Hydrodynamics and mass transfer characteristics of co-current downflow packed tube columns

Hydrodynamics and effective interfacial area in a 25 mm i.d. packed tube column were studied over a wide range of operating conditions for demister pad packings (DPP). Flow maps have been prepared. Values of effective interfacial area as high as 1880 m^?1 in the spray flow regime were obtained. Data on pressure drop and effective interfacial area have been correlated for different flow regimes. Values of liquid side volumetric mass transfer coefficient, k_L a , were measured by absorption and desorption of oxygen in different packed tube columns containing Pall rings (standard and low height to dia. ratio), multifilament wire gauge packings (MFWGP) and DPP. kL a was found to vary from 0.017 to 0.34 s^?1 for DPP. Values of wall side solid-liquid mass transfer coefficient, k_SL, were obtained in a 25 mm i.d. copper tube column packed with MFWGP by the dissolution of copper in acidic dichromate solutions. Values of wall side heat transfer coefficient could be obtained by analogy.     

Liquid mixing by large gas bubbles in bubble columns

Liquid mixing by large gas bubbles of spherical cup was investigated for co- and counter-current contact of air-water system with bubble columns of 5 and 10 cm dia. The results obtained are that for the column of 5 cm dia., the longitudinal dispersion coefficient ranges from 5 to 20 cm²/sec for superficial gas velocity from 0·07 to 8 cm/sec and that for the one of 10 cm in diameter it ranges from 9 to 45 cm²/sec for that from 0·035 to 8 cm/sec. Liquid mixing under the coexistence of large and small bubbles was also investigated and it was found that the gas holdup was fairly well explained by an equation derived on the assumption that the mixture of small bubbles and liquid behaves independently of large bubbles. The expansion model was applied to the experimental results on the longitudinal dispersion coefficient and it was observed that there should be the lower limit in the holdup of small bubbles where this model can be applied.     

Hydrodynamics in a pilot‐scale cocurrent trickle‐bed reactor at low gas velocities

Hydrodynamic data obtained from laboratory‐scale trickle‐beds often fail to accurately represent industrial‐scale systems with high packing aspect ratios and column‐to‐particle diameter ratios. In this study, pressure drop, liquid holdup, and flow regime transition were investigated in a pilot‐scale trickle‐bed column of 33 cm ID and 2.45 m bed height packed with 1.6 mm × 8.4 ± 1.4 mm cylindrical extrudates for air‐water mass superficial velocities of 0.0023 – 0.094 kg/m²s and 4.5 – 45 kg/m²s, respectively, at atmospheric pressure. Significant deviation was observed from pressure drop and liquid holdup correlations at low liquid flows rates, corresponding to gravity‐driven flow limit. Likewise, liquid saturation is overestimated by correlations at high liquid flow rates, owing to significantly reduced wall effects. Lastly, trickle‐to‐dispersed bubble flow and trickle‐to‐pulsing flow regime transitions are reported using a combination of visual observations and analysis of the magnitude of local pressure fluctuations within the column. © 2018 American Institute of Chemical Engineers AIChE J, 64: 2560–2569, 2018     

Counter-current absorption using wire gauze packings

The dynamic liquid hold-up, ?_LD, effective interfacial area, a, and the liquid side mass transfer coefficient k_La were determined for 0.1 m and 0.2 m multifilament wire gauze packings, 0.0125 m double walled wire gauze partition rings and 0.025 m wire gauze saddle packings in columns operated countercurrently. The theory of gas absorption accompanied by fast pseudo mth order reaction was used to determine the effective interfacial area. The values of liquid side mass transfer coefficient for the multifilament wire gauze packings were obtained by absorbing lean carbon dioxide in a buffer solution of sodium carbonate and sodium bicarbonate. K_La values for the other packings were obtained by absorbing pure carbon dioxide in tap water. The values of a and k_La for multifilament wire gauze packings were found to be two to four times higher as compared to the conventional ring or saddle packings. Further, the superficial liquid velocity was found to have marginal effect on a. The double walled wire gauze partition rings offered a values which were 1.5–2.0 times higher than that offered by 0.016 m s.s. Pall rings at low values of superficial liquid velocity (<3 × 10^{?3 m/s}.     

