Dispersion-Free Solvent Extraction and Stripping of Gold Cyanide with

Ind. Eng. Chem. Res. , 2002, 41 (3), pp 613–623. DOI: 10.1021/ie010141p. Publication Date (Web): January 11, 2002. Copyright © 2002 American Chemic...
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Ind. Eng. Chem. Res. 2002, 41, 613-623

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Dispersion-Free Solvent Extraction and Stripping of Gold Cyanide with LIX79 Using Hollow Fiber Contactors: Optimization and Modeling Anil Kumar,† R. Haddad, G. Benzal, and A. M. Sastre* Chemical Engineering Department, Universitat Polite` cnica de Catalunya, ETSEIB, Av. Diagonal 647, E 8028 Barcelona, Spain

This paper reports experimental data on hollow fiber dispersion-free membrane extraction (DFSX) of gold(I) from alkaline cyanide media using microporous hydrophobic polypropylene hollow fiber contactor. A recently developed new solvent extraction reagent, namely, LIX79 (Henkel Corporation), was employed as an extractant. The DFSX operation was carried out with 12% LIX79 in n-heptane by passing alkaline feed containing gold through the tube side and organic extractant through the shell side. Extraction studies were performed under different hydrodynamic conditions and mass correlations for the tube and shell side were presented for cocurrent and countercurrent flow conditions. A model is presented that identifies the different ratecontrolling steps under various chemical and hydrodynamic experimental conditions in feed, organic, and membrane phases. The mass-transfer resistance resulting from a chemical extraction reaction at the surface was observed to be dominant during the entire extraction process. The validity of this model was evaluated with experimental data and found to tie in well with theoretical values. It was possible to separate and concentrate Au(I) (up to 400 mg/L) in the presence of other metal cyanide salts such as Fe(II), Cu(I), Ni(II), Ag(I), and Zn(II) (NaCN ) 1000 mg/L) by employing this technique. The stripping of gold from LIX79 was accomplished efficiently by using NaOH, which flowed through the tube side (11.11 cm3/s) and the organic was passed through the shell side at (6.94 cm3/s). Introduction The commonly used pack towers, mixer settlers, extraction columns, and so forth are employed to carry out traditional liquid-liquid extraction processes, which afford high interfacial area for contacting and results in high mass transfer.1 The intimate mixing that often occurs in these devices can lead to the formation of stable emulsions of the two phases, thereby inhibiting phase separation and product recovery. In addition, the flooding and loading limits in continuous countercurrent devices, the need for density difference between the phases, and the high initial, operating, and maintenance costs of centrifugal devices are the major disadvantages encountered during this process. Furthermore, scaleup is always difficult. In view of this, recently developed membrane extraction processes using microporous hollow fibers are of particular interest because of their versatility and the fact that they overcome problems encountered in conventional liquid-liquid extraction. Dispersion-free membrane extraction (DFSX) is simply a liquid-liquid extraction in a hollow fiber contactor in which aqueous and organic streams flow through the capillary shell side and contact each other in the pore mouth without dispersion, which minimizes the possibility of forming emulsion/third phase or crud formation with extractant. * To whom correspondence should be addressed. Fax: (34)93-4016600. E-mail: [email protected]. † On leave from PREFRE, Nuclear Recycle Group, Bhabha Atomic Research Centre,Tarapur 401502, India. Fax: +91252-572866. E-mail: [email protected].

Hollow fiber modules may be connected in series or in parallel, and the length and diameters of fibers and modules can be varied to provide the required interfacial area. Such membrane processes not only remove the required compounds from the streams but also can concentrate these species simultaneously on the product side for possible further processing.2,3 DFSX is also characterized by being rapid in separation, high in efficiency, low in power consumption, and adaptable to diverse uses in addition to being suitable for industrial applications due to a large area-to-volume ratio that can approach a value of 104 m2/m3.4,5 Furthermore, DFSX techniques have been extensively deployed in separation science applications such as metal recovery from leaching and wastewaters, extraction of precious and strategic metals from neutral waters, and treatment of large volumes of the effluents including toxic and hazardous wastes generated by industries.6-12 Reed et al.13 have successfully carried out pilot-scale evaluation of microporous membrane-based solvent extraction of a wide range of organic contaminants from industrial wastewater at two industrial plants sites in The Netherlands. The process economics appeared quite competitive with conventional technologies. Kathios et al.14 have demonstrated the utility of the membranebased solvent extraction modules for the extraction of actinides. The performance of this technique was stated to be promising. In the late 1980s, DFSX of Au(III) from chloride media was reported by Alexander and Callahan15 using a hollow fiber contactor using diethylene glycol dibutyl ether (DGDE) as an extractant. The stripping of gold from the loaded organic phase was reported to be poor under studied experimental condi-

10.1021/ie010141p CCC: $22.00 © 2002 American Chemical Society Published on Web 01/11/2002

