltage signals to the computer, and these signals are then processed to get more bubble information. Sampling frequency is set to 20 kHz for each sensor, and at least 1000 bubbles are sampled for statistics analysis under these conditions. First, the original signal was binarized with a threshold. Since noise commonly corrupts the ideal output signal, it becomes necessary to determine a proper threshold. The value of the threshold is typically determined by calibrating the sensor probe to the gas hold-up of the test section by pressure measurement33,34 or by photographic method.29,35 We found that the threshold of 10% of the pulse amplitude gives the nearest result of gas hold-up to that measured by the pressure measured. Therefore, threshold of 10% was used to process original conductivity signals, and this value was in agreement with that employed in many literatures.30,36,37 Then, binarized signals and the dwell time τ1 and τ2 which are the time that a bubble spends at tips 1 and 2, respectively, are shown in Figure 1. Local gas hold-up is estimated from the ratio of dwell time to the total data acquisition time.38 The bubble chord length, lb, were calculated by the following equations:

lb = ubτ1 = Lpτ1/Δt

(2)

where ub is the bubble velocity which is obtained from the distance between the two probe tips (Lp) and the time lag (Δt) for two different probe tips hit the same bubble. The signals obtained by distinct sensor are not perfectly identical, it is necessary to adopt a set of filtering criteria to

(1) 5943



Article

distinguish the pair of signals resulting from identical bubbles.39,40 In the analysis, it is assumed that the movements of the bubbles were in the vertical. Signal pairs are accepted if they satisfy the following conditions:30,41 1. trise,1 < t rise,2 or tdown,2 < t down,1. 2. trise,1 + τ1 < trise,2 + τ2 or tdown,2 + τ2 < tdown,1 + τ1. 3. 0.75 < 2τ1/(τ1 + τ2) < 1.25 or 0.75 < 2τ2/(τ1 + τ2) < 1.25. The unsatisfied pairs of the signals are removed from further data analysis procedures. A chord length will be smaller than the diameter of the bubble due to the probe will not always intersect the bubble at its center: In this study, the reconstruction method of Liu and Clark42 is employed to transform the chord length distribution into the bubble diameter distribution. The dual-tip conductivity probe can measure bubble velocity in the vertical direction, in order to get local bubble velocity in two directions, we tried to design a triple-tip conductivity probe (Figure 1). Probes 1 and 2 were used to measure axial velocity (see above) and probes 2 and 3 to measure velocity in radial direction. Probes 2 and 3 are located in the vertical plane through the axial of the tank with plane of probes 1 and 2 perpendicular to it. The horizontal distance between the needle tips of probes 2 and 3 is 0.52 mm. For the horizontal direction, signal pairs were accepted if they satisfy the following conditions: 1. trise,2 < t rise,3. 2. trise,2 + τ2 < trise,3 + τ3. 3. 0.75 < 2τ2/(τ2 + τ3) < 1.25. The unsatisfied pairs of the signals were removed from further radial velocity analysis.

3. RESULTS AND DISCUSSION 3.1. Flow Pattern Generated by Different Impeller Combinations. Time-averaged liquid flow fields for the 3RT, HBT+2WHd, HBT+2WHu, and 3WHu configurations by PIV technique under ungassed conditions are shown on Figure 3. The velocity vectors shown are normalized by the impeller tip speed, Utip. Figure 3a shows the typical flow pattern for the 3RT. The fluid is discharged radially with a maximum velocity magnitude of ∼0.6Utip. Each RT impeller produces its characteristic upper and lower loops, so six stable loops are observed, similar to that of Nienow.43 Figure 3b illustrates the flow pattern for the HBT+2WHd. Fluid is discharged from beneath each WHd and down pumping flow cascades between the two WHd impellers, which form a whole axial loop (loop2) with a maximum velocity magnitude of ∼0.5Utip. There are two loops created by the HBT impeller, but the upper loop is merged into the loop2. At the upper part (z/H = 0.8) that close to the impeller shaft, the flow is less affected by the flow loop and relatively quiescent. For the HBT+2WHu (see Figure 3c), there are two loops observed by HBT impeller, the upper loop2 is bigger than the lower loop1. The impeller discharges of upper two WHu impellers pump upward and extend to a height of z/H = 0.8 at the wall, then it splits into two loops, one a relatively strong downward flow at the wall, which then recirculates back into the middle impeller. Different to HBT+2WHd, loop4 was formed in the upper part of the vessel for z/H > 0.85. Velocity magnitudes in the upper flow loop are less intense, typically less than 0.1Utip. For the 3WHu (see Figure 3d), the impeller

Figure 3. Flow fields for all impeller combinations under ungassed condition.

discharges of WHu impellers pump upward and extend to a height of z/H = 0.9 at the wall, then it also splits into two loops. It is very difficult to obtain the gas and liquid flow field for PIV measurement in the gas−liquid system under high gas hold-up conditions. We tried to get the velocity distributions of bubbles by triple-tip conductivity probes technique, and results for the tested four impeller combinations are shown in Figure 4. There are also two circulation loops generated by each Rushton turbine under gassed conditions, but the lower loop move toward to the wall. Though HBT+2WHd can produce a whole circulation loop at upper part of the tank in singular phase, it cannot get the bubbles circulated in a whole loop. In contrast, the HBT+2WHu combination produces a whole loop under gas−liquid phases conditions instead of two circulation loops in single phase. More bubbles join the gas circulation loops created by the WHu impellers. The gas circulation loops created by the 3WHu are similar to the liquid circulation loops. 3.2. Mixing Efficiency. There are two kinds of correlations for mixing time. One is based on bulk flow mode, which correlates the mixing time with flow number, Fl44 or power number, Po45 as 5944