Mass transfer in packed columns: co-current operation

The theory of gas absorption accompanied by fast pseudo-m^th order reaction was used to obtain values of effective interfacial area in 10,15 and 20 cm i.d. packed columns which were operated co-currently. A variety of ceramic, metal and plastic packings were used. The range of superficial gas and liquid velocities was 50–300 cm/sec, and 0·1–3·5 cm/sec, respectively. Values of gas side mass transfer coefficients for some of the packings were also obtained. In addition some data were obtained for the counter-current mode of operation.     

Axial mixing in a 15 cm diameter reciprocating plate bubble column

Axial dispersion coefficients (E) in the liquid phase have been measured by unsteady tracer response methods in a 15 cm internal diameter reciprocating plate bubble column, using air and water in countercurrent and cocurrent flow. The operating variables studied were amplitude (0.6–1.27 cm) and frequency (0–5 Hz) of reciprocation, and the superficial velocities of the liquid and gas phases and the spacing between plates. Three types of plate were studied; conventional Karr-type plates with perforation diameters 1.43 cm, plates with smaller (0.635 cm) perforations, and single-perforation (doughnut) plates with internal diameter 7 cm. Measured values of E ranged from about 1 cm²/s to a maximum of 116 cm²/s. In general, the plates with 0.635 cm perforations gave the smallest values of E while the largest values of E were obtained with the doughnut plates, due to vortex ring shedding. The single liquid phase data for the three types of plate were approximately consistent with the correlation of Stevens and Baird (1990). The gas-liquid flow results were interpreted in terms of several different hydrodynamic effects.     

Dispersion in trickle and two-phase flow in packed columns

The coefficients of axial and radial dispersion in both trickle and two-phase flow have been measured for a range of spherical packings, and for $\text{case1}\text{2}$ in., 1 in. and 1$\text{case1}\text{2}$in. Raschig rings packed in 0.10 m and 0.30 m dia. columns. The coefficients were estimated from the dispersion of a pulse of tracer injected from a small tube on the axis of the column and measured before the trace material had dispersed to the wall. Under these conditions the dispersion coefficients so found are not affected by wall flow of liquid, and therefore the coefficients of axial dispersion are generally smaller than those found by other investigators who estimated coefficients for overall dispersion and so included the effect of wall flow in their estimates.Liquid hold-up within the column was measured from the change in the first moment of the injected dispersed pulses. Liquid hold-up for conditions of trickle flow is given in the form of a correlation of dimensionless groups.     

Absorption in packed bubble columns

The mass transfer characteristics of packed bubble columns were studied by employing various packings of different sizes and shapes in 10–38·5 cm i.d. columns. The theory of absorption accompanied by pseudo-mth order reaction was used to obtain the values of effective interfacial area. The values of liquid side mass transfer coefficient were obtained by using the theory of absorption accompanied by slow chemical reaction. The superficial gas velocity was varied from 5 to 25 cm/sec. The packed bubble columns showed a considerable improvement in the performance over empty bubble columns. A criterion for the scale-up of these columns has been suggested.     

Hydrodynamic and mass transfer characteristics of bubble and packed bubble columns with downcomer

The hydrodynamic and mass transfer characteristics of bubble and packed bubble columns with downcomer were investigated. The contactor consisted of two concentric columns of 0.11 and 0.2 m i.d., with the annulus acting as the downcomer. The packing used in this investigation was standard 16 mm stainless steel Pall rings. The superficial gas and liquid velocities, V_G and V_L, were varied from 0.01 to 0.09 and 1 × 10^?3 to 8.8 × 10^?3 m s^?1 respectively. Two flow patterns, namely the bubble and pulse flows were observed in the packed bubble column with downcomer, as shown by a flow map. The liquid circulation velocity in both the contactors was observed to be constant throughout the ranges of V_G and V_L covered in this work. The effect of liquid viscosity (0.8 to 9.5 mPa ? s) and surface tension (45 to 72 mN m^?1) on the flow pattern, liquid circulation, gas hold-up and pressure drop was investigated. The pressure drop characteristics across the two contactors have been compared with those across a bubble column. Values of the effective interfacial area, a, and the volumetric mass transfer coefficient, k_L a, were measured by using chemical methods. Values of a as high as 180 and 700 m^?1 and k_L a as high as 0.075 and 0.22 s^?1, in the bubble and packed bubble columns with downcomer, respectively, were obtained. The values of true liquid-side mass transfer coefficient, k_L, were found to be independent of V_G and were of the order of 5.5 × 10^?4 and 3.5 × 10^?4 m s^?1, respectively, in the two contactors.     