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tions as the mass-transfer coefficient for stripping was reported to be 2 orders of magnitude lower than that of extraction. Recently, the DFSX technique for real hydrometallurgical solutions was accomplished using LIX860 (5-dodecylsalicyclaldoxime) by Valenzuela et al.16-18 for copper recovery from acid leach residual solutions and water from the El Teniente mine in Chile. To the best of our knowledge, very few studies have been made of conventional liquid-liquid extraction to recover gold from cyanide media. In the early 1980s, Mooiman and co-workers19-21 performed liquid-liquid extraction of Au(I) from cyanide media with amines and modified amines. To develop a complete pilot plant campaign, Riveros22 performed gold recovery from real cyanide solutions using commercial quaternary amines (Aliquat-336 dissolved in Solvesso-150). This process seems complicated for further applicability in industry because acidic thiourea was recommended as the stripping agent. Moreover, another alternative for stripping was recommended as the incineration of gold-loaded quaternary amine salt, which again, on a plant scale, could be more expensive as it would result in a demand for more organic extractant after one cycle of operation. Kordosky et al.23 developed a conventional solvent extraction (SX) process for gold(I) recovery from cyanide media using SX circuits with reagent N,N′-bis(2-ethylhexyl)guanidine and N,N′-bis(tridecylguanidine) (similar to LIX79 containing guanidine functionality). The selectivity of gold(I) against copper and iron was stated to be good. Stripping was also very efficient with dilute NaOH (0.25 M) and (0.1 M) NaCN. Despite good performance, third-phase and crud formation was reported by the authors. Similarly, in further studies Virnig24 performed SX studies with a real mine solution of Au(I) from cyanide media using LIX79, which afforded good selectivity and satisfactory performance when a circuit configuration using three stages of extraction and two stages of stripping was employed. Third-phase formation was reported by the authors and attempts were made to eliminate this problem by adding 50 g/L of tridecanol. This problem can be more severe when plant-scale experiments are performed. Hence, to eliminate these problems, DFSX could be more appropriate for obtaining clean extraction-separation of Au(I) from alkaline cyanide media. The DFSX studies of Au(I) from cyanide media has not been attempted so far. On the basis of the previous studies, LIX79 containing a guanidine functionality (Henkel Co.) was chosen as the extractant, which has a good extractability at pH > 9 and stripping of Au(I) in highly basic (pH > 12) solutions.23,25,26 In previous studies, the hollow fiber supported liquid membrane configuration was checked with LIX79-Au(I) from cyanide media.12 Although the results were consistent with respect to the operating conditions and stability of the system, high flow rates (to obtain high mass transfer) could not be used because of the impregnation mode, which could force out the organic extractor from the pore. Therefore, to obtain a high mass-transfer coefficient, the dispersion-free hollow fiber liquid-liquid extraction mode was selected. The aim of this paper is to evaluate design parameters to optimize performance of hollow fiber modules for the DFSX of Au(I) from alkaline cyanide media. The overall mass-transfer coefficient (KEAu) for extraction and stripping (KSAu) of Au(I) was calculated. The recovery of

gold(I) from synthetic hydrometallurgical solutions was studied to optimize the conditions for selective separation of gold from other base metals under the same experimental conditions. Experimental Procedure Reagents. A stock solution of Au(I) (5 g/L) was prepared from pure solid KAu(CN)2, and 1 g/L each of cyanide salts such as KAg(CN)2, Zn(CN)2, KNi(CN)4, CuCN [Johnson Mattey Chemicals (Karlsruhe, Germany)], and K4Fe(CN)6‚3H2O (Merck, Darmstadt, Germany) were dissolved in NaCN (Merck). KNi(CN)4 and CuCN salts were dissolved in excess NaCN in deionized water. The organic solvent used in DFSX studies was n-heptane, which is a commercially available solvent. All chemicals were used as received. LIX79 (N,N′-bis(2-ethylhexyl)guanidine) was kindly provided by Henkel Co. The HFM is manufactured by Hoechst Celanese, Charlotte, NC (Liqui-Cel, 8 × 28 cm 5PCG-259 contactor and 5 PCS-1002 Liqui-Cel laboratory LLE) and specified as below. type of module number of fibers module diameters (cm) module length (cm) active interfacial area (m2)

5PCG-259 (contactor) 10 000 8 28 1.4

Hollow Fiber Membrane Details fiber i.d. (cm) 24.0 × 10-3 fiber o.d. (cm) 30.0 × 10-3 fiber wall thickness (cm) 3.0 × 10-3 fiber length (cm) 15 porosity (%) 30 pore size (µm) 0.03 polymeric material polypropylene hydraulic diameter (cm) 0.1 area per unit volume (cm2/cm3) 29.3

Partition Coefficients of Au(I). Equal volumes of 10 mL each containing Au(I) at pH 10.3-10.5 (pH was adjusted by small additions of NaOH) and an organic extractor LIX79 (12% v/v) dissolved in n-heptane were pipetted into a 30-mL glass-stoppered equilibration vial and stirred mechanically for 30 min at room temperature (25 ( 1 °C). Equilibrium was generally reached in 20 min with the extractor LIX79.25,27 After phase separation, the equilibrium pH was measured and the metal concentration was determined by atomic absorption spectrometry techniques.25,27 The partition coefficient (H) of Au(I) for extraction and stripping, defined as the ratio of concentration of the ion in the organic phase or strip phase at equilibrium and in the aqueous phase, respectively, was calculated. For the study of back extraction of Au(I), 5 mL of loaded organic extractor was drawn into an equilibration tube and backextracted for about 30 min with the same volume of the strippant (0.5 M NaOH). The Au(I) content after the stripping of the metal ion was analyzed. All the measurements were performed at least in duplicate and the agreement of the H values obtained was within (2%, with a good mass balance (>95%). Density and Viscosity Measurements. The density and viscosity of the LIX79/n-hepane solutions were measured to relate the Au(I) extraction with their physical properties in the DFSX process. The viscosity was measured at 25 ( 1 °C using a Cannon-Ubelhode semimicro 100 viscosimeter (Cannon Instruments Co.,

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Figure 1. Schematic view of the membrane-based extraction process of Au(I) from cyanide media using a hollow fiber contactor: (1) hollow fiber contactor; (2,3) organic extractor and feed; (4) feed and organic pump; (5,6) inlet and outlet pressure gauge respectively for organic and feed; (7) flowmeters for feed and organic.

