Experimental Investigation of Flow Development in Large-Scale

Feb 26, 2016 - State Key Laboratory of Multi-phase Complex System, Institute of Process ... scale effect indicates that for the large-scale industrial...
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Experimental Investigation of Flow Development in Large-Scale Bubble Columns in the Churn-Turbulent Regime Xiaoping Guan,† Ning Yang,‡ Zhaoqi Li,† Lijun Wang,† Youwei Cheng,† and Xi Li*,† †

College of Chemical and Biological Engineering, Zhejiang University, Hangzhou, Zhejiang 310027, P. R. China State Key Laboratory of Multi-phase Complex System, Institute of Process Engineering, Chinese Academy of Sciences, P.O. Box 353, Beijing 100190, P. R. China

‡

ABSTRACT: Flow development in the bubble columns is of importance for the sparger design and reactor scale up. In the present study, axial development of local gas holdup and liquid velocity in two bubble columns (0.5 and 0.8 m in diameter, respectively) operated in the churn-turbulent regime was experimentally investigated. The results show that gas distribution type, superficial gas velocity and column diameter have remarkable influences on the flow development. Compared with the near-wall aeration, the uniform and central aeration will generate more rapid flow development. The observed strong influence of the sparger type on flow development shows an opposite trend to the reported findings. Furthermore, increasing superficial gas velocity or decreasing column diameter will reduce the entrance region length. The scale effect indicates that for the large-scale industrial bubble column reactors, a large portion of flow region may be covered by the distributor region rather than the well-developed region.

1. INTRODUCTION Bubble columns have been widely used as multiphase reactors or contactors in chemical, biochemical, petrochemical, and wastewater treatment industries.1 At the column bottom, the gas, introduced via a gas sparger, experiences complex changes through bubble formation, growth, detachment, and mutual interactions. Hence, the gas sparger configuration directly determines the initial bubble size and the initial distribution of flow parameters. For high bubble columns, the flow patterns can be roughly compartmentalized into three flow regions from bottom to top: distributor region, well-developed region, and gas−liquid disengagement region.2 The distributor region is the transition section from initial flow pattern to fully developed flow region, and is also called the entrance region. Knowledge of its length and range in the bubble columns contributes to understanding the importance of the sparger design and configuration. It has been reported that the gas sparger configuration has a significant effect on flow behavior in a homogeneous bubbly flow regime.3−6 Michele et al.4 measured the liquid velocity in a 0.63 m diameter bubble column equipped with plate sparger, ring sparger, or central nozzle sparger. The operated superficial gas velocity was in the range of 0.02−0.09 m·s−1. The experimental results indicated that the sparger geometry significantly impacted the liquid circulation structure, and the ring sparger even created abnormal downflow in the center part of the column bottom. Harteveld et al.5 studied the influence of gas distribution on coherent structures in a 0.15 m diameter bubble column with superficial gas velocity in the range of 0.015−0.023 m·s−1. The different gas distribution types were © 2016 American Chemical Society

generated through closing or opening the rings in a multiring aeration system. They found that nonuniform aeration created dynamic circulation cells in the entrance region, and no coherent structures were observed above the entrance region. The entrance region length was highly sensitive to aeration patternsit was very short for uniform aeration, four column diameters for the two wall rings closed, and three column diameters for the center region closed. Kulkarni et al.6 used a laser doppler anemometer (LDA) to measure the axial evolution of liquid velocity and gas holdup in a 0.15 m diameter bubble column with multipoint and single point spargers under a superficial gas velocity of 0.02 m·s−1, and found that the entrance region was about four column diameters for both spargers and the radial distributions of the flow parameters in fully developed region showed insignificant difference between them. Nonetheless, the effect diminishes in the churn-turbulent regime and the length of the distributor region is confined to about two column diameters for the air−water system.7−11 Wilkinson et al.7 has shown that gas holdup is independent of sparger layout when the column diameter is larger than 0.15 m, the column height to diameter ratio is in excess of 5, and the hole diameter of the sparger is larger than 2 × 10−3 m. Krishna et al.8 studied gas holdup in a 0.1 m diameter bubble column with sintered glass plate and sieve plate spargers, and revealed Received: Revised: Accepted: Published: 3125