Onset of pulsing in gas-liquid trickle bed filtration

When liquid suspensions containing low concentration of fine solids are treated in catalytic packed bed gas-liquid-solid reactors, which are operated in trickle flow or near the transition between trickle and pulse flow, plugging develops and increases the resistance to two-phase flow. Also due to obstruction, such accumulation of fines in the catalyst bed shifts progressively the flow pattern from trickling to pulsing flow. The progressive onset of pulsing flow along the packed bed was estimated using a sequential approach based on combining a “large time-scale” unsteady-state filtration solution of two-phase flow with a “short time-scale” solution of a linear stability analysis of two-phase flow. Space-time evolution and two-phase flow of the deposition of fines in trickle bed reactors under trickle flow regime was described using a one-dimensional two-fluid model based on the volume-average mass and momentum balance equations and volume-average species balance equation for the fines. The model hypothesized that plugging occurred via deep-bed filtration and incorporated physical effects of porosity and effective specific surface area changes due to the capture of fines, inertial effects of phases, and coupling effects between the fines filter rate equation and the interfacial momentum exchange force terms. The transition between trickle flow and pulse flow regimes was described from a stability analysis of the solution of the transient two-fluid model around an equilibrium state of trickle flow under pseudo steady state conditions. The impact of liquid superficial velocity, viscosity and surface tension, gas superficial velocity and density, feed fines concentration, and fines diameter on the transition between trickle and pulse flows in the presence of fines deposition was analyzed.     

Bubble formation in flowing liquid

Initial bubbles in flowing liquid from a nozzle were observed from two mutually perpendicular directions. Two nozzles of 0.086 cm and 0.305 cm in diameter were used. The gas flow rate and the superficial liquid velocity ranged from 0.33 cm³/s to 36.2 cm³/s and from 0 cm/s to 154.9 cm/s, respectively. The bubble size formed in flowing liquid decreased with decreasing gas flow rate and with increasing superficial liquid velocity. Three types of bubble formation, i.e. single bubbles, coalescent bubbles and gas jets, were observed depending upon the gas rate and the liquid velocity. Two empirical equations of the bubble sizes are given.     

Flow regimes in a gas–liquid–solid three-phase moving bed

The gas–liquid–solid three-phase moving beds could supply a potential solution for multiphase reactions with catalyst easily deactivated, and the flow regimes in it were studied by optical method and pressure drop measurement. Results showed that taking the trickle flow as the initial flow regime, the flow channels were more obvious as the particle velocity increased. When the initial flow regimes were pulse flow and bubble flow respectively, the pulse-to-trickle and bubble-to-pulse flow transitions mainly occurred at moderate-to-high particle velocities (0.01–0.04 m s⁻¹ under conditions used in this work). Moreover, the flow regime map in the three-phase moving bed was constructed and shown that the region of trickle flow increased and the region of bubble flow decreased. Finally, the application of three-phase moving beds was discussed, and it could be suitable for those reactions, which had to operate in the pulse flow, bubble flow, and transition zone.     

On flow regime transition in trickle bed: Development of a novel deep-learning-assisted image analysis method

An image analysis method was developed based on deep-learning algorithms to extract phase fractions quantitatively in a rectangular trickle bed, and the average identification error was lower than 5%. Furthermore, the flow regime transition in the trickle bed was studied. In trickle-to-pulse flow transition, the trickle flow could be further classified into the stable trickle flow and accelerated one. The SD of liquid fractions and the peak width at half-height of the probability density curve of liquid fractions were close to zero in stable trickle flow, increased rapidly in accelerated trickle flow, and remained approximately constant in pulse flow. In bubble-to-pulse flow transition, dispersed bubbles in bubble flow induced the outliers outside the upper boundary of the boxplot of gas fraction, while alternative appearance of gas-rich zone and liquid-rich zone in pulse flow induced outliers outside both the upper and lower boundaries of the boxplot of gas fraction.     

