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Lithium Enrichment of High Mg/Li Ratio Brine by Precipitation of Magnesium via Combined CO2 Mineralization and Solvent Extraction Peng Chen, Siyang Tang,* Hairong Yue, Changjun Liu, Chun Li, and Bin Liang Multi-Phases Mass Transfer and Reaction Engineering Laboratory, School of Chemical Engineering, Sichuan University, Chengdu 610065, China S Supporting Information *

ABSTRACT: A new method coupling CO2 mineralization and solvent extraction was proposed and investigated to precipitate magnesium from high Mg/Li ratio brine. Organic amine was employed to extract HCl formed during the CO2 mineralization process and to realize continuous conversion of MgCl2 to precipitated MgCO3, thereby removing the magnesium contained in the brine. The effects of brine concentration, extractant, diluents, extractant concentration, phase ratio, and reaction temperature were systematically investigated. A blend of trioctylamine and isoamyl alcohol was used as the extractant and diluent, and the reaction parameters were optimized. The mechanism was proposed, and a stripping process was designed. Under the optimal conditions, the conversion of magnesium was as high as 67.41%, and the Mg/Li ratio of the raffinate decreased from 20 to 5.4. Many separation methods, such as adsorption,7,8 solvent extraction,5,9 nanofiltration,10 salting out,11 roasting−leaching separation,12 etc., were reported to overcome separation of Li from high Mg/Li brine. Wang et al.8 and Ö zgür7 used spinel lithium manganese oxide ion-sieves as Li inorganic adsorbent with a remarkable capacity for Li (41.26 mg/g), but the granulation process is challenging for commercialization. Triisobutyl phosphate5 and tetraphenylborate/tributyl phosphate9 mixtures were used as extractant to extract lithium from brine, respectively, while the triisobutyl phosphate single-step extraction efficiency of lithium was 83.71%. Somrani et al.10 introduced nanofiltration and low-pressure reverse osmosis to separate lithium from lake brines; the results showed that all magnesium could be rejected by the membrane NF90 under pressure of 14 ± 2 bar, but 50% of the initial flow rate remained after 6 h of filtration resulting from the fouling of the membrane. Zeng et al.11 introduced an extended BET model to predict the phase equilibrium of system in which the HCl was used as salting-out agent to separate MgCl2 from a HCl−LiCl− MgCl2−H2O system. Yang et al.12 proposed roasting−leaching separation; the mixture of MgCl2·4H2O and LiCl was roasted at high temperature (over 700 °C), forming MgO and HCl gas, and LiCl was leached with water. However, this process is of high energy consumption and of high corrosion risk. Even though many separation methods have been reported and some

1. INTRODUCTION The use of lithium ion batteries continues to increase exponentially with an annual growth rate of 64% due to the increasing energy demand of new technologies such as electric automobiles and portable electronic devices;1 thus, lithium resources have become more and more important. Most of the lithium currently consumed comes from lithium ores as well as salt lake resources.2 According to the United States Geological Survey (USGS) (2016),3 global lithium reserves contain about 40.5 million tons of lithium and more than 85% of the world’s recoverable lithium resources are in the liquid state.4 Because the lithium extracted from rock minerals is normally more expensive, more and more lithium is being extracted from salt lake water and subsurface brines.5 In the extraction of lithium from salt lake resources, the salty water is naturally evaporated and concentrated under solar light, where both KCl and NaCl are precipitated by crystallization. The magnesium and lithium are concentrated and remain in the mother brine. When the Mg/Li ratio in the brine is below 6, lithium can be economically separated by a chemical precipitation method.6 However, the cost of the precipitation separation process increases significantly at high Mg/Li ratio. In the traditional precipitation separation process using soda, a high Mg/Li ratio results in high soda consumption for precipitation of magnesium and high loss of lithium by coprecipitation, with formation of a large amount of magnesium precipitate. Unfortunately, the Mg/Li ratio of most brine resources is higher than 6, especially in China,6 which makes the separation of lithium from magnesium by precipitation more difficult. © XXXX American Chemical Society

Received: Revised: Accepted: Published: A

December 18, 2016 April 10, 2017 April 22, 2017 April 23, 2017 DOI: 10.1021/acs.iecr.6b04892 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

Article

Industrial & Engineering Chemistry Research

lithium chloride in deionized water to simulate the mother brine; the Mg/Li ratio of the simulated brine was 20. 2.2. ECM Experiments. ECM experiments (Figure 1) were conducted in a 250 mL three-necked glass flask prefilled with the simulated brine, organic amine, and diluent. The flask was immersed in a water bath to curtail fluctuation of the reaction temperature to below ±1 °C. When the temperature reached the given value, CO2 was introduced by bubbling with a pure CO2 stream under mechanical agitation to begin the reaction, and reaction time was set at 120 min according to the effect of reaction time as shown in Figure S1. Under the extraction conditions, magnesium was gradually converted to MgCO3 and precipitated from the solution. The final slurry was filtered with a glass filter. The precipitate was washed with deionized water and anhydrous alcohol, and the filter cake was then dried at 110 °C for 24 h. The wash water was separately collected. The filtrate was settled for 120 min and separated in a separatory funnel. The organic phase containing the extractant was recovered, and the raffinate (the aqueous phase) was collected. The Mg2+ content in both the washing water and the raffinate were analyzed by titration with a calibrated EDTA solution. Li+ contents were analyzed by inductively coupled plasma analysis (ICP-AES, IRIS Advantage, Thermo Jarrell Ash, United States). The conversion of magnesium in mineralization, xMg, is calculated as follows:

of these techniques have been implemented on the commercial scale, the utilization of high Mg/Li brines is still a challenge in terms of economical lithium extraction. Qinghai is the base for potassium fertilizer production in China, where more than 120 million cubic meters of mother brine is discharged annually in the production of KCl. This brine contains lithium and magnesium at concentrations of about 3−7 g·L−1 and 80−120 g·L−1, respectively, where the Mg/Li ratio can be as high as 20. Instead of soda, CO2 is used as the precipitant to mineralize magnesium, where MgCl2 is converted to MgCO3 precipitate by bubbling CO2 into the brine according to the following equation: MgCl2(aq) + CO2 (aq) + H 2O(aq) ↔ MgCO3↓ + 2HCl(aq) (1)