October 25, 2015 February 17, 2016 February 26, 2016 February 26, 2016 DOI: 10.1021/acs.iecr.5b04015 Ind. Eng. Chem. Res. 2016, 55, 3125−3130

Article

Industrial & Engineering Chemistry Research

in height Plexiglas column with uniform aeration type (the related structure parameter is listed in Table 1) at UG = 0.12 m· s−1, and the corresponding gas−liquid dispersion was 3.7 m in height. In this paper, the conductivity probe and Pavlov tube were applied to measuring local gas holdup and axial liquid velocity, respectively, and the related fundamentals of the measuring techniques have been discussed in the literature.15,18,19 During experiments, the sample frequency and sample time for local gas holdup and liquid velocity were 2 kHz and 20 s, and 65 Hz and 60 s, respectively, while more than three samples for each radial location were repeated for statistical analysis. To evaluate the measured data uncertainty, volume average of the measured local gas holdup was compared with the overall gas holdup, and the area-averaged liquid flow rates were compared with liquid flux (zero for batch mode). The comparison suggested that the mean deviations for local gas holdup and liquid velocity were 10% and 5%, respectively.20 In the 0.8 m diameter column, local gas holdup and liquid velocity at seven axial positions (H = 0.7 m, 1.2 m, 1.7 m, 2.2 m, 2.7 m, 3.2 m and 3.7 m) were measured. For each axial position, 10 dimensionless radial locations (r/R = 0, 0.1, 0.24, 0.37, 0.5, 0.63, 0.76, 0.84, 0.92, and 0.96) were taken for local gas holdup measurement, and the same locations except r/R = 0.96 for liquid velocity measurement. For the 0.5 m diameter bubble column, five axial positions (H = 0.7 m, 1.2 m, 1.7 m, 2.2 m, and 2.7 m) were selected, and the measured radial locations were the same as those in the 0.8 m column. The operated superficial gas velocity was in a range of 0.08−0.62 m·s−1 as shown in Table 1, which was in the churn-turbulent flow regime to reflect industrial operating conditions.

that total gas holdup and large bubble holdup in the churnturbulent regime were insignificantly influenced by the gas sparger type. Ueyama et al.9 measured the gas holdup profile with superficial gas velocity up to 0.3 m·s−1, and found the entrance region length was about two column diameters. George et al.10 measured axial evolution of gas holdup profile by gamma densitometry tomography in a 0.483 diameter bubble column with superficial gas velocity from 0.05 m·s−1 to 0.3 m·s−1 and pressure from 0.1 to 0.5 MPa. The measured data indicated that the development length was about two column diameters. Ong et al.11 used computed tomography to study gas holdup profiles at axial positions above two column diameters in a 0.162 m diameter bubble column with superficial gas velocity in the range of 0.14−0.3 m·s−1, and observed insignificant influence of sparger type and design. Since most studies concluded that the length of the distributor region in the churn-turbulent bubble columns was very limited, open literature mainly focused on the flow behavior in the well-developed region,12−17 and the flow characteristics in the distributor region received less attention. In effect, the conclusion has been drawn from the experimental data of small-scale bubble columns (D < 0.5 m), and whether it is applicable to large-scale bubble columns remains an open question. Moreover, in the churn-turbulent regime, previous literature employed the cross-sectional holdup or gas holdup profile to characterize the development length, and liquid velocity has not been employed to determine the length for the difficulty in measuring liquid velocity at high gas holdup.18 Finally, it is unknown whether the development lengths based on different flow parameters are identical. Thereby, in the present study, local gas holdup and liquid velocity are measured at different axial positions in two bubble columns (outer diameter of 0.8 and 0.5 m, respectively) operated in the churnturbulent regime, and the effects of gas distribution, superficial gas velocity, and column diameter on the flow development are assessed. Furthermore, the entrance region lengths determined by the local gas holdup and liquid velocity are compared.

