Zinc Chloride and Hydrochloric Acid Coextraction from Galvanizing

Feb 14, 2014 - ABSTRACT: Spent pickling acid from the industrial galvanizing process currently requires specialist disposal, yet there is the potentia...
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Zinc Chloride and Hydrochloric Acid Coextraction from Galvanizing Pickling Waste in the Presence of Iron(II). Results with Hollow Fiber Membrane Contactors Kwan H. Lum, Samuel J. Cook, Geoff W. Stevens, Jilska M. Perera, and Sandra E. Kentish* Department of Chemical & Biomolecular Engineering, The University of Melbourne, Victoria, Australia, 3010 ABSTRACT: Spent pickling acid from the industrial galvanizing process currently requires specialist disposal, yet there is the potential to recover the zinc as a valuable metal. In this work, we compare the extraction of zinc using tributyl phosphate from real galvanizing effluent and laboratory prepared solutions. Both commercial and handmade hollow fiber membrane contactors are used for this purpose. While the zinc is effectively extracted, care must be taken to prefilter the solution upstream of the membrane contactor to minimize crud formation and interfacial effects. The extraction of both zinc and HCl from the effluent is satisfactorily described using mathematical models for both the multicomponent equilibria and mass transfer balances. In development of the multicomponent equilibrium model, the thermodynamic equilibrium constants obtained from the ZnCl2− NaCl−HCl aqueous system were found to be applicable to the ZnCl2−FeCl2−HCl system.

1. INTRODUCTION Spent pickling acid from the Australian hot dip galvanizing industry is currently neutralized to precipitate a mixture of iron and zinc hydroxides, which is discarded into specialized landfill. With the increasing costs of disposal, tighter environmental legislation, and greater need for water conservation, the development of a sustainable process for the selective recovery of the metals warrants further investigation. With the effluent composed of 5 to over 200 g/L zinc, 60 to 150 g/L iron, and 10 to 80 g/L HCl,1−3 the highly contaminated liquors lend themselves to the use of solvent extraction for the recovery of these species. In particular, the zinc is of high value if it can be recovered in a suitable form. However, conventional solvent extraction equipment, such as mixer−settlers and columns, often occupies significant space, requires constant supervision, and has the potential to contaminate the solution with organics. Another option is to use hollow fiber membrane contactors (HFM). HFM contactors are well suited for use as an on-site treatment option. They have a very high surface area-to-volume ratio, leading to a reduced footprint. A number of studies have been conducted on the use of these HFM contactors; with respect to zinc extraction from spent pickling liquor, published works include both experimental and modeling results.4−9 We have also previously demonstrated successful selective zinc recovery using membrane-based solvent extraction at a bench scale level.10 However, these studies focused only on zinc extraction even though tri-n-butyl phosphate (TBP) is known to also extract hydrochloric acid (HCl).11−13 Earlier studies have shown HCl coextraction is small at such conditions,12,13 and our intent is not to recover this acid as a valuable byproduct. However, the coextracted acid should still be considered, especially when the stripped product is to be further processed. For example, total chloride levels become important if the product is to be used for zinc electrowinning, as electrodes used in electrowinning plants are susceptible to corrosion by chloride ions. © 2014 American Chemical Society

A useful summary of the mathematical modeling of HFM contactors is provided in Bringas et al.14 In terms of the extraction of zinc from chloride solutions, the HFM modeling work by Torz et al.,7,8 Samaniego et al.,15 and Polyga et al.9 relied on mass transfer coefficients fitted to experimental data. On the other hand, Bringas et al.4 derived their model from first principles, and hence no fitted parameters were required other than equilibrium constants. However, the computational approach involved solving of simultaneous partial differential equations and did not account for the baffled geometry and flow in the commercial HFM that was used. Alternatively, a simpler method using mass transfer correlations can be used,16 and this forms the basis of our present work. The use of such correlations implicitly includes the flow patterns in the HFM and can be better utilized to extend the results beyond the immediate experiments to predict future performance. Therefore, the aim of this paper is to investigate the extraction of both zinc chloride and hydrochloric acid from spent galvanizing pickling liquor by TBP using an HFM contactor and to establish a model for the coextraction of both HCl and ZnCl2 without further fitted parameters.