However, this reaction cannot go to completion when the generated HCl remains in the liquid phase because of the equilibrium limitation. The solvent extraction technique has been used to extract HCl from the mineralization system, thereby driving the reaction forward by eliminating the equilibrium limitation. Ye et al.13 used tributylamine to extract HCl in the mineralization of CaCl2 with CO2 in order to separate CaCl2 from KCl solution. For the dilute salt solution, the extraction requires a highly basic extraction solvent like trioctylamine (TOA). Solvent extraction of acid is widely applied in acid purification14 and recovery.15−17 Extractants such as Cyanex 923,18 trioctylphosphine oxide,19 Alamine 336,20 Aliquat 336,21 tributyl phosphate (TBP),22 and tris(2ethylhexyl) amine (TEHA)23 are typical extractants used in commercial operation. In this work, acid extraction is combined with CO2 mineralization (ECM) to separate magnesium from high Mg/ Li brine. Compared with soda precipitation, ECM using CO2 has a greatly reduced chemical cost. The effect of operation parameters such as the concentration of brine, reaction time, extractant, diluent, extractant concentration, phase ratio (O/A), and temperature are systematically investigated, and the extraction mechanisms are discussed.

xMg =

yLi =

concentration (g·L )

Mg

Li

B

S(SO42−)

Cl

K

96

5

3.46

3

290

1

× 100% (2)

′ + V ″C Li ″ V ′C Li × 100% VC Li

(3)

where CLi is the initial concentration of Li+ in brine, CLi ′ the concentration of Li+ in raffinate, and CLi ″ the concentration of Li+ in washing water. The organic phase product was characterized by Fourier transform spectroscopy (FT-IR, Spectrum Two Li10014, PerkinElmer, United States), and the solid product was characterized by X-ray diffraction (XRD, DX-2700, Dandong, China). 2.3. Stripping Experiments. All stripping experiments were carried out in a 100 mL three-necked glass flask. For each test, 50 mL of the loaded organic phase and required volumes aqueous ammonia were brought into contact at 100 rpm for 20 min, where the stripping could be done under these conditions. The flask was immersed in a water bath at 25 ± 1 °C. After stripping, the liquid were transferred to a separating funnel and allowed to settle for 20 min; then, the organic phase and aqueous phase were separated. The organic phase before and after stripping were titrated by calibrated NaOH solution (0.1 mol/L) to determine the stripping ratio of the organic phase separately. The stripping ratio is expressed as

Table 1. Chemical Composition of Typical Mother Brine24 −1

VCMg

where CMg is the initial concentration of Mg2+ in brine, CMg ′ the concentration of Mg2+ in raffinate, C″Mg the concentration of Mg2+ in washing water, V the volume of brine, and V′ the volume of rafffinate. V″ is the volume of washing water. The recovery ratio of lithium, yLi, is defined as follows:

2. EXPERIMENTAL SECTION 2.1. Materials. The organic amines, triethylamine (99.0%) and trioctylamine (TOA, 95.0%), were purchased from Aladdin Chemical Reagent. Tripropylamine (99.0%) was purchased from Macklin; diisobutylamine (99.5%) was purchased from Sinopharm Chemical Reagent Co., Ltd.; triamylamine (97%) was purchased from Tokyo Chemical Industry Co., Ltd. Isoamyl alcohol (98.5%), magnesium chloride (98.0%), and lithium chloride (98.5%) were purchased from Kelong Chemicals Co., Ltd. Ethylenediamine tetraacetic acid disodium salt (EDTA, 99.5%) was purchased from Tianjin Guangfu Technology Development Co., Ltd. The typical chemical composition of the mother brine from West Taijinar Salt Lake of China is presented in Table 1. Because other anion and cation content are much lower than that of Cl−, Mg2+, and Li+, MgCl2 and LiCl were used to simulate the brine in all ECM experiments. The experimental brine was prepared by dissolving magnesium chloride and

component

′ − V ″CMg ″ VCMg − V ′CMg

* ·HCl − C NR ·HCl)/C NR * ·H+ × 100% xr = (C NR 3 3 3 B

(4)

DOI: 10.1021/acs.iecr.6b04892 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Figure 1. Schematic of ECM. 1, CO2 gas; 2, gas flow meter; 3, gas buffer vessel; 4, stirring motor; 5, water bath; 6, constant-temperature magnetic agitator; 7 and 9, glass filters; 8 and 10, filtering flasks; 11, separating funnel.

where C*NR3·HCl is the concentration of trioctylamine hydrochloride in the organic phase before stripping and CNR3·HCl is the concentration of trioctylamine hydrochloride in the organic phase after stripping. The concentration of NR3·HCl was determined on the basis of the volume of NaOH solution was consumed.

k7

CO2 (g) ⇄ CO2 (aq) k8

CO2 (aq) + H 2O ⇄ H 2CO3(aq)

(8)

k9

H 2CO3(aq) ⇄ HCO3−(aq) + H+(aq)

(9)

k10

3. RESULTS AND DISCUSSION 3.1. Thermodynamic Calculations. 3.1.1. Using Carbonate as Precipitant to Separate Mg2+ and Li+. Reaction equilibrium constants of reaction of Mg2+ or Li+ with CO32− in aqueous solution at 25 °C are as follows:25

HCO3−(aq) ⇄ CO32 −(aq) + H+(aq)

(10)

where k7, k8, k9, and k10 are the reaction equilibrium constants. According to the literature,25,28 the values of k7 and k10 are 3.43 × 10−2 and 4.69 × 10−11, respectively. The equilibrium constant of reactions 8 and 9 could be reviewed as combined apparent equilibrium value k8 × k9 of 4.45 × 10−7.25,28 Combined with reaction 4, the global equation can be expressed as

k5

CO32 −(aq) + Mg 2 + ⇄ MgCO3(s)↓ k5 = 1/ksp(MgCO3) = 1.47×105

(7)

(5)

k11

CO2 + Mg 2 + + H 2O ⇄ MgCO3 + 2H+

k6

CO32 −(aq) + 2Li+ ⇄ Li 2CO3(s)↓ k6 = 1/ksp(Li 2CO3) = 1.23 × 103

k11 = k5k 7k 8k 9k10 = 1.05 × 10−13 (6)

(11)

It is obvious that the reaction is thermodynamically unfavorable. In order to form MgCO3, the conditions of expression 12 should be satisfied.