3. RESULTS AND DISCUSSION 3.1. Overall Gas Holdup. The overall gas holdup was calculated by measuring the dynamic gas−liquid dispersion height and the static liquid height. H − Hs εG = d Hd (1)

2. EXPERIMENTAL WORK Most experimental runs were performed in a Plexiglas column of 0.8 m in diameter and 5 m in height. Compressed air was introduced into the column through a perforated distributor with drilled holes of 2.5 × 10−3 m, and the tap water was operated in a batch mode. Three gas distribution types (uniform aeration, central aeration, and near-wall aeration) were created through regionally blocking holes, and the corresponding configurations are listed in Table 1. The dynamics height of the gas−liquid dispersion was kept at 4 m above the distributor while the static liquid height varied with the superficial gas velocity. To assess the scale effect, another experimental run was performed in a 0.5 m in diameter and 4 m

where Hd is the dynamic height and Hs is the static height. In the churn-turbulent regime, the bubbling height shows strong oscillation. Hence, to reduce the error with visual observation, both maximum and minimum heights were recorded and the average value was used to evaluate over all gas holdup. The runs were repeated more than 3 times to compute the mean and the standard deviation of overall gas holdup. The effect of gas distribution on overall gas holdup is illustrated in Figure 1. The

Table 1. Gas Distribution Configuration gas distribution type

aeration region

uniform aeration

0 < r/R < 1

pitch/m

opening ratio

0.062

0.14%

0.031

0.53%

central aeration

0 < r/R < 0.58

0.062

0.19%

near-wall aeration

0.82 < r/R < 1

0.062

0.16%

UG/m·s−1 0.08, 0.12, 0.19 0.31, 0.47, 0.62 0.08, 0.12, 0.19 0.08, 0.12, 0.19

Figure 1. Influence of gas distribution type on overall gas holdup in a 0.8 m diameter bubble column. 3126

DOI: 10.1021/acs.iecr.5b04015 Ind. Eng. Chem. Res. 2016, 55, 3125−3130

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Industrial & Engineering Chemistry Research

more remarkable increase in the annulus and near the wall makes the profile become slightly flatter. For the uniform aeration, the gas holdup profile is peaked in the annulus and is reduced on both sides at H = 0.7 m. With increasing axial position, the gas holdup is gradually increased in the core region (r/R = 0−0.63) and is almost unchanged near the wall. The mean deviation of local gas holdup at H = 2.7 m and H = 3.2 m is less than 5%, and the gas holdup profile is fully developed at the axial position of 2.7 m. For the central aeration, increasing axial position tends to increase local gas holdup both in the core region and near the wall, and a more significant change in the column core leads to the increasing steepness of the gas holdup profile. Furthermore, similar to the uniform aeration, the gas holdup profile reaches the well-developed region at the axial position of 2.7 m, as the mean deviation of local gas holdup at H = 2.7 m and H = 3.2 m is less than 5% . For the near-wall aeration, the gas holdup tends to increase in the column core and to decrease near the wall with increasing axial position. In comparison with the other two gas distribution types, the difference between the axial positions of 2.7 and 3.2 m is relatively larger, and no well-developed region is thus observed for the gas holdup profile. Except in the gas− liquid disengagement region (H = 3.7 m), the gas holdup peaks in the annulus for all axial positions, whereas the well-developed gas holdup profile is parabolic. Furthermore, the gas holdup peak position tends to move to the column core with the increase of axial position. From the above discussion, reducing the difference between the initial flow pattern (determined by the gas distribution) and the well-developed flow characteristics will accelerate the flow development and decrease the length of the distributor region. The development length for the gas holdup profile is 2.7 m (about 3.5 column diameters) for uniform and central aeration, and is much longer for near-wall aeration inferred from the measured data. On the basis of the available experimental data, Wu et al.22 developed a correlation for gas holdup profiles in the fully developed region, and the predictions are also shown in Figure 2. The gas holdup predicted by the correlation shows a linear decrease with dimensionless radius, since the calculated power index is about 1. The application of the correlation to largescale bubble column is limited. The axial evolution of liquid velocity profile for different gas distribution types is displayed in Figure 3. With increasing axial position (except at H = 3.7 m), the liquid circulation is strengthened as a result of a steeper gas holdup profile in Figure 2. As the axial position increases from 3.2 to 3.7 m, the liquid velocity profile becomes flatter as a result of the decrease of steepness of the gas holdup profile in Figure 2. Different liquid flow behavior is observed at H = 0.7 m for the three gas distribution types: upward flow in the annulus and downward flow on both sides for the uniform aeration, typical upward flow in the core and downward flow near the wall for central aeration, and abnormal downward flow in the core and upward flow near the wall for near-wall aeration. This is coherent with the different gas holdup profiles at H = 0.7 m in Figure 2, and the initial flow pattern is strongly dependent on the gas distribution type. Furthermore, the liquid velocity at H = 2.2 m is closely proximate to that at H = 2.7 m for uniform and central aeration, while a salient gap is found for the liquid velocity between H = 2.7 m and H = 3.2 m for the near-wall aeration. Consequently, the liquid velocity profile is fully developed at H = 2.2 m for

