Switchable polarity solvents as draw solutes for forward osmosis

Desalination 312 (2013) 124–129

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Switchable polarity solvents as draw solutes for forward osmosis Mark L. Stone, Cathy Rae, Frederick F. Stewart, Aaron D. Wilson ⁎ Idaho National Laboratory, P.O. Box 1625 MS 2208, Idaho Falls, ID 83415-2208, USA

H I G H L I G H T S ► ► ► ► ►

Switchable polarity solvents (SPS) are novel draw solutes for forward osmosis. SPS draws provide high osmotic strengths, > 13 Osm/kg. SPS draws provide high FO fluxes. SPS draws provide positive flux even at high feed concentrations (5.0 mol/kg NaCl). SPS draws are easily recycled and trace amounts of SPS removed with reverse osmosis.

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Article history: Received 18 April 2012 Received in revised form 4 July 2012 Accepted 21 July 2012 Available online 16 August 2012 Keywords: Forward osmosis Desalination Switchable polarity solvents Draw solution Osmotic pressure

a b s t r a c t Switchable polarity solvents (SPS), mixtures of carbon dioxide, water, and tertiary amines, are presented as viable forward osmosis (FO) draw solutes allowing a novel SPS FO process. In this study substantial osmotic strengths of SPS are measured with freezing point osmometry and were demonstrated to induce competitive fluxes at high salt concentrations on a laboratory-scale FO unit utilizing a flat sheet cellulose triacetate (CTA) membrane. Under the experimental conditions the SPS degrades the CTA membrane; however experiments with polyamide reverse osmosis (RO) membranes display stability towards SPS. Once the draw is diluted the major fraction of the switchable polarity solvent can be mechanically separated from the purified water after polar to nonpolar phase shift induced by introduction of 1 atm carbon dioxide to 1 atm of air or nitrogen with mild heating. Trace amounts of SPS can be removed from the separated water with RO in a process that avoids solution concentration polarization. The separated nonpolar phase can be regenerated to a full strength draw and recycled with the re-addition of 1 atm of carbon dioxide. © 2012 Elsevier B.V. All rights reserved.

1. Introduction Some geographic locations do not have sufficient native supplies of fresh water to meet the steadily increasing demands of residential expansion and agricultural, thermoelectric, mining, and energy production. This increased demand also highlights the need to recover water from sewage, agricultural waste, industrial, mining, and landfill leachates, and from oil and gas operations [1–3]. For all these water recovery processes there is an added desire to remove trace amounts of residential and agricultural pharmaceuticals and other persistent organic pollutants (POP) [4]. Other means of meeting fresh water demands could include an effective and economic process for recycling waste streams and accessing unrealized water sources such as seawater and saline aquifers. Nonincremental advances in water recovery processes are therefore required to successfully address these challenges. Several technologies are currently used to address these challenges. One is reverse osmosis (RO), but it has problems with fouling, requires pressures greater than 50 atm, and recovers only 35–50% of fresh water ⁎ Corresponding author. Tel.: +1 208 526 1103; fax: +1 208 526 8511. E-mail address: [email protected] (A.D. Wilson). 0011-9164/$ – see front matter © 2012 Elsevier B.V. All rights reserved. http://dx.doi.org/10.1016/j.desal.2012.07.034

from seawater while producing large volumes of waste brine [5,6]. Another is distillation, which is mature but heavily dependent on the cost of energy and thus not expected to advance further. Most distillation systems are estimated to cost twice as much as RO, except for mechanical vapor compression, which is comparable [7]. Forward osmosis (FO) has been considered for the industrial purification of water [8]. This technology moves water from a feed solution across a semipermeable membrane to a draw solution based on an osmotic pressure differential [8]. FO is a low energy method that draws a large portion of water from a feed, making it suitable for water contaminated with particulate and biological agents. The implementation of FO water purification has been limited by a few technical challenges, not the least of which is the draw solute. When using conventional draw solutes, an FO system must either incorporate the draw solute into the final product [9] or remove it through RO or a related process. For example Hydration Technology Innovations (HTI) sells desalination pouches that contain sugar and salts which act as a draw solute to move fresh water across a FO membrane from seawater, producing a potable energy drink. These methods which include the draw solute in their products are not paths to fresh water. A standing challenge to the success of industrially relevant FO is the development of a next generation draw solute that