where k5 and k6 are the equilibrium constants for reactions 5 and 6, respectively. ksp(MgCO3) and ksp(Li2CO3) are the solubility product constants for MgCO3 and Li2CO3, respectively. On the basis of the equilibrium constants, the precipitation of Mg2+ should occur prior to that of Li+. If the brine contains 120 g·L−1 of magnesium and 6 g·L−1 of lithium, no lithium should be coprecipitated under the ideal conditions until the concentration of magnesium decreases to 0.15 g·L −1 . Theoretically, carbonate is effective as a precipitating agent for separating Li+ and Mg2+. However, commercial operation using soda as a precipitating agent requires excess soda to remove magnesium completely; thus, the loss of lithium occurs in the initial period of precipitation because of the high initial supersaturation of Li2CO3. For example, if the brine contains 120 g·L−1 of magnesium and 6 g·L−1 of lithium, about 4.93 mol of soda is required per liter of brine; the initial CO32− concentration of 4.93 mol·L−1 is over the critical concentration of 1.07 × 10−3 mol·L−1 for lithium. 3.1.2. Use of CO2 to Precipitate Mg2+ from Li+ Solution. When CO2 is used as the precipitant, the reactions include the hydration of CO2 and the ionization of H2CO3. These reactions can be expressed as26−28

[H+]2 /[Mg 2 +][PCO2 ] < 1.05 × 10−13 2−

(12)

+

The activities of CO 3 and H are equal to their concentrations in dilute solution. For the brine containing 120 g·L−1 of magnesium and 6 g·L−1 of lithium, expression 13 can be obtained. [H+] < 2.31 × 10−7mol ·L−1 or pH > 6.64

(13)

3.1.3. Enhancement via Organic Amine Extraction. Organic amines are strong bases that extract HCl from the aqueous phase to form amine salts: k14

HCl(aq) + NR3(o) ⇄ R3NH+·Cl−(o)

(14)

This promotes the forward reaction in expressions 9 and 10 by reducing the proton concentration in the aqueous phase.29 The global reaction is expressed as MgCl2(aq) + 2 NR3(o) + CO2 (g) + H 2O(aq) k15

⇄ 2R3NH+·Cl−(o) + MgCO3(s)↓ C

(15)


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Concomitantly, more water entered the organic phase because of the increasing solubility of water in the organic phase. On the other hand, more water was consumed by reaction 15 and in the form of water of crystallization; the aqueous phase completely disappeared when the initial magnesium concentration was more than 90 g·L−1. This results in difficulty in separating the solid and organic phases. Considering the recovery ratio of lithium, the conversion of magnesium, and the separating process, the initial concentration of magnesium of 50 g·L−1 was selected as the optimal condition. 3.2.2. Effect of Organic Phase. 3.2.2.1. Organic Amine. The efficacy of different tertiary amines in the experiments was compared. The experiments were conducted under the same conditions as described above by using isoamyl alcohol as the diluent. The initial brine contained 50 g·L−1 magnesium and 2.5 g·L−1 lithium. The results presented in Figure 3 show that the

The pH of the aqueous phase is associated with the apparent basicity of the organic amine, and is determined by the equilibrium limitation of reaction 14. The equilibrium concentration of magnesium ions depends on the pH of the aqueous phase.30−33 Wang et al.34 calculated the equilibrium concentrations of magnesium with variation of the pH when an ammonia/ammonium chloride buffer was employed under a CO2 pressure of 1 atm, demonstrating values of 39.03 ppm at pH 8 and 10 ppm at pH 9. The results showed that CO2 mineralization was facilitated during the extraction process by using a suitable solvent with reasonable basicity. 3.2. Effects of EMC Experimental Conditions. 3.2.1. Effect of Initial Concentration of Brine. The ECM reaction was conducted using 25 mL of brine with an initial Mg/Li ratio of 20 as the reactant, 50 mL of TOA as the extractant, and 50 mL of isoamyl alcohol as the diluent; the reaction was conducted at 25 °C for 120 min using a stirring speed of 300 rpm and bubbling with a CO2 stream at a flow rate of 60 mL·min−1. The effects of varying the initial Mg2+ concentration from 25 to 90 g·L−1 on the recovery ratio of lithium, the conversion of magnesium, as well as the Mg/Li ratio were measured, and the results are shown in Figure 2.

Figure 3. Effect of organic amine on the recovery ratio of lithium, the conversion of magnesium, and the Mg/Li ratio. TEA, triethylamine; TPA, tripropylamine; DIBA, di-isobutylamine; TAA, triamylamine; TOA, tri-n-octylamine; isoamyl alcohol was used as the diluent; initial brine concentration, Mg (50 g·L−1) and Li (2.5 g·L−1); 25 °C; 300 rpm; 120 min; CO2, 60 g·L−1.

Figure 2. Effect of initial magnesium concentration on the recovery ratio of lithium, the conversion of magnesium, and the Mg/Li ratio. Initial Mg/Li ratio, 20; TOA and isoamyl alcohol, 50; 25 °C; 300 rpm; 120 min; CO2, 60 mL·min−1.

recovery ratio of lithium was consistently over 93%, while the conversion of magnesium and the Mg/Li ratio varied significantly. The conversion of magnesium with different extractants followed the order triethylamine > di-isobutylamine > tripropylamine > triamylamine > TOA > N235, which is ascribed to the inductive effect and steric hindrance.35 Triethylamine, which was most effective in terms of CO2 mineralization, gave rise to a magnesium precipitation ratio of 97.4%, and the Mg/Li ratio declined from 20 to 0.54. Extraction of HCl from the aqueous solution by the organic amines should be regarded as a reaction involving a proton, and the organic amine extraction mechanism includes anion exchange, ion-pairing, hydrogen bonding, and/or solvation.36 Ion-pair formation is the major pathway for HCl extraction in this system. The basicity of an organic amine depends on its substituent groups, which influences the inductive effect and the steric hindrance effect. Increasing the number of substituent groups and the alkyl chain length increases the electron density on the nitrogen atom, thereby increasing the potential to attract protons (inductive effect), but also increases the resistance to contact with the protons (steric hindrance effect). Triethylamine, tripropylamine, triamylamine, TOA, and N235, as

With the increase of initial Mg2+ concentration, the recovery ratio of lithium first kept at about 99% then decreased linearly at magnesium concentration over 50 g·L−1. The recovery ratio of lithium was 84.39% at initial Mg2+ concentration of 90 g·L−1. The lost lithium entered the solid phase as Li2CO3, and it cannot be recovered. The magnesium conversion gradually increased in response to increasing the initial Mg 2+ concentration up to about 50%. The final Mg2+ content was limited by the CO32−. According to the equilibrium values of H2CO3, HCO3−, and CO32− with proton, the pH of the raffinate determined the concentration of magnesium. The Mg/Li ratio decreased at first and then slowly increased with increasing initial Mg2+ concentration. Even though magnesium was continuously precipitated in the high magnesium concentration zone, the accompanying lithium loss resulted in a minimum Mg/Li ratio of 9.79 at an initial magnesium concentration of 50 g·L−1. With increasing initial magnesium concentration, more HCl was generated and extracted into the organic amine phase. D