Article

Because power input per unit volume controls the turbulence in the microscale range, eq 6 is indeed consistent with the conclusion by Nienow.46 This eq 6 was used in this work to fit the experimental measured mixing time using different impeller combinations. The fitted equation for different impeller is shown in Table 1. It Table 1. Correlations for Prediction of Mixing Time under Ungassed Conditions configuration 3RT HBT+2WHu HBT+2WHd 3WHu

equations −1/3

t0.05 = 176.52(P/V) t0.05N = 5.2Po−1/3(D/T)−2(H/T)2.90 t0.05 = 53.31(P/V) −1/3 t0.05N = 5.2Po−1/3(D/T)−2(H/T)1.22 t0.05 = 32.63(P/V) −1/3 t0.05N = 5.2Po−1/3(D/T)−2(H/T)0.33 t0.05 = 24.78(P/V) −1/3 t0.05N = 5.2Po−1/3(D/T)−2(H/T)0.01

std. dev, %

R2

3.70 11.63 4.42 4.23 5.89 5.57 6.04 6.87

0.987 0.986 0.981 0.994 0.979 0.990 0.982 0.993

can be clearly seen that K1 is strongly dependent on impeller configuration while K2 is independent of it. Vasconcelos et al.14 reported K1 of 150 considering the radial dual-impellers (T/D = 3, H/T = 2, the basic Rushton turbine and the modifications with respect to blades). In this work the K1 is 176.52 for the triple Rushton turbines (T/D = 2.33, H/T = 1.79). Among the four impeller combinations tested in this work, the 3RT gives the longest mixing time at the same power consumption (Figure 5a). HBT+2WHu gives the second longest mixing time with 69.8% lower than that of 3RT; however, HBT+2WHd provides better mixing performance with 81.5% lower than the mixing time of 3RT. 3WHu produces the shortest mixing time

Figure 4. Flow field of bubbles for all impeller combinations under gassed condition.

t0.05N = 3.9(D/T )−3 Fl −1

(3)

t0.05N = 3(D/T )−3 Po−1/3

(4)

The other is based on turbulence mode. Nienow46 has shown based on turbulence theory47 that, with single impellers, the mixing time is independent of the impeller type and can be calculated from t0.05 = 5.9T 2/3(P /ρV )−1/3 (D/T )−1/3

(5)

Vasconcelos et al.14 reported the mixing time correlation in terms of the specific power input is given as

t0.05 = K1(P /V )K2

Figure 5. Mixing time for all impeller combinations under (a) ungassed and (b) gassed conditions.

(6) 5945



Article

significantly in the case of flooding. These facts can be explained with the hampering of the gas bubbles on liquid mixing at high rotational speed (agitator-controlled regime) and enhancement of gas bubbles on axial liquid mixing at low speeds (gas-controlled regime).52 Pinelli and Magelli9 have studied the effect of aeration rate on the axial dispersion coefficient and have reported that mixing time extent increases with an increase in QG at high agitation speeds, whereas a partial reversal of this effect take place at low agitation speeds. Therefore, data under the agitator-controlled regime is used to correlate with power input per volume.

at the same power dissipation with 85.9% lower than that of 3RT. The great difference of mixing efficiency for the four impeller combination is the direct results of the distinct flow patterns generated by each impeller combination. Correlated K1 values rank is 3RT > HBT+2WHu > HBT+2WHd > 3WHu. The turbulence-based correlations by Nienow46 established for blending in single-impeller systems give a K1 value of 33.49 in eq 6 in this work. The K1 values for 3RT and HBT+2WHu are larger than 33.49, the K1 value for HBT +2WHd gets pretty close to 33.49, and the K1 value for 3WHu is lower than 33.49. The mixing time for the four impeller combinations can be correlated by the relationship in terms of power input per volume and power number under ungassed condition −1/3

t0.05 = 9.692(P /V )

1.124

(Po)

t0.05 = K1(PG/V )K2

Correlated parameters of eq 10 are given in Table 2. Correlated exponent with a value of −0.244 is clearly higher

(7)

Table 2. Correlations for Prediction of Mixing Time under Gassed Conditions

with standard deviation is 14.78% and correlation coefficient is 0.979. Grenville and Nienow48 carried out a large amount of work on the mixing of liquids within the turbulent regime for Newtonian fluids and correlated the dimensionless mixing time in the turbulent regime to eq 8. t0.05N = 5.2Po−1/3(D/T )−2 (H /T )1/2

configuration 3RT

HBT+2WHu

(8)

But the data for this correlation was mainly collected on vessels with H/T < 1, and it is not really applicable to large aspect ratios. Cooke49 found that the exponent of H/T was to be 2.43 for three turbines. Rodgers et al.50 explored the exponent on the H/T term used the following correlation: −1/3

t0.05N = 5.2Po

−2

α

(D / T ) (H / T )

(10)

HBT+2WHd

3WHu

(9)

They found that α varies with agitator type, and the value of α is 1.72 for the two Rushton turbines. This correlation was also used to fit the experimental data of mixing time in this work. the exponent of H/T, α, is 2.9 for 3RT, 1.22 for HBT+2WHu, 0.33 for HBT+2WHd, and the α is nearly zero for 3WHu. The value of α for 3RT in this work is larger than that obtained by Cooke,49 this may be caused by a different measured point or the size of ellipsoid base. The four impeller combinations have the distinct exponent α is due to the different behavior of radial and axial impellers on liquid homogenization. Generally, a radial impeller generates one upper loop and one lower loop, an axial flow impeller generates one loop. when the H/T increases may also generate secondary flow loops.50 For impeller combinations, the loops formed by impellers are also affected by impeller spacing. The specific flow pattern generated by the four impeller combinations are shown in Figure 3. There are six circulation loops for the 3RT, four loops for the HBT+2WHu, and two loops for HBT+2WHd and 3WHu. Zoning has been shown to significantly increase the mixing time in vessels, and the improvement in mixing is mostly due to the reduction in the amount of zoning or axial barriers.8 Gassed mixing time results are plotted in Figure 5b against specific power input PG/V. The mixing time increases or stays constant with an increase in the specific power input at the low speed, while the mixing time dramatically decreases as the power input increases. These observations were found to be in good agreement with previous results.10,51 The mixing time decreases with the introduction of the gas into the tank at low speed (see Figure 5b); however, this only true at high stirring speed. Vrábel et al.6 has shown that the mixing time is reduced

equations

std. dev, %

R2

t0.05 = 99.49(PG/V)−0.244 t0.05 = 269.80(PG/V)−1/3(εG)0.143 t0.05N = 5.2Po−1/3(D/T)−2(H/T)2.89 t0.05 = 56.36(PG/V) −0.301 t0.05 = 112.84(PG/V)−1/3(εG)0.175 t0.05N = 5.2Po−1/3(D/T)−2(H/T)1.65 t0.05 = 67.61(PG/V)−0.336 t0.05 = 200.65(PG/V)−1/3(εG)0.364 t0.05N = 5.2Po−1/3(D/T)−2(H/T)1.69 t0.05 = 26.00(PG/V)−0.287 t0.05 = 72.03(PG/V)−1/3(εG)0.249 t0.05N = 5.2Po−1/3(D/T)−2(H/T)0.40