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can provide significant osmotic pressure but be separated from water by means other than RO. Materials explored as next generation draw solutes include dissolved gasses, hydroscopic organics, and inorganic salts. Each of these can each be removed by distillation or precipitation with the addition of salt mixtures, pH adjustments, or temperature adjustments [10–17]. Much of this work has been previously reviewed [17]. More recently, highly hydrophilic nanoparticles that can be magnetically isolated have been investigated [18,19]. A mixture of ammonium hydrogen carbonate has been demonstrated as a draw solute that can achieve high water flux and subsequently decomposed (65 °C) into ammonia and carbon dioxide. The ammonia and carbon dioxide can then be regenerated to ammonium hydrogen carbonate for further use as a draw [17,20–22]. Each of these systems has serious drawbacks, such as limited osmotic pressure and flux rate, limited recyclability, loss of the draw solute through reverse diffusion, and overall system cost that have hindered their implementation. This work investigated switchable polarity solvents (SPSs) [23–26] as next generation FO draw solutes. The transition of SPSs from water miscibility to water immiscibility is dependent on the presence or absence of carbon dioxide, respectively, at ambient pressures as shown in Fig. 1. This remarkable SPS phenomenon had evaded discovery until recently, despite work with relevant chemicals at the commodity scale for many decades. The work that follows is the first attempt to apply SPS to forward osmosis.

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autosampler) GC with a column (Restek – Bellefonte, PA) XTI-5 (Crossbond 5% diphenyl–95% dimethyl polysiloxane) 30 m×0.25 mm ID×0.5 μm df. The sample was introduced with a: split ration of 30:1, pressure of 6.78 psi, total flow of 19.3 mL/min, column flow of 5.44 mL/min, initial oven temperature of 40 °C for 2 min, ramp of 8 °C/min to 75 °C, ramp of 70 °C/min to 275 °C hold for 1 min, injector temperature of 150 °C, and detector temperature of 250 °C. 2.2. Production of SPS draw solute Deionized water (206.1 g) and N(Me)2Cy (180.1 g, 1.42 mol) are placed in a bottle equipped with a gas diffuser and exposed steady stream of carbon dioxide for 3 h while being stirred with a magnetic stirring bar. The resulting product is a clear homogenous solution with a mass of 443.8 g, a solution viscosity of 20.5 cP at 21.3 °C, pH of 8.8 at 22 °C, and density of 1.08 g/mL at 21 °C. The concentration of 60.5 wt.% dimethylcyclohexylammonium hydrogen carbonate ([HN(Me)2Cy HCO3]) was calculated from the integrals of a quantitative 1H NMR in CD3OD using the alkyl peaks relative to the combination water and carbonate peak, with the ambient water concentration accounted for through the measurement of a CD3OD blank. The measured concentration is within error of the 59 wt.%, 7.6 m, or 3.4 M based on the one-to-one volume of water to N(Me)2Cy formulation upon which all further calculations are based. 2.3. Forward osmosis experiments

Fig. 1. A general example of an SPS from its nonpolar water immiscible form to its polar water miscible form.

