Investigation of Local Mass Transfer in a Packed Bed of Pall Rings

The average mass-transfer rate with the multipoint distributor was about 14% ...... Chau, P. C. Local Liquid−Solid Mass Transfer Measurement in A Tr...
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Investigation of Local Mass Transfer in a Packed Bed of Pall Rings Using a Limiting Current Technique Jeffrey Gostick,† Huu D. Doan,*,‡ Ali Lohi,‡ and Mark D. Pritzker§ Teck Cominco Metals Ltd., Product Technology Centre, Sheridan Park, Mississauga, Ontario, L5K 1B4 Canada, Department of Chemistry, Biology and Chemical Engineering, Ryerson University, Toronto, Ontario, M5B 2K3 Canada, and Department of Chemical Engineering, University of Waterloo, Waterloo, Ontario, N2L 3G1 Canada

The mass-transfer coefficients at varied axial, radial, and angular positions in a packed bed of Pall rings were measured using the limiting current technique. The experimentation was carried out in a 0.30-m-diameter column with a bed height of 5.5 column diameters. Ferri-/ferrocyanide ions were used as the redox couple for the mass-transfer measurement. The local mass-transfer rate was found to be strongly dependent on the liquid feed distribution, especially at the top section of the packed bed. The local mass-transfer rate also varied significantly in the axial and radial directions in the packed bed. A maximum mass-transfer rate was observed at an axial distance of 1.5 column diameters (x/D ) 1.5) when the multipoint liquid distributor was used while the highest transfer rate occurred at x/D ) 2.5 for the single-point liquid distributor. However, for both types of liquid distributors the mass-transfer rate became steady at x/D > 4.5. The wall flow also reached steady-state x/D ) 3.5. The average mass-transfer rate with the multipoint distributor was about 14% higher than that with the single-point distributor. 1. Introduction The packed bed is one of the most common contactors in the chemical process industries. This unit operation is very mature and several well-established design methods are available to size these columns.1 In general, these design methods assume a constant mass-transfer coefficient at all locations in the column. However, this is known to be untrue, particularly in gas-liquid contacting columns where the liquid flows in the form of channels through the packing and tends to accumulate on the wall of the column. Considering the maturity of the packed bed as a chemical engineering unit, there is a surprisingly sparse amount of literature available that considers local phenomena in a truly systematic fashion. This is especially true of the irrigated liquid downflow mode. Many authors point out that they expect spatial variations to occur and accordingly measure the masstransfer coefficient at several locations to obtain an average value.2-4 However, most do not present the local variations in mass-transfer coefficients, although some exceptions are to be found. Using the electrochemical method, Gabitto and Lemcoff5 studied the radial dependence in a trickle-bed reactor but did not report results at different axial positions. Further, since they did not mention the axial position at which the data were obtained, it is difficult to fully interpret their results. Marcandelli et al.6 attempted to measure the local heat-transfer coefficients in a packed bed only at one axial position. They found that the heat-transfer * To whom correspondence should be addressed. Department of Chemistry, Biology and Chemical Engineering, Ryerson University, 350 Victoria Street, Toronto, Ontario, M5B 2K3 Canada. Tel.: (416) 979-5000, ext. 6341. Fax: (416) 9795044. E-mail: [email protected]. † Teck Cominco Metals Ltd. ‡ Ryerson University. § University of Waterloo.

coefficient was again highest at the column center and wall, similar to the findings of Gabitto and Lemcoff.5 Some very detailed profiles are given by Stichlmair and Stemmer.7 They determined the gas-liquid mass transfer in a large-scale packed column by measuring the temperature profiles as hot water and air were run countercurrently. Several hundred thermocouples were placed in the column to measure high-resolution profiles, which were presented in the form of isotherms. Very convincing profiles were presented with a variety of packing materials and liquid distributors. However, because of latent heat effects, changing driving forces, and almost certain humidity saturation in the air at the top of the column, these data are difficult to interpret accurately. The present paper describes the use of an electrochemical method to systematically and directly measure local mass-transfer coefficients throughout an entire packed bed and study variations in the axial, radial, and angular directions. This method is applied to elucidate the effects that liquid distribution and flow patterns have on local mass-transfer coefficients in a large laboratory-scale packed bed with stainless steel Pall rings. This approach can be easily extended in future work to investigate the effects of geometric factors (e.g., packing size, shape, and material as well as column size). 2. Experimental Section The system investigated consisted of a 0.30-mdiameter PVC column randomly packed with 0.025-m stainless steel Pall rings. A bed height of 5.5 column diameters was used to ensure that development of flow patterns could be observed in the bed. An “X” type liquid distributor with 16 nozzles in total was used, giving a nozzle density of 219.3 per m2. A single-point distributor was also used to investigate the effects of a high degree of liquid feed maldistribution. The process flow configuration is shown in Figure 1. This setup allowed both