2.90 0.83 10.60 6.36 0.86 5.96 19.31 0.85 7.84 9.07 0.87 7.72

0.978 0.997 0.991 0.932 0.997 0.993 0.765 0.995 0.998 0.871 0.996 0.990

than the theoretical ungassed value −1/3 for 3RT, which is in good agreement with that reported by Vasconcelos et al.14 However, correlation coefficient is low with only the specific power included in the relationship. When we included the gas hold-up into the mixing time date correlation as follows: t0.05 = K1(PG/V )K2 (εG)K3

(11)

The correlation coefficients for the four impeller combinations are larger than 0.95 given in Table 2. The results show the correlations agree with experimental data very well. The coefficient, K1, represents the intensity of turbulent exchange, and the value of K2 is also −1/3. The value of K1 under gassed conditions are larger than that under ungassed conditions. This means the reduction of turbulence exchange, so mixing time increases when the gas introduced into the tank. The exponent, K3, related to the gas hold-up indicates the effect of the gas hold-up on the mixing time. Different impeller combinations have been checked to unveil the principles behind this adjustable parameter. In order to include the impeller configurations effect, we included the power number into the mixing time date correlation. t0.05 = 73.66(PG/V )−1/3 (εG)0.444 (Po)1.072

(12)

The standard deviation of the resulting correlation is 0.75%. The correlation fits well with the experimental results obtained for the geometries in this work. We also correlated mixing time according to eq 9 under gassed condition, but the power number was changed to gassed power number, PoG. The fitted results are given in table 2. For 5946



Article

3RT the exponent of H/T, α, is 2.89, which is similar to the value obtained from the ungassed condition. This agrees well with results reported by Cooke49 who found a similar expression for Rushton turbines systems, whether gassed or ungassed. The exponent α is 1.65 for HBT+2WHu, 1.69 for HBT+2WHd, and 0.40 for 3WHu. The variation of α between gassed and ungassed condition could be explained by the bubbles flow field (see Figure 4). Under the gassed condition, there are also six loops for 3RT, but the loops increase compared to ungassed condition for HBT+2WHd. For the HBT+2WHu and 3WHu, the numbers of loops are the same, but the top loop is larger than that under the ungassed condition. Therefore, the exponent on the H/T is affected by agitator type and gassed condition. The loops formed by different impeller combinations have a considerable effect on the mixing time. 3.3. Gas Hold-Up Properties. The overall average gas hold-up as a function of the impeller power consumption and superficial gas velocity was expressed by the following relationship: εG = K1(PG/V )K2 (VG)K3

(13)

where K1, K2, and K3 are parameters which are specific for each impeller combination. The constants and the exponents in the equation also affect by various geometrical parameters, like the D/T ratio, diameter of the impeller, clearance from the bottom, etc. Correlations obtained by the experimental data regression for multiple-impeller systems are given in Table 3.

Figure 6. Gas hold-up comparison at superficial gas rate 1.625 and 8.124 mm s−1 for various impeller types.

Table 3. Correlations for Prediction of Gas Hold-Up configuration 3RT 3RT 2RT 4RT HBT +2WHd HBT +2WHu 3WHu

equations εG = 0.0948(PG/ V)0.397(VG)0.618 εG = 0.278(PG/ V)0.217(VG)0.662 εG = 0.1(PG/ V)0.37(VG)0.65 εG = 0.0835(PG/ V)0.375(VG)0.62 εG = 0.276(PG/ V)0.311(VG)0.711 εG = 0.202(PG/ V)0.286(VG)0.60 εG = 0.101(PG/ V)0.341(VG)0.52

ref

std. dev, %

R2

8.93

0.981

Pinelli et al.55

30.58

Vasconcelos et al.14 Nocentini et al.53

24.92

velocity 3WHu configuration produces by up to 64% higher gas hold-ups than HBT+2WHd and the gas hold-up data for HBT +2WHu lie between them. 3RT and HBT+2WHd give almost the same gas hold-ups as power input below 800 W/m3 while 3RT generate higher gas hold-up than HBT+2WHd when power input above 800 W/m3. At high gas velocity, the effect of impeller configuration on the gas hold-up is diminished as 3WHu, HBT+2WHu, HBT+2WHd give almost the same gas hold-ups and a little higher than that generated by 3RT. The different gas hold-ups for various combinations can be explained by the flow pattern generated by the impeller combination. At low gas flow rate, the stirring speed for 3RT is lowest with the same specific power input, which results in bad gas liquid dispersion and recirculation, hence lower gas hold-up. For HBT+2WHd, many bubbles dispersed by HBT impeller directly rise to the liquid surface along the loop2 (Figure 3). For the HBT+2WHu and 3WHu, an extremely large number of bubbles are entrained in the circulation loop 3 for HBT+2WHu and loop 1 for 3WHu (Figure 3) and the residence time of the bubbles are very long, so the gas hold-ups are higher than other combinations. For 3RT and HBT+2WHd, a continual decrease in PG/P from 1 to 0.5 is noticed as the gas flow rate increases. For the 3WHu and HBT+2WHu, on the other hand, the PG/P is from 1 to 0.8 with increasing gas flow rate. At high gas flow rate, power drop for HBT+2WHd and 3RT are large, so the gas hold-up for them increases faster than other combinations at the same power consumption. At low power input, 3RT still shows bad gas dispersion capacity at high gas flow rate. Figure.7 shows that gas hold-up values in this work are much higher than those predicted using correlations from literatures.53−55 These differences might attribute to different D/T ratio. The value of D/T ratio in the literatures is 0.333, while our studied configuration with a ratio of D/T = 0.43. This may

23.87 9.43

0.985

13.40

0.987

13.53

0.984

The exponents in eq 13 on (PG/V) and (VG) for 3RT are very similar to those previously published by Vasconcelos et al.14 with dual-Rushton turbines and by Nocentini et al.53 with quadruple-Rushton turbines. A typical correlation in eq 1345 for such conditions gives exponents −1/3 and −2/3, respectively,. There are many correlations for different impeller configuration in the gas hold-up data, so it is not easy to make general correlation about multi-impeller systems from the gas hold-up data. In order to include the impeller configuration effect we add the power number in the gas hold-up data correlation. The resulting correlation as follows: εG = 0.160(PG/V )0.328 (VG)0.592 (PoG)−0.0916

(14)

with standard deviation of 16.29% and correlation coefficient of 0.971. The gas hold-ups for the four multiple-impeller configurations at superficial velocities 1.625 (0.2 vvm) and 8.124 mm s−1 (1 vvm) are given in Figure 6a,b, respectively. At low gas 5947



Article

The gas hold-ups distribution in stirred reactor is directly affected by the liquid flow field generated by the given configuration. And it was measured by conductivity probe method. Gas hold-up profiles under gas flow rate of 1.0 vvm and the same rotating speed of 350 rpm are given in Figure 8.