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magnesium precipitation degree increases with polarity of diluent for different alcohol increases. The alcohols with short chain alkyls show polarity and lower steric hindrance to make solvation easier and the solubility of the amine salt larger. However, the polarity of these species also causes loss of the organic phase because of their high solubility in water. For isoamyl alcohol, as diluent, a conversion of magnesium more than 50% was acceptable. Considering the solubility of the diluent in water, isoamyl alcohol was chosen as the diluent for the following experiments. 3.2.2.3. Amine/Alcohol Ratio. The total volume of TOA and isoamyl alcohol was kept at 100 mL, and the influence of the amine/alcohol ratio on the ECM process was investigated. Figure 5 shows the conversion of magnesium, Li recovery ratio,

tertiary amines, have the same number of amine substituent groups on the nitrogen atoms; triethylamine exerted the lowest steric hindrance effect because of its short chains. On the other hand, triethylamine absorbed more water from the aqueous phase because of its strong water solvation effect. The longer the substituent of tertiary amine, the lower the solubility in water and vapor pressure. Because TOA, the only commercial application tertiary amine, has three C8 substituents, its loss was less than that of other teritary amine in each EMC experiments. Therefore, considering the volatility, solubility in water, and the cost of the amine, TOA was the optimal extractant in the EMC process with the lithium recovery ratio of 99%, Mg conversion ratio of 50.9%, and the Mg/Li ratio decrease from 20 to 9.79. 3.2.2.2. Diluent. Diluents were used to modify the physical and chemical properties of the organic phase. The diluent was used to reduce the viscosity and avoid the formation of a third phase by adjusting the solubility.37 The diluent also modified the basicity, capacity, as well as the state of the extraction complex and the distribution of the solute.37 Figure 4 shows the recovery ratio of lithium, conversion of magnesium, and Mg/Li ratio for all tested diluents. The

Figure 5. Effect of amine/alcohol ratio on the recovery ratio of lithium, the conversion of magnesium, and the Mg/Li ratio. Initial brine concentration, Mg (50 g·L−1) and Li (2.5 g·L−1); 25 °C; 300 rpm; 120 min; CO2, 60 mL·min−1.

and the Mg/Li ratio as a function of different amine/alcohol ratios, where the abscissa of the graph is the volume percentage of TOA in the organic phase. The Li recovery ratio was consistently over 95% for various extractant concentrations. However, the magnesium conversion increased and then decreased with the volume percentage of TOA, and the maximum magnesium conversion reached 50.9% at TOA 50 vol %; the Mg/Li ratio decreased and then increased with the volume percentage of TOA, and the minimum Mg/Li ratio reached 9.79 also at TOA 50 vol %. The amine is the active component for extraction of protons from the aqueous phase; the extraction capacity increased with the TOA concentration. However, excess alcohol may prevent the formation of R3NH+, thereby reducing the proton extraction efficacy. On the other hand, a high TOA content also influenced the solvation of R3NH+.41 The alcohol improved the mobility of the organic phase and enhanced the extraction because the pure amine was very viscous and phase transfer of the protons limited the extraction rate. Therefore, the optimal TOA/alcohol ratio was about 50 vol %. 3.2.3. Effect of O/A Ratio. The volume ratio of the organic phase to the aqueous phase (O/A ratio) is also an important parameter in the ECM process. Figure 6 shows the results obtained with different O/A ratios ranging from 2 to 12. The recovery ratios of lithium were over 98% in all O/A ratio variance experiments, whereas the conversion of magnesium increased monotonically with the O/A ratio. The maximum

Figure 4. Effect of diluents on the recovery ratio of lithium, the conversion of magnesium, and the Mg/Li ratio. PA, 1-propanol; IPA, isopropyl alcohol; BA, butyl alcohol; IAA, isoamyl alcohol; HA, n-hexyl alcohol; CA, octanol; EHA, 2-ethylenehexanol; CHA, cyclohexane; MIBK, methyl isobutyl ketone; initial brine concentration, Mg (50 g· L−1) and Li (2.5 g·L−1); 25 °C; 300 rpm; 120 min; CO2, 60 mL·min−1.

recovery ratios of lithium with different diluents are consistently over 93%, while the Mg/Li ratio and magnesium conversion varies with different diluents. Although nonpolar diluents are widely employed in many industrial processes,38 nonpolar diluents, for example, cyclohexane (CHA) and methyl isobutyl ketone (MIBK), resulted in poor magnesium conversions of 2.87% and 2.37%, respectively. The poor solubility of the polar extraction complex in the organic phase is in accord with the similarity-intermiscibility theory.39,40 The alcohols exhibited the better performance in terms of mineralization of magnesium. The conversion of magnesium decreased from 65.32% (1propanol) to 36.75% (2-ethylenehexanol) varied with different alcohols, while the Mg/Li ratio showed an opposite trend, increasing from 6.88 (1-propanol) to 13 (2-ethylenehexanol). The conversion of magnesium decreased with the length of the carbon chain of the alcohol, which suggested that the E