USA), and densities were determined by weighing a known volume of solution using a pycnometer. The Nondispersive Membrane Extraction Setup. A schematic view of the membrane-based extraction process of Au(I) using a hollow fiber contactor in recirculation mode is shown in Figure 1. Both aqueous and organic phases were contacted in a hollow fiber module in a cocurrent or countercurrent for extraction or stripping run under recirculating mode. In the extraction module, the feed aqueous phase flows (4.1715.28 cm3/s) through the lumen of the fibers, while the organic phase circulates (2.78-8.33 cm3/s) through the shell side, wetting the wall of the hydrophobic fibers. The pressure of the aqueous phase was held at 0.2-0.5 bar higher than the pressure in the organic phase. While these conditions are maintained, the appearance of the organic phase on the other side of the membrane can be prevented if an immobilized phase is maintained on this side at a pressure equal to or greater than that of the organic phase.4,6 In the stripping run, loaded organic extractor (LIX79 with gold complex) flowed through the shell side (4.72-6.94 cm3/s), whereas NaOH flowed through the tube side (4.72-6.94 cm3/s) cocurrently or countercurrently in recirculation mode. The organic phase is a 1000-mL solution of 12% LIX79 in n-heptane; 1000 mL of aqueous feed solution of the desired concentration of Au(I) was prepared by taking a suitable aliquot from the stock solution. Furthermore, the desired feed pH was adjusted by adding 1 M NaOH solution. The Au(I) concentration in the feed varied between 10 and 200 mg/L. The solution of 0.4-1 M NaOH was used as the stripping phase when back extraction of Au(I) was carried out in the hollow fiber module. Small aliquots of the stripping stream and organic solution were taken at the selected time for the analysis of the metal concentration by standard atomic

absorption spectrometry (2380 Perkin-Elmer absorption spectrometer). In the experiments dealing with the separation of Au(I) from base metals in the hollow fiber contactor, the mixture containing Au(I), Cu(I), Fe(II), and Zn(II) was analyzed by ICP (Spectroflame by Spectro Analytical Instruments) to determine each metal concentration. It was observed that in all cases the pH of raffinate decreases slightly, which is against the proposed extraction mechanism as H+ ions in the feed are continuously consumed by forming the Au(I)-LIX79 complex. The probable reason, as suggested by Virnig et al.,24 is that LIX79 was being loaded with bicarbonate through exposure to the air during the experimental run. To overcome this problem, the LIX79/n-heptane phase was preloaded with a carbonate/bicarbonate buffer with a pH of 10. This resulted in very precise control of the pH of the raffinate leaving the extraction at a constant pH. Theoretical Background Extraction Equilibrium. The extraction mechanism for gold with LIX79 is based on the ion pair type.26 The extraction of Au(I) from the process stream (further referred to as the removal of Au(I) from the feed phase) takes place at the interface of the feed and membrane. The Au(I) ions in alkaline cyanide media (present as Au(CN)2-) form a complex with the extractant LIX79 (N,N′-bis(2-ethylhexyl)guanidine, R), which is protonated between pH 10.0 and pH 10.5, and is expressed as

Rorg + H+aq + Au(CN)2-aq S [HAu(CN)2R]org Kex (1)

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Figure 2. Extraction mechanism of gold cyanide with LIX79 in a hollow fiber contactor.

and similarly for stripping, the chemical reaction is presented as

[HAu(CN)2R]org + NaOHs S Au(CN)2-s + Rorg + H2O + Na+ (2) The extraction equilibrium can be described by the following reactions and extraction constants,

Kex )

[HAu(CN)2R]org -

+

[Au(CN)2 ][H ][R]org

(3)

where R is the organic extractant. The value of Κex for Au(I) with LIX79 was found to be 2.19 ( 0.15 × 1011 27 and is related to the partition coefficient (H) through the following expression:

log H ) log Kex + log[H+] + log[R]org

(4)

Nondispersive Membrane Extraction. The extraction mechanism of gold with LIX79 in a hollow fiber contactor is shown in Figure 2. The following assumptions were made for developing the model: (1) The system is at steady state. (2) Equilibrium exists at the fluid/fluid interface. (3) Pore size and wetting characteristics are uniform throughout the membrane. (4) The curvature of the fluid/fluid interface does not significantly affect the rate of mass transfer, the equilibrium solute distribution, or the interfacial area. (5) No bulk flow correction is necessary; that is, mass transfer is described adequately by simple film-type mass-transfer coefficients. (6) No solute transport occurs through the nonporous parts of the membrane. (7) The two fluids are virtually insoluble in each other.