2. Experimental methods 2.1. General House deionized water was used for this experiment. N, N-dimethylcyclohexylamine 99% (N(Me)2Cy) and anhydrous toluene 99.8% were obtained from Aldrich and used as received. All equipment was used in accordance with manufacturer specification unless stated otherwise. Freezing point depression osmometry was performed using an Advanced Instruments Inc. Model 3250 Osmometer. Viscosity measurements were made using the falling bob method with a Cambridge Applied systems VL4100 viscometer. Nuclear magnetic resonance (NMR) spectra were acquired on a Bruker Avance III 600 MHz spectrometer with a magnetic field strength of 14.093 T, corresponding to operating frequencies of 600.13 MHz (1H), and 152.92 MHz (13C). Gas chromatography was conducted according to two methods: • Method 1. The Solid Phase Microextraction (SPME) method was used for low concentration sample collection of dimethylcyclohexylamine (N(Me)2Cy). A 5 mL aliquot of sample was taken and put into a 20 mL serum bottle. Sodium hydroxide (~90 mg) was added to drive the amine into the headspace. The samples were allowed to sit for a minimum of 30 min to come to equilibrium. An 85 μm polyacrylate coated SPME fiber developed by Supelco (Bellefonte, PA) was exposed to the headspace for 5 min before injection into Hewlett Packard 5890 Series II GC fitted (Restek — Bellefonte, PA) XTI-5 (Crossbond 5% diphenyl–95% dimethyl polysiloxane) 30 m×0.25 mm ID×0.5 μm df column. The sample was introduced through a splitless injection at 15 psi column flow of 5.44 mL/min at an initial oven temperature of 40 °C for 2 min, a ramp of 8 °C/min to 100 °C, a ramp of 70 °C/min to 250 °C hold for 1 min, injector temperature of 250 °C, and detector temperature of 250 °C. • Method 2. A FID liquid injection method for higher concentrations of N(Me)2Cy and toluene, used a Hewlett Packard 6890 (with

A feed solution (~500 mL) of varying salt concentrations and a draw solution (~100 mL, 7.6 mol/kg [HN(Me)2Cy HCO3]) were passed across opposing sides of an FO membrane (HTI cartridge membrane 81118-ES-2, 0.00190 m2) in a channeled cross flow pattern cell; the membrane's active layer was oriented towards the feed solution and the support layer oriented towards the draw solution. The feed and draw solutions were maintained at a constant temperature of 30 °C and flow rate of 0.4 L/min. The draw solution was bathed in ambient flow of carbon dioxide (30 mL/min). The draw solution was placed on a scale and its mass measured with a computer every 20 s. Fresh membranes were soaked in deionized water for at least 30 min prior to their introduction to the flow cell. The membrane was exposed to the flowing feed solution and a small volume of pure water to bring it to the equilibrium temperature. To begin an experiment, an already flowing and temperature controlled draw solution was redirected to the membrane surface opposite the feed solution. The mass of the feed solution was given 1 to 3 min to stabilize before readings were obtained every 20 s over a 10 minute period. This 10 minute period was then used to calculate the initial flux. For example, with a draw solution (450 mL, 1 molal NaCl) between 60 and 660 s, the mass of the draw solution increased by 5.25 g for an initial flux of 16.6 L/M2 h. Reverse flow of the SPS into the feed was monitored through GC–FID SPME method. Diffusion of NaCl into the draw during long runs was measured by evaporating the water and [HN(Me)2Cy HCO3] from the solution in a vacuum oven and weighting the residual material. 2.4. Extraction of simulated persistent organic pollutant Toluene (95 μL, 82.5 mg) was dissolved in water (202.6 g) to which was added [HN(Me)2Cy HCO3] solution (113.3 g, 7.6 mol/kg, 0.35 mol) to form a 275 μg/mL solution of toluene. This solution was mixed for several minutes after which sodium hydroxide (19.2 g, 0.48 mol) was added, resulting in a phase separation into an aqueous (~296 g) and organic phase (~36 g). The organic phase represents 82% of the total N(Me)2Cy in the system with 8.2 g still contained in the aqueous phase. Samples were taken from each phase and analyzed with GC–FID liquid injection method, the aqueous phase contained 110 μg/mL toluene and the organic phase contained 1200 μg/mL toluene. This yields a partition coefficient logK(organic/aqueous) of 1.1.

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2.5. Separation of simulated persistent organic pollutant

3.5

Toluene (1.84 g) was added to N(Me)2Cy (6.86 g, 54 mmol) and was completely miscible. Toluene (1.83 g) was added to 7.6 mol/kg [HN(Me)2Cy HCO3] (17.11 g, 53 mmol) and was generally immiscible. The aqueous [HN(Me)2Cy HCO3] fraction contained ~0.3 wt.% toluene based on quantitative 1H NMR.