Figure 7. Comparison of gas hold-up calculated from correlation (13) for multiple Rushton impeller with literature data at VG = 1.625 mm s−1.

indicate that K1 value in eq 13 is function of D/T ratio. Smith56 have reported another type of correlation that including the effect of D/T ratio for the air−water system as follows εG = K1(ReFrFl)K2 (D/T )1.25

(15)

57

Murugesan also proposed a generalized correlation in terms of fundamental operating variables for the direct estimation of dispersed gas phase hold-up which are proportional to (D/ T)0.65. According to these reports, the higher gas hold-up value in our work is reasonable. Arjunwadkar et al.2 studied eight impeller combinations using Rushton turbine (RT), pitchedblade turbine pumping up (PBU) and down (PBD) and reported that the RT+PBD combination gives the highest gas hold-up at equivalent power input .However, the gas hold-up of HBT+2WHd gotten in this work are almost the same as that of 3WHu and HBT+2WHu. The reason that cause the difference may be the different impeller spacing. Pinelli et al.55 used three impellers of the same type (Rushton turbines, axial flow impellers A310 and A315). They also concluded that the impellers with the axial flow pattern produce higher gas holdup. These results are consistent with our results at high gas flow rate, but the trends of gas hold-up in this work are different at low gas flow rate. Moucha et al.7 studied the gas hold-up produced by the Rushton Turbines, six Pitched Blade impellers pumping down and hydrofoil impellers Techmix 335 pumping up (TXU) or down (TXD) and their combinations, and showed at low gas velocity RT+2TXU configuration produces by up to 40% higher gas hold-ups than RT+2TXD, which is agree with our results. At high gas velocity, the two configurations give the same gas hold-ups, but 3TXU produces lowest gas hold-up, the gas hold-up data for 3RT lies between them. However, in this work 3WHu, HBT+2WHu, and HBT +2WHd give almost the same gas hold-ups and a little higher than that generated by 3RT. Bao et al.58 investigated five impeller combinations using a deep hollow blade (semiellipse) disc turbine (HEDT) and four-wide-blade hydrofoil impellers pumping-up (WHu) and pumping-down (WHd). At the low gas velocity, the WHd+2WHu produces the highest gas hold-up and 3WHu produces the lowest gas hold-up, HEDT+2WHu and HEDT+2WHd give the highest gas hold-up at the high gas velocity. However, the 3WHu gives the highest gas hold-up at the low gas velocity in this work. All in all, the distinct gas holdups were caused by the flow patterns generated by different impeller types, impeller spacing opening parameters and etc. Local gas hold-ups are also very important in many industrial aerobic processes, such as polymerization and fermentation.

Figure 8. Local gas hold-up (%) distributions for all impeller combinations.

These gas hold-up profiles clearly depict the flow patterns prevailing in the stirred vessel under gassing conditions. For 3RT configuration, each Rushton turbine disperses the bubbles radially outward with the impeller discharge stream. In the bottom impeller discharge stream, the hold-up increases considerably away from the agitator. The highest local gas hold-ups are located in the center of the upper circulation loop of the bottom Rushton. For the middle and top impellers the gas hold-ups also increase gradually toward the wall. For the HBT+2WHd combination, the gas hold-ups generated by the bottom impeller is similar to the Rushton turbine, while upper two down pumping propellers produce lower gas hold-up around them, since the middle and top impellers are down flow type. 5948



Article

the lower specific area (see Figure.12) and longer mixing time. The power number also added to correlate gas hold-up data for the all tested impeller combinations as follows:

For HBT+2WHu agitated vessel, a completely different gas hold-up picture is obtained compared to that of HBT+2WHd combination. With the circulation pattern formed by the up pumping propeller, a higher gas hold-up value is observed located between the two upper impellers around the shaft and above the top impeller near the wall, but with very low gas hold-up above the top impeller elsewhere. Without the bottom radial impeller, local gas hold-up formed by 3WHu at the bottom impeller is extremely low, while gas hold-up at the upper part is the same as that of HBT+2WHu. From these observations, it can be concluded that bottom radial impeller is crucial for the gas dispersion when multiple impeller combination is used. 3.4. Mass Transfer Capacity. The experimental mass transfer coefficient data was correlated as a function of the power input and gas superficial velocity by using the model proposed by Cooper et al.59 kLa = K1(PG/V )K2 (VG)K3

kLa = 0.0296(PG/V )0.430 (VG)0.438 (PoG)−0.0467

(17)

with standard deviation of 15.65% and correlation coefficient of 0.968. The mass transfer coefficient obtained for the four tripleimpeller configurations at superficial velocities 1.625 and 8.124 mm s−1 are plotted in Figure 9a,b, respectively. At low

(16)

Model parameters for different impeller combinations are given in Table 4. The theoretical results of Kawase and MooTable 4. Correlations for Prediction of Volumetric Mass Transfer Coefficient configuration Any 3RT 3RT 2RT 2RT 4RT 4RT HBT +2WHd HBT +2WHu 3WHu

equations kLa = 0.0126(PG/ V)0.65(VG)0.5 kLa = 0.0115(PG/ V)0.560(VG)0.448 kLa = 0.00487(PG/ V)0.719(VG)0.497 kLa = 0.0083(PG/ V)0.62(VG)0.49 kLa = 0.0053(PG/ V)0.590(VG)0.274 kLa = 0.00861[(PG+ρgVG)/ V]0.637(VG)0.54 kLa = 0.015(PG/ V)0.59(VG)0.55 kLa = 0.0506(PG/ V)0.383(VG)0.511 kLa = 0.0302(PG/ V)0.452(VG)0.479 kLa = 0.0258(PG/ V)0.438(VG)0.436

ref

std. dev, %

R2

10.38

0.976

Kawse and Moo-Yong60

Fujasová et al.12 Vasconcelos et al.14 Gezork et al.67

19.38

Linek et al.68

22.99

Nocentini et al.53

19.28

24.02 52.20

Figure 9. Mass transfer coefficient comparison at superficial gas rate 1.625 and 8.124 mm s−1 for various impeller types.