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The experimental error of kL was estimated using the method given by Kline and McClintock (1953) for singlesample experiments in the present study.8 The experimental error was found to be at 7% and used for error bars in all graphs presented in the Results section. The reaction system selected for the present study was the ferri-/ferrocyanide redox couple.

Fe(CN)63- + e- T Fe(CN)64-

Figure 1. Process flow diagram of the experimental apparatus.

full-liquid upflow and irrigated downflow modes to be investigated. Upflow mode was used as a diagnostic tool to ensure that the electrodes were functioning properly. The mass-transfer coefficient profiles presented herein were obtained in the irrigated downflow mode. The liquid flow rate was varied to yield particle Reynolds numbers between 80 and 550. The Reynolds number was calculated as

Re )

deU ν

(1)

where de is the diameter of a sphere having the same surface area as the packing, U is the superficial liquid velocity, and ν is the kinematic viscosity of the liquid. The Schmidt number was between 1400 and 1600 depending on the temperature of the electrolyte. The bed was initially flooded to wet the particles and ensure that flow patterns developed randomly and were not influenced by pre-existing bed conditions prior to a run, such as partially wetted sections. The liquid flow rate was increased from low to high values as well as at random for some runs. At each flow rate, sufficient time was allowed for steady-state conditions to be reached before the mass-transfer measurement was made. For varied flow rates used, the dynamic liquid holdup in the column was estimated between 0.051 and 0.10 m3‚m-3 or 5.4-10.5% of the void space in the packed bed with the packing porosity of 0.95 for Pall rings. Stainless steel 304 Pall rings, plated with pure nickel, were used as the cathode. These packing pieces were placed strategically throughout the bed to provide profiles of the mass-transfer coefficient in the axial, radial, and angular directions. Stainless steel 304 Pall rings were also used as the anode. These electrodes were separated from each other, and the anode, by polymercoated SS304 Pall rings. Measurements of the limiting current ie at each cathode enabled the determination of the mass-transfer coefficient according to eq 2. where

kL )

ie azFCb

(2)

a is the cathode area, z is the number of electrons transferred in the electrode reaction, F is the Faraday constant, and Cb is the bulk concentration of the reactant.

(3)

Ferri- and ferrocyanide concentrations of 3.8 and 4.0 mol‚m-3 respectively, were used. This system was chosen because of its many advantages. The reactions do not involve solid reactants or products that would alter the electrode surface. The electrode surface conditions thus remain intact, which is an important factor in irrigated flow. Another advantage is that the bulk concentration of ferricyanide remains constant with this reacting system since ferrocyanide is oxidized back to ferricyanide at the anode. This ensures that the driving force for mass transfer is uniform throughout the bed. The presence of a parasitic side reaction at either electrode would skew the 1:1 reaction of the redox couple, causing the concentrations to vary with time. The steps taken to prevent this are discussed below. A large excess (500 mol‚m-3) of NaOH was employed to reduce the effect of the ionic migration on the transport of the reactants (ferricyanide and ferrocyanide ions); hence, the transfer process was mass-transfercontrolled. When measuring the limiting reduction of ferricyanide at the cathode, migration has the effect of decreasing the observed mass-transfer coefficient. Since the ferricyanide ion is negatively charged, it must diffuse against the gradient of electrical potential between the electrodes. Adding an excess of a supporting electrolyte increases the conductivity of the solution and reduces the potential gradient. Therefore, the effect of the ionic migration on the transport of the reactants to the electrodes becomes negligible, and the transport of the reactants to the electrodes is mainly due to convection mass transfer.9 Physical property data for this system were taken from reported literature.10 Concentrations were measured before and after each run. The concentrations did not drift noticeably during the course of a run, which typically lasted 4-5 h. Light and oxygen were rigorously excluded to prevent the degradation of the electrolyte.11 All translucent parts were covered with aluminum foil to block out light since it is known to cause decomposition of both ferricyanide12,13 and ferrocyanide.14 Nitrogen was bubbled into the electrolyte tank to strip oxygen from the liquid while the whole system including the packed column was kept under vacuum.15 During the experiment, there was no gas in the column. The experiment was performed with a liquid flow only. The presence of oxygen not only accelerates the decomposition of the reactants but also reacts at the cathode, introducing a parasitic current into the measurements.16 However, even with these effects, the electrolyte had a useful lifetime of approximately a month. Fouling of the electrodes in the form of a white surface film occurred after long operation time. The exact nature of this film is not fully understood. Moggi et al.17 investigated the composition of the film and found it to consist of many ferrous-based compounds, indicating that ferri-/ferrocyanide decomposition products were precipitating onto the surface. This film reduced the limiting current since it rendered a portion of the active surface inert. We found that these effects could be reversed by soaking the electrodes in a