2. METHODS 2.1. Reagents. On the basis of our prior work and the composition of the waste solution, TBP is the most suitable extractant for the removal of zinc from Australian galvanizing effluent.1 In the present study, 50 vol % TBP (Ajax Chemicals, 99.9%) in Shellsol 2046 (Shell Chemicals, 69 wt % paraffin, 14 wt % naphthenes, naphthalenes, 17 wt % aromatics) was used. Hydrochloric acid (Sharlau, 37 wt % HCl) and ZnCl2 (Ajax Finechem Pty. Ltd., 95.0%) were used to prepare model solutions while sodium chloride (Chem Supply, 99.0%) was Received: Revised: Accepted: Published: 4453

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titration with methyl orange indicator. Chloride concentration was determined by potentiometric titration with 0.02 mol/L AgNO3 using a Metrohm Titrando 809 autotitrator with a Ag/ AgCl electrode and software package Tiamo 1.0. Iron(II) concentration was also determined using a Metrohm Titrando 809 but with a combined platinum ring electrode, titrated against 0.02 mol/L potassium dichromate. The organic phase zinc, iron, and acid concentrations were calculated from a mass balance of the aqueous phase before and after equilibration. The viscosity of the loaded organic phase was also measured using a Canon-Fenske Routine Viscometer (size 75). 2.2. Shake Tests. Shake tests were initially conducted to study the distribution of zinc and HCl between the aqueous industrial waste and organic phase (50 vol % TBP). These experiments were carried out in an orbital-shaker/incubator (Ratek OM15C) to provide consistent mixing of solutions with temperature control. The organic phase was presaturated with water before use in experiments, as TBP solutions have been reported to extract water.18−20 The organic-to-aqueous volume ratio (O:A) was varied between 1:1 and 1:6 (i.e., 100 mL organic:100−600 mL aqueous). The aqueous and organic phases were contacted in a conical flask and allowed to preheat to 30 °C in the orbital-shaker/ incubator for 0.5 h. The mixture was then shaken for 1 h at 150 rpm and allowed to stand within the shaker/incubator for a further 0.5 h to achieve complete phase separation. This time is sufficient for complete equilibration to be achieved.13 The aqueous phase was collected and analyzed while the organic phase concentrations were determined via mass balance of the aqueous phase. 2.3. HFM Contactor Setup. The experimental setup was constructed in-house (Figure 1). A commercially available hollow-fiber membrane contactor manufactured by Membrana (Liqui-cel Extra flow 2.5 in. × 8 in.) was selected as the contactor. The key specifications of this membrane contactor are provided in Table 2.16,21 The aqueous effluent was supplied through a variable speed peristaltic pump (Watson Marlow 520SN). Organic solvent was pumped in a countercurrent direction using a variable speed magnetic gear pump (Model 1805/IEC71, Micropump, Vancouver, WA). The aqueous and organic phase piping was constructed from PTFE and 316stainless steel, respectively. The pressure of the aqueous phase was adjusted using a PFA needle valve and measured using 0− 100 kPa bourdon tube pressure gauges which were isolated from the process fluid using Plast-O-Matic gauge guards with a PTFE diaphragm (PI-3 and PI-4 in Figure 1). The organic phase pressure on either side of the contactor was measured

added to adjust total chloride concentration to 5 mol/L. Sodium hydroxide (Chem Supply, 97.0%), disodium ethylenediaminetetraacetic acid (EDTA) (Chem Supply, 99.0%), sodium acetate (Chem Supply, 99.0%), glacial acetic acid (Chem Supply, 99.7%), silver nitrate (Merck, 99.5%), and potassium dichromate (Ajax Finechem, 99.8%) were used for sample analysis via titration. The industrial waste was obtained from the stripping acid bath in the Industrial Galvanizers Australia (Melbourne) plant as required. This bath is a secondary acid bath and is used predominantly to strip zinc layers from ineffectively galvanized steel and the various hooks and jigs used to carry the steel items throughout the process. The concentrations of zinc and iron are significantly higher than those in the primary pickling bath; conversely, the hydrochloric acid concentration is lower due to its consumption during the stripping process. The concentrations of components within this bath varied from day to day, but the average composition is shown in Table 1. Table 1. Average Spent Pickling Liquor Composition

a

component

concentration (g/L)

hydrochloric acid zinc iron(II) iron(III) chloride chromium, cadmium, lead, copper, and nickel grease, bismuth