11.96

0.988

10.24

0.994

7.77

0.988

superficial gas rate 3WHu produces 53% higher mass transfer coefficient than HBT+2WHd and the mass transfer coefficient for HBT+2WHu lies between them. When 3RT compared with HBT+2WHu, they are 30% lower than HBT+2WHu at lower power inputs (300 W m−3), and are almost equal at higher power inputs (700 W m−3). At high superficial gas rate, however, all four impeller combinations give approximate mass transfer coefficient value. Volumetric mass transfer coefficient depends on liquid side mass transfer coefficient, kL, and interfacial area. Based on isotropic turbulence theory,64−66 the kL value relies on specific power consumption. While interfacial area relies on local gas hold-up and bubble size, so the trend of mass transfer coefficient is similar to that of gas hold-up under the same specific power consumption. For 3RT the lowest gas hold-up but similar kLa to other combinations was obtained, this may be due to its higher break-up effectiveness than that of axial flow impeller under identical specific power consumption. Comparison of data measured in this work and literature data is given in Figure 10. Good agreement was found between our data and data of Vasconcelos et al.14 and Nocentini et al.53 However, a large discrepancy was found to that of Gezork et al., 67 who measured the k L a by using the peroxide

Young60 are higher than the empirical and experimental ones, but the empirical results are very close to the experimental data. The exponent related to the power input is 0.65 in eq 16 Kawase and Moo-Young,60 derived from Kolmogorov’s theory. The value of K2 for 3RT is the biggest, and the rising bubbles can be broken by the blades in their discharge, resulting in the effectiveness of the break-up process. However, the value of K2 for HBT+2WHd is lower than 0.4, the accumulation of bubbles below the impeller as well as low break-up effectiveness. Bubble break-up is due to its deformation under the impeller. The break-up effectiveness of HBT+2WHu and 3WHu is middle, bubbles coalescent occurs in the tank, but when bigger bubbles are generated, break-up processes are easier. The available specific area and the scattering of the bubbles determine the experimental value of K3.61,62 This value has been reported with typical value is 0.5 as in eq 16.63 In this work, only the value of K3 for HBT+2WHd is higher than 0.5, it is because of 5949



Article

Figure 10. Comparison of mass transfer coefficient calculated from correlation (16) for multiple Rushton impeller with literature data at VG = 1.625 mm s−1.

decomposition steady-state technique with manganese dioxide as the catalyst. Linek et al.68 and Fujasová et al.12 showed that this method resulted in higher values of kLa than other methods because of spontaneous nucleation of bubbles. Gogate and Pandit69 found that the dynamic pressure method is suitable for large scale bioreactors with errors less than 10%. Arjunwadkar et al.25 performed experiments with different impeller combinations and reported that DT-PTD combination gives the best mass transfer performance due to the maximum holdup obtained from this combination. In this work, when HBT +2WHd compared with 3RT at low gas flow rate, the kLa values of HBT+2WHd are 10% higher than that of 3RT at lower power inputs (400 W m−3), but are 20% lower than that of 3RT at higher power inputs. Bouaifi and Roustan3 measured various dual-impeller combinations using A310, A315, PB impellers, and Rushton turbines and they found all the tested configurations give similar kLa values for a given gas flow and power input, which agrees with our results at high gas flow rate. Moucha et al.7,11 showed among the triple-impeller configurations, 3RT gives the best mass transfer performance. The RT+2PBD and RT+2PBU configurations produce 15% lower kLa values at the same specific power dissipation as 3RT. The 3TXU configuration gives the poorest volumetric mass transfer coefficients. In this work 3WHu gives the best mass transfer performance at low gas flow rate, but HBT+2WHu gives the better mass transfer performance, being 10% higher than that achieved by 3RT at high gas flow rate. The mass transfer coefficient for the 3RT is similar to the other impeller combinations, whereas the gas hold-up for the 3RT is smaller than that of other three impeller combinations. This can be explained by the smaller bubble size resulted in higher specific area in 3RT stirred tank. Figure 11 shows maps of local Sauter mean bubble diameters, d32, for the four tripleimpeller combinations obtained using the conductivity probes technique. The 3RT shows the smallest bubble size compared to other combinations. The smallest bubble size in the reactor was found in the discharge stream region for the 3RT combination. High rate of bubble breakage in impeller discharge stream was caused by the high level of turbulence generated by the impeller. Most of the coalescence is completed by the time the discharge stream reaches the wall, which already reported in the literatures.23,70 For HBT+2WHd combination, the small bubble size is also found in the impeller discharge stream. In the bulk region no large variations in the Sauter mean diameter values with spatial position are observed. The bubble coalescence mainly occurs near the wall in the radical direction along the HBT impeller and close to the top

Figure 11. Local sauter mean diameter distributions for all impeller combinations.

WHd impeller. The biggest bubble mean diameter was observed above the top impeller and near the wall for the HBT+2WHu combination, this because that the rising bubbles encounter down flowing fluid and high gas hold-up promote bubble collision and coalescence. However, the Sauter mean diameter is small in loop 4 as only small bubbles with low buoyancy follow the liquid circulation stream. For the 3WHu combination the bubble sizes leaving the bottom WHu impeller are very large because of its badly dispersion performance. Then bubbles entrain into loop 1 similar to HBT+2WHu combination. The overall mass transfer coefficient is a very useful parameter for reactor design. However, it does not reveal all of the information about local mass transfer. In order to acquire a detailed and complete picture of the internal structure of the gas−liquid mass exchanging, the actual local mass transfer should also be considered. In the stirred reactor, the kL values depend on specific power consumption based on isotropic turbulence theory.64−66 Locations with high local specific interface area will definitely get high mass transfer capacity. 5950



Article

different zones of higher interfacial area are also due to the flow pattern formed by different impeller combinations.

Knowing local bubble size and gas hold-up allows evaluating the local specific interfacial area: a = 6εG /d32

(18)

4. CONCLUSION The objective of this work was to study the flow pattern, mixing, gas hold-up, and mass transfer coefficient of different triple impeller configurations. These studies showed that at low superficial gas rate the up pumping impeller shows good mass transfer capacity compared to the down pumping impeller. 3WHu generates 53% higher mass transfer coefficient than HBT+2WHd and the mass transfer coefficient of HBT+2WHu and 3RT lie between them. At high superficial gas rate, however, all of the tested configurations give similar mass transfer coefficient under identical specific power input. The agitator type and gassed condition affect the exponent of the H/T, which increases under gassed condition except 3RT. The loops formed by different impeller combinations have a considerable effect on the mixing time under gassed or ungassed conditions. The axial impellers combination (3WHu) provides more effective homogenization performance than the impellers combined radial and axial flow (HBT +2WHu and HBT+2WHd), and mixing of the radial impellers combination (3RT) are the worst. For 3RT configuration, highest holdup is in the bottom impeller discharge stream, the higher holdup was observed near the wall for the middle and top impellers. For the HBT+2WHd combination, the gas hold-ups generated by the bottom impeller is similar to the Rushton turbine. In the bulk region no large variations in the gas hold up values. For HBT+2WHu and 3WHu agitated vessel, the big circulation loop generated by the up-pumping impellers entrains the gas bubbles for a longer time, however the liquid velocity in upper smaller circulation loop is low. This results in high gas hold-up between the two up pumping impellers, moderately low gas hold-up above the top impeller. There are three zones of higher interfacial area for the HBT +2WHu and 3RT combination, two zones of higher interfacial area for 3WHu, and one zone of higher interfacial area for HBT+2WHd combination. The good fit of mixing time, gas hold-up and kLa values correlated with power consumption, superficial gas velocity, and power number (characterizing impeller type) are presented Hence, axial flow impeller pumping up combination 3WHu and the combination of a bottom radial flow impeller (HBT) and combined axial flow impellers pumping up (WHu) in the upper stages are recommended as the most efficient impeller combinations for the mass transfer and mixing performance.