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Figure 2. Placement of electrodes in the column. (A) Radial locations of electrodes at each layer. (B) Axial locations of electrode layers in the packed bed.

for 20 min. The finished surface of polymer-coated Pall rings was wettable. A simple liquid-drop test was done, and the observed contact angle was similar to that of the uncoated stainless steel Pall rings. The column was packed with all cathodes in place but only one layer was active at a given time. Once data were acquired for all flow rates at a given layer, the circuitry was switched to activate the next layer. Only the anode section below the active cathode layer was active so that the bulk concentration of the electrolyte entering the cathode layer was not affected by the ferricyanide production at the anode. The surface area of an anode section was much larger than that of a cathode section to ensure that the anodic reaction never limited the extent of the cathodic reaction in any way. 3. Results 3.1. Full Liquid Upflow. A series of experiments using full liquid upflow mode was conducted to assess whether the column was functioning acceptably. Aside from enabling a comparison with the performance of other columns reported in the literature, it also provided a baseline for analyzing the behavior of the gas-liquid contacting mode. Much literature has been published for full-liquid flow in packed beds.20-22 However, no studies were found that utilized Pall rings. The study by Nosier et al.23 on the diffusion-controlled dissolution of copper Raschig rings in acidified dichromate solutions provided the closest match. For a range of Reynolds numbers between 5 and 246 with the particle diameters between 0.006 and 0.012 m, they proposed a correlation for the Sherwood number in the form of Sh ) ARebScc as below:

Figure 3. Layout of cathode and anode section in the packed bed.

5% acetic acid solution for several minutes.18 The hydrogen evolution treatment was found to be insufficient for this task, despite being often recommended.19 Although an improvement did occur after the hydrogen evolution treatment, it did not provide a full recovery of surface activity; furthermore, performing this treatment after the acetic acid treatment did not yield any additional improvement. The electrochemically active Pall rings were placed so that measurements could be made at three radial positions at each of six elevations (Figure 2). Four electrodes were placed at the inner and outer radii to account for random deviations in the angular direction. The two electrodes placed at the center of the column were operated on the same circuit. This was done to alleviate variations at the center electrode since it was not averaged between four electrodes as were the inner and outer radial locations. By placing two electrodes here, a larger cross section of the bed was measured, providing a more representing picture of the flow conditions. It was experimentally verified that two electrodes at a given location reduced the variability (average deviation) dramatically to about 0.075 compared to 0.22 for one electrode. Beyond two electrodes the variability was not reduced significantly.18 The layout of the cathode-anode sections in the bed is shown in Figure 3. As shown, each section was isolated from an adjacent section by a layer of insulating polymer-coated Pall rings. Powder coating, which is commonly used to apply a protective coating on automotive parts, was used to coat Pall rings. In the powdercoating process, polymer powder was melted and cured on the metal surface under a high temperature (149 °C)

Sh ) 2.3Re0.43Sc0.33

(4)

This compares favorably to the results obtained in the present study. For the Reynolds number range of 80400, the mass-transfer correlation was found to be