uniform aeration gives the largest overall gas holdup, and the near-wall aeration creates the smallest value. However, the mean deviation between uniform and near-wall aeration is about 5%, and the effect of gas distribution types on overall gas holdup is thus insignificant. Thorat et al.21 reported when the ratio of gas−liquid dispersion (Hd) and column diameter (D) is larger than 5, the equilibrium bubble size can be attained, and the gas sparger configuration shows little influence on overall gas holdup. Wilkinson et al.7 demonstrated that for column diameter larger than 0.15 m, aspect ratio (Hd/D) larger than 5 and gas hole size larger than 2 × 10−3 m, the overall gas holdup is independent of the column size and the gas sparger. In our experiments, the aspect ratio is slightly larger than 5 and gas hole diameter is 2.5 × 10−3 m. Therefore, Figure 1 shows insignificant effect of gas distribution on overall gas holdup, and this is in accord with the conclusion in the literature.7,8,21 3.2. Influence of Gas Distribution. Figure 2 provides axial evolution of a gas holdup profile under different gas distribution

Figure 2. Influence of gas distribution type on gas holdup profiles at UG = 0.12 m·s−1 in a 0.8 m diameter bubble column: (a) uniform aeration, (b) central aeration, (c) near-wall aeration.

types at UG = 0.12 m·s−1. The gas holdup profile becomes steeper (except at H = 3.7 m) with the increase of axial position for the three gas distribution types. At H = 3.7 m (4 m for gas− liquid dispersion height), the bubbly flow lies in the gas−liquid disengagement region. Thereby, increasing axial position from 3.2 to 3.7 m raises the gas holdup at all radial locations, and 3127

DOI: 10.1021/acs.iecr.5b04015 Ind. Eng. Chem. Res. 2016, 55, 3125−3130

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Figure 4. Gas holdup profiles with uniform aeration at UG = 0.47 m· s−1 in a 0.8 m diameter bubble column.

