Environ. Sci. Techno/. 1994, 28, 1090-1098
Emulsion Liquid Membranes for Wastewater Treatment: Equilibrium Models for Some Typical Metal-Extractant Systems Bhavani Raghuraman, Neena Tirmizi, and John Wiencek’ Department of Chemical and Biochemical Engineering, Rutgers University, Piscataway, New Jersey 08855
Extraction and recovery of heavy metals from wastewater is more attractive than methods such as precipitation that result in sludges that have to be disposed in landfills. Emulsion liquid membranes are capable of extracting metals from dilute waste streams to levels much below those possible by equilibrium-limited solvent extraction. Binary equilibrium data are reported for copper, nickel, and zinc with extractants that can be used in emulsion liquid membrane formulations. Predictive models that incorporate aqueous-phase nonidealities and all aqueousphase ionic reactions have been developed. For the CuLIX and Zn-D2EHPA systems, a single value for the equilibrium constant ( K )is accurate over a large range of pH and ionic strength. For the Ni-D2EHPA system, an average single K value was obtained for loadings (fraction of D2EHPA dimers complexed to the metal in the organic phase) less than 0.1. At higher loadings, the organic-phase nonidealities become significant, and the equilibrium constant was fitted as an exponential function of the loading. Emulsion liquid membrane extractions of Cu, Ni, and Zn from aqueous phases were successfully carried out. The effects of various parameters such as mixing speed, stripping reagent concentration, and extractant concentration on the extraction process are reported. Introduction Aqueous streams contaminated with heavy metal ions are frequently encountered. Such streams may be produced as effluents from various industrial processes (such as mining and smelting, metal plating, and metal finishing) or during attempts to remediate solids loaded with heavy metals (e.g., soil washingileaching). Removal of heavy metal ions from aqueous solutions has traditionally meant the precipitation of the ion. Such practice is now unpopular as this produces a sludge that has to be disposed in a landfill. Electrochemical metal recovery is promising; however, waste streams are often dilute in metals and have low plating efficiencies. Solvent extraction of metals has been extensively used in hydrometallurgical operations. Because metal salts usually are not soluble in organic solvents, the process requires the introduction of an extractant that will combine with the metal ion to form an organic soluble species. An effective extracting agent must (a) provide a very high distribution coefficient from dilute solutions; (b) react reversibly with the metal so that the metal can be recovered by a stripping reaction; (c) have reasonably fast extraction and stripping kinetics; and (d) have very low solubility in the aqueous phase for economic and environmental reasons. The major disadvantage of solvent extraction is that the extraction is limited by equilibrium considerations. Emulsion liquid membranes (ELMs) remove equilibrium limitations of solvent extraction by combining
* To whom correspondence should be addressed; e-mail address: JWIENCEKa ZODIAC.RUTGERS.EDU. I090
Environ. Sci. Technol., Vol. 28, No. 6, 1994
R E C E M N G PHASE (LowpH to Strip)
MICRODROPS
cu FEED PHASE (High pH t o Extract) MA~RODROPS
‘
MEMBRANE PHASE [OIL)
Figure 1. Schematic representation of copper ion extraction with an emulsion liquid membrane. Copper(1I ) is transported to the emulsionfeed phase interface and reacts with the complexing agent (RH) to form a soluble copper complex (CUR,). This complex diffuses to the interior of the emulsion droplet until it encounters a microdroplet of the internal phase where the metal ion is exchanged for a hydrogen ion. The net effect is a unidirectionalmass transport of the cation from the original feed to the receiving phase with countertransport of hydrogen ions. The dispersion is then allowed to settle, and the lower aqueous stream is withdrawn for discharge. The upper emulsion phase is then demulsified to split the membrane and the enriched stripping phases.
extraction and stripping in a single operation. ELMShave been successfully used to treat aqueous streams contaminated with heavy metal ions like copper, zinc, cadmium, nickel, mercury, lead, and chromium (1-6). ELMs, first are made by forming an emulsion invented by Li (3, between two immiscible phases. Usually stabilized by surfactants, the water-in-oil emulsion contains the ex. tracting agent in the oil phase and the stripping reagent in the aqueous receiving phase. This emulsion is then dispersed by mechanical agitation into a feed phase containing the metal to be extracted. Figure I is a schematic representation of an emulsion liquid membrane extraction of copper(I1). Combining the extraction and stripping processes removes equilibrium limitations and reduces metal concentrations in the feed to very low levels. Demulsification by application of high-voltage electric fields has proven to be most efficient (8). Heavy metals concentrated in the receiving phase can be recovered by electroplating or crystallization (as a single pure salt).The oil phase can be recycled. A different class of ELMs, microemulsion liquid membranes, have been used to successfully extract copper and mercury from aqueous streams (9, 10). Unlike the 0013-936X/94/0928-1090$04.50/0
0 1994 American Chemical Society