the profiles of local interfacial area for the four impeller combinations are shown in Figure 12. It can be seen that

■

Figure 12. Local interfacial area distributions for all impeller combinations.

AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Phone: +86-021-64251946.

distribution of local interfacial area for each impeller combination is almost identical to the distribution of local gas hold-up. The highest local interfacial area were found in the impeller discharge stream for 3RT combination, which is due to relatively high gas hold-ups and small bubble size in these regions. The high interfacial area locates in the bottom impeller discharge stream for the HBT+2WHd combination. There are three zones of higher interfacial area for the HBT +2WHu combination: one is in the bottom impeller discharge stream, the other one is above the middle impeller, and the last one is near the wall above the top impeller. There are two zones of higher interfacial area in 3WHu combination vessel. The

Notes

The authors declare no competing financial interest.

■

ACKNOWLEDGMENTS This work was supported by State Key Development Program of Basic Research of China (973 Program, Grant No. 2013CB733600), National High Technology Research and Development Program of China (863 Program, Grant No. 2012AA021201), China Ministry of Science and Technology under Contract (Grant No. 2012BAI44G00), and Open Funding Project of the State Key Laboratory of Bioreactor Engineering is gratefully acknowledged. 5951



■

Article

radial or radial and axial up-pumping impellers: modelling and measurements. Chem. Eng. Sci. 2000, 55 (23), 5881−5896. (9) Pinelli, D.; Magelli, F. Analysis of the Fluid Dynamic Behavior of the Liquid and Gas Phases in Reactors Stirred with Multiple Hydrofoil Impellers. Ind. Eng. Chem. Res. 2000, 39 (9), 3202−3211. (10) Shewale, S. D.; Pandit, A. B. Studies in multiple impeller agitated gas-liquid contactors. Chem. Eng. Sci. 2006, 61 (2), 489−504. (11) Moucha, T.; Linek, V.; Erokhin, K.; Rejl, J. F.; Fujasová, M. Improved power and mass transfer correlations for design and scale-up of multi-impeller gas-liquid contactors. Chem. Eng. Sci. 2009, 64 (3), 598−604. (12) Fujasová, M.; Linek, V.; Moucha, T. Mass transfer correlations for multiple-impeller gas-liquid contactors. Analysis of the effect of axial dispersion in gas and liquid phases on ″local″ kLa values measured by the dynamic pressure method in individual stages of the vessel. Chem. Eng. Sci. 2007, 62 (6), 1650−1669. (13) Sardeing, R.; Aubin, J.; Xuereb, C. Gas-Liquid Mass Transfer: A Comparison of Down- and Up-pumping Axial Flow Impellers with Radial Impellers. Chem. Eng. Res. Des. 2004, 82 (12), 1589−1596. (14) Vasconcelos, J. M. T.; Orvalho, S. C. P.; Rodrigues, A. M. A. F.; Alves, S. S. Effect of Blade Shape on the Performance of Six-Bladed Disk Turbine Impellers. Ind. Eng. Chem. Res. 2000, 39 (1), 203−213. (15) Pinelli, D.; Bakker, A.; Myers, K. J.; Reeder, M. F.; Fasano, J.; Magelli, F. Some Features of a Novel Gas Dispersion Impeller in a Dual-Impeller Configuration. Chem. Eng. Res. Des. 2003, 81 (4), 448− 454. (16) Cooke, M.; Heggs, P. J. Advantages of the hollow (concave) turbine for multi-phase agitation under intense operating conditions. Chem. Eng. Sci. 2005, 60 (20), 5529−5543. (17) Jüsten, P.; Paul, G. C.; Nienow, A. W.; Thomas, C. R. Dependence of Penicillium chrysogenum growth, morphology, vacuolation, and productivity in fed-batch fermentations on impeller type and agitation intensity. Biotechnol. Bioeng. 1998, 59 (6), 762−775. (18) Heo, J.-H.; Ananin, V.; Park, J.-S.; Lee, C.-R.; Moon, J.-O.; Kwon, O.; Kang, H.-A.; Kim, C.; Rhee, S. Impeller types and feeding modes influence the morphology and protein expression in the submerged culture ofAspergillus oryzae. Biotechnol. Bioprocess Eng. 2004, 9 (3), 184−190. (19) Li, X.; Zhang, J.; Tan, Y. L.; Li, Z. H.; Yu, X. F.; Xia, J. Y.; Chu, J.; Ge, Y. Q. Effects of flow field on the metabolic characteristics of Streptomyces lincolnensis in the industrial fermentation of lincomycin. J. Biosci. Bioeng. 2013, 115 (1), 27−31. (20) Xia, J.-Y.; Wang, Y.-H.; Zhang, S.-L.; Chen, N.; Yin, P.; Zhuang, Y.-P.; Chu, J. Fluid dynamics investigation of variant impeller combinations by simulation and fermentation experiment. Biochem. Eng. J. 2009, 43 (3), 252−260. (21) Gabriele, A.; Nienow, A. W.; Simmons, M. J. H. Use of angle resolved PIV to estimate local specific energy dissipation rates for upand down-pumping pitched blade agitators in a stirred tank. Chem. Eng. Sci. 2009, 64 (1), 126−143. (22) Adrian, R. J. Image shifting technique to resolve directional ambiguity in double-pulsed velocimetry. Appl. Opt. 1986, 25 (21), 3855−3858. (23) Alves, S. S.; Maia, C. I.; Vasconcelos, J. M. T.; Serralheiro, A. J. Bubble size in aerated stirred tanks. Chem. Eng. J. 2002, 89 (1−3), 109−117. (24) Gupta, P.; Al-Dahhan, M. H.; Dudukovic, M. P.; Mills, P. L. A novel signal filtering methodology for obtaining liquid phase tracer responses from conductivity probes. Flow Meas. Instrum. 2000, 11 (2), 123−131. (25) Arjunwadkar, S. J.; Sarvanan, K.; Kulkarni, P. R.; Pandit, A. B. Gas-liquid mass transfer in dual impeller bioreactor. Biochem. Eng. J. 1998, 1 (2), 99−106. (26) Chisti, Y.; Jauregui-Haza, U. J. Oxygen transfer and mixing in mechanically agitated airlift bioreactors. Biochem. Eng. J. 2002, 10 (2), 143−153. (27) Garcia-Ochoa, F.; Gomez, E. Bioreactor scale-up and oxygen transfer rate in microbial processes: an overview. Biotechnol. Adv. 2009, 27 (2), 153−76.