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recovery ratio of lithium and the minimum Mg/Li ratio appeared at 5 °C with the value of 90.81% and 4.39, respectively. The lithium recovery ratio was an approximately constant 99% over 25 °C. The maximum Mg/Li ratio was 7.23 at 45 °C. Whereas the conversion of magnesium decreased with the reaction temperature increasing, the minimum conversion of magnesium was 64.90% at 45 °C and the maximum conversion of magnesium was 80.05% at 5 °C. With increasing temperature, the apparent basicity of the organic phase decreased and the molecular vibration became stronger, which should lead to a weaker N−H bond. At 5 °C, the conversion of magnesium was high, but the lithium loss was also as high as 9.19% (wt). Low temperature led to high apparent basicity and lower solubility of the carbonates; thus, the lithium loss was higher. Low vapor pressure and the low solubility of the organic phase in the aqueous phase at low temperature was favorable for recovery of the organic phase.14 The results suggest that the optimal temperature was 15−25 °C. 3.3. Mechanism of ECM. The ECM process (Figure 8) can be summarized in the following three steps: (1) CO2 was dissolved in the aqueous solution and was ionized in water to form H+ and CO32−. (2) HCl was extracted from the aqueous phase into the organic phase and promoted the reaction in the aqueous phase. (3) CO32− and Mg2+ formed MgCO3, which then precipitated. The second step is the key step of the ECM and limits the CO32− concentration. This step is influenced by the extractant as well as the brine composition. The organic amine extracts HCl rather than H2CO3 because the charge density of Cl¯ is higher than that of CO32−. Furthermore, CO32− has a thermochemical radius (0.189 ± 0.019 nm) that is larger than that of Cl¯ (0.168 ± 0.019 nm);25 the steric hindrance toward Cl¯ entering the organic phase was smaller than that of CO32−. The ion-pair effect is the major factor influencing extraction of HCl from the aqueous solution. When TOA is used as an extractant, the nitrogen atom connects with three C8 alkyls and the electron density of this nitrogen becomes larger because of the inductive effect. TOA extracts protons from the aqueous phase to form R3NH+, and Cl¯ enters the organic phase because of the ion pair effect. The diluent, isoamyl alcohol, forms a hydrogen bond with the nitrogen atom of TOA. The FT-IR data for isoamyl alcohol, TOA, and a blend of these components in equilibrium with water clearly shows the presence of a hydrogen bond (Figure 9). The FT-IR peaks at 3368 and 3351 cm−1 are ascribed to the O−H bond. Compared with the spectrum of pure isoamyl alcohol, the peak of the blend was shifted to 17 cm−1 lower, which suggests hydrogen bond transfer from H−O···H−O to N···H−O. When HCl enters the organic phase, the FT-IR peaks of the O−H bond shift to higher wavenumbers (Figure 10). Increasing the HCl concentration from 0.15 to 1.18 mol·L−1 resulted in the peak of the O−H bond moving from 3363 to 3401 cm−1. This proves the presence of a hydrogen bond between R3NH+ and O−H. In this process, the hydrogen bond is transferred from N···H−O to R3NH+···O−H. Comparison of Figures 9 and 10 shows that the R3N···H− O−R hydrogen bond could be considered to be substituted by

Figure 6. Effect of O/A on the recovery ratio of lithium, the conversion of magnesium, and the Mg/Li ratio. Initial brine concentration, Mg (50 g·L−1) and Li (2.5 g·L−1); TOA concentration, 50 vol %; 25 °C; 300 rpm; 120 min; CO2, 60 mL·min−1.

magnesium conversion was 70.01% at O/A ratio of 12. The Mg/Li ratio showed a monotonic decreasing trend, and the minimum Mg/Li ratio was 5.95 at O/A ratio of 12. A high O/A ratio means that the system possesses high proton extraction capacity, but the conversion tends to be constant in the high O/A ratio zone. On the other hand, the utilization ratio of the organic phase decreased with increasing O/A. Considering the utilization ratio of the organic phase, the Mg/Li ratio, and the magnesium conversion, the O/A ratio of 8 was the optimized condition; while under this ratio, the conversion was 67.41% and the Mg/Li ratio was 6.42 (5.40 of the raffinate). 3.2.4. Effect of Reaction Temperature. Temperature influences the CO2 dissolution, mass transfer, and reaction equilibrium. Figure 7 shows the influence of the temperature in the range of 5−45 °C on the conversion of magnesium and lithium recovery. The recovery ratio of lithium and the Mg/Li ratio increased below 25 °C and remained approximately constant higher than 25 °C with increasing reaction temperature; the minimum

Figure 7. Effect of reaction temperature on the recovery ratio of lithium, the conversion of magnesium, and the Mg/Li ratio. Initial brine concentration, Mg (50 g·L−1) and Li (2.5 g·L−1); TOA concentration, 50 vol %; O/A = 8; 300 rpm; CO2, 60 mL·min−1. F


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Figure 8. Mechanism of ECM.

Figure 9. FT-IR spectrum of isoamyl alcohol, TOA, and mixture of isoamyl alcohol and TOA.

atom in ammonium and the oxygen atom in the alcohol. This greatly increases the solubility of the extracted complex in the organic phase. Thus, alcohols are superior to ketones and alkyls as diluents. The FT-IR peaks in the range of 2700−2500 cm−142 (Figure 10) were assigned to the stretching bands of the N−H vibration. Although R3NH+ has the same atomic components, N−H undergoes hydrogen bonding with other molecules in a different manner based on the FT-IR peaks.43 The hydrogenbond interaction between R3N+ and water or isoamyl alcohol results in broadening of the N−H vibration bands. Therefore, with increasing H+ concentration, the intensity and breadth of the peaks became greater.

a hydrogen bond between the H in the amine and the O in the alcohol (R3NH+···O−H) during the extraction process. Three steps occur as follows: (1) breakage of the R3N···HO-R hydrogen bond; (2) formation of the R3N+-H ion by accepting a proton; (3) solvation in the organic phase. The first step is an endothermic process, whereas the second step is an exothermic process. The thermal effect of the third process depends on the solvent and solute. When nonpolar cyclohexane is used as the diluent, the solvation of R3NH+ is endothermic and nonspontaneous. Consequently, the nonpolar diluent did not enhance the ECM process. In contrast, the solvation of R3NH+ in the diluent is an exothermic process because of the formation of a R3NH+···O−H hydrogen bond between the hydrogen G


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Figure 10. FT-IR spectrum of the organic phase with different concentrations of HCl.

The peaks at 1633 cm−1 were assigned to the H−O−H bending vibration.44 With increasing H+ concentration, the intensity of the peak increased. From the literature, the bending vibration of H−O−H occurs at 1650 cm−1. The peak shift to 1633 cm−1 indicates the influence of the hydrogen bond between water and R3NH+. The peak at 1095 cm−1 is assigned to the C−N stretching vibration;45 with increasing H+ concentration, the peak of the C−N stretching vibration shifted from 1095 to 1122 cm−1. As shown Figure 10, the peak at 1095 cm−1 weakened, and that at 1122 cm−1 became more intense with the formation of the amine salt, which hinders the C−N stretching vibration. 3.4. Process Analysis. 3.4.1. Mineralization. An amine is a base and may react with MgCl2 to form Mg(OH)2 as a precipitate. To clarify the effect of CO2, a blank experiment was conducted to react brine with the organic extractant without CO2 bubbling. Under the employed conditions (O/A ratio of 4; TOA as an extractant; isoamyl alcohol as a diluent) using aqueous phase Mg2+ (50 g·L−1), a Mg/Li ratio of 20, and a temperature of 25 °C, no precipitate was observed in 120 min. However, under the same conditions, 50.9% magnesium was precipitated by bubbling CO2 at a rate of 60 mL·min−1. This demonstrates that precipitation occurred along with the mineralization of CO2 and that the effect of the organic phase was simply extraction of HCl from the aqueous phase. The precipitation solid of the ECM process was characterized by XRD as shown in Figure 11. When the results are compared with the standard XRD card from JCPDS, it demonstrates that the solid is mainly MgCO3·3H2O and a little of MgCO3. These results suggested that Li+ remained and was enriched as Mg2+ precipitated with the water. In this ECM process, the conversion of magnesium at the optimal condition was 67.41%, which is much lower than values reported in the literature (over 90%).34,46 According to reaction 5, the concentration of magnesium and the conversion of magnesium is decided by the concentration of CO32−. According to reactions 8, 9, and 10, H2CO3, HCO3, and CO32− coexist in aqueous solution and the major species are determined by the pH of the aqueous solution. The concentration of CO32− is too low to precipitate Mg2+ at pH less than 7. At the end of the reaction, the pH value was 6.58. According to equilibrium values of H2CO3, HCO3−, and CO32− with proton, and the reaction equilibrium constants25,28 k7, k8,