emulsions described earlier (also known as coarse emulsions), which are formulated by blending, microemulsions form spontaneously when the various constituents are brought in contact (11). Microemulsions have been shown to be more efficient because of their greater stability (leading to lower leakage rates of internal phase into the feed phase) and lower interfacial tension (which leads to smaller macrodrops and hence faster extraction rates). This study represents the initial phase of an effort to develop coarse and microemulsion extraction processes for the removal of copper, nickel, and zinc from aqueous waste streams. Extraction agents selected for these systems must, in addition to the criteria discussed earlier, be capable of forming stable emulsion membranes. Equilibrium data are reported for these systems, and the equilibrium has been modeled for the more promising systems. Binary equilibrium data and a predictive model are necessary for developing an overall model for the emulsion extraction process. While there is ample published dataon copper extraction equilibrium (12-14), they are essentially with older extracting agents (Henkel hydroxyoximes, such as LIX 64N, LIX 63, and LIX 65) that are no longer marketed commercially. Better extractants such as LIX 860 (5dodecylsalicylaldoxime)and LIX 984 (equivolumemixture of 5-dodecylsalicylaldoxime and 5-nonylacetophenone oxime) which show faster extraction kinetics, easier phase disengagement, and strong extraction of copper even at low pH values of 1-2 are now available. Strong selectivity is reported for copper over other metals like ferric, ferrous, and calcium around a pH of 2 (15). Data on these newer, more efficient reagents are lacking. Other chelating extractants that have been used are Kelex 100 (8hydroxyquinoline) (14), Shell Chemical's SME 529 (2hydroxy-5-nonylacetophenoneoxime) and ICI's Acorga P5000 Series (5-nonylsalicylaldoxime). Acid extractants like D2EHPA [bis(2-ethylhexyl)ester of phosphoric acidl have been used in copper extractions in the pH range 3-5 (16). Carboxylic acids like heptanoic, capric, and valeric acids have also been used (17). The disadvantage with these acids is their increased water solubility a t high pH levels. Equilibrium models reported for copper extraction ignore nonidealities and aqueous-phase ionic equilibria of other salts and buffers that are present, and only apparent concentration equilibrium constants are reported [ CuKelex 100 (18);Cu-LIX 860 (19,201; CU-D~EHPA(16)l. Semiempirical and empirical techniques have been used for Cu-Kelex 100 (21) and Cu-2-hydroxy-5-nonylbenzophenone oxime (the main component of the Acorga P5000 Series reagents) (22). The extractant most commonly used for nickel is D2EHPA at a pH of 4-6 (1,23). LIX 84 (5-nonylacetophenone oxime) has been reported to extract nickel from ammonical solutions (15). Carboxylic acids like naphthenic acid, capric acid, and versatic acid have also been used in nickel extractions from ammonical solutions (24). D2EHPA has been reported to be a good extractant for zinc in the pH range 1.5-3 (15,251. Phosphinic acids like Cyanex 272 [bis(2,4,4-trimethylpentyl)phosphinicacidl and Cyanex 302 [bis(2,4,4-trimethylpentyl)monothiophosphinic acidl marketed by American Cyanamid are also reported to be good extractants for zinc (26). The selectivity with respect to ferric ions is poor but is high for calcium. Cyanex 302 is reported to be a stronger extractant than Cyanex 272; however, stripping is more difficult. Since
zinc can form chloride complexes, it can also be extracted with amine extractants like Alamine 336 (triisooctylamine) (15). Stripping is accomplished with inorganic salt solutions like NaCl or NaOH. Apparent concentration equilibrium constants have been reported for nickel (16,27) and zinc (25, 28, 29) extractions. Nonidealities and aqueous-phase ionic equilibria of other salts and buffers that are present were not considered. This study aims to develop predictive equilibria models for copper extraction with LIX 860 and LIX 984, the extractants which are commerciallyavailable and for which there is very little information reported in literature. For the Ni-D2EHPA and the Zn-D2EHPA systems, we propose to develop more complete predictive models that would account for all ionic equilibria involving the metal salts and buffers in the aqueous phase as well as for nonidealities. Successful emulsion extractions of these metals with the selected extracting agents are also reported, and their advantage over equilibrium-limited solvent extraction is demonstrated. Equilibrium Model. Typical metal (M2+)complexation with LIX extractants (RH) is described by M2++ 2RH
-
MR,
+ 2H'
(1)
The thermodynamic equilibrium constant, K , is defined as
Metal complexation with n dimers of DBEHPA [(RH121 is described by M2'
+ n(RH),
MR2(n - l)(RH),
+ 2H'
(3)
and
K=
(MR& - 1)(RH)2){H')2
W2+WW,Jn
(4)
where terms in braces (( )) represent activities that are products of concentrations and activity coefficients. The activity coefficients for the ionic species are calculated using the modified Guggenheim equation (30): (5) where I is the ionic strength of solution cizi2 i=l
zi is the ionic charge; Ci is the concentration of ionic species;
A is the constant = 0.509; B is equal to 0.15 for the unmodified equation valid for I < 0.1 M and is an adjustable parameter in the modified equation for higher ionic strengths. However for all the predictive models developed here, the standard deviation in the estimated equilibrium constant did not appreciably vary withB value. Hence a value of 0.15 was used throughout. For neutral species (31):
log yi = 0.11 Envlron. Sci. Technol., Vol. 28, No. 6, 1994
(7) I091
Activity ( a ) for water is given by (31)
5000.-
n
aHZO= 1- 0 . 0 1 7 c ci
(8)
i=l
The organic-phase nonidealities are initially ignored. The equilibrium constants estimated from experimental data are then examined to see if the deviations in K show any trend with respect to loading (which defines both metal complex concentration and unreacted extractant concentration in organic phase). The absence of any trend would indicate that either the organic phase is ideal or the ratio of activity coefficients as it appears in the expression for the equilibrium constant (eqs 2 and 4) is a constant that is accounted for in the value of K itself. If K varies with loading, then the organic-phase nonidealities are lumped together and accounted for by fitting an equation for K as a function of loading. This was found to be necessary only for the Ni-D2EHPA system above a loading of 0.1.