NOTATION a local interfacial area, m2 m−3 (=m−1) B width of baffle, m C* saturation oxygen concentration C0 initial oxygen concentration d32 Sauter mean bubble diamter, mm D impeller diameter, m H height in the absence of gas HG height in the presence of gas kLa volumetric mass transfer coefficient, s−1 K1−K3 empirical constants lb the bubble chord length, mm Lp distance between the two probe tips, mm P power input under ungassed conditions, W PG power input under gassed conditions, W Po power number under ungassed conditions PoG power number under gassed conditions QG Gas-flow rate, m3 s−1 Re Reynolds number t time, s Δt lag time between the bubble hitting probes, s t0.05 mixing time for 95% mixing, s T tank diameter, m U liquid velocity, m s−1 Utip impeller tip speed, m s−1 V volume of liquid, m3 VG gas superficial velocity, m s−1 Greek Letters

εG θ ρ τ

gas hold-up (volumetric fraction of gas) temperature, °C density of solution dwell time of a bubble, s

Acronyms

HBT WHd WHu RT

■

hollow blade turbine wide-blade hydrofoil impellers pumping down wide-blade hydrofoil impellers pumping up Rushton turbine

REFERENCES

(1) Gogate, P. R.; Beenackers, A. A. C. M.; Pandit, A. B. Multipleimpeller systems with a special emphasis on bioreactors: a critical review. Biochem. Eng. J. 2000, 6 (2), 109−144. (2) Arjunwadkar, S. J.; Saravanan, K.; Pandit, A. B.; Kulkarni, P. R. Optimizing the impeller combination for maximum hold-up with minimum power consumption. Biochem. Eng. J. 1998, 1 (1), 25−30. (3) Bouaifi, M.; Roustan, M. Bubble size and mass transfer coefficients in dual-impeller agitated reactors. Can. J. Chem. Eng. 1998, 76 (3), 390−397. (4) Puthli, M. S.; Rathod, V. K.; Pandit, A. B. Gas-liquid mass transfer studies with triple impeller system on a laboratory scale bioreactor. Biochem. Eng. J. 2005, 23 (1), 25−30. (5) Lara, A.; Galindo, E.; Ramírez, O.; Palomares, L. Living with heterogeneities in bioreactors. Mol. Biotechnol. 2006, 34 (3), 355−381. (6) Vrábel, P.; Van der Lans, R. G. J. M.; Cui, Y. Q.; Luyben, K. C. A. M. Compartment Model Approach: Mixing in Large Scale Aerated Reactors with Multiple Impellers. Chem. Eng. Res. Des. 1999, 77 (4), 291−302. (7) Moucha, T.; Linek, V.; Prokopová, E. Gas hold-up, mixing time and gas-liquid volumetric mass transfer coefficient of various multipleimpeller configurations: Rushton turbine, pitched blade and techmix impeller and their combinations. Chem. Eng. Sci. 2003, 58 (9), 1839− 1846. (8) Vrábel, P.; van der Lans, R. G. J. M.; Luyben, K. C. A. M.; Boon, L.; Nienow, A. W. Mixing in large-scale vessels stirred with multiple 5952



Article

(28) Jackson, M. L.; Shen, C.-C. Aeration and mixing in deep tank fermentation systems. AIChE J. 1978, 24 (1), 63−71. (29) Zhao, D.; Guo, L.; Lin, C.; Zhang, X. An experimental study on local interfacial area concentration using a double-sensor probe. Int. J. Heat Mass Transfer 2005, 48 (10), 1926−1935. (30) Rampure, M. R.; Kulkarni, A. A.; Ranade, V. V. Hydrodynamics of Bubble Column Reactors at High Gas Velocity: Experiments and Computational Fluid Dynamics (CFD) Simulations. Ind. Eng. Chem. Res. 2007, 46 (25), 8431−8447. (31) Bombač, A.; Ž un, I.; Filipič, B.; Ž umer, M. Gas-filled cavity structures and local void fraction distribution in aerated stirred vessel. AIChE J. 1997, 43 (11), 2921−2931. (32) Bao, Y.; Chen, L.; Gao, Z.; Chen, J. Local void fraction and bubble size distributions in cold-gassed and hot-sparged stirred reactors. Chem. Eng. Sci. 2010, 65 (2), 976−984. (33) Revankar, S. T.; Ishii, M. Theory and measurement of local interfacial area using a four sensor probe in two-phase flow. Int. J. Heat Mass Transfer 1992, 36 (12), 2997−3007. (34) Wang; Mao, Z.-S.; Yang, C. Experimental and Numerical Investigation on Gas Holdup and Flooding in an Aerated Stirred Tank with Rushton Impeller. Ind. Eng. Chem. Res. 2006, 45 (3), 1141−1151. (35) Hibiki, T.; Hogsett, S.; Ishii, M. Local measurement of interfacial area, interfacial velocity and liquid turbulence in two-phase flow. Nucl. Eng. Des. 1998, 184 (2−3), 287−304. (36) Hogsett, S.; Ishii, M. Local two-phase flow measurements using sensor techniques. Nucl. Eng. Des. 1997, 175 (1−2), 15−24. (37) Young, M. A.; Carbonell, R. G.; Ollis, D. F. Airlift bioreactors: Analysis of local two-phase hydrodynamics. AIChE J. 1991, 37 (3), 403−428. (38) Lo, C.-S.; Hwang, S.-J. Local hydrodynamic properties of gas phase in an internal-loop airlift reactor. Chem. Eng. J. 2003, 91 (1), 3− 22. (39) Dias, S. G.; França, F. A.; Rosa, E. S. Statistical method to calculate local interfacial variables in two-phase bubbly flows using intrusive crossing probes. Int. J. Multiphase Flow 2000, 26 (11), 1797− 1830. (40) Matsuura, A.; Fan, L.-S. Distribution of bubble properties in a gas-liquid-solid fluidized bed. AIChE J. 1984, 30 (6), 894−903. (41) Yasunishi, A.; Fukuma, M.; Muroyama, K. Measurement of behavior of gas bubbles and gas hold-up in a slurry bubble column by a dual electro-resistivity probe method. J. Chem. Eng. Jpn. 1986, 19 (5), 444−449. (42) Liu, W.; Clark, N. N. Relationships between distributions of chord lengths and distributions of bubble sizes including their statistical parameters. Int. J. Multiphase Flow 1995, 21 (6), 1073−1089. (43) Nienow, A. W. Hydrodynamics of Stirred Bioreactors. Appl. Mech. Rev. 1998, 51 (1), 3−32. (44) Nienow, A. W. Agitators for mycelial fermentations. Trends Biotechnol. 1990, 8 (0), 224−233. (45) Van’t Riet, K.; Tramper, J. Basic Bioreactor Design; Marcel Dekker, lnc.: New York, 1991. (46) Nienow, A. W. On impeller circulation and mixing effectiveness in the turbulent flow regime. Chem. Eng. Sci. 1997, 52 (15), 2557− 2565. (47) Corrsin, S. The isotropic turbulent mixer: Part II. Arbitrary Schmidt number. AIChE J. 1964, 10 (6), 870−877. (48) Grenville, R. K.; Nienow, A. W. Handbook of Industrial Mixing: Science and Practice; John Wiley & Sons: Hoboken, NJ, 2004. (49) Cooke, M. Processing of Solid-liquid Suspensions; ButterworthHeinemann: London, 1988. (50) Rodgers, T. L.; Gangolf, L.; Vannier, C.; Parriaud, M.; Cooke, M. Mixing times for process vessels with aspect ratios greater than one. Chem. Eng. Sci. 2011, 66 (13), 2935−2944. (51) Abardi, V.; Rovero, G.; Sicardi, S.; Baldi, G.; Conti, R. Sparged vessels agitated by multiple turbines. Proc. Sixth Eur. Conf. Mix. 1988, 329−336. (52) Nienow, A. W. Gas-liquid mixing studies: A comparison of Rushton turbines with some modern impellers. Chem. Eng. Res. Des. 1996, 74 (A), 417−423.