Figure 11. XRD patterns of the precipitate product of ECM.

k9, and k10, the calculated concentration of CO32− is 1.5 × 10−5 mol/L and the equilibrium concentration of magnesium is 0.45 mol/L (10.8 g/L). The concentration of magnesium in raffinate of ECM is tested to be 22.30 g/L (0.93 mol/L) at the optimal condition, and the activity coefficients of MgCl2 solution is about 0.544 according to the literature;25 therefore, the calculated activity of magnesium is 12.13 g/L (0.51 mol/L). The calculated magnesium concentration is close to the experimental value, and the difference could be attributed to experimental error. Theoretical and experimental results proved that pH does determine the magnesium concentration in raffinate in the ECM process. 3.4.2. Lithium Enrichment. Water, as reactant, was important but consumed in the ECM process (reaction 15), and it could be coextracted into organic phase47 and precipitated by MgCO3 as crystal water (MgCO3·3H2O). In the experimental process, the volume of raffinate after ECM was less than that of brine before ECM, so the lithium in raffinate was enriched. Under the optimal ECM conditions (25 °C, TOA 50 vol %, O/A ratio of 8, 50 g·L−1 Mg, and 2.5 g·L−1 Li), the volume of raffinate was about 12.1 mL and about 80% lithium stayed in raffinate; therefore, the concentration of lithium in raffinate increased from 2.5 to 4.13 g/L. H


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Industrial & Engineering Chemistry Research 3.4.2.1. Stripping. Regeneration of the loaded organic extractant was also evaluated. The HCl extracted by the amine is difficult to strip using water. The loaded organic phase was regenerated with aqueous ammonium. The stripping liquor could be directly used to produce NH4Cl. For the loaded organic phase obtained under the experimental conditions employing 50 g·L−1 Mg, a Mg/Li ratio of 20, an O/A ratio of 8, and temperature of 25 °C, stripping was conducted with different amounts of ammonium solution. The stripping reaction converted R3NH+Cl− to R3N and NH4Cl, and ammonium chloride was separated by stripping the aqueous phase. Different concentrations of aqueous ammonia (0.1, 0.35, 1, and 3 mol·L−1) and differrent phase ratios (0.1, 0.4, and 1) were investigated. Considering the utilization ratio of aqueous ammonium, 0.35 mol·L−1 aqueous ammonia was used as the stripping liquor under the following stripping conditions: A/O ratio of 0.4 and temperature of 25 °C; the stripping ratio of NH4Cl was 99.25%. 3.4.2.2. Subsequent Evaluation. When the Mg/Li ratio is less than 6, the soda precipitation technique can effectively be used to separate lithium and magnesium. For the mother brine, after magnesium precipitation by the ECM process, the raffinate meets the requirement of having a Mg/Li ratio of less than 6, and the recovery ratio of lithium was 99%. The traditional soda precipitation technique was also used to precipitate magnesium from liquors with Mg/Li ratio of 20 to 6, and the recovery ratio of lithium was 84%. After the ECM process, the soda consumption was reduced by 70.25% compared with that of the traditional soda precipitation technique. We introduced the following technological process (Scheme 1) to precipitate magnesium from brine to enrich the lithium content based on the above investigation.

precipitation, ECM with a proper stripping agent has a greatly reduced chemical cost.

4. CONCLUSION Mineralization with CO2 was introduced to reduce the magnesium content of high Mg/Li brine as an effective approach for purifying lithium from high magnesium salt lake mother brines. Thermodynamic analysis and experimental tests showed that magnesium in the brine can be removed by precipitation in the mineralization reaction employing extraction with organic amines. The precipitate was confirmed to be MgCO3·3H2O and MgCO3 by XRD analysis. The optimal mineralization parameters are 25 °C, TOA 50 vol %, and O/A ratio = 8 for the brine with 50 g·L−1 Mg and 2.5 g·L−1 Li. Under these conditions, the recovery ratio of lithium was 99%, the conversion of magnesium was 67.41%, and the Mg/Li ratio was reduced from 20 to 6.42 (5.40 of the raffinate). The mother brine processed by the ECM method can be directly used in the traditional soda precipitation technique, where the soda consumption was reduced by 70.25%. The mechanism of organic phase extraction of HCl involved the ion pair mechanism, where the loaded organic extractant can be regenerated by stripping with dilute ammonium solution. The ECM process includes three steps as follows: (1) CO2 reacts with water and produces H2CO3. (2) H2CO3 reacts with MgCl2 and produces MgCO3 and HCl. (3) HCl is extracted into the organic phase and enhances step 2. The water balance and stripping were also evaluated.

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ASSOCIATED CONTENT

* Supporting Information S

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.iecr.6b04892. Effect of reaction time on ECM process; material prices used in the evaluation (PDF)

Scheme 1. Schematic of ECM Process for Lithium Enrichment from Mother Brine

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AUTHOR INFORMATION

Corresponding Author

*Tel.: 8628-85997677. E-mail: [email protected]. ORCID

Siyang Tang: 0000-0003-4757-3992 Hairong Yue: 0000-0002-9558-0516 Notes

The authors declare no competing financial interest.