Mass transfer in bubble columns packed with motionless mixers

The cocurrent upward mode was employed to absorb pure oxygen into water in bubble columns packed with Koch (Sulzer) motionless mixers. The liquid-side volumetric mass transfer coefficient, K_La, in the packed bubble column was found to be always larger than that in the unpacked bubble column. In the range of liquid velocities from 6.7 cm/sec to 39.9 cm/sec, the value of K_La in the packed bubble column increased with the increasing liquid velocity while that in the unpacked bubble column was almost independent of the liquid velocity. The equation of the formK_La= mν_l^β? was successfully adopted to correlate the K_La data.     

Flow regime transition identification in three phase co‐current bubble columns

Bubble columns have wide applications in absorption, bio‐reactions, catalytic slurry reactions, coal liquefaction; and are simple to operate, have less operating costs; provide good heat and mass transfer. Experiments have been performed for identifying transition regime in a 15 cm diameter bubble column with liquid phase as water and air as the gas phase. Glass beads of mean diameter 35 µm have been used as solid phase. The superficial gas velocity is in the range 0 ≤ U_g ≤ 16.3 cm/s and superficial liquid velocity in the range of 0 ≤ U_l ≤ 12.26 cm/s. Solid loading up to 9% (w/v) has been used. Pressure signals have been measured using differential pressure transducers (DPTs) at four different axial locations. Classical analysis (Wallis approach and Zuber–Findlay approach), Statistical analysis and Fractal analysis have been used for regime transition identification. Statistical analysis and Fractal analysis have shown almost the same transition points for all the liquid and gas velocities. Effect of solid concentration, liquid velocity and gas velocity over transition regime has also been studied. As the solid concentration is increased it has insignificant effect over transition regime for lower values (<1%), while transition values decrease for higher solid concentration (>1%). © 2012 Canadian Society for Chemical Engineering     

Holdup and flow pattern of solid particles in freeboard of gas—solid fluidized bed with fine particles

The behavior of solid particles suspended in the freeboard of a 12 cm i.d. fluidized bed was studied using FCC particles of which the weight-mean diameter was 65 – 68 μm. The height of the freeboard was 180 cm, and the superficial gas velocity was changed from 15 to 50 cm s^?1.The weight-mean diameter and holdup of solid particles in the freeboard below the TDH vary with the axial position and the gas velocity. The lateral distribution of the solids holdup measured by a capacitance probe is almost flat in the region of r ? 5 cm, but the relative pulse frequency detected by a fiber optic probe indicates that the solids holdup increases greatly in the immediate vicinity of the column wall. The flow patterns of solid particles were measured by a fiber optic probe and a thermal response probe. The steady internal circulation of solid particles is formed in the freeboard below the TDH, and the circulating solids flow rate is much larger than the net entrainment rate of solid particles.     

Mass transfer in packed liquid—liquid extraction columns

The mass transfer characteristics of 3·5, 7·3, 10·16 and 15·6 cm i.d. packed liquid—liquid extraction columns were studied with a variety of packings such as, $\text{3}\text{8}$, $\text{1}\text{2}$ and 1 in. ceramic Raschig rings, $\text{5}\text{8}$ in. stainless steel Raschig rings, $\text{1}\text{2}$ and 1 in. ceramic Intalox saddles, $\text{5}\text{8}$ and 1 in. stainless steel Pall rings, and 1 in. polypropylene Intalox saddles and Pall rings. Some data were also obtained for the co-current mode of operation (up-flow) for packed columns, and without packings. In addition the mass transfer charactersitics of a packed extraction column with film flow were studied.The theory of extraction accompanied by a fast pseudo-first order reaction was employed to measure the values of effective interfacial area. The values of overall (continuous and/or dispersed phase) mass transfer coefficient were measured by the Colburn—Welsh technique. A fairly wide range of physical properties of the two phases was covered.The values of overall (continuous phase) mass transfer coefficient and effective interfacial area for new packings such as Pall rings and Intalox saddles, under otherwise similar conditions, are only about 10 and 35 per cent higher, respectively, than those provided by the conventional packings of the same nominal size. However, the flooding velocities for the newer packings are as much as 80 per cent higher than those for the conventional packings of the same nominal size.     