2.5

Osm/Kg

3

1 0

3.1. Osmotic strength and flux experiments In their polar forms, switchable polarity solvents form highly concentrated ionic solutions with substantial osmotic pressures that are well suited to act as a draw solute. An equal volume two phase N(Me)2Cy and water solution, when exposed to carbon dioxide (1 atm), forms a single phase 7.6 mol/kg (59 wt.%) solution of [HN(Me)2Cy HCO3] with a reasonable viscosity of 20.5 cP at 21.3 °C. While this high concentration is impressive, there are a number of SPSs that can achieve higher concentrations [25,27]. However, N(Me)2Cy is produced commercially for use as a polyurethane polymerization catalyst and more competitively priced as a fine chemical than any of the other known high concentration SPSs with good switching behavior. For the current demonstration, N(Me)2Cy is more than adequate, but the selection of SPSs for industrial FO application must consider a combination of material costs, maximum achievable osmotic pressure (concentration), SPS viscosity, propensity to reverse diffuse, toxicity, ease and completeness with which the SPS switches polarities, interaction with membranes, and other equipment used in the overall process. A fully concentrated 7.6 mol/kg (59 wt.%) [HN(Me)2Cy HCO3] solution is too concentrated to measure its osmolality through freezing point depression; however, diluted solutions can be measured as illustrated in Fig. 2. Extrapolation from this dilute data predicts that a fully concentrated 7.6 mol/kg [HN(Me)2Cy HCO3] solution has an ionic strength of 13.3 Osm/kg or ~325 atm [28]. The maximum solubility of sodium chloride is 6.14 mol/kg, assuming a van't Hoff factor of 2, a saturated NaCl solution would have an osmotic strength of 12.3 Osm/kg. In theory, the 7.6 mol/kg [HN(Me)2Cy HCO3] could extract water from a fully saturated brine solution resulting in the precipitation of NaCl. FO experiments were conducted to verify the substantial osmotic force of SPS in its water miscible form. Feed solutions of varying NaCl concentrations exposed to a 7.6 mol/kg [HN(Me)2Cy HCO3] solution draw the solution across an HTI cellulose triacetate (CTA) FO membrane. Fig. 3 presents the initial flux measured at each concentration. These initial results are promising with favorable flux observed through 5.0 mol/kg NaCl feed solution, which is 10-times the concentration of 35 salinity seawater. Extracting water from a 5.0 mol/kg NaCl feed solution corresponds to removal of 90% of the water from a typical seawater feed [17].

0

0.5

1

1.5

2

mol/Kg Fig. 2. Osmotic pressure of [HN(Me)2Cy HCO3] measured through freezing point depression experiments.

All experiments were run between 2 and 6 h, during that time immediate flux decreased at a rate that is greater than can be explained through the dilution of the draw or concentration of the feed. This decrease in flux is attributed to the degradation of the membrane, which increased with increasing concentration of NaCl feed. This degradation process may also influence the initial flux, causing relative reduction in the flux performance at higher NaCl feed than would be expected for system using an undamaged membrane. Based on these assumptions, studies with concentrations >5.0 mol/kg feed solution presumably damaged the membranes before a flux reading could be made. The salt rejection and reverse diffusion was only measured at the end of experiments, which lasted between 2 and 6 h. This makes it impossible to correlate the flux performance measured in the first 15 min of the experiment with salt rejection measured at the end of the experiment. Signs of membrane degradation and loss of salt rejection over the course of an experiment cast doubt on the reported flux performance. However, magnitudes of the initial observed fluxes are consistent with what would be expected from the draw solution's osmotic potential. Ultimately, the initial fluxes reported here may be found to be too low or too high once further studies are conducted with more robust membranes, but these studies provide a starting point for further work. Reverse diffusion, where the draw solute undesirably moves across the FO membrane into the feed solution, was also monitored. Experiments involving NaCl feed concentrations of 0.0 to 0.75 mol/kg were found to have less than b 0.5 wt.% N(Me)2Cy in the feed solution. Solutions with NaCl concentrations of ≥1.0 mol/kg were found to contain approximately 1.8 wt.% of N(Me)2Cy in the final feed solution after runs lasting 2 to 6 h. This data supports that the membrane degradation increases with the NaCl concentrations. The 1.8 wt.% is also near the solubility limit of N(Me)2Cy in water, which means experiments with the higher NaCl concentrations appear to have reached an equilibrium concentration of N(Me)2Cy. The CTA FO membrane restricts the movement of [HN(Me)2Cy HCO3] but allows the reverse diffusion of N(Me)2Cy until the feed solution is saturated. Fully concentrated [HN(Me)2Cy HCO3] solutions have a pH of 8.8 and the HN(Me)2Cy ion has an approximate pKa of 10.5. Hydrolysis of CTA membranes occurs outside the 4 to 7 pH range, at operating temperatures 35 30