(8) The equilibrium solute distribution is constant over the concentration range of interest. As derived by D’Elia et al.,28 the key equation for the calculation of KEAu or KSAu for cocurrent flow is

[ ][

]

1 1 + Qf Qe/sH C0e/s/H - C0f ) ln 0 ) 1 1 (Ce/s - C0f ) + (Vf/HVe/s)(C0f - Cf) + Vf Ve/sH t[1 - exp(-4KEAuVm/d{1/Qf + 1/Qe/sH})] (5)

For countercurrent contact, the following equation was used to calculate KEAu or KSAu:28

ln t

[

(C0e/s/H - C0f )

(C0e/s/H - C0f ) + (Vf/HVe/s)(C0f - Cf)

]

)

[1 - exp(-4KEAuVm/d(1/Qf - 1/Qe/sH))][1/Vf + 1/Ve/sH]

[

E

(

4KAuVm 1 1 1 1 exp Qf Qe/s d Qf Qe/sH

)]

(6)

where Qf and Qe/s are the feed and extract/strip flows; Vf and Ve/s are the feed and extract/strip volumes; C0f and C0e/s are the concentrations of the solute in the feed and in the extract/strip solutions at time zero; Cf is the concentration of the solute in the feed at time t; Vm is the volume of all the hollow fibers; and d is the diameter of one fiber. Model Development The system consists of an aqueous phase containing Au(CN)2- flowing in the tube side of microporous hollow fiber membranes, the pores of which are filled with the

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organic extractant, which flows cocurrently or countercurrently in the shell side (Figure 1). The reaction takes place at the inside wall of the membrane where the phase interface is located. The various steps in the extraction process are assumed to be as follows:29-31 Step 1. The gold cyanide in the aqueous phase (tube side) diffuses from the bulk to the aqueous-organic interface (the inside wall of the fiber) through the boundary layer. Step 2. At the aqueous-organic interface, Au(CN)2reacts with LIX79 [(RH)+, which is protonated at a higher pH, i.e., 10.3] in the organic phase in the membrane pore to form the gold cyanide-LIX79 complex. Step 3. Gold cyanide complex diffuses from the aqueous organic interface to the outside wall of the fiber through the organic-filled membrane pore; free LIX79 diffuses in the opposite direction from the shell side into the pore. Step 4. Gold cyanide complex diffuses from the outside fiber wall to the organic-phase bulk (the shell side), flowing cocurrently or countercurrently to the aqueous phase. Therefore, the expression for the overall mass-transfer coefficient, KEAu, can be written as32

( )

ri 1 1 1 1 ) + + + + E k k r KAu i f lm Kex[H ][Rorg]km ri 1 (7) ro k K [H+][R ] s ex org

()

where

1 1 ) ki k [H+][R ] e org

(8)

and

lumen and shell fluid, respectively, and correlations are available in the literature expressing these dependencies. As described in previous work, correlations for mass transfer in a hollow fiber or small tubes were established for the tube side and the shell side by Prasad and Sirkar33 and Dahuron and Cussler,34 respectively. For the tube-side mass-transfer coefficient, the following correlation was given:

( )( )

kfdi Dt di2υt NSh ) ) 1.64 Dt di LDt

Thus,

[H+][Rorg] ) H/Kex Furthermore, upon suitable substitution in eq 8,

1/ki ) 1/(ke[H/Kex]) and ki is the effective rate of interfacial reaction at the surface and kf, km, and ks are the mass-transfer coefficients in the aqueous feed, membrane, and organic solvent, respectively. The term rlm denotes logarithmic mean radius. The first term on the right-hand side of eq 7 represents the resistance due to interfacial reaction; the second term indicates the mass-transfer resistance in the aqueous phase; the third term is the membrane resistance; and the fourth term is the resistance of the organic extraction solvent. The partition coefficient, H ) [H+][Rorg]Kex, appears in the membrane resistance because in our experiments the organic solvent wets the membrane but water does not. The overall mass-transfer coefficient can be calculated from the individual transfer coefficients ki, kf, km, and ks. The tube and shell side mass-transfer coefficients are known to depend on the flow conditions in the fiber

(9)

In general, the above correlation predicts mass-transfer coefficients with reasonable accuracy for NGz > 4 but overestimates them for NGz < 4.35 Under the present experimental conditions, NGz ranges from 5.5 to 8.2. Similarly, for the shell-side mass-transfer coefficient, the following correlation was given,7,36,37

NSh )

ksDh 0.33 ) β[Dh(1 - φ)/L]N0.6 Re NSc Ds

(10)

where the Reynolds number NRe ) υsDh/ηs and the Schmidt number NSc ) ηs/Ds and β is 5.85 for hydrophobic membranes and 0 < NRe < 500 and 0.04 < φ < 0.4. In the present system, both the conditions are met as NRe ranges between 3 and 11.0 and φ ) 0.35 and the notations are Dh ) hydraulic diameter, Dt ) diffusion coefficient of the solute in the tube side; di ) inner fiber diameter; L ) fiber length, Ds ) diffusion coefficient of solute on the shell side, υt and υs are the velocity of the liquid inside the fiber and shell side, respectively, and φ is the packing fraction. The membrane mass-transfer coefficient can be determined from the following expression:4,38

km ) H ) Kex[H+][Rorg]

1/3

Dm tmτ

(11)