36a 134 74 5 MΩ·cm). After back-extraction experiments, the organic phase was regenerated by contact with excess water. A bench-scale HFM was also used, with the experimental setup as described in earlier work.10 This setup used a HFM constructed by hand using polypropylene fibers (Memtec, Australia). The specifications of this handmade unit are also included in Table 2. In preliminary experiments, droplets of the aqueous phase were observed in the organic outlet stream when the contactor was operated at transmembrane pressures above 30 kPa. At operating pressures below 15 kPa, the organic phase was observed in the aqueous phase. An optimal differential pressure of 18 kPa was thus utilized throughout further experiments.

Figure 2. Comparison of zinc extraction modeling with experimental results of zinc extraction from spent pickling liquor.

Four modeling options for the mixing parameters were assessed (Figure 2). Model 1 calculated the zinc extraction assuming that no Fe(II) was present, but this approach underestimated zinc extraction, with 15.3% mean absolute percent error with respective to percent zinc extraction (MAPE). Model 2 accounted for the presence of Fe(II) using the available single salt parameters for FeCl2 and H−Fe(II)−Cl parameters, but no mixing parameters for the Fe(II)−Zn−Cl system were used. In this case, zinc extraction was overestimated (21.6% MAPE). Model 3 substituted the concentration of Fe(II) with sodium, as the Na−Zn−Cl mixing parameter is available in the literature (Table 3). However, to keep the initial total chloride concentration equal to that of Fe(II), the sodium concentration used was double that of Fe(II). This gave the worst fit (32.7% MAPE). The last case, Model 4, proved the best approach with 5.6% MAPE. In this case, the mixing parameters of Na−Zn−Cl were used in lieu of those for Fe(II)−Zn−Cl. The results here indicate that the mixing parameters cannot be ignored. Also, the effect of Fe(II) on zinc extraction is different from that of sodium (as indicated by model 3). While it is possible to determine the Fe(II)−Zn−Cl mixing parameters from this experimental data, a more rigorous process and larger database of experimental data is required for these to be accurate. Therefore, in the present work, the Fe(II)−Zn−Cl mixing parameters are not fitted, but Na−Zn− Cl mixing parameters are used instead. Also included in this analysis are the literature data of Mansur et al.23 who also studied zinc extraction from spent pickling liquor at two different TBP concentrations and varying A:O ratio; the initial metal and acid concentrations for Mansur et al.23 are listed in Table 4.

3. RESULTS 3.1. Modeling of the HCl−FeCl2−ZnCl2−TBP−Diluent System. In our previous work, we developed a fundamental model for the liquid−liquid multicomponent equilibria for ZnCl2 and HCl coextraction by TBP using the Pitzer model for the aqueous phase and the Hildebrand−Scott model for the organic phase12 (see Appendix for a summary of the key equations). However, the earlier work was modeled based on the HCl−NaCl−ZnCl2 aqueous system. As the aqueous system used here contains Fe(II) instead of NaCl, the applicability of the model for the HCl−FeCl2−ZnCl2 system needs to be examined. The Pitzer parameters for FeCl2 are available from the literature (see Table 3) but the Fe(II)−Zn−Cl mixing parameters are not. Shake tests using the industrial effluent are used to estimate these parameters. Other parameters follow those used in our earlier work.12,22 4455

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Table 4. Industrial Waste Composition for Shake-Test Experiments reference

Zn (g/L)

Fe(II) (g/L)

HCl (mol/L)

Cl (mol/L)

TBP (vol %)

Mansur et al.23a this workb

70.2 135.8

92.2 91.8

0.25 0.5

5.7 7.9

100, 50 50

a

Organic diluent used: Exxsol D-80 (aliphatic kerosene). bOrganic diluent used: ShellSol 2046 (aliphatic kerosene).

Figure 4. Zinc and iron coextraction from industrial waste using a HFM contactor and 50 vol % TPB in full recycle mode. Operating conditions are as follows: Vaq = 1 L, Vorg = 10 L, Qaq = 48.5 mL/min, Qorg = 213 mL/min, T = 32 °C.