Sh ) 4.1Re0.39Sc0.33

(5)

with a correlation coefficient R2 ) 0.997, an average deviation of 0.54%, and a maximum deviation of 1.13% between the predicted and the experimental Sherwood numbers. Differences between the correlations given in eqs 4 and 5 may be due to the differences in particle type. The Pall rings used in the present study have perforations in their sides, which would promote higher solid-liquid interaction in the inner surface. This may explain the larger value of the coefficient, A, observed with Pall rings. Furthermore, Pall rings have a higher void fraction than Raschig rings, 0.95 versus 0.75, respectively. Consequently, the interstitial velocity would be lower with Pall rings for the same Reynolds number, which is based on the superficial velocity. This may explain the reduced exponent, b, obtained in the present study since a similar increase in flow rate would result in a smaller increase in interstitial velocity for the Pall rings. 3.2. Irrigated Liquid Downflow. The irrigated liquid downflow mode was the focus of the present study. When this mode of operation was used, profiles of the mass-transfer coefficient were obtained in the radial and axial directions. From the averaged data of the entire bed, the overall mass-transfer coefficient was also obtained.

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Figure 4. Comparison of the correlations for the overall mass transfer in a packed bed with a liquid downflow using the electrochemical method.

3.2.1. Overall Mass-Transfer Coefficient. Knowledge of the overall mass-transfer coefficient is important to put the profiles of local mass-transfer coefficients into context. Comparison between the overall value and the individual values obtained locally allows determination of errors introduced when only a single overall value is used. The overall value was obtained from the local values by calculating A and b values of the masstransfer correlation for each electrode and taking an average. The Reynolds number was calculated using the superficial liquid velocity in the column, which is the average liquid velocity since the local liquid velocity at individual electrodes was not able to be measured. In addition, according to the analysis of the boundary layer under a laminar flow, the Sherwood number is proportional to the Schmidt number to the power of 1/3. This proportionality has been adopted in many reported pieces of literature; hence, it was used in the present study as well. The equation obtained was

Sh ) 5.1Re0.44Sc0.33

(6)

with a correlation coefficient R2 ) 0.989, a maximum deviation of 13.3%, and an average deviation of 7.8%. The exponent b is relatively constant when the data of each individual electrode at different locations in the packed column is fitted. The range of b values is only (5%. On the other hand, the values of the coefficient A can change by as much as (60%. This result may be expected since the value of A is affected by the particle wetting that changes significantly with the location in an irrigated bed. The value of b should remain constant for a given flow regime. Comparing the correlation for irrigated liquid downflow in eq 6 with the correlation for full liquid flow in eqs 4 and 5 shows that irrigated liquid downflow provides a higher mass-transfer coefficient for the same liquid flow rate. A high degree of bulk mixing occurs when the liquid streams impact the solid surfaces of the packing in irrigated liquid downflow. In full liquid upflow bulk convection occurs with much less intensity through liquid eddies and swirling flow. Also, the liquid velocity is higher with segregated flow for the same volumetric flow rate since segregated flow occupies less cross-sectional area of the packed bed. The results obtained for the overall mass-transfer coefficient compare favorably with other work that used the electrochemical method to study irrigated downflow with a whole bed section acting as a cathode. Figure 4 shows a plot of several comparable correlations found