For the membrane and solvent here, tm ) membrane thickness ) 30 µm; τ ) tortuosity ) 3 (value obtained from Celgard GmbH); Dm ) diffusion coefficient of the gold complex in the membrane ) 2.5 × 10-6 cm2/s 12 (this value was determined by plotting 1/permeability coefficient vs 1/Kex[H+][Rorg] for different extractant concentrations and at varying pH with fixed LIX79 concentration while performing modeling of Au(I) from aqueous cyanide media across a hollow fiber supported liquid membrane);  ) 0.30; and H ) 4.44. The membrane tortuosity was also determined by the WakaoSmith relation, expressed as the inverse of the membrane porosity, which almost matches the value suggested by Celgard GmbH.39 Results and Discussion Gold Cyanide-LIX79 Extraction System. The physical properties and equilibrium parameters required to calculate the mass-transfer coefficient are presented in Table 1. The equilibrium constant, Kex, used in eq 3 was 2.19 ( 0.5 × 1011, which was determined from the batch extraction earlier.25,27 The forward reaction kinetic constant (ke) in eq 1 was evaluated while mathematical modeling was performed across SLM for the gold cyanide-LIX79 system and was found to be ke ) 6.57 × 1010 cm/s.40 This value could

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Table 1. Physical Conditions and Equilibrium Parameters symbol

property

value

H

partition coefficient

Dm Ds Dt

pore liquid diffusivity (cm2/s) shell liquid diffusivity (cm2/s) aqueous diffusion coefficient of Au(I) (cm2/s) shell fluid viscosity (g/cm‚s) extraction equilibrium constant forward reaction rate constant of eq 8 (cm/s) effective rate of interfacial reaction (cm/s)

ηs Kex ke ki a

4.44 5.00a 2.5 × 10-6 2.5 × 10-5 7.2 × 10-6 4.47 × 10-3 2.19 ( 0.5 × 1011 6.57 × 1010 6.18 × 10-7 Figure 4. Effect of feed flow rate on extraction of gold cyanide with LIX79 in cocurrent mode.

Partition coefficient for stripping.

Figure 3. Concentration courses of Au(I) (Y ) left-hand side of eq 5) extraction obtained from eq 5 vs time (t).

Figure 5. Effect of mode of contact on extraction of Au(I) with LIX79.

Table 2. Mass-Transfer Coefficients for Extraction of Au(I) from Aqueous Alkaline Cyanide Media (pH ) 10.3) with 12% LIX79 in n-Heptane Using the DFSX Technique in Cocurrent Mode as a Function of Aqueous and Organic Flow Rates run

aqueous flow rate (cm3/s)

LIX79 flow rate (cm3/s)

KEAu (×10-6 cm/s)

1 2 3 4 5 6 7 8 9 10 11

4.72 5.54 6.95 9.17 11.11 12.77 15.28 11.11 11.11 11.11 11.11

4.72 4.72 4.72 4.72 4.72 4.72 4.72 2.78 6.94 8.33 11.11

0.2 0.4 0.9 2.9 2.3 4.5 5.5 1.0 4.5 6.5 6.8

not be compared with earlier studies as this is the first ever attempt with the gold cyanide-LIX79 system. Also, Henkel developed this organic extractant very recently. In the DFSX technique, it is well-established4,41 that to keep the interface within the pores of a hydrophobic membrane, it is necessary to maintain a higher local pressure in the aqueous phase. Figure 3 shows a typical plot of concentration differences versus time as predicted by eq 5 for extraction of Au(I) into LIX79. In this plot, the Au(I) concentration varies semilogarithmically with time. From such plots, the overall mass-transfer coefficient, KEAu, can be calculated. The results of these calculations are given in Table 2. Mass-transfer coefficients KEAu for the extraction varied from 2.0 × 10-7 to 6.5 × 10-6 cm/s. In the extraction module, the Au(I) extraction ranged from 90 to 95% in around 1-2 h for cocurrent and 0.5-1 h for countercurrent contact (Figures 4 and 5).

Figure 6. Effect of feed flow rate on mass-transfer coefficient of gold cyanide with LIX79 in an extraction module in cocurrent mode.

As seen in Figure 6, the overall mass-transfer coefficient increased with the feed flow rate and its value was maximum at 15.2 cm3/s. For example, when linear flow velocity (vf) varied by a factor of 2, the overall masstransfer coefficient (KEAu) increased by 4 times. The theoretically calculated values of kf from eq 9 were 2-3 orders of magnitude higher than KEAu. Hence, the feed resistance (1/kf) does not dominate the overall masstransfer coefficient. The Au(I) removal from the feed phase was found to be faster in countercurrent mode (Figure 5). This is due to the fact that the concentration driving forces at the two ends of the extractor in the countercurrent mode are at the maximum, which influenced the gold extraction rate. On the other hand, this was not the case in the cocurrent mode where both flows are in one direction. Figure 7 shows the effect of changing the shell-side organic-phase flow rate and keeping the aqueous flow rate constant for 12% v/v LIX79/n-heptane. Module

Ind. Eng. Chem. Res., Vol. 41, No. 3, 2002 619 Table 5. Mass-Transfer Coefficients for Extraction of Au(I) from Aqueous Alkaline Cyanide Media as a Function of Initial NaCN Concentration with 12% LIX79 in n-Heptane Using the DFSX Technique in Cocurrent Mode run

initial NaCN concentration in aqueous feed (mg/L)

KEAu (10-6 cm/s) (cocurrent)

1 2 3 4 5 6a

0 100 200 300 1000 1000

1.0 1.0 1.0 1.0 0.1 2.2

a

Figure 7. Effect of organic flow rate on mass-transfer coefficient of gold cyanide with LIX79 in an extraction module in cocurrent mode. Table 3. Mass-Transfer Coefficients for Extraction of Au(I) from Aqueous Alkaline Cyanide Media as a Function of Initial Feed pH with 12% LIX79/n-heptane in Countercurrent Mode Using the DFSX Technique KEAu (×10-6 cm/s) run

initial pH of aqueous feed

countercurrent

1 2 3 4

9.1 10.0.3 11.0 12.1

2.2 1.0 0.8 0.05

cocurrent 1.0

Table 4. Mass-Transfer Coefficients for Extraction of Au(I) as a Function of Initial Au(I) Concentration in Feed Using 12% LIX79 in n-Heptane Using the DFSX Technique in Countercurrent Mode run

initial Au(I) concentration in aqueous feed (mg/L)