(53) Nocentini, M.; Fajner, D.; Pasquali, G.; Magelli, F. Gas-liquid mass transfer and holdup in vessels stirred with multiple Rushton turbines: water and water-glycerol solutions. Ind. Eng. Chem. Res. 1993, 32 (1), 19−26. (54) Vasconcelos, J. M. T.; Rodrigues, J. M. L.; Orvalho, S. C. P.; Alves, S. S.; Mendes, R. L.; Reis, A. Effect of contaminants on mass transfer coefficients in bubble column and airlift contactors. Chem. Eng. Sci. 2003, 58 (8), 1431−1440. (55) Pinelli, D.; Nocentini, M.; Magelli, F. Hold-up in low viscosity gas-liquid systems stirred with multiple impellers. Comparison of different agitator types and sets. Inst. Chem. Eng. Symp. Ser. 1994, 81− 88. (56) Smith, J. M., Simple Performance Correlations for Agitated Vessels. In Proceedings of 7th European Congress on Mixing, Bruges, Belgium, 1991; pp 233-241. (57) Murugesan, T. Dispersed phase hold-up in mechanically agitated gas−liquid contactors. J. Chem. Technol. Biotechnol. 1998, 72 (3), 221− 226. (58) Bao, Y.; Hao, Z.; Gao, Z.; Shi, L.; Smith, J. M.; Thorpe, R. B. Gas dispersion and solid suspension in three-phase stirred tank with multiple impellers. Chem. Eng. Commun. 2006, 193 (7), 801−825. (59) Cooper, C. M.; Fernstrom, G. A.; Miller, S. A. Correction Performance of Agitated Gas-Liquid Contactors. Ind. Eng. Chem. Res. 1944, 36 (9), 857−857. (60) Kawse, Y.; Moo-Yong, M. Volumetric mass transfer coefficients in aerated stirred tank reactors with Newtonian and non-Newtonian media. Chem. Eng. Res. Des. 1988, 66 (3), 284−288. (61) Martín, M.; Montes, F. J.; Galán, M. A. Bubbling process in stirred tank reactors II: Agitator effect on the mass transfer rates. Chem. Eng. Sci. 2008, 63 (12), 3223−3234. (62) Martín, M.; Montes, F. J.; Galán, M. A. Physical explanation of the empirical coefficients of gas-liquid mass transfer equations. Chem. Eng. Sci. 2009, 64 (2), 410−425. (63) Sideman, S.; Hortaçsu, Ö .; Fulton, J. W. Mass transfer in gasliquid contacting systems. Ind. Eng. Chem. Res. 1966, 58 (7), 32−47. (64) Kawase, Y.; Moo-Young, M. Mathematical models for design of bioreactors: Applications of: Kolmogoroff’s theory of isotropic turbulence. Chem. Eng. J. 1990, 43 (1), B19−B41. (65) Lamont, J. C.; Scott, D. S. An eddy cell model of mass transfer into the surface of a turbulent liquid. AIChE J. 1970, 16 (4), 513−519. (66) Bouaifi, M.; Hebrard, G.; Bastoul, D.; Roustan, M. A comparative study of gas hold-up, bubble size, interfacial area and mass transfer coefficients in stirred gas-liquid reactors and bubble columns. Chem. Eng. Process. 2001, 40 (2), 97−111. (67) Gezork, K. M.; Bujalski, W.; Cooke, M.; Nienow, A. W. Mass Transfer and Hold-up Characteristics in a Gassed, Stirred Vessel at Intensified Operating Conditions. Chem. Eng. Res. Des. 2001, 79 (8), 965−972. (68) Linek, V.; Moucha, T.; Sinkule, J. Gas-liquid mass transfer in vessels stirred with multiple impellersI. Gas-liquid mass transfer characteristics in individual stages. Chem. Eng. Sci. 1996, 51 (12), 3203−3212. (69) Gogate, P. R.; Pandit, A. B. Survey of measurement techniques for gas−liquid mass transfer coefficient in bioreactors. Biochem. Eng. J. 1999, 4 (1), 7−15. (70) Barigou, M.; Greaves, M. Bubble-size distributions in a mechanically agitated gasliquid contactor. Chem. Eng. Sci. 1992, 47 (8), 2009−2025.

5953


Flow Pattern, Mixing, Gas Hold-Up and Mass ... - ACS Publications

Recommend Documents