Experimental Section Materials. The sources for copper, nickel, and zinc were copper sulfate, nickel nitrate, and zinc sulfate, respectively (Fisher Scientific). The LIX reagents used were supplied as approximately 46 wt % in kerosene by Henkel and were used without further purification. Other extracting agents used were oleic acid (Fisher Scientific), D2EHPA (Sigma Chemical Co.), and Cyanex 272 (Cynamid Co.). Solvent tetradecane (99 % ) was obtained from Humphrey Chemicals. Surfactants used for the emulsions were ECA 5025 (polyamine supplied by Exxon) and DNP-8 [poly(8)ethyleneoxide dinonylphenoll (Henkel). Sodium acetate (buffer) and sulfuric acid (stripping agent in emulsions) were purchased from Fisher Scientific. Procedures. Aqueous solutions of metals were prepared by dissolving the salt in deionized water. For NiD2EHPA, Cu-D2EHPA, and Cu-oleic acid systems, the pH was controlled by using a 0.1 M sodium acetate-acetic acid buffer. The organic phases containing the extracting agent in tetradecane solvent were prepared on a weighti weight basis. A 10 w t % LIX solution refers to 10 wt % of the LIX reagent (as received from Henkel) in tetradecane. Equilibrium distribution experiments required contacting known volumes of aqueous phase and organic phase in test tubes on a tube rotator for about 2 h at 20 OC. A laboratory centrifuge was used to ensure complete disengagement of the two phases a t the end of the extraction. The metal concentrations were measured using flame atomic absorption spectroscopy on a Perkin-Elmer Model 3030 spectrophotometer. The wavelengths used for copper, nickel, and zinc were 324.8, 232.1, and 213.9 nm, respectively. The precision of multiple readings of the same sample was generally within 5 % while that of duplicated experiments ranged from 3 to 10%. Copper concentrations in the ppb range were measured by Zeeman corrected graphite furnace atomic absorption (PerkinElmer Model 4100 ZL). Organic-phase metal concentration was calculated by material balance. Material balances were checked by back-extraction of the metal from a few typical samples of the loaded organic phase with 6 N Hz-
sod.
Coarse emulsions were formulated by blending the organic membrane solvent (tetradecane) containing the extractant and surfactant (ECA 5025) and the internal aqueous stripping phase in a high-speed blender for 2 min. 1092 Envlron. Sci. Technol., Vol. 28,No. 6,1994
5
5
(.
I
l
l
,
'
c
LIX 984;ON H,SO,
4000I
U
v
5 l F
.-C
1
LIX 860;0.88N H2S0,
2000
I1
L
0
200
400 600 800 Copper in aqueous (ppm)
1000
Flgure 2. Copper equilibrium with LIX 860 and LIX 984. Comparing experimental data with model predictions (solid lines) as a function of sulfuric acid concentration (0.0 N, 0.88 N, and 3.93 N) in the aqueous phase. Initial copper concentration in aqueous phase = 1000 ppm. Organic phase has 10 wt % LIX extractant in solvent tetradecane. Extraction occurs in the high pH range, and stripping takes place in concentrated sulfuric acid solutions.
Microemulsions were formulated by equilibrating solvent tetradecane containing the extractant and surfactant (10 wt % DNP-8) with the stripping phase (6 N H2SO4). Equilibration time was 1 2 h. The exact formulation of the emulsions for the different systems are reported with the results of those runs. Aqueous content of the emulsion was measured using a Mettler Model DL18 Karl Fischer titrator. Extraction was carried out in a stirred contactor using a turbine impeller at a constant stirring speed in the range of 300-400 rpm. Samples were collected periodically, and the aqueous-phase metal concentration was measured. Leakage of the internal phase into feed phase was monitored by analyzing the S042-ion concentration in the feed phase by ion chromatography on a Dionex 4000i instrument.
Results and Discussion Copper. Preliminary experiments in our laboratory with the LIX reagents indicated significant copper extraction even at pH values as low as 1. Because of the strong copper extraction at low pH, the stripping solution employed was a strong acid. The strongest extractant, LIX 860, requires 6 N H2SO4 (-300 g/L) for complete stripping of copper. Inoue et al. (19) have measured the distribution constant for LIX 860 between toluene or hexane and aqueous solutions of nitric acid and ammonium nitrate. The distribution constant is reported to be independent of pH. Typically, the LIX 860 partitioning to aqueous phase is 0.3-3 ppm for organic-phase concentrations of LIX 860 in the range of 3-20 wt 9%. LIX 860 has been reported to form stable microemulsion and coarse emulsions containing 6 N H2S04as the internal phase and successfully extract copper from aqueous solutions (9). However, the water content of the microemulsion is low (4-6 wt %), and this limits the amount of copper that can be solubilized in the internal phase. Figure 2 shows plots of copper distribution between the aqueous phase and the organic phase for LIX 860 and LIX 984 as a function of sulfuric acid strength. For the runs where no sulfuric acid was added (0 N), the concentration of the hydrogen ions liberated by the extraction reaction was obtained by
Table 1. Equilibrium Constants for Metal Extraction system
aqueous phase 1M (H,NH.dNOs 1M (H,NHd)N03
CU-LIX 860
HC1 CU-LIX 860 CU-LIX 860 CU-LIX 984 Ni-DPEHPA (loading < 0.1)
0-3 M &So4 0-2 M H2S04 1M (Na,H)NOs 0.5 M (Na,H)N03 0.5 M (Na,H)N03 0.5 M Na2S04 0.1 M NaAc 0.4-1 M (H2,Naz)SOd 0.5 M (Na,H)S04 0.05 M HC104 04.3 M
Zn-D2EHPA
organic solvent
Ka
hexane toluene toluene toluene C10-C13 n-hydrocarbons tetradecane tetradecane n-dodecane n-heptane n-heptane toluene or benzene toluene/benzene kerosene tetradecane n-heptane kerosene Escaid 100 tetradecane
110 7.9 4.2 1.0 5.4 x 104 328.0 75.48 4.59 x loa 8.0 X 4.5 x loa 2.0 x 10-6 1.2 x 10-6 6.6 X lo6 6.95 x 105 7.35 x 10-3 9.45 x 10-3 0.101 0.1084
source ref 20 ref 20
ref 20 ref 20 ref 21 this study this study ref 36 ref 16 ref 16 ref 16 ref 16 ref 27 this study ref 33 ref 31 ref 34 this study
a Units for K are dimensionless, (L/gmol)-', and (gm~l/L)O.~ for Cu, Ni, and Zn systems, respectively. Except for this study, Kvalues reported are apparent concentration equilibrium constants.