in the literature.3-5,24-26 The differences between the correlations in Figure 4 can be attributed to several parameters. The most important factor is that different particle types and diameters were used in each study. All the plots for various reported correlations in Figure 4 were for a packed column with spheres of a diameter ranging from 2.0 to 7.8 mm. In addition, different liquid distributor designs were employed in each study. This can have a profound effect on the magnitude of the coefficient A since it is sensitive to the liquid wetting of the particle. It has also been found in the present study that the location of the active packing in the bed has an effect on the mass-transfer behavior due to local phenomena. This is discussed in more detail later. The spread of the correlations in Figure 4 is reasonable when the different experimental factors are considered. The correlation obtained in the present study is on the upper end of the data shown in Figure 4, perhaps indicating that the design of Pall rings provides superior contact efficiency over spheres and cylinders. It is relevant to note that the particle diameter used in the present study was considerably higher than the other works shown. In addition, the configuration of Pall rings is much different than that of solid spheres or cylinders. The dimensionless numbers such as the Reynolds number and the Sherwood number do not completely account for this variable, especially the shape and configuration of the particles. 3.2.2. Local Mass-Transfer Coefficient. 3.2.2.1. Axial Profiles. Values of the mass-transfer coefficient were obtained at 54 locations in the bed for two types of liquid distributors. Axial, radial, and angular profiles were constructed from these data. Figure 5 shows the axial profiles of the mass-transfer coefficient obtained using a multipoint liquid distributor (MPLD). Each point represents the arithmetic average of all 10 electrodes at a given axial position (or a given elevation in the packed bed). From Figure 5, a definite dependence on position can be observed. The initial rise in the masstransfer coefficient can be attributed to liquid spreading over the packing. Although the liquid distribution is relatively uniform as the liquid first exits the multipoint liquid distributor, it is still improved by flow over the packing down the column. Better liquid spreading to the electrode region occurs, which leads to a higher local Reynolds number and a higher effective transfer area. The mass-transfer rate and, hence, the mass-transfer coefficient increase. After the maximum mass-transfer rate is reached at x/D ) 1.5, a slow degradation of the mass-transfer performance occurs as the liquid progresses down the column. This behavior can be explained by a decreasing local velocity of the liquid as it flows down the column. A decrease in velocity would occur due to a loss of momentum in the liquid as it experiences viscous drag from flowing over the packing down the column. Due to the tortuous pattern of the opening structure of Pall rings in the packed bed, the liquid stream experienced a flow path that was much longer than the vertical displacement. The potential energy change of the flow might not be enough to make up for the viscous drag. The liquid velocity thus decreased, and hence, the measured mass-transfer coefficient decreased. Eventually, the liquid reaches a steady velocity at long bed lengths, which can be seen to be developing at x/D > 4.5 in Figure 5. At a low liquid flow rate, that is, a low Reynolds number, liquid maldistribution tends to be more severe and leads to a poor mass transfer in a packed bed.

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Figure 5. Axial profile of the mass-transfer coefficient at varied liquid Reynolds numbers using a multipoint liquid distributor.

Figure 6. Axial profile of the mass-transfer coefficient for each radial section with a multipoint liquid distributor.

Consequently, the profile of the local mass-transfer coefficient is expected to be significantly skewed. However, this is not observed in the profiles at lower Reynolds numbers presented in Figure 5. Although the variation of the mass-transfer coefficient with the axial location changes slightly with the Reynolds number, the axial profiles of the local mass-transfer coefficient are similar for different Reynolds numbers. The lowest liquid flow rate used in the present study was about 3844 lb‚h-1‚ft2. This flow rate was much higher than a critical liquid flow rate (about 1250 lb‚h-1‚ft2 27), above which liquid maldistribution became less severe, and the liquid maldistribution coefficient was leveled out; that is, no further improvement in liquid distribution would be realized with increases in the liquid flow rate. An alternative explanation for the decrease in masstransfer coefficient at long bed lengths is that the development of wall flow removes liquid from the bulk of the bed, resulting in lower flow rates in the center. The column in the present experiment was designed to test this hypothesis by allowing the mass-transfer coefficient to be measured at various radii in the bed, including the wall region at the outer electrodes as shown in Figure 2A. Figure 6 shows the axial profiles

of the mass-transfer coefficient for each of the three radial sections at a Reynolds number of 528. Although a considerable amount of mass transfer at the wall region does exist, as shown by the high mass-transfer coefficients at the outer section in Figure 6, it does not increase appreciably with axial position. It would be expected otherwise if the wall flow grew at the expense of the bulk bed. It is reasonable, then, to conclude that decreasing velocity is the dominant factor in reduction of mass-transfer coefficients in the axial direction. The outer section of the column does not display the same profile as the inner and center sections because the liquid flowing at the wall is not subject to the same viscous drag and liquid spreading conditions. The liquid at the wall flows downward through the high void section adjoining the wall and does not experience the tortuous path followed by liquid in the bulk of the bed, resulting in much less viscous friction. Also, once liquid reaches the wall, it no longer undergoes the same random-walk spreading and so it cannot spread out further. The amount of liquid wall flow reaches equilibrium where the amount of liquid joining the wall flow is equal to the amount heading back into the bulk of

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Figure 7. Axial profile of the mass-transfer coefficient for the outer electrodes at varied Reynolds numbers using a single-point liquid distributor.