Figure 3. Zinc extraction from spent pickling liquor using TBP as a function of the organic to aqueous volume ratio. Symbols represent experimental data both from this work and Mansur et al.;28 the lines represent the zinc extraction predicted by the model with Na−Zn−Cl mixing parameters used as Fe(II)−Zn−Cl.

The zinc extraction predicted using model 4 is shown in Figure 3 for experimental data of both this work and Mansur et al.;23 Exxsol-D80, the diluent used by Mansur et al.,23 was modeled as ShellSol 2046. The figure shows that the model is able to predict zinc extraction not only with different initial zinc content and A:O volume ratios but also at varying TBP concentration. 3.2. HFM Experiments. 3.2.1. On-Site Trials. In initial onsite trials, the spent pickling was passed through an inline strainer to remove larger particulate matter and an inline 5 μm polypropylene gradient filter prior to membrane extraction. In these trials, the focus was to determine the amount of Fe(II) coextracted. In the first trial, 1 L of aqueous industrial waste and 10 L of organic phase were both circulated in recycle mode for 2 h of extraction. During the extraction step (Figure 4), iron coextraction was found to be negligible while zinc concentration dropped to 54 g/L at the end of 2 h (an extraction yield of 53%). These results are comparable to similar work by Ortiz et al.24 who used an identical contactor and recorded extraction yields of 50−70% from similar solutions in 1 h in a recycle mode. The waste was then replaced with fresh water as a stripping solution and stripping of the organic phase conducted in recycle mode for a further 2 h. After stripping (Figure 5), less than 0.3 g/L of Fe(II) was detected in the stripping solution and the final zinc concentration reached 11.5 g/L. Because the Fe(II) concentration in the stripping solution remained relatively constant throughout the stripping process, the iron content in this stripping solution is likely to be residual iron remaining within the rig from the extraction process (after flushing of the rig). During this trial, a loss of aqueous phase of 8 mL/h was observed due to breakthrough into the organic phase. Despite the relatively small quantity, there is a potential of aqueous

Figure 5. Stripping of loaded TBP with service water, following extraction from industrial waste using a HFM contactor and 50 vol % TPB in full recycle mode. Operating conditions are as follows: Vaq = 1 L, Vorg = 10 L, Qaq = 48.5 mL/min, Qorg = 213 mL/min, T = 32 °C.

phase build-up in the solvent loop over time if the transmembrane pressure is not strictly controlled. In a second trial, the extent of potential fouling was investigated by passing over 100 L of industrial waste through the tube-side of the membrane contactor in a once-through mode. The organic phase was operated in a full recycle mode and periodically regenerated by direct contact with service water. The majority of the suspended solids in the industrial effluent were removed via the upstream strainer and filter. However, there was still visible evidence of membrane fouling observed through the transparent wall of the membrane contactor. During this long-term operation, it became increasingly difficult to control the loss of aqueous phase to the organic phase by controlling transmembrane pressure. These losses increased beyond those noted in the full recycle mode. This is consistent with the membrane becoming hydrophilic due to extended contact with surfactants in the effluent. It is thus recommended that a more hydrophobic membrane material or one with a smaller pore size be investigated to increase resistance to aqueous phase breakthrough. At the end of the extended trial, a dark viscous deposit containing suspended particles (crud) was observed at the 4456

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increase in organic flow rate causes a reduction in zinc stripping due to the lower organic phase residence time in the contactor. With respect to HCl coextraction (Figure 7), an average of 6% of HClaq was extracted from the model solution and 26% of HCl in the loaded TBP was stripped. Comparing HCl extraction to that of zinc, zinc has a higher extraction percentage which is to be expected. However, the stripping experiments showed that HCl is easily stripped; despite the low percentage of acid extraction, the percentage of stripped acid is comparable to that of zinc.

surface of the final TBP solution. Analysis of the deposit after filtration showed that it contained significantly more zinc and iron than the bulk organic phase, indicating that this phase may contain organic-insoluble metal species. Thus, while the interfacial deposit appeared substantially organic in nature, it contained higher concentrations of both zinc and iron. While it did not appear to disrupt the extraction during the relatively short pilot scale study, it may become a concern in a full scale process. More work is required to further identify the nature of the deposit and to observe the impact it has on extraction during more extended trials. 3.2.2. Laboratory-Based Trials. Following these onsite trials, the rig was returned to the laboratory, but work continued with both industrial waste and laboratory-prepared model solutions (Table 5). A fresh commercial HFM contactor (Liqui-Cel 2.5 Table 5. Initial Concentrations for Laboratory-Based HFM Experiments experiment industrial waste (extraction) model solution (extraction) loaded TBP (Stripping) a