titrating the aqueous-phase samples before and after equilibration with NaOH. As expected, it was found that for every 1mol of copper extracted, 2 mol of hydrogen ion were produced. Equilibration experiments were carried out to determine the selectivity of the LIX reagents for copper in the presence of ferric ions. The aqueous phase contained 1000 ppm each of copper ions (from copper sulfate) and ferric ions (from ferric nitrate) and was contacted with a stoichiometric excess of LIX 860 in the organic phase. The equilibrium pH was measured to be 1.5. The ferric extraction, which was 80-85% in the absence of copper, decreased to about 30% when copper was present. The copper was, however, completely extracted. The extraction-stripping reaction of copper with LIX reagents (RH) can be described by eq 13. This reaction has been widely studied with the now discontinued LIX 64N and reported to be interfacial with a L1X:Cu stoichiometry of 2 (13,14). The LIX solution supplied by Henkel is reported to contain approximately 46 wt % of LIX in kerosene. The exact initial molar concentration of the LIX extractant in the organic phase was obtained by completely loading the organic phase with copper and measuring the saturating copper concentration by stripping with an excess of 6 N H2S04. Assuming a 2:l stoichiometry of LIX 860:Cu, the initial molar concentration of LIX 860 for a 10 w t % solution was calculated to be 0.126 M. Ionic equilibria in the aqueous phase are described below:
-
Cu2++ SO,H,S04
-
CuSO,
K , = 229
(9)
+ H+ + H+
K , = 95.49
(10)
K , = 0.0105
(11)
HSOL
HSO;
-
H,O
H+ + OH-
Cu2++ 2RH
-
SO,*-
CUR,
K4 = 1.0 X
(12)
+ 2H+
K = 328.0 (LIX 860, calculated) = 75.48 (LIX 984, calculated) (13)
Hydrolysis of Cu is negligible below a pH of 6.9 (30). The equilibrium constants for the ionic equilibria were
taken from Lindsay (30). The above equilibria together with mass balance equations for copper and sulfur and a charge balance equation were solved numerically using the IMSL subroutine NEQNF. Experimental results were used as input to predict the equilibrium constant ( K )for the extraction-stripping equilibria. Average K values for the Cu-LIX 860 and Cu-LIX 984 systems were estimated to be 328.0 (13.8% standard deviation) and 75.48 (13.3% standard deviation), respectively. Some data points a t the extremes of the extraction curves (close to 0% and 99% extractions) were excluded as even small errors in aqueous-phase concentration measurement could lead to large errors in K estimates. Solid lines in Figure 2 are model predictions. It can be seen that the single K value is able to give satisfactory prediction over a large range of pH and ionic strength. Table 1compares our values for K with those reported by other workers. The latter are actually concentration equilibrium constants and were estimated by graphical techniques and ignore aqueousphase ionic equilibria and nonidealities. This might explain the variations in K for the different aqueous media used in their study. Figure 3a compares extraction of copper from aqueous phase (initial concentration of 1000 ppm) with ELMS containing 10 wt % LIX 860 and 10 wt % LIX 984. Simultaneous extraction and stripping by coarse emulsions reduces copper concentrations 2 orders of magnitude lower than the equilibrium value obtained by solvent extraction. The small increase in copper concentration a t the end of the run is due to leakage from the internal-phase droplets of the emulsion. LIX 860 is a stronger extractant, and hence the extraction rates are faster. The ratios of residual copper concentration in the feed phase after an ELM extraction to the equilibrium concentration are 36 and 45 for the LIX 860 and LIX 984 systems, respectively. Even though the equilibrium concentration of Cu for LIX 860 extraction is lower than that for LIX 984 by a factor of 2, the LIX 984 ELM is more efficient. This is because the stripping of copper from the stronger binding LIX 860 is more difficult than that from LIX 984. Figure 3b shows extraction of copper from a very low concentration solution (initial concentration of 1000 ppb) with a 10 wt % LIX 860 coarse emulsion. The ratio of residual copper concentration in the feed phase after the ELM extraction to Environ. Scl. Technol., Vol. 28, No. 6, 1994
1093
8
.
.
Equilibrium
.c
a
LIX860
U
f 1
2 c
0001
t 0
c
--
-+-+-----------c+
10
20
30 40 Time (min)
60
50
70
1---C---t-*
'----I---
Emulsion
10-4-
0
- _ i _5
----+
10
j
+ .
15
20 25 Time (rnin)
30
35
40
Flgure 3. Coarse emulsion extraction of copper with LIX 860 and LIX 984. Coarse emulsion formulation: 82 wt % organic phase (10 wt % LIX reagent 3 wt % ECA 5025 surfactant in tetradecane) and 18 wt % aqueous phase (6 N sulfuric acid). Speed = 350 rpm; treat ratio (volume of feed phase/volume of emulsion) = 5. (Panel a, top) Emulsion extraction lowers copper concentration in aqueous feed phase 2 orders of magnitude below that possible by equilibrium-limited solvent extraction. LIX 860 is a stronger extractant with faster extraction kinetics. Initial copper concentration in feed = 1000 ppm; initial pH = 4.5; final pH = 1.6. (Panel b, bottom) LIX 860 coarse emulsion extraction of copper from very dilute solutions. Initial copper concentration in feed = 1000 ppb; initial pH = 3.6;final pH = 3.1.