KEAu (10-6 cm/s) (cocurrent)

1 2 3

10 50 100

1.0 0.9 0.7

extraction efficiency changes dramatically below 2.78 cm3/s of organic flow rate. This means that most gold cyanide can be easily transferred into the organic phase through the interfacial reaction when the concentration of the LIX79 is in excess due to an increase in the organic phase flow rate because the aqueous inlet concentration of gold is only 10-50 mg/L and the aqueous stream flow rate, Qf, is around 11.1 cm3/s (υf ) 2.5 cm/s). Thus, the shell-side resistance is negligible beyond an organic flow rate of 11.1 cm3/s ((υs ) 0.65 cm/s). The effect of pH on the overall mass-transfer coefficient is presented in Table 3. The initial pH of 10.010.5 was observed to be optimum for mass transfer when experiments were performed in countercurrent mode, maintaining the aqueous flow rate at 6.94 cm3/s. The overall mass-transfer coefficient decreased to 0.05 × 10-6 cm/s at pH 12, as suggested by the extraction mechanism. The effect of the Au(I) concentration on the overall mass-transfer coefficient is presented in Table 4. The KEAu was observed to be constant when the concentration of Au(I) in the feed was maintained at 50 mg/L and slightly decreased with an increase in the gold concentration during experiments in countercurrent mode (aqueous flow rate at 6.94 cm3/s). This is in accordance with the partition coefficient, which decreased with an increased Au(I) concentration in the feed, as shown in our earlier studies.12 The overall mass-transfer coefficient of gold plummeted from 1.0 × 10-6 to 0.1 × 10-6 cm/s when NaCN

LIX79 concentration was 18% LIX79/n-heptane.

in the feed increased from 0 to 1000 mg/L. This is probably due to the high level of free cyanide, which competes with aurocyanide for the reagent. To minimize this effect, a higher concentration of LIX79, that is, 18% (v/v), was tested, and recovery of Au(I) was found to be closer to the data indicated in Table 5. The same behavior was also observed when Virnig and Wolfe26 performed conventional liquid-liquid extraction studies with gold mine solutions containing Au(I) (1.3 mg/L) from alkaline cyanide media (5000 mg/L NaCN) using 15% (v/v) LIX79 in Exxon Aromatic 150 with 50 g/L of tridecanol as the modifier. The first term in eq 7 presents the local value of the total resistance. This resistance is in turn the sum of the mass-transfer resistances inside the fiber, across the fiber wall, and in the solution surrounding the fiber. The value of [R]org[H+]aq ) θ is 10-16-10-18. The value of ki was around 6.18 × 10-7 cm/s and ks was (6.5 × 10-6)(10.3 × 10-5) cm/s, whereas km was about 8.33 × 10-5. The value of kf was on the order of 10-3 cm/s. Therefore, on the right-hand side of eq 7, the first term is around 2 × 106, the second term is around 103, the third term ranges from 2000 to 3000, and the fourth term ranges from 1.7 × 103 to 2.7 × 104 s/cm. The overall resistance in the experiments calculated from eq 5 or eq 6 was observed to be 2-5 × 106 s/cm, as compared to an overall resistance value of 2-4 × 106 s/cm estimated from the model, which shows that the resistance due to reaction at the interface is dominant at high flow rates of extractant and feed. The second, third, and fourth terms are thus negligible. In this respect, the fractional resistance of each step of the overall process, Rom, Rof , Ros, and Roi , can be calculated; for example, Rom can be calculated by the following equation:

Rom ) Rm/Ri + Rm + Rs + Rf

(12)

where Ri, Rm, Rs, and Rf are mass-transfer resistances due to interfacial reaction, membrane, organic extractor, and feed. Under the present experimental conditions, the values of %Rom, %Rof , %Ros, and %Roi were 0.12, 0.05, 1.33, and 98.5. This clearly indicated that the ratecontrolling step was the interfacial reaction on the membrane surface. The differences in the theoretical and the experimentally observed NSh and NRe may be due to several causes. First, the boundary condition of constant wall concentration is not strictly valid, especially where a significant amount of solute is extracted.7 It has been observed42 that Newman’s extension of the Leveque solution with only the first term almost perfectly fits the mass-transfer data obtained in single microporous hydrophobic hollow fiber studies with fibers that are not

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Ind. Eng. Chem. Res., Vol. 41, No. 3, 2002

Figure 8. Concentration courses of Au(I) stripping obtained from eq 6 vs time (t) (U ) left-hand side of eq 6). Table 6. Mass-Transfer Coefficients for Stripping of Au(I) with 12% LIX79 (Loaded) in n-Heptane Using the DFSX Technique run

aqueous flow rate (cm3/s)

1 2

4.72 6.95

Cocurrent 4.72 4.72

0.9 3.8

1

6.95

Countercurrent 4.72

5.8

LIX79 flow rate (cm3/s)