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The organic amine after extraction could be recycled by stripping, so the cost of this process includes power, equipment, and stripping costs. Because the ECM process was carried out under ambient temperature and pressure, without the strong acid consumed and generated, the power and equipment costs were low. Therefore, the stripping cost mainly constitutes the EMC cost. In this work, we employed aqueous ammonium as stripping agent, while other alkaline such as NaOH and CaO46 could also be used as stripping agent instead of aqueous ammonium. To dispose 67.41% of the MgCl2 in one cubic meter mother brine, 336.71 and 167.78 kg of the stripping agents aqueous ammonia (25%) and CaO (90%) were consumed, respectively. The costs are 319.87 RMB and 43.6 RMB, respectively, which is much lower than the cost of the traditional method using Na2CO3 (99%, 288.70 kg, 519.66 RMB) as precipitating agent (the prices of raw materials are listed in Table S1). Therefore, compared with soda

ACKNOWLEDGMENTS This work was financially supported by National Natural Science Foundation of China (NSFC 21336004). REFERENCES

(1) Zhou, P.; Tang, J.; Xiang, R. Prospect of Lithium Resources Supply and Demand. Acta Geol. Sin. (Engl. Ed.) 2014, 88, 287−288. (2) Choubey, P. K.; Kim, M.; Srivastava, R. R.; Lee, J.; Lee, J. Advance Review on the Exploitation of the Prominent Energy-storage Element: Lithium. Part I: From Mineral and Brine Resources. Miner. Eng. 2016, 89, 119−137. (3) Jaskula, B. W. Lithium. In U.S. Geological Survey, Mineral Commodity Summaries. https://minerals.usgs.gov/minerals/pubs/mcs/ (accessed December 8, 2016). (4) Harben, P. W.; Edwards, G. H. The Global Lithium Industry: A Portrait of Rapid Flux. JOM 1997, 49, 21−22.

I


Article

Industrial & Engineering Chemistry Research (5) Gao, D.; Yu, X.; Guo, Y.; Wang, S.; Liu, M.; Deng, T.; Chen, Y.; Belzile, N. Extraction of Lithium from Salt Lake Brine with Triisobutyl Phosphate in Ionic Liquid and Kerosene. Chem. Res. Chin. Univ. 2015, 31, 621−626. (6) Zhao, Z.; Si, X.; Liu, X.; He, L.; Liang, X. Li Extraction from High Mg/Li ratio Brine with LiFePO4/FePO4 as Electrode Materials. Hydrometallurgy 2013, 133, 75−83. (7) Ö zgür, C. Preparation and Characterization of LiMn2O4 Ion-sieve with High Li+ Adsorption rate by Ultrasonic Spray Pyrolysis. Solid State Ionics 2010, 181, 1425−1428. (8) Wang, L.; Ma, W.; Liu, R.; Li, H.; Meng, C. Correlation between Li+ Adsorption Capacity and the Preparation Conditions of Spinel Lithium Manganese Precursor. Solid State Ionics 2006, 177, 1421− 1428. (9) Kuz’min, V. I.; Gudkova, N. V. Extraction of Lithium Using TBP and the Noncoordinating Cation Exchanger Tetraphenylborate: Principles of Selectivity from Sodium and Higher-Valent Cations. Solvent Extr. Ion Exch. 2015, 33, 183−195. (10) Somrani, A.; Hamzaoui, A. H.; Pontie, M. Study on Lithium Separation from salt Lake Brines by Nanofiltration (NF) and Low Pressure Reverse Osmosis (LPRO). Desalination 2013, 317, 184−192. (11) Zeng, D.; Liu, H.; Chen, Q. Simulation and Prediction of Solubility Phase Diagram for the Separation of MgCl2 from LiCl Brine Using HCl as a Salting-out Agent. Hydrometallurgy 2007, 89, 21−31. (12) Yang, J.; Cheng, W.; Deng, T.; Gao, J. A Research on the Comprehensive Utilization of Intercrystalline Brines from East Taijinaier Salt Lake (Roasting-leaching separation). Inorg. Chem. Ind. 1996, 29−32. (13) Longpo, Y.; Shuang, L.; Hairong, Y.; Chun, L.; Bin, L.; Jiahua, Z.; Heping, X. Process Intensification by Coupling Reaction and Extraction for CO2 Mineralization in Ca2+-Rich solution. CIESC J. 2015, 66, 3511−3517. (14) Chen, F.; Wang, X.; Liu, W.; Liang, B.; Yue, H.; Li, C. Selective Extraction of Nitric and Acetic Acids from Etching Waste Acid Using N235 and MIBK mixtures. Sep. Purif. Technol. 2016, 169, 50−58. (15) Kesieme, U. K.; Aral, H.; Duke, M.; Milne, N.; Cheng, C. Y.; Kesieme, U. K.; Aral, H.; Milne, N. Recovery of Acid by Solvent Extraction from Mining Waste and Process Solutions. ALTA Nickel/ Cobalt/Copper Uranium and Gold, Perth, Australia, May 25−June 1, 2013. (16) Agrawal, A.; Kumari, S.; Ray, B. C.; Sahu, K. K. Extraction of Acid and Iron Values from Sulphate Waste Pickle Liquor of a Steel Industry by Solvent Extraction Route. Hydrometallurgy 2007, 88, 58− 66. (17) Kesieme, U. K.; Aral, H.; Duke, M.; Milne, N.; Cheng, C. Y. Recovery of Sulphuric Acid from Waste and Process Solutions Using Solvent Extraction. Hydrometallurgy 2013, 138, 14−20. (18) Jiang, X.; Yang, Y.; Sun, S.; Yin, Z.; Wang, X.; Bao, M. Microstructure Studies of Primary Amine N1923/n-heptane/Alcohol/ Water Reverse Micelles. J. Phys. Chem. B 1999, 103, 8657−8662. (19) Jiang, J.; Li, W.; Gao, H.; Wu, J. Extraction of Inorganic Acids with Neutral Phosphorus Extractants Based on a Reverse Micelle/ Microemulsion Mechanism. J. Colloid Interface Sci. 2003, 268, 208− 214. (20) Manaa, E.-S. A.; Lasheen, T. A. Liquid−Liquid Extraction of Sulfuric Acid from Aqueous Sulfate Waste Solution using Alamine 336/Kerosene/TBP Solvent. J. Dispersion Sci. Technol. 2016, 37, 137− 143. (21) Sarangi, K.; Padhan, E.; Sarma, P. V. R. B.; Park, K. H.; Das, R. P. Removal/Recovery of Hydrochloric Acid using Alamine 336, Aliquat 336, TBP and Cyanex 923. Hydrometallurgy 2006, 84, 125− 129. (22) Lum, K. H.; Stevens, G. W.; Kentish, S. E. The Modelling of Water and Hydrochloric Acid Extraction by Tri-n-butyl Phosphate. Chem. Eng. Sci. 2012, 84, 21−30. (23) Eyal, A. M.; Hazan, B.; Bloch, R. Recovery and Concentration of Strong Mineral Acids from Dilute Solutions through Llx. Ii.Reversible Extraction with Branched-Chain Amines. Solvent Extr. Ion Exch. 1991, 9, 211−222.