tion of copper from aqueous solutions at a pH of 5.4 over equilibrium-limited solvent extraction. The microemulsion extracted the copper a t a faster rate as expected, but the coarse emulsion showed more stability. It is possible that this microemulsion is more stable at the lower pH (2.2-2.8) used in mercury extraction than at the higher pH (5.4) used for copper extraction. Nickel. DBEHPA showed good extraction of nickel from aqueous phase containing acetate buffers in the pH range 4-5 (Figure 4). Stable coarse emulsion formulations with D2EHPA required the addition of 10 w t 9% light mineral oil to solvent tetradecane to increase the membrane viscosity and hence reduce internal phase leakage. Microemulsions employing DBEHPA were unstable and displayed high leakage rates during the extraction. D2EHPA solubility in the aqueous phase was measured by ion chromatography and by a total organic carbon analyzer. It was observed to be a function of the pH, the DBEHPA concentration in the organic phase, and the organic:aqueous phase volume ratios. Typically, the D2EHPA partitioning to water was about 1000 ppm for 10 wt % D2EHPA and 600 ppm for 5 wt % D2EHPA in tetradecane for a phase volume ratio of 1. However, for Vaq/Vorgequal to 10, the concentration fell by an order of magnitude. For the coarse emulsion extraction using 5 wt 7% D2EHPA and a treat ratio of 5, the concentration in the aqueous phase was measured to be about 100 ppm. DBEHPA is reported to exist as a dimer in the organic phase and complexes as a dimer with the metal (16,27). Graphical techniques applied to experimental data for low loadings of DBEHPA predict that three dimers complex with one Ni ion. The ionic equilibria for this system can be described by the following equations:
+
the equilibrium concentration is 12.5. This is lower than the ratios obtained for the ELM extraction from a 1000 ppm aqueous solution. This is due to the very low loading of LIX 860 in the membrane phase at the low copper concentrations used in the feed, which makes the contribution of the stripping step to the overall extraction process negligible. The role of stripping would probably become more obvious at lower concentrations of LIX 860 in the membrane phase. A 10 w t 5% solution of D2EHPA in tetradecane gave more than 80% extraction of copper from buffered solutions (0.1 M sodium acetate-acetic acid, 1000 ppm initial copper) in the pH range 4-5 with an aqueous:organic volume ratio of 1. Oleic acid (10 wt 7% in tetradecane) was effective in the 4.8-5 pH range giving 50-80 % extractions under similar conditions. Both these extractants were able to form stable coarse emulsions. Oleic acid coarse and microemulsions have been used for mercury extraction (10). The microemulsion was reported to extract the mercury faster and reduce it to lower levels and was also more stable. Oleic acid coarse and microemulsions gave 71- and 41-fold improvements, respectively, in the extrac1094
Environ. Sci. Technol.. Vol. 28,No. 6,1994
Ni2++ NO;
+ Ni2++ OHNi2++ 20HNi2+ 2NO;
Ni2++ Ac-
+
Ni2+ 2Ac-
HA^
H,O Na'
-
- + - + -
+ 2(RH),
Ni2++ 3(RH),
K5 = 2.51
NiN03+
K , = 0.25
Ni(NO,),
(14) (15)
NiOH'
K , = 12589
(16)
Ni(OH),
Kg = 1 X 10'
(1'7)
NiAc'
K , = 13.18
(18)
Ni(Ac),
K,, = 64.56
(19)
H+
AC-
K,, = 1.75 x 10-~ (20)
H'
OH-
K,, = 1.0 X lo-'*
NaR.3(RH)
(21)
+ H+ K,, = 2.03 X lo4 (22)
NIR2-2(RH), + 2H' K = 6.95 X loT5(calculated) (23) The equilibrium constants for the ionic equilibria were taken from the literature (32,33). Sodium ions are also extracted by D2EHPA, and equilibrium data for this has been reported by Li et al. (29). The above equilibria together with four mass balance equations for nickel, sodium, nitrate, and acetate and a charge balance equation were solved numerically using the IMSL subroutine NEQNF. Experimental results were used as input to predict the equilibrium constant ( K ) for the extraction-
0.3
100
L
I
I
I
t
80
l5
1 60 0
U
E
predicted experimental
02Y
w
0.05
0
2
1
4
3
5
Flgure 4. Equilibriumextraction of nickel and zinc as a function of pH. Ni aqueous feed phase: 1000 ppm Ni and 0.1 M sodium acetate buffer: acetic acid used to adjust pH. Zinc aqueous feed phase: 1000 ppm Zn; sulfuric acid usedto adjust pH. Organic phase: wt % extractant in tetradecane. Vo,,lVa, = 1.
0.001
4:
0.0001
i
I
I
I
I
1
1
1
1
I
I
005
0.1
0.15 Loading
0.2
0.25
0.3
1
J Mean K =6.95*105 (Loading c 0.1)
I __ 0
20
I
40
60
80
100
Nickel in aqueous phase (ppm)
Equilibrium pH
I
0
t
t
Figure 5. Variation in predicted equiilbriumconstant (4 for Ni-D2EHPA system with loading. Equilibrium constants predicted by the model for different experimental data sets are plotted as a function of loading. Loading is defined as the ratio of Ni concentration in organic phase to initial DPEHPA dimer concentration. For loadings less than 0.1, an average value of K = 6.95 X may be used. For loadings greater than 0.1, the equilibrium constant is fitted as an exponential function of loading.
stripping equilibria. For loadings less than 0.1 (loading = [organic Ni complexl/[initial D2EHPA dimer]), the K value was reasonably constant with an average value of 6.95 X (standard deviation = 22%) (Figure 5). Table 1 summarizes K values estimated by other workers. However, in these studies nonidealities were ignored, and aqueous phase ionic equilibria for metal-metal salts and buffer salts were not considered. This might explain the variation in the equilibrium constant with ionic strength. Sodium buffers used for pH control can result in some sodium ion extraction, which may give erroneous equilibrium constant values if ignored. At higher loadings, K was observed to increase with the loading. Such behavior has been observed by other workers (16, 34) and is attributed to the polymerization of the Ni-DBEHPA complex in the organic phase. Brisk and McManamey (34)report that the organic-phase viscosity rises sharply above a loading of 0.08 and attribute it to polymer formation in the organic phase (Note: The loading in their paper is defined as a ratio of organic nickel complex
Flgure 6. Comparing predicted and experimental equilibrium data for Ni-DPEHPA. Aqueous phase: initial Ni concentration in feed is varied from 0 to 1000 ppm; 0.1 M NaAc buffer added. Organic phase: 5 wt % DPEHPA in tetradecane; VwalVaq= 1. Solid lines are model predictions.