Zn (g/L)

Fe (g/L)

HCl (mol/L)

Cl (mol/L)

133

59.8

1.0

7.2

82 5.5

0 0

1.0 0.025

5a 0.19

Chloride concentration adjusted by addition of NaCl. Figure 7. Outlet aqueous HCl concentration following extraction and stripping using a Liqui-Cel 2.5 in. × 8 in. contactor with 50 vol % TBP in ShellSol 2046. Symbols represent experimental data, while the lines are the modeling results.

in. × 8 in.) was used. To reduce crud formation, three prefiltration units were used in place of two. First, large particulates were removed using polypropylene filter bags (Cole-Parmer, 5 μm pore size). A second unit with an oiladsorbing cartridge (Pentek, OAC-20BB Oil Adsorbing Cartridge) was employed to remove organic material (grease and surfactants). Finally a hollow fiber cartridge filter (Minntech, 0.05 μm pore size) removed the smallest particles. In these experiments, both aqueous and organic streams were passed through the HFM contactor in once-through mode. The average zinc extraction achieved is 16% for industrial waste and 18% for the model solution (Figure 6). As expected, increasing the organic flow rate increases zinc extraction. For the stripping experiments, an average of 35% of zinc in the loaded organic phase was stripped (Figure 6). In this case, an

Figure 8. Modeling of zinc extraction by 50 vol % TBP in ShellSol 2046 using a handmade HFM contactor. Aqueous inlet: 2−19 g/L Zn, 1 mol/L HCl, 5 mol/L Cl−; Qorg = 2 mL/min.

The effect of increasing organic flow rate on HCl extraction and stripping shows a similar trend to zinc: increasing organic flow rate increases HCl extraction but causes a decrease in acid stripping. The degree of increase in HCl extraction is significantly less than zinc due to preferential extraction of the metal. 3.2.3. Hollow Fiber Membrane Contactor Modeling. The data presented in Figures 6−8 were modeled based on a standard model for hollow fiber contactors (see Appendix). The viscosity of the organic phase was measured at varying degrees of organic loading, and these values were used in the calculation of the relevant diffusion coefficients (see Appendix).

Figure 6. Outlet aqueous zinc concentration following zinc extraction and stripping using a Liqui-Cel 2.5 in. × 8 in. contactor with 50 vol % TBP in ShellSol 2046. Symbols represent experimental data, while the lines are the modeling results. 4457

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model of the ZnCl2−NaCl−HCl aqueous system. In particular, it was shown that this model is applicable to ZnCl2−FeCl2− HCl aqueous systems if the Zn−Na−Cl mixing parameters are used as Zn−Fe−Cl parameters. The model was able to satisfactorily fit the data for both the extraction of zinc and acid in both commercial and handmade contactors without the use of any additional fitting parameters. The ratio of the residence times of aqueous and organic phase as the A:O ratio was found to produce the best predictions.