KSAu (×10-5 cm/s)

very long. The problem then seems to lie in the fiber bundle itself. Other reasons for deviations of these data from theoretical values are discussed in another section, dealing with gold stripping in a hollow fiber contactor. Gold Cyanide-LIX79 Stripping System. Stripping was performed separately in a single module by flowing loaded organic solution (LIX79 with metal complex) on the shell side and NaOH solution on the tube side. Linear plots were obtained from eqs 5 and 6 for the stripping of Au(I) from LIX79 into aqueous NaOH. Figure 8 shows a typical plot of concentration differences vs time as predicted by eq 6 for the stripping of Au(I) from LIX79-gold cyanide complex by 1 M NaOH. Mass-transfer coefficients for stripping (KSAu) were calculated by multiplying with the reversed partition coefficient (H), as the mass-transfer driving force is in the opposite direction of the forward extraction. The overall mass-transfer coefficients (KSAu) for stripping ranged from 1.0 × 10-5 to 5.8 × 10-5 cm/s. The flow rate of the stripping solution influenced the mass transfer as KSAu increased, as shown in Table 6. Also, the KSAu value was always observed to be greater than KEAu throughout the stripping experiments. The order of magnitude of KSAu values in our experiments was observed to be KSAu ) 3.8 × 10-5 cm/s in cocurrent mode. The values of KSAu were found to be of the same order of magnitude regardless of the contacting mode (cocurrent or countercurrent) for stripping experiments. The rate of the stripping reaction at the surface was not considered in the stripping experiments. The reaction is instantaneous and fast, which was also confirmed when the overall mass-transfer coefficient was estimated to be of the same order as that predicted by the model using kf, km, and ks. The following equation could be used to calculate the overall mass-transfer coefficients for stripping,15

1 1 1 1 ) + + S k k k KAu s m stH

(13)

where ks, km, kst, H, 1/KSAu, 1/ks, 1/km, and 1/kstH are defined in the Notation section. For gold extraction, the partition coefficient appears in the first and second term, but it does not for gold stripping. This occurs because of the wetting situation for the membrane. The membrane is wetted by the organic solvent in the case of gold extraction. Moreover, the solvent has a higher solubility for solute. For gold stripping, the membrane is also wetted by the organic solvent, but in this case it is the solvent that has the lower solubility for the solute.15 Equation 12 indicates that the overall resistance for stripping is the sum of local values of the individual resistances. The value of ks (precision given with a minimum of 10% error) for organic feed varied between 6.5 × 10-6 and 8 × 10-5 cm/s (calculated from eq 10), whereas km was around 8.33 × 10-5 cm/s (calculated from eq 11), as determined earlier for extraction studies. The value of kst was on the order of 10-3 cm/s (calculated from eq 9). Therefore, on the right-hand side of eq 12, the first term is around (1.3 × 104)-(1.5 × 105) s/cm, the second term is around 1.2 × 104 s/cm, and the third term is around 200 s/cm. The second and third terms are thus negligible. The overall resistance in the experiments calculated from eq 5 or eq 6 was observed to be (2 × 104)-(1 × 105) s/cm, as compared to the value of the overall resistance, (3.2 × 104)-(1.6 × 105) s/cm, estimated from the model (eq 12), which shows that the resistance due to organic mass transfer is dominant at low organic flow rates. In this respect, the fractional resistances of each step in the overall process, Rom, Ros, and Rost, can be calculated; for example, Rom can be calculated by the following equation,

Rom ) Rm/Rs + Rm + Rst

(14)

where Rm, Rs, and Rf are the mass-transfer resistances due to membrane, organic extractant, and feed. Under the present experimental conditions, the values of %Rom, %Ros, and %Rost were 0.5, 6, and 93.5. This clearly indicated that the rate-controlling step was in the organic solution. On the other hand, when the organic flow rate is high, then both membrane and organic mass-transfer resistances are dominant as values of fractional resistances, that is, %Rom, %Ros, and %Rost, were estimated to be 37.5, 62, and 0.5, respectively. The deviation of mass-transfer coefficients in extraction was not observed significantly as the interfacial reaction was dominant in the entire extraction process and mass-transfer correlations were of not much use because the flow rate did not influence the mass transfer considerably. On the other hand, in stripping experiments, this deviation was more pronounced as stripping is more influenced by the organic flow rate on the shell side. This behavior could be due to an irregular flow on the shell side caused by the existence of stagnant zones, preferential pathways, and deficient mixing due to a nonuniform distribution of fibers and to their possible deformation by the action of organic solvent.33,43,44 These effects are important mainly when commercial modules are used because the fibers are not uniformly spaced. Also, the poor distribution of the shell-side flow due to the random packing of fibers in the module bundle probably contributed to the inaccuracy of shell-side mass-transfer coefficients, as observed by Rogers and Long.45 In a single-fiber module contactor which did not show such a poor distribution, Takeuchi et al.46 also

Ind. Eng. Chem. Res., Vol. 41, No. 3, 2002 621 Table 7. Separation Factor Values of Gold against Different Cyanide Metal Salts Using LIX79/n-Heptane in Countercurrent Modea metal metal concentration % recovery in separation factor, ions (mg/L) stripping phase SFA,B Au(I) Fe(II) Ni(II) Cu(I) Zn(II) Ag(I)