to initial D2EHPA monomer concentration and has to be multiplied by 2 to conform to the definition used here.) They postulated a series of polymer-forming reactions for loading greater than 0.2 in the organic phase. By assuming the equilibrium constant (Kr) for the formation of the polymer with an aggregation number equal to r to be the same as that for the polymer with an aggregation number equal to r - 1,the equilibrium partitioning was modeled for high loadings. This assumption is valid only for high loadings, which implies a high r value and is justified by the principle of equal reactivity of all functional groups in linear chain polymers. The Kr value is reasonably constant a t higher loadings (greater than 0.4) but shows large deviations in the 0.2-0.4 loading range. The loading range in our experimental work is less than 0.26, and hence the Brisk and McManamey model was not applied. As per eq 4 for the equilibrium constant, where the activity is the product of the species concentration and its activity coefficient, organic-phase nonidealities would be reflected in a ratio of activity coefficients for the organic-phase species. Most activity coefficient correlations are exponentially dependent on the species concentrations. Thus, an exponential curve for K as a function of loading was fitted for loadings greater than 0.1 with a regression coefficient of 0.97 (Figure 5). Attempts to fit the entire range of data (loading 0-0.26) with a single exponential curve were not successful as the model significantly underpredicted loading in the loading range of 0-0.05. Model predictions are compared with experimental data in Figure 6 for an aqueous:organic phase volume ratio of 1and an initial concentration of nickel varying from 0 to 1000 ppm. The model did not work for the low pH region where the concentration of undissociated acetic acid in the aqueous phase becomes significant. This is possibly because of water-acetic acid interactions that were not incorporated in the model. The effect of stripping reagent concentration on nickel extraction from aqueous solutions with a 5 wt % D2EHPA coarse emulsion is shown in Figure 7a. A 10 wt % light mineral oil was added to the membrane solvent (tetradecane) to increase its viscosity and to make the emulsion more stable. The extraction with 6 N HzS04 is more efficient and faster as expected. A higher stripping reagent concentration implies more efficient stripping and hence Environ. Scl. Technol., Vol. 28, No. 6, 1994
1095
,
EquilibriLrn
U
Y
a,
I'
J ~
0.1
0.01
a
loo-a
U a, L
c c a P
10
L
Stage 1
e 1
01
I 0
-
2 s -
v
1
IO
20
30
1 40
-
50
60
70
Time (min)
Flgure 7. Coarse emulsion extraction of nickel with D2EHPA. Coarse emulsion formulation: 82 wt % organic phase (5 wt % D2EHPA 3 wt YOECA 5025 surfactant 10 wt % light mineral oil in tetradecane) and 18 wt % aqueous phase (2-6 N sulfuric acid). Aqueous phase: 1000 ppm Ni and 0.1 M sodium acetate buffer; initial pH = 6.88; speed = 350 rpm; treat ratio (volume of feed phaselvolume of emulsion) = 5. (Panel a, top) Effect of stripping reagent (sulfuric acid) concentration on extraction. Coarse emulsion extraction with 6 N H2S04in the internal phase is most efficient because of effective stripping and, hence, higher driving force for the extraction process. (Panel b, bottom) Twostage extraction of nickel with 6 N H2S04as internal phase. Stage 1 reduces nickel concentration from 931 to 7.46 ppm in 15 min; stage 2 reduces it from 7.46 to 0.4 ppm in 7 min.
+
+
lower nickel concentration in the membrane phase. This increases the driving force for nickel extraction from the feed phase, and hence, the extraction is faster. The emulsion stability is also observed to be a function of the sulfuric acid strength as seen from the rise in nickel concentration in the latter stages of the extraction, which is an indication of internal-phase leakage. This was checked by analyzing for sulfate in the feed phase. The leakage (= 100 X ratio of actual sulfate concentration in feed phase to that present if all sulfate from internal phase were to leak) was calculated to be 3.4%, 5.676, and 60% for 6 N, 4 N, and 2 N HzSO4 emulsions after 60 min of extraction, respectively. It is possible that the higher viscosity of the 6 N HzS04solution (2.30 cP) as compared to the 4 N HzSO4 (1.72 cP) and 2 N HzS04 (1.47 cP) solutions may be the reason for the higher stability of the ELM with 6 N HzSO4 as the internal phase. Figure 7b is a two-stage extraction of nickel from an aqueous solution with an initial nickel concentration of 931 ppm. Since 99 % of nickel extraction with a 6 N HzSO4 coarse emulsion is completed in 15 min (Figure 7a), the first stage was run 1096
Environ. Sci. Technol., Vol. 28, No. 6,1994
for 15 min and the two phases were separated. 'The nickel concentration in the feed was reduced from 931 to 7 ppm. The raffinate from the first stage was then contacted with a fresh coarse emulsion for the second stage, and the concentration was further reduced to 0.4 ppm in 7 min. The slight rise in nickel concentration is due to internalphase leakage. Oleic acid was observed to be less effective than D2EHPA in nickel extraction from acetate buffers and hence was not investigated further. Experiments using LIX 84 showed good extraction of nickel from solutions in the 3.8-8.25 pH range. The pH was adjusted using nitric acid and ammonium hydroxide. However, emulsion extractions from ammonia buffers render the stripping acid ineffective due to the diffusion of ammonia from the external phase to the internal phase and neutralization of the acid in there; thus, the acid was unavailable for stripping nickel from the organic phase. Proof of this explanation is the rapid fall in the pH of the external phase as compared to the control experiment (no internal phase) even when there is no leakage of the acid from the internal phase. Hence, this system was not explored further. LIX 84 performed well both in equilibrium extractions and in coarse emulsion extractions when acetate buffers were used. The kinetics of extraction was very slow (typically 50% extraction as compared to 98% extraction with D2EHPA in 