Figure 8. Axial profile of the mass-transfer coefficient at varied Reynolds numbers using a single-point liquid distributor.

the bed. This behavior was shown quite clearly by Kouri and Sohlo.28 Since the arms of the multipoint liquid distributor extend to the outer section of the column, a portion of the liquid is distributed very near the wall. The wall flow in the column thus becomes pronounced almost immediately. However, the values of the mass-transfer coefficient in the three sections approach each other and one section does not dominate. One might expect the mass-transfer coefficient in the outer section to rise above the inner and center sections as the wall flow developed. However, this is not seen, indicating that the wall flow reaches a steady amount at some axial position down the column (x/D > 3.5). As shown by the profiles of the mass-transfer coefficients at the outer radius with a single-point liquid distributor (SPLD) in Figure 7, the development of the wall flow was in contrast to the flow behavior when the multipoint liquid distributor was used. For the case with the single-point liquid distributor, at an axial position of x/D ) 0.5 the mass-transfer coefficient is insignificant at the wall of the column under low Reynolds numbers, and it becomes noticeable at high Reynolds number. At

sufficient length along the bed, there appears to be a leveling off of the coefficient, again showing steady wall flow. In a study of liquid distribution in a bed of Pall rings geometrically similar to the packing used in the present study, Kouri and Sohlo28 observed a similar trend of liquid flow at the wall of the column. Specifically, they noticed that steady wall flow occurred after a bed length of 3 column diameters (x/D > 3). The axial profiles of the mass-transfer coefficients for all radii obtained using the single-point liquid distributor are shown in Figure 8. The average values shown in this figure were obtained in the same manner as in Figure 5, taking the arithmetic average of all electrodes on each layer. The mass-transfer coefficients obtained at long bed lengths for both multipoint and single-point liquid distributors match closely, indicating that steady flow is reached in the column regardless of the type of liquid feed distribution (at x/D ) 5.5). This was seen at all radial positions. Similar behavior of the local mass transfer in the bed can be seen with both the single-point and multipoint liquid distributors; however, the values of the masstransfer coefficient obtained at comparable locations are

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Figure 9. Radial profile of the mass-transfer coefficient at varied axial positions with a multipoint liquid distributor.

Figure 10. Radial profile of the mass-transfer coefficient at varied axial positions with a single-point liquid distributor.

quite different. In the case of the single-point distributor, the value of Sh/Sc0.33 at the top of the column is much lower, between 15 and 40 versus 40 and 55, and the maximum value is also lower, between 40 and 60 versus 45 and 70 for the multipoint distributor. However, the values at the bottom of the column are identical for both distributors. The deterioration of the mass-transfer conditions is slower when a single-point liquid distributor is used; however, since the multipoint liquid distributor yields higher mass-transfer coefficient in the top section of the column, it provides better overall performance. 3.2.2.2. Radial Profiles. Figure 9 shows the radial profiles of the mass-transfer coefficient obtained at several axial positions in the bed using the multipoint liquid distributor at a liquid Reynolds number of 264. The arrows in the figure indicate the positions of the nozzles on the liquid distributor. The decline of the mass-transfer coefficient as the liquid flows down the column is clearly shown in this figure. The outer section changes very little while the inner section shows a larger drop. The change in the

inner section can be attributed to loss of velocity in this section as the liquid spreads out over the packing. The central section remains quite stable. The liquid spreading conditions stabilize in this section very quickly because of its relatively small area compared to the inner section, which requires more axial distance to achieve steady conditions. A similar radial profile of the mass-transfer coefficient in an irrigated bed was reported in the literature.5 The radial profiles obtained with the single-point liquid distributor are shown in Figure 10. In contrast to the radial profiles obtained with the multipoint liquid distributor, these profiles show a much stronger radial dependence, as expected. The profile at x/D ) 5.5 is very similar to the profile at the same location obtained with the multipoint distributor, showing that a steady flow pattern develops in the bed at x/D ) 5.5 regardless of which type of liquid feed distribution is used. 3.2.2.3. Angular Profiles. Although no angular dependence of the mass-transfer coefficient is expected due to the axial symmetry of the column, these data could provide some information about the nature of the