10 30 10 30 5 3

72 Cu(CN)43Under these circumstances, significant discrimination between the extraction of Au(CN)2- and other cyano ions can be achieved, as shown in Figure 9. Despite these fluctuations, it is apparent that LIX79 exhibits a marked selectivity for gold over copper and iron. These results agree with previous work, indicating that LIX79 has a higher affinity for univalent ions, such as Au(CN)2-, than for multivalent ions, such as Cu(CN)43-, Fe(CN)64-, and Cu(CN)32-.23,26 For anions of the same type, particularly Au(CN)2- and Ag(CN)2-, the larger Au(CN)2anion is extracted in preference to Ag(CN)2-. And other highly charged cyano anions were not amenable to extraction, except zinc complexes. Similar results were found in the extraction of cyano ions by quaternary amines.49

Figure 9. Separation of Au(I) from Fe(II), Cu(II), Ag(I), Ni(II), and Zn(II) using 18% (v/v) LIX79. Au(I): 10.0 mg/L; Fe(II): 30 mg/L; Cu(II): 30 mg/L; Ni(II): 10 mg/L; Ag(I): 3 mg/L; and Zn(II): 5 mg/L in feed at pH 10.5; stripping solution: 1.0 M NaOH.

The extraction of these complexes largely depends on the metal coordination number, and from the results shown in Table 7 it is seen that the extraction order follows the series Me(CN)2- > Me(CN)4n- > Me(CN)6n-. Those complexes with a lower coordination number were extracted preferably over those with a higher number, and in general the same deduction can be made with respect to the complex charge: lower charge complexes are extracted in preference to higher charge complexes.50 The cyano ions of zinc and silver compete with the gold during extraction, but the stripping of these anions in a back-extraction agent was Zn(CN)42- > Ag(CN)2- > Ni(CN)42- > Fe(CN)64- > Cu(CN)43The results obtained allow one to envisage promising behavior of the process subject to its practical applications. The use of stripping solutions containing NaOH provided efficient and fast back extraction of Au(I) from the LIX79/n-heptane organic phase. A model is presented that identifies the different rate-controlling steps under various chemical and hydrodynamic experimental conditions in the feed, organic, and membrane phases. The mass-transfer resistance, which resulted from a chemical extraction reaction at the surface, was observed to be dominant during the entire extraction process, whereas the stripping reaction was fast and instantaneous and the mass-transfer coefficient for stripping was found to be more than that of extraction. The validity of this model was evaluated with experimental data and found to tie in well with theoretical values.

reaction, and stripping, respectively Rm, Rf, Rs, Ri, Rst ) resistance due to membrane, feed, solvent, interfacial reaction, and stripping, respectively (s/cm) tm ) thickness of the fiber membrane (cm) Ve ) organic tank volume (cm3) Vf ) feed tank volume (cm3) Vs ) stripping tank volume (cm3) Vm ) volume of hollow fibers (cm3) Subscripts in ) inside the fiber out ) along the fiber outside i ) for inner radii o ) for outer radii e/s ) for extract/strip t ) tube side s ) strip side or shell side f ) feed m ) membrane Superscripts 0 ) refers to the concentration at time zero

Acknowledgment This work was supported by the CICYT (QUI 99-0749) and CIRIT (SGR-98-0082). Dr. Anil Kumar acknowledges the financial support from the Comisio´n Interministerial de Ciencia y Tecnologı´a, Spain, through the award of the Visiting Scientist Fellowship. R. Haddad acknowledges the support from ICMA for a fellowship. G. Benzal acknowledges the support from FOMEC (Fondo para la mejora de la ensen˜anza y calidad universitaria) of the Universidad Nacional de Tucuma´n, Tucuma´n, Argentina, for the fellowship. The authors would also like to thank R. Ninou for performing the experiments. The authors gratefully acknowledge the Henkel Corporation, U.S.A., for supplying the LIX79. Notation C ) metal concentration (g/cm3) d ) diameter of one fiber (cm) Dt ) diffusion coefficient of the solute on the tube side (cm2 s-1) Ds ) diffusion coefficient of the solute on the shell side Dm ) diffusion coefficient of the metal complex in the membrane Dh ) hydraulic diameter (cm) di ) inner fiber diameter kf ) mass-transfer coefficient of the aqueous feed (cm/s) ki ) rate of interfacial reaction at the surface (mol2/(cm.5s) km ) membrane mass-transfer coefficient (cm/s) ks ) organic mass-transfer coefficient (cm/s) kst ) mass-transfer coefficient for stripping (cm/s) E S KAu , KAu ) overall mass-transfer coefficient for extraction and stripping, respectively (cm/s) L ) fiber length (cm) Q ) flow rate (cm3/s) NRe ) Reynolds number for feed (NRe ) υd/η) H ) partition coefficient of extraction and stripping NSc ) Schmidt number (NSc ) η/D) NGz ) Graetz number (NGz ) d2υ/(DL)) NSh ) Sherwood number for the tube side, NSh ) kfdi/Dt, and the shell side, NSh ) ksDh/Ds, respectively r ) hollow fiber radius (cm) rlm ) logarithmic mean radius of lumen (cm) Rom, Rof , Ros, Roi , Rost ) fractional resistance due to membrane, feed, solvent, interfacial

Greek Letters τ ) tortuosity of the membrane ηt ) viscosity of aqueous feed/stripping solution, cP or g/(cm‚s) ηs ) viscosity of organic solution, cP or g/(cm s) υt and υs ) velocities of liquid on the fiber side and shell side (cm/s) φ ) packing fraction of HF module  ) porosity

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Received for review February 12, 2001 Revised manuscript received October 9, 2001 Accepted October 24, 2001 IE010141P