(24) Huang, H. The Technology Research on Magnesium-Lithium Separation of the Acidified Brine in West Taijnar Salt Lake of Qinghai. Master degree, Chengdu University of Technology, 2009. (25) Lide, D. R. CRC Handbook of Chemistry and Physics, 90th ed. (Internet Version 2010); CRC Press/Taylor and Francis: Boca Raton, FL, 2009. (26) Druckenmiller, M. L.; Maroto-Valer, M. M. Carbon Sequestration using Brine of Adjusted pH to form Mineral Carbonates. Fuel Process. Technol. 2005, 86, 1599−1614. (27) Zhou, Z.; Liang, F.; Qin, W.; Fei, W. Coupled Reaction and Solvent Extraction Process to form Li2CO3: Mechanism and Product Characterization. AIChE J. 2014, 60, 282−288. (28) Ellis, A. J. The System Na2CO3‑NaHCO3-CO2-H2O at Temperatures up to 200-Degrees. Am. J. Sci. 1959, 257, 10. (29) Hu, L.; Adeyiga, A. A. Kinetic Model of Gas−Liquid−Liquid Decomposition Extraction. Ind. Eng. Chem. Res. 2002, 41, 5584−5593. (30) Soong, Y.; Goodman, A. L.; McCarthy-Jones, J. R.; Baltrus, J. P. Experimental and Simulation Studies on Mineral Trapping of CO2 with Brine. Energy Convers. Manage. 2004, 45, 1845−1859. (31) Liu, Q.; Maroto-Valer, M. M. Studies of pH Buffer Systems to Promote Carbonate Formation for CO2 Sequestration in Brines. Fuel Process. Technol. 2012, 98, 6−13. (32) Azdarpour, A.; Asadullah, M.; Mohammadian, E.; Hamidi, H.; Junin, R.; Karaei, M. A. A Review on Carbon Dioxide Mineral Carbonation through pH-swing Process. Chem. Eng. J. 2015, 279, 615−630. (33) Steel, K. M.; Alizadehhesari, K.; Balucan, R. D.; Bašić, B. Conversion of CO2 into Mineral Carbonates using a Regenerable Buffer to Control Solution pH. Fuel 2013, 111, 40−47. (34) Wang, W.; Hu, M.; Zheng, Y.; Wang, P.; Ma, C. CO2Fixation in Ca2+-/Mg2+-Rich Aqueous Solutions through Enhanced Carbonate Precipitation. Ind. Eng. Chem. Res. 2011, 50, 8333−8339. (35) Balucan, R. D.; Steel, K. M. A Regenerable Precipitant-Solvent System for CO2 Mitigation and Metals Recovery. Int. J. Greenhouse Gas Control 2015, 42, 379−387. (36) Eyal, A. M.; Canari, R. pH Dependence of Carboxylic and Mineral Acid Extraction by Amine-Based Extractants: Effects of pKa, Amine Basicity, and Diluent Properties. Ind. Eng. Chem. Res. 1995, 34, 1789−1798. (37) Bízek, V.; Horácě k, J.; Koušová, M. Amine Extraction of Citric Acid: Effect of Diluent. Chem. Eng. Sci. 1993, 48, 1447−1457. (38) Nguyen, T. H.; Lee, M. S. Separation of Vanadium and Tungsten from Sodium Molybdate Solution by Solvent Extraction. Ind. Eng. Chem. Res. 2014, 53, 8608−8614. (39) Schunk, A.; Maurer, G. Distribution of Hydrochloric, Nitric, and Sulfuric Acids between Water and Organic Solutions of Tri-noctylamine. Fluid Phase Equilib. 2003, 207, 1−21. (40) Shan, X.; Qin, W.; Wang, S.; Dai, Y. Dependence of Extraction Equilibrium on Apparent Basicity of Extractant. Ind. Eng. Chem. Res. 2006, 45, 9075−9079. (41) Li, Y.; Song, X.; Chen, G.; Sun, Z.; Xu, Y.; Yu, J. Preparation of Calcium Carbonate and Hydrogen Chloride from Distiller Waste Based on Reactive Extraction−crystallization Process. Chem. Eng. J. 2015, 278, 55−61. (42) Huang, K.; Chen, J.; Zhang, C.; Li, Z.; Pan, Q.; Sun, Y.; Liu, S.; Liu, X.; Li, W.; Yang, Z.; Weng, S.; Hu, T.; Xu, Y.; Wu, J. Formation of Super-concentrated Hydrochloric Acid in the Third Phase in Tertiary Amine N235-PtCl62−-HCl System and its Influences on the Pt Microemulsion Extraction. Sci. China, Ser. B: Chem. 2009, 52, 1825− 1834. (43) Qiu, X.; Chang, Z.; Li, W.; Zhou, H.; Dong, B.; Qiao, L. Structural Analysis of N H···O in Viscoelastic Scum Formation during Solvent Extraction of Sulfuric Acid with Trioctylamine. Sep. Purif. Technol. 2012, 95, 196−201. (44) Qiu, X.; Chang, Z.; Zhou, H.; Li, W.; Dong, B. Effect of Hydroxyl on Formation of Viscoelastic Scum during Solvent Extraction of Sulfuric Acid with Trioctylamine. Sep. Purif. Technol. 2012, 86, 137−142. J


Article

Industrial & Engineering Chemistry Research (45) Tamada, J. A.; Kertes, A. S.; King, C. J. Extraction of Carboxylic Acids with amine Extractants. 1. Equilibria and Law of Mass Action Modeling. Ind. Eng. Chem. Res. 1990, 29, 1319−1326. (46) Wang, W.; Wang, M.; Liu, X.; Wang, P.; Xi, Z. Experiment and Optimization for Simultaneous Carbonation of Ca2+ and Mg2+ in A Two-phase System of Insoluble Diisobutylamine and Aqueous Solution. Sci. Rep. 2015, 5, 10862. (47) Li, Y.; Song, X.; Sun, S.; Xu, Y.; Yu, J. Extraction Equilibrium of Hydrochloric Acid at Low Concentrations between Water and N235 in Isoamyl Alcohol Solution: Experiments and Simulation. J. Chem. Eng. Data 2015, 60, 3000−3008.

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