flux(L/(M2h)

3. Results and discussion

2 1.5 0.5

2.6. Reverse osmosis removal of trace N(Me)2Cy N(Me)2Cy (68.4 g) was added to deionized water (4.095) kg and stirred overnight to form a homogenous 16.7 g/L solution. This solution was run through a Katadyn Survivor-6 seawater RO unit containing a polyamide thin-film composite DOW FILMTEC™ membrane. During the 4 h experiment, 3.925 kg of purified water was produced and divided into 17 samples of approximately 230 g each. The feed solution became turbid when the RO process started, and an organic layer was formed within the removal of the first 200 g of liquid. When the separation was complete, the remaining feed solution and first wash of the system contained organic layers totaling 67.5 mL (59.4 g). On average, the purified water contained less than 0.06 g/L N(Me)2Cy as measured by the GC–FID SPME method. The concentration did not significantly increase or decrease during the course of the experiment. This is a 99.65% rejection rate of the dissolved N(Me)2Cy.

y = 1.744x - 0.012 R² = 1.000

25 20 15 10 5 0

0

1

2

3

4

5

6

NaCl (mol/Kg) Fig. 3. Water flux using [HN(Me)2Cy HCO3] draw solution against different concentrations of sodium chloride at 30 °C.

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greater than 30 °C and through exposure of the membrane to a number of chemicals [6,29,30]. This well known failure route and limited salt rejection performance has motivated the RO community to explore materials beyond CTA such as cross-linked aromatic polyamide composites [30]. A second possible mechanism of failure, which could facilitate the reverse diffusion of N(Me)2Cy, is that it may dissolve micro supports and cause membrane swelling [30]. The polyamide thin-film composite (TFC) membranes being developed for FO systems [31–37] will likely solve many problems with CTA membranes for the SPS FO system; the successful use of N(Me)2Cy with polyamide TFC RO membranes is discussed below. Also, increasing the carbon dioxide pressure on the draw side would lower the solution pH and ensure that more of the N(Me)2Cy is protonated; these shifts would make the system more amenable to any membranes with sensitivity to organics and high pHs. Despite the ultimate failure of CTA membranes, this work demonstrates that substantial osmotic force and, presumably, water flux that can be achieved by the water miscible form of SPS. Fig. 4 provides the water flux rates for a desalination process that can be estimated by plotting the initial water flux data against the volume fraction of water removed from 0.5 molal NaCl, a simulant for 35 salinity seawater, at an experimentally measured concentration. Fitting a polynomial to this data allows for the calculation of definite integral and the determination of the weighted average water flux for the extraction of a total fraction of water, assuming the constant application of full concentration SPS draw. Using this weighted average method, it was found that the purification of 75% of the water from a seawater sample can be accomplished at an average flux of 18.8 L/(M2 h) which compares favorably with the 14–20 L/(M2 h) used in commercial seawater RO for the purification of a smaller fraction of the water (b50%) [6]. 35

flux (L/(M2h)

30

y = -55.184x 3 + 43.287x2 - 16.936x + 21.871 R² = 0.9966

25 20 15 10 5 0

0

0.2

0.4

0.6

0.8

1

vol fraction of water removed from a NaCl 0.5 mol/Kg solution Fig. 4. Water flux using HN(Me)2Cy HCO3 (7.6 m) draw solution against different concentrations solution of sodium chloride at 30 °C plotted against the fraction water removed from a 0.5 m NaCl solution.