Density was calculated following the method outlined in our earlier work.12 The pores were assumed to be filled by the organic phase because the membrane is hydrophobic. Based on the inlet concentration of the aqueous and organic stream, the equilibrium concentration and subsequently the distribution ratios (m) can be predicted (see Appendix). However, an appropriate A:O is required to determine the outlet concentrations. The A:O can be obtained either from the ratio of the total liquid volume on the tube-side (0.15 L) to the shell-side (0.4 L), the respective flow rates, cross-sectional area, or residence time. We found that the ratio of aqueous to organic phase residence time gave the best fit to the data and was used, as it accounts for both the effect of phase volume in the HFM and also the flow rate. The option of dividing the HFM into a number of segments to accommodate changes in various properties along the contactor length, such as density, viscosity, and distribution ratio, was also examined. However, it was found that the number of segments used only affected the fitting slightly. One of the reasons for this is because the length of the HFM is short such that the changes in properties along the contactor are small. Furthermore, because the flow and the geometry in the shell-side HFM is complex, especially with the baffle and central distributor in the Liqui-Cel module, the mass transfer coefficient will vary along the contactor. This variation is implicitly included in the model. However, dividing the contactor without explicitly accounting for these differences is partly the reason why the use of larger numbers of segments does not affect the fitting significantly. Hence, a single segment was used. The model fits well to the data for the extraction of zinc from both the model solution and the industrial waste feed (Figure 6). However, the stripping of zinc is less well predicted. This is likely due to the low concentration of zinc in both the organic and aqueous phase, making the experimental data susceptible to analytical errors especially when the organic phase content is determined via mass balance. The acid extraction/stripping modeling also showed errors greater than that of zinc modeling (Figure 7). These errors are likely to be due to analytical error related to the small concentrations of HCl being extracted and stripped. The model was also tested on data from the handmade HFM module, as Shen et al.16 indicated that the shell-side mass transfer coefficient model used here (see Appendix) is applicable to both laboratory-made and commercial modules. It is also useful to observe whether the model is appropriate at lower initial zinc concentrations. Figure 8 shows that the model provides a good fit to the experimental data at 2 and 10 g/L initial Znaq. The data fit is not as good at 19 g/L initial Znaq, possibly reflecting experimental error.

■

APPENDIX: MODELING OF THE HOLLOW FIBER MEMBRANE CONTACTOR

A.1. Membrane Mass Transfer Coefficient

The membrane is assumed to be completely filled with either the organic phase for a hydrophobic membrane or the aqueous phase for a hydrophilic membrane. The equation used in determining the membrane mass transfer coefficient is given in eq 1. k m,i =

2εDi τ(d t,out − d t,in)

(1)

where km,i is the membrane mass transfer coefficient, ε the porosity of the membrane, Di the diffusivity of species i, τ the tortuosity of the membrane pore, dt,out the tube outer diameter, and dt,in the tube inner diameter. The diffusivity of solutes, Di, can be estimated using the Wilke−Chang equation:25 Di = 7.4 × 10−8 ×

(xM w,i)0.5 T (μ × 103)Vi 0.6

(2)

where x is 1 for nonassociated solvents and 2.6 for water, Mw,i the molecular weight of species i (g/mol), T the temperature (K), μ the dynamic viscosity (Pa·s), and Vi the molar volume of species i at normal boiling point (cm3/mol). A.2. Tube-Side Mass Transfer Coefficient

The Lévêque correlation has been used in calculating the tubeside mass transfer coefficient.16,26 Depending on the Graetz number, Gz, the tube-side mass transfer coefficient can be calculated using eqs 3 or 4. Sht,i =

Sht,i =

k t,id t,in Di k t,id t,in Di

1/3 ⎛ d t,in ⎞ = 1.62⎜Sc iRe ⎟ ; L ⎠ ⎝

⎛ d t,in ⎞ = 0.5⎜Sc iRe ⎟; L ⎠ ⎝

Gz > 6 (3)

Gz < 6 (4)

where Sht,i is the tube-side Sherwood number of species i, kt,i the tube-side mass transfer coefficient of species i, dt,in the tube inner diameter, Di the diffusivity of species i, Sci the Schmidt number of species i, Re the Reynold number, L the length of tube, and Gz the Graetz number. For the tube-side, the diameter used in the calculation of Re is the tube inner diameter, din, giving the cross-sectional area as: π 2 AX,t = N × (d t,in ) (5) 4

4. CONCLUSIONS Onsite experiments with HFM contactors for zinc extraction from industrial pickling effluent showed the importance of proper filtering of the industrial waste to avoid third phase formation and stable operation of the HFM contactor. Careful transmembrane pressure control is also necessary to eliminate breakthrough of one phase into another. It is unclear how this control could be maintained in a larger module where significant cross-membrane pressure drop also occurs; therefore, this issue requires further investigation. A model for the coextraction of zinc and HCl from this industrial pickling effluent was developed based on a prior

where AX,t is the tube-side cross-sectional area, and N the number of tubes. The equations for Reynold and Schmidt number calculation are given in eqs 6 and 7, respectively. 4458