Figure 5. Liquid velocity profiles with uniform aeration at UG = 0.47 m·s−1 in a 0.8 m diameter bubble column.

respectively. Similar to those in Figure 2a and Figure 3a, gas holdup and liquid velocity profiles both become steeper with increasing axial position. And at H = 2.2 m, gas holdup and liquid velocity are both fully developed. Therefore, when superficial gas velocity increases from 0.12 m·s−1 to 0.47 m·s−1, the development length of the gas holdup profile decreases from 2.7 to 2.2 m, and higher superficial gas velocity diminishes the entrance region length owing to the strengthened liquid turbulence intensity and enhanced bubble radial mitigation. Furthermore, in the gas−liquid disengagement region (H = 3.7 m), the gas holdup profile is flatter and the liquid velocity profile is steeper in comparison with those at H = 3.2 m, which is different from the measured data at low superficial gas velocity in Figure 3a. This is in accord with the experimental observation of strong oscillations of free surface and intensive splashing of the gas−liquid mixture at high superficial gas velocity. The conflicting liquid velocity behavior in the gas− liquid disengagement region at low and high superficial gas velocity may come from two opposite sides: the flatter gas holdup profile reduces the driving force for liquid circulation and tends to decrease the axial liquid velocity; on the other hand, the dispersed bubble phase becomes a continuous gas phase at the free surface, and releases the surface tension energy into the liquid, which tends to accelerate liquid movement. The dual effects give rise to a deceleration at low gas velocity and acceleration at high gas velocity when the liquid flows from the bulk to the disengagement region. 3.4. Influence of Column Scale. The preceding data were measured in the 0.8 m diameter bubble column. Figures 6 and 7 show the axial evolution of gas holdup profile and liquid velocity profile at UG = 0.12 m·s−1 with uniform aeration in a 0.5 m diameter bubble column. The flow parameter profiles are nearly overlapped except at H = 0.7 m. Thereby, the length of

Figure 3. Influence of gas distribution type on liquid velocity profiles at UG = 0.12 m·s−1 in a 0.8 m diameter bubble column: (a) uniform aeration, (b) central aeration, (c) near-wall aeration.

uniform and central aeration, whereas it is still developing in the bubble column for near-wall aeration, similar to the gas holdup profile in Figure 2c. It is indicated once again that the gas distribution type has a profound effect on flow development. In addition, for uniform and central aeration, compared with gas holdup profile, the liquid velocity profile is more liable to become well-developed since the length of the distributor region is reduced from 2.7 m (about 3.5 column diameters) for gas holdup to 2.2 m (about 3 column diameters) for liquid velocity. This may come from the different transfer speed: gas holdup is determined by the conservation of mass and convection is the main transport mechanism, while the liquid velocity is decided by the conservation of momentum, and convection, turbulent diffusion, and interface interactions contribute to the liquid velocity development. The velocity profiles in the fully developed region predicted by the correlation of Wu and Al-Dahhan23 are shown in Figure 3, and the measured data in the fully developed region agree well with the predictions. For the near-wall aeration, the differences between the correlation and the measured data for all axial positions demonstrate that the flow is not fully developed, and it is consistent with foregoing discussion. 3.3. Influence of Superficial Gas Velocity. Figures 4−5 show the axial evolution of gas holdup profile and liquid velocity profile at UG = 0.47 m·s−1 with uniform aeration, 3128

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Industrial & Engineering Chemistry Research

profiles requires a certain axial distances from the pipe inlet,24−27 and the development length is predicted by L /D ≈ 0.06Re

(2)

It means that the ratio of development length to pipe diameter is proportional to fluid velocity and pipe diameter. Durst et al.28 reported the convection and viscous diffusion contribute to momentum transfer and flow development in the laminar flow. Increasing pipe diameter reduces the velocity gradient and the driving force for diffusion, and thus extends the entrance region. In the bubble column, similar to the laminar pipe flow, the liquid velocity evolves from initial distributions to fully developed parabolic profiles, and gas holdup transits from initial distributions to the well-developed center-peak profiles. It is revealed from the measured data in the present study that the increase of superficial gas velocity or the decrease of the column scale or the difference between initial flow pattern and fully developed flow characteristics tends to accelerate the flow development and reduce the entrance region length. For the gas−liquid bubbly flow in the bubble columns, besides convection and viscous diffusion, turbulent diffusion and interface interactions are responsible for the momentum transfer. Consequently, the increase of superficial gas velocity intensifies the liquid turbulence and the interface interactions and reduces the development length. Furthermore, the experimental results indicated that the development of local gas holdup and liquid velocity is not synchronous, and the liquid velocity develops slightly more rapidly than local gas holdup, which may come from their different transfer speeds in the bubble flow. Krishna et al.29,30 found that the off-center gas holdup peak can also cause the parabolic liquid velocity profile in the 2D simulations, and it indirectly demonstrates that the development length for gas holdup and that for liquid velocity may not be identical. In addition, the effect of column scale on flow development found in this paper illustrates the importance of designing a suitable gas sparger for large-scale industrial bubble column reactors, as the main part of the column may be the distributor region rather than the generally accepted fully developed region.2,31,32