10 rnin). Efforts to increase the separation rate by varying buffer concentration from 0.1 M to 0.5 M acetate, using LIX 860 instead of LIX 84 or reducing the mineral oil content in the organic phase, were not successful. Zinc. D2EHPA extracts zinc from aqueous solutions in the pH range 1-2 (Figure 4). The pH was varied by the addition of sulfuric acid. Various workers have modeled the Zn-D2EHPA system equilibrium (25,28,29) (Table 1). The stoichiometry isgenerallyreported to be 1.5dimers of D2EHPA per zinc ion when aliphatic hydrocarbons like heptane, decane, and kerosene are used as the solvent for D2EHPA. The species in the organic phase is reported to be ZnRZ-RH, where R is the D2EHPA anion. Studies done with solvents like benzene and chloroform report a stoichiometry of 2 (29). The references cited here all report apparent or concentration equilibrium constant for a specific ionic strength. Ionic equilibria and nonidealities are ignored. The ionic equilibria for this system can be described by the following equations: ZnSO, Zn2++ H,O
-
+
Zn2+ 2H,O
-
Zn2" + SO,'-
K,, = 211.8
+ H+ Zn(OH), + 2H'
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-
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-
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Hf
-
ZnZf + 1.5(RH),
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K,, = 95.5
(27)
+
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(28)
§04,- Hf
+ OH-
-
K,, = 1.0 X
(29)
ZnR,.(RH) + 2H+ K = 0.1084 (calculated) (30) The equilibrium constants for the ionic equilibria were taken from Lindsay (30). The above equilibria together
0.40
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/b= -/ 1
0.01 0
0.1
0.2
0.3
0.4
0.5
1
I
I
0
10
20
1 I
I
Loading Flgure 8. Variation in predicted equilibriumconstant (K)for Zn-D2EHPA system with loading. Equilibrium constants predicted by the model for different experimental data sets are plotted as a function of loading. Loading is defined as the ratio of Zn concentration in organic phase to initial D2EHPA dimer concentration. An average Kvalue of 0.1084 was obtained for the entire range.
1 A
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I
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I
I
I
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60
70
I
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10 wt% DPEHPA
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-
,
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, I
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, I
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0
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,
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100
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,
,
150
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200
250
300
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100
Zinc in aqueous phase (ppm)
i
L a
Flgure 9. Comparing predicted and experimental equilibrium data for Zn-D2EHPA. Aqueous phase: initial Zn concentration in feed was = 1; H2S04added = 0.0 M. varied from 0 to 1000 ppm; V,,,lV,,
with mass balance equations for zinc and sulfate and a charge balance equation were solved numerically using the IMSL subroutine NEQNF. Experimental results were used as input to predict the equilibrium constant ( K ) for the extraction-stripping equilibria. The variation of K with loading is observed to be random (Figure 8). The average K value was calculated to be 0.1084 (standard deviation = 23.5%). Figure 9 compares predicted (solid lines) and experimental data for 5 and 10 w t % DBEHPA. The effect of mixing speed on coarse emulsion extraction of zinc from aqueous feed is shown in Figure loa. The extraction rate a t 350 rpm is observed to be faster than that at 300 rpm, indicating that the mass transfer of zinc from the feed phase into the membrane phase (which is a function of surface area and mass transfer-coefficient, both of which increase with mixing speed) is important. Increasing the speed above 350 rpm may further increase extraction rate, but the higher shear would also increase the leakage rate from the internal phase making the extraction inefficient. The extracting agent (DBEHPA) concentration did not seem to effect the extraction rate for zinc (Figure lob). This is because of the simultaneous extraction and stripping in a coarse emulsion which keeps the loading low in the membrane phase. Hence, the actual
Stage I
3
0.01
0
10
20
30
40
50
60
70
Time (min) Flgure 10. Coarse emulsion extraction of zinc with DPEHPA. Coarse emulsion formulation: 82 wt % organic phase (5 wt % D2EHPA -t 3 wt % ECA 5025 surfactant 4- 10 wt % light mineraloil in tetradecane) and 18 wt % aqueous phase (6 N sulfuric acid). Aqueous phase: 1000 ppm Zn; initial pH = 4.34; treat ratio (volume of feed phase/voiume of emulsion) = 5.(Panel a, top) Effect of speed of stirring on extraction. Faster extraction at 350 rpm indicates mass-transfer resistance for this system. (Panel b, middle) Effect of extractant concentration on extraction rate. Speed = 350 rpm. (Panel c, bottom) Two-stage extraction of zinc with 5 wt % D2EHPA coarse emulsion and 350 rpm stirring speed. Stage 1 reduces zinc concentration from 1049 to 16 ppm in 10 min; stage 2 reduces it from 10 to 0.07 ppm in 7 min.
concentration of D2EHPA is not important as long as it is above a certain critical concentration. The extraction reaction at the interface is not rate controlling. This observation further confirms the importance of masstransfer resistances for this system. Figure 1Oc is a twostage extraction process where the raffinate from stage Envlron. Sci. Technol., Vol. 28, No. 6 , 1994
1087
one is contacted with a fresh coarse emulsion in stage two. In the first stage, the zinc concentration is reduced from 1049 to 16 ppm in 10 min. Stage two further reduces it to 0.05 ppm in 15 min. Cyanex 272 also displayed good extraction of zinc (Figure 4). Coarse emulsion extractions with Cyanex 272 were stable; however, the extraction rate was slow compared to D2EHPA (typically 80% extraction with Cyanex 272 as compared to 97% extraction with DBEHPA in 10 min).