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Table A1. Equilibrium Constants for the Overall Reaction Scheme equation

extracted species

log K°

solubility parameter

references

12

HCl(H 2O)7 ·3TBP ZnCl2(H2O)3·3TBP HZnCl3(H2O)3·3TBP

−1.176

18.70

Lum et al. (2012)22

1.76 1.10

17.85 18.73

Lum et al. (2013)12 Lum et al. (2013)12

13 14

Table A2. Single Salt Pitzer Parameters

a

salt

β(0)

β(1)

Cϕ

max. I (molal)

references

HCl FeCl2 ZnCl2a

0.21049 0.35011 0.041312

0.06035 1.40092 1.80161

−0.00419 −0.01412 0.010082

16 2 23

Lum et al. (2012)22 Kim and Frederick27 Anstiss and Pitzer28

α1 = 1.53469; ω(Zn, 3Cl) = −1.77325 × 10−4; ω(Zn, 4Cl) = 1.30677 × 10−6.

Table A3. Pitzer Mixing Parameters θij

Ψijk

max. I (molal)

references

0.0368 0.00 0.0533 −0.366

−0.0033 0.012 −0.1728 −0.0082

3 3 2 18

Kim and Frederick29 Marion et al.30 Tialowska-Morchala et al.31 Anstiss and Pitzer28

S

mixed electrolyte H−Na−Cl H−Fe(II)−Cl H−Zn−Cl Na−Zn−Cla a

Na−Zn−Cl values used as Fe(II)−Zn−Cl mixing parameters, as no literature values were available.

Re =

Qd υAX

(6)

Sc i =

υ Di

(7)

A.4. Overall Mass Transfer Coefficient

The overall mass transfer coefficient can be calculated by the summation of the inverse of the individual mass transfer coefficients through the boundaries (tube-side, membrane, and shell-side) as shown in eq 11.16,26 The equation takes into account the distribution of solutes between the two phases using the distribution ratio, mi. For this modeling, the solutes of interest are zinc and HCl.

where Re is the Reynold number, Q the volumetric flow rate (m3/s), d the diameter (m), υ is the kinematic viscosity (m2/s), AX the cross-sectional area (m2), and Sci the Schmidt number of species i.

d t,in d t,in 1 1 = + + K t,i k t,i m i k m,idlm m i ks,id t,out

A.3. Shell-Side Mass Transfer Coefficient

The model proposed by Shen et al.16 is used for the shell side coefficient(eq 8). Shen et al.16 has shown that this model can accommodate both laboratory-made (parallel flow) and small scale commercial HFM contactors (crossflow). Shs,i =

ks,id t,out Di

= 0.055Re 0.72Sc i0.33

where Kt,i is the overall mass transfer coefficient of species i for mass transfer from tube-side to shell-side, mi the distribution ratio of species i, and dlm the log mean tube diameter. The distribution ratio used in eq 1112,22 (either zinc or acid) is defined as the ratio of the total concentration of either of these species in the organic phase over that of the aqueous phase. Earlier modeling work by the present authors has shown that the main complexes present in the organic phase are ZnCl 2(H 2O)3 ·3TBP, HCl(H 2O)7 ·3TBP and the acido-metal complex HZnCl3(H 2O)3 ·3TBP.12 The activity coefficients for these complexes are determined from the Hildebrand−Scott model for the organic phase12,22 while the activity coefficients in the aqueous phase utilize the Pitzer model. The relevant equations, equilibrium constants, and modeling parameters are summarized in eqs 12−14 and Tables A1−A4.

(8)

where Shs,i is the shell-side Sherwood number of species i, ks,i the shell-side mass transfer coefficient of species i, Re the Reynold number, and Sci the Schmidt number of species i. Due to the presence of the central distributor tube and the hollow fiber tubes, the shell-side hydraulic diameter is used instead in the calculation of Re.16 Equations 9 and 10 are used to calculate the shell-side hydraulic diameter and the crosssectional area, respectively. ds,hyd =

AX,s =

2 2 ds,in − dc2 − Nd t,out

Nd t,out

π 2 2 (ds,in − dc2 − Nd t,out ) 4

(11)

H+ + Cl− + 4H 2O + 3H 2O·TBP ⇌ HCl(H 2O)7 ·3TBP (12)

(9)

Zn 2 + + 2Cl− + 3H 2O·TBP ⇌ ZnCl 2(H 2O)3 ·3TBP

(10)

(13)

where ds,hyd is the shell-side hydraulic diameter, ds,in the shellside inside diameter, dc the central distributor diameter, dt,out the tube-side outside diameter, and AX,s the shell-side crosssectional area.