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Figure 11. (a) Angular variation of the mass-transfer coefficient at varied Reynolds numbers. (b) Angular variation of the mass-transfer coefficient at varied axial positions.

liquid flow in a random packed bed. Figure 11a shows the angular variation of the mass-transfer coefficient with the liquid flow rate. At the two lower flow rates, the mass-transfer profiles vary similarly. When the flow rate is increased further, the shape of the profile alters. This possibly indicates the effect that increasing the liquid load has on the flow pattern inside the column. For small flow rates, the existing flow channels in the bed remain stable. When the liquid load increases, different flow channels develop to accommodate the additional liquid and consequently the same electrode behaves much differently. Figures 11b shows the angular variations at different levels in the bed for a constant flow rate. There is no discernible pattern in this figure, which is understandable. Although these electrodes are directly lined up in the column, there is 30 cm between them axially. The tortuous path of the liquid channels in the packing make it highly unlikely that two corresponding locations on adjacent layers would experience the same flow conditions. The results show that the local behavior of the masstransfer coefficient is closely related to liquid flow conditions in the column. Virtually all of the observed local mass-transfer behavior can be understood in terms

of local liquid conditions. This strong relationship is expected and the behavior observed in the present experiment match expectations. 4. Conclusions The local measurements provide valuable information regarding the nature of the liquid flow in a packed bed and its effect on the local mass-transfer coefficient. Although some of the conclusions from the data could be drawn hypothetically, this work has involved a unique quantitative measurement of local effects. (1) It was seen that regardless of which type of liquid feed distribution is used, the liquid flow evolves to the same steady state after traversing a sufficient distance down the column (i.e., at x/D ) 5.5). Also, the amount of wall flow was observed to stabilize after the same length of the packed bed (at x/D of about 3.5) for both types of liquid feed distribution. The radial profiles at the bottom of the column (x/D ) 5.5) also showed this effect. Both the profile and the magnitude of the local mass-transfer coefficient were similar. (2) The multipoint liquid distributor provided a higher overall performance than the single-point liquid dis-

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tributor, as expected. On the average, the overall masstransfer coefficient with the multipoint liquid distributor was 14% higher than that with the single-point liquid distributor. (3) At similar radial and axial positions variations of the mass-transfer coefficient with the angular position were observed. The trend of the angular variation also changed with the axial position. This angular variation illustrates that segregated channel flow is the primary form of liquid flow in the packed column. The results of this study provide a starting point in an effort to reduce the discrepancies between laboratory data and full-scale performance. The reliability of the scale-up process depends on the accuracy of the masstransfer data since the height of a transfer unit, and hence the required column height, is inversely proportional to the mass-transfer coefficient for a given column diameter and packing material. It remains to be seen how varying the dominant geometric factors (packing shape and configuration) will affect the local behavior. The effect of packing size and column diameter must be studied so that local behavior can be predicted in large-scale columns. Acknowledgment Financial support from the Natural Sciences and Engineering Research Council of Canada (NSERC) to this project is greatly appreciated. A deep gratitude also goes to Koch-Glitsch Canada, Uxbridge, Ontario, for donating Pall rings. Nomenclature a ) area of the cathode (m2) A ) coefficient of mass-transfer correlation Cb ) bulk concentration of ferricyanide (mol‚m-3) de ) equivalent diameter of the packing (m) D ) column diameter (m) D ) diffusivity of ferricyanide in solution (m2‚s-1) F ) Faradays constant (96487 C/equiv) i ) current (A) kL ) liquid-phase mass-transfer coefficient (m‚s-1) U ) superficial velocity of liquid (m‚s-1) x/D ) dimensionless axial distance from liquid distributor z ) number of electrons transferred in the electrode reaction Greek Letters υ ) kinematic viscosity of solution (m2‚s-1) Dimensionless Groups Re ) Reynolds number Sh ) Sherwood number (Sh ) (kLde/D)) Sc ) Schmidt number (Sc ) (ν/D))

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Received for review November 5, 2002 Revised manuscript received May 8, 2003 Accepted May 16, 2003 IE020881D