Figure 6. Gas holdup profiles with uniform aeration at UG = 0.12 m· s−1 in a 0.5 m diameter bubble column.

Figure 7. Liquid velocity profiles with uniform aeration at UG = 0.12 m·s−1 in a 0.5 m diameter bubble column.

the distributor region for the 0.5 m diameter bubble column is less than 1.2 m (about two column diameters), which agrees with the findings in the publications.9−11 Compared with the measured results in the 0.8 m diameter bubble column in Figures 2a and 3a, it is revealed that the development length is extended from about 2 column diameters to 3 column diameters for liquid velocity when the column size increases from 0.5 to 0.8 m. Consequently, the increasing of column size slows the flow development in the bubble column, and the flow becomes fully developed at an axial position of larger aspect ratio. Hitherto the finding has not been reported in the literature. It is inferred that in large-scale industrial bubble column reactors, the region controlled by gas sparger expands and designing a suitable gas sparger is of profound significance. For the 0.5 m bubble column, the correlations22,23 show agreement with the measured data in the fully developed region. Hence, the velocity profile correlation23 can be applied more widely than the gas holdup profile correlation,22 since the power index of velocity profile correlation is less sensitive to column size and superficial gas velocity.

5. CONCLUSIONS Flow development in large-scale bubble columns was experimentally investigated in the present study. Local gas holdup and liquid velocity were measured to determine the development length, and it is strongly influenced by the gas distribution type, superficial gas velocity, and column scale. The flow develops more rapidly for uniform and central aeration than near-wall aeration. The finding of a strong influence of the sparger type on flow development is opposite to the previous observations in the literature owing to the different column scales. Increasing the superficial gas velocity or decreasing the column size tends to accelerate the flow development and reduce the entrance region length. Furthermore, the experimental data indicated that the liquid velocity is more liable to become fully developed in comparison with local gas holdup, and it may be related to the different transfer speed. In addition, the finding about the effect of column scale on flow development reveals the importance of designing a suitable sparger for large-scale industrial bubble column reactors, as the distributor region, rather than the well-developed region, may occupy the main part of the column.

4. DISCUSSION For the fluid flow, the entrance region is the transition section from the initial flow pattern to the well-developed region, and its length is dictated by the transfer speed. For the flow behavior in the bubble columns, most literature focuses on the fully developed state, and limited publications accounted for the range of the entrance region. This paper studies the length of the distributor region in the bubble columns and found that it is significantly influenced by the gas distribution type, superficial gas velocity, and column scale. For the single phase laminar pipe flow, the evolution from initial flat velocity profiles to fully developed parabolic velocity 3129

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS This work is funded by the National Natural Science Foundation of China (U1361112) and National High-tech R&D Program (NO.2011AA05A205).

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NOMENCLATURE D = column or pipe diameter, m H = axial position, m Hd = dynamic height, m Hs = static height, m L = development length, m r = radial position, m R = radius of column, m Re = Reynolds number, Re = Duρ/μ u = velocity, m·s−1 UG = superficial gas velocity, m·s−1 uL = axial liquid velocity, m·s−1

Greek Symbols

εG = local gas holdup, dimensionless εG = overall gas holdup, dimensionless μ = viscosity, Pa·s ρ = density, kg·m−3

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REFERENCES

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