H+ + Zn 2 + + 3Cl− + 3H 2O·TBP ⇌ HZnCl3(H 2O)3 ·3TBP 4459

(14)

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⎛ ∑ V iφc i ⎞ ⎟+ ρsoln = ρsolv ⎜1 − i 1000 ⎠ ⎝

Table A4. Organic Species Parameters Used in the Hildebrand−Scott Model organic species

molecular weight, Mw (g/mol)

molar volume, Vi,org (cm3/mol)

solubility parameter, δi (MPa0.5)

density, ρ (g/cm3)

266.318 152.0d

273.85 188.59d

18b 15.95

0.9725 0.806

136.282 36.458 18.015

60.54e 25.025f 16.259f

NA NA NA

NA NA NA

TBPa ShellSol 2046c ZnCl2 HCl H2O

∑ M r,ic i i

(20)

where ρsoln is the density of solution (g/L), ρsolv the density of solvent (water or ShellSol 2046), Vφi the apparent molar volume of solute i (cm3/mol), ci the concentration of species i (mol/L), and Mr,i the molecular weight of species i (g/mol)

■

AUTHOR INFORMATION

Corresponding Author

*Tel: +61 3 8344 6682. Fax: +61 3 8344 4153. E-mail: [email protected].

a Schulz et al.32 bBrandrup et al.33 cShell.34 dLyford.35 eLum et al. (2013).12 fLum et al. (2012).22

Notes

The authors declare no competing financial interest.

A.5. Outlet Concentrations of HFM Contactors

■

To predict the outlet concentrations of solutes from HFM contactors, the analytical solutions by Prasad and Sirkar36 were used (eqs 15−17). These equations are written with the aqueous phase running in the tube-side and the organic in the shell-side of the HFM contactor. For a detailed derivation of these equations, refer to Prasad and Sirkar.36 out Caq,i =

ri =

in Caq,i (1 − ri) −

in Corg,i

mi

ACKNOWLEDGMENTS The authors acknowledge the Australian Research Council and Industrial Galvanisers (Australia) for financial support. We also thank the Particulate Fluid Processing Centre (PFPC), an ARC special research center, for infrastructure support. The technical support given by Dr. Peter Hetherington of Industrial Galvanizers (Australia) is also greatly appreciated. We are also grateful for the assistance of Xiaotian Edward Li in the laboratory.

(1 − e i)

e i − ri

(15)

■

Q aq m i Q org

(16)

⎛ πK aq,id t,inN ⎞ ⎟ e i = exp⎜⎜L(1 − ri) × ⎟ Q aq ⎠ ⎝

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(17)

in where Cout aq,i and Caq,i are the aqueous phase concentration of species i in the outlet and inlet, respectively, Cinorg,i is the inlet organic phase concentration of species i, Qaq and Qorg are the aqueous and organic phase flow rate, and N is the number of hollow fibers in the HFM. The outlet organic phase concentrations can be determined via mass balance:

out in in out ) Corg,i = Corg,i + (Caq,i − Caq,i

Q aq Q org

REFERENCES

(18)

From these sets of equations (eqs 1−18), no parameters were fitted to enhance the prediction of the HFM model. A.6. Physical Property Prediction

To assist with this modeling, the viscosity of the organic phase was first measured at varying degrees of organic loading. Equation 19 describes the change in organic viscosity with increasing organic phase zinc concentration up to 17 g/L. μorg = 8.61 × 10−5 × (65.384 × [Zn]) × 2.74 × 10−3 (19)

where μorg is the organic phase dynamic viscosity of loaded 50 vol % TBP in ShellSol 2046 (Pa·s) and (Zn) the concentration of organic zinc (mol/L). This viscosity was used in calculations of the relevant diffusion coefficients. The density of the aqueous and organic phase was calculated using eq 20; this equation follows the method outlined in our earlier work.12 For the aqueous phase, the apparent molar volume was estimated with the Masson equation37 while a constant is used for the organic phase. 4460

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