X-ray computed tomography in large bubble columns

X-ray computed tomography (CT) is used to explore the differences in a semi-batch bubble column operated at superficial gas velocities of U_g=3, 10, and 18 cm/s. Air-water or air-water-cellulose fiber systems comprise the multiphase flow, and the bubble column has a 32.1 cm internal diameter. A CT image of a phantom object composed of several air-filled tubes immersed in water is used to identify several characteristic features of the X-ray CT system. CT images are then compared between air-water and air-water-cellulose fiber systems. When the fiber mass fraction is 0.1%, gas holdup is slightly higher than that of the air-water system in the column center and near the column wall. In 1.0% cellulose fiber slurries, gas holdup is lower than that of air-water results at all radial positions.     

The effect of gas‐liquid counter‐current operation on gas hold‐up in bubble columns using electrical resistance tomography

BACKGROUND: In order to improve the performance of a counter‐current bubble column, radial variations of the gas hold‐ups and mean hold‐ups were investigated in a 0.160 m i.d. bubble column using electrical resistance tomography with two axial locations (Plane 1 and Plane 2). In all experiments the liquid phase was tap water and the gas phase air. The superficial gas velocity was varied from 0.02 to 0.25 m s^?1, and the liquid velocity varied from 0 to 0.01 m s^?1. The effect of liquid velocity on the distribution of mean hold‐ups and radial gas hold‐ups is discussed. RESULTS: The gas hold‐up profile in a gas–liquid counter‐current bubble column was determined by electrical resistance tomography. The liquid velocity slightly influences the mean hold‐up and radial hold‐up distribution under the selected operating conditions and the liquid flow improves the transition gas velocity from a homogeneous regime to a heterogeneous regime. Meanwhile, the radial gas hold‐up profiles are steeper at the central region of the column with increasing gas velocity. Moreover, the gas hold‐up in the centre of the column becomes steeper with increasing liquid velocity. CONCLUSIONS: The value of mean gas hold‐ups slightly increases with increasing downward liquid velocity, and more than mean gas hold‐ups in batch and co‐current operation. According to the experimental results, an empirical correlation for the centreline gas hold‐up is obtained based on the effects of gas velocity, liquid velocity, and ratio of axial height to column diameter. The values calculated in this way are in close agreement with experimental data, and compare with literature data on gas hold‐ups at the centre of the column. Copyright © 2010 Society of Chemical Industry     

Trickle bed hydrodynamics and flow regime transition at elevated temperature for a Newtonian and a non-Newtonian liquid

Despite the hydrodynamics of trickle beds experiencing high pressures has become largely documented in the recent literature, trickle bed hydrodynamic behavior at elevated temperatures, on the contrary, largely remains terra incognita. This study's aim was to demonstrate experimentally the temperature shift of trickle-to-pulse flow regime transition, pulse velocity, two-phase pressure drop, liquid holdup and liquid axial dispersion coefficient. These parameters were determined for Newtonian (air-water) and non-Newtonian (air-0.25% Carboxymethylcellulose (CMC)) liquids, and the various experimental results were compared to available literature models and correlations for confrontation and recommendations. The trickle-to-pulse flow transition boundary shifted towards higher gas and liquid superficial velocities with increasingly temperatures, aligning with the findings on pressure effects which likewise were confirmed to broaden the trickle flow domain. The Larachi-Charpentier-Favier diagram [Larachi et al., 1993, The Canadian Journal of Chemical Engineering 71, 319-321] provided good predictions of the transition locus at elevated temperature for Newtonian liquids. Conversely, everything else being kept identical, increasingly temperatures occasioned a decrease in both two-phase pressure drop and liquid holdup; whereas pulse velocity was observed to increase with temperature. The Iliuta and Larachi slit model for non-Newtonian fluids [Iliuta and Larachi, 2002, Chemical Engineering Science 46, 1233-1246] predicted with very good accuracy both the pressure drops and the liquid holdups regardless of pressure and temperature without requiring any adjustable parameter. The Burghardt et al. [2004, Industrial and Engineering Chemistry Research 43, 4511-4521] pulse velocity correlation can be recommended for preliminary engineering calculations of pulse velocity at elevated temperature, pressure, Newtonian and non-Newtonian liquids. The liquid axial dispersion coefficient (D_ax) extracted from the axial dispersion RTD model revealed that temperatures did not affect in a substantial manner this parameter. Both Newtonian and power-law non-Newtonian fluids behaved qualitatively similarly regarding the effect of temperature.