Conclusions Binary equilibrium data for selected heavy metalextractant systems used in emulsion extractions have been reported. This is the first stage of an overall project to develop emulsion extraction processes for heavy metal recovery from dilute waste streams. Predictive models that include aqueous-phase nonidealities and all aqueous ionic equilibria have been developed for the more promising systems. For the Cu-LIX systems, a single K value could be used over a large range of pH and ionic strength. For the Ni-D2EHPA system, a single Kvalue was obtained for loadings less than 0.1. At higher loadings, the organicphase nonidealities become significant, and the equilibrium constant was fitted as an exponential function of the loading. For the Zn-D2EHPA system, a single K value was used for the entire range. Emulsion liquid membrane extractions of Cu, Ni, and Zn from aqueous phases were successfully carried out. The potential advantages of using ELMS over equilibriumlimited solvent extractions for metal recovery from wastestreams, containing metals in the ppm and ppb range, have been demonstrated. The effects of various parameters such as mixing speed, stripping reagent concentration, and extractant concentration were studied. Future research will concentrate on demulsification of the spent emulsion to recover the loaded internal phase and modeling of the extraction and stripping kinetics and the overall emulsion extraction process. Alternate contactors for emulsion liquid membrane extractions that will minimize leakage and water uptake by emulsions and improve overall efficiencies are also being examined (35).
Acknowledgments This work has been funded by the Hazardous Substance Management Research Center of New Jersey, Grant PHYS-28 (a NSF Industry/University Cooperative Center and a New Jersey Commission on Science and Technology Advanced Technology Center); by New Jersey Water Resources Research Institute (USGS Grant G-1577-02); and by Merck and Company, Inc. The authors appreciate the help of Barry Hu, Gang Chang, Shilpa Bhagat, and Nai-Jen Hsueh in collecting some of the data.
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(5) Kitagawa, T.; Nishikawa, Y.; Frankenfeld, J.; Li, N. N. Environ. Sci. Technol. 1977, 11, 602-605. (6) Izatt, R. M.; Bruening, R. L.; Geng, W.; Cho, M. H.; Christensen, J. J. Anal. Chem. 1987,59, 2405-2409. ( 7 ) Li, N. N. U.S. Patent 3,410,794, 1968. (8) Draxler, J.; Furst, W.; Marr, R. J . Membr. Sci. 1988, 38, 281-293. (9) Wiencek, J.; Qutubuddin, S. Sep. Sei. Technol. 1992, 27, 1407-1422. (10) Larson, K. A.; Wiencek, J. M. Znd. Eng. Chem. Res. 1992, 31, 2714-2722. (11) Miller, C. A.; Qutubuddin, S. Interfacial Phenomena in Non-Aqueous Media; Eicke, H. F., Parfitt, G. D., Eds.; Dekker: New York, 1986, Chapter 4. (12) Fisher, J. F. C.; Notebaart, C. W. Handbook of Solvent Extraction; Lo, T. C., Baird, M. H. I., Hanson, C., Eds.; Krieger Publishing Co.: Malabar, FL, 1991; Chapter 25.1. (13) Cox, M.; Flett, D. S. Handbook of Solvent Extraction;Lo, T. C., Baird, M. H. I., Hanson, C., Eds.; Krieger Publishing Co.: Malabar, FL, 1991; Chapter 2.2. (14) Ritcey, G. M.; Ashbrook, A. W. Solvent Extraction, Principles and Applications to Process Metallurgy, Part I; Elsevier: New York, 1984; Chapter 3. (15) Henkel Corporation. Minerals Industry Division Red Book; Kordosky, G. A,, Ed.; Henkel Corp.: Tucson, AZ, 1990. (16) Komasawa, I.; Otake, T.; Higaki, Y. J . Znorg. Nucl. Chem, 1981,43, 3351-3356. (17) Kojima, I.; Uchida, M.; Tanaka, M. J . Inorg. Nucl. Chem. 1970,32, 1333-1340. (18) Hoh, Y. C.; Bautista, R. G. Metall. Trans. 1978,923,69-75. (19) Inoue, K.; Arita, H.; Baba, Y.; Yoshizuka, K. Proc. Symp, Solvent Extr. 1987, 155-160. (20) Lazarova, Z.; Boyadzhiev, L. J.Membr. Sci. 1993, 78, 239245. (21) Bauer, G. L.; Chapman, T. W. Metall. Trans. 1976, 7R, 519-527. (22) Bogacki, M. B.; Szymanowski, J. Znd. Eng. Chem. Res. 1990, 29,-601-606. (23) Gu, Z. M.; Wasan, D. T.; Li, N. N. Sep. Sei. Technol. 1985, 20, 599-612. (24) Jacobs, J. J.; Allard, M.; Behmo, S.; Moreau, J. Nickel and Cobalt Extraction Using Organic Compounds; Pergamon Press: New York, 1985. (25) Huang, T.; Juang, R. J. Chem. Eng. Jpn. 1986,19,379--386. (26) Cynamid Company Technical Bulletin, 1991. (27) Huang, T. C.; Tsai, T. H. Acta Chem. Scand. 1991,45,383391. (28) Ajawin, L. A.; Perez de Ortiz, E. S.; Sawistowski, H. Chem. Eng. Res. Des. 1983, 61, 62-66. (29) Li, Z. C.; Furst, W.; Renon, H. Hydrometallurgy 1986,16, 23 1-24 1. (30) Lindsay, W. Chemical Equilibria in Soils; John Wiley and Sons: New York, 1979. (31) Allison, J. D.; Brown, D. S.; Novo-Gradac, K. J. Minteqa2/ Prodefa2, A Geochemical Assessment Model for Environmental Studies: Version 3; U.S. EPA: Athens, GA, 1990. (32) Lange’s Handbook o f Chemistry, 13th ed.; Dean, J. A., Ed.; McGraw Hill: New York, 1985. (33) Smith, R. M.; Martell, A. E. Critical Stability Constants; Plenum Press: New York, 1974; Vol. 4. (34) Brisk, M. L.; McManamey, W. J. J . Appl. Chem. 1969,19, 103-108. (35) Raghuraman, B.; Wiencek, J. AIChE J . 1993, 39, 18851889. (36) Grimm, R.; Kolarik, Z. J . Znorg. Nucl. Chem. 1974,36,189192.
Received for review August 3, 1993. Revised manuscript received February 25, 1994. Accepted March 2, 1994. @
@
Abstract published in Advance ACS Abstracts, April 1, 1994.