3.2. Draw solute recycling Conventional FO systems act primarily as pretreatment methods, homogenizing a water stream for ultimate purification through RO,

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or they must include their solute in the final product. An SPS FO system is significantly different as shown in Fig. 5, where FO is the primary water purification method in which a miscibility shift liberates the majority of the water from the draw. Once a draw solution is saturated with water or reaches a desired dilution, the carbon dioxide is removed from the solution driving a shift in the SPS to its nonpolar immiscible form, which phase separates from the water. The two phases then can be mechanically separated, resulting in water purified to greater than 98 wt.% in a system using N(Me)2Cy (solubility 18 g/L). This switching process was demonstrated in Jessop's initial report with the best result found using low grade heat of 60 °C and a nitrogen purge [25]. Further investigation into the process has demonstrated that polar to nonpolar switch is a complex process that depends on the specific SPS, reaction chamber surface area, temperature, purge gas, pressure, and a number of other variables. However, the 60 °C with nitrogen purge uses significant energy, and the switching may be achievable at a much lower cost. The work associated with this research will be published at a later date. Trace amounts of SPS draw solute can be removed from the separated water through an RO process. This RO step is very different than the RO process involving a feed containing conventional draw solutes. During conventional salt draw RO, the volume of the feed solution is reduced as pure water is forced through the membrane, the reduction of volume can correspond to concentration polarization and fouling, which ultimately limits how much water can be obtained from the feed despite the application of pressures between 50 and 70 atm [6]. With SPS, when the solute is saturated and the volume of the RO feed solution reduced, the solubility limit is exceeded, resulting is increased solute phase separation from water as illustrated in Fig. 6. At its 18 g/L solubility limit, N(Me)2Cy has a maximum osmotic pressure of 0.14 Osm/kg. While this is a slight under-estimate of the actual osmotic pressure, residual amounts of the protonated N(Me)2Cy create far lower osmotic pressure as compared to such conventional draw solutes as NaCl (12.3 Osm/kg), KCl (9.2 Osm/kg), MgCl2 (17.12 Osm/kg), CaCl2 (20.1 Osm/kg), etc. Thus, low osmotic pressure aids RO by minimizing required energy and pressure. An RO experiment conducted on a 16.7 g/L solution of N(Me)2Cy with a polyamide thin-film composite membrane achieved a stable 99.65% rejection of the organic over a 4‐hour period to produce a solution with 0.06 g/L N(Me)2Cy concentration. A 0.06 g/L N(Me)2Cy concentration may be acceptable for some beneficial uses such as agriculture, microbial (algae) growth, solution mining, hydraulic fracturing water, and thermoelectric makeup water. There is no immediate toxicity concern from a 0.06 g/L solution at the oral LD50 rate of 425 mg/kg, but it is likely still too high for human consumption over the long term. The RO treated water will need to be purified through an additional polishing step. In large scale applications it may be possible to apply standard or specially adapted biological nitrogen removal methods, but there are many situations where biological treatment is not possible or optimal. For these later situations an inline polishing step will be required.

Fig. 5. Forward osmosis (FO) water purification scheme with an SPS draw solute.

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Fig. 6. Reverse osmosis (RO) removal of trace amounts of switchable polarity solvents (SPS) avoiding concentration polarization through phase separation of the SPS.

Fig. 7. Scheme for the removal of trace organics, such as POP, using the switchable polarity solvent (SPS) forward osmosis (FO) system.

3.3. Extraction of organics The work described thus far addresses high saline feeds, but a SPS FO system with the proper membrane is expected to also be effective against water containing trace or higher concentrations of organics with or without inorganic salts. Though the FO membrane is expected to block the majority of organics, any organic material that passes the FO membrane has the potential to be removed by the SPS draw cycling process. Throughout the proposed SPS FO process, the nonpolar water immiscible SPS is removed from an aqueous phase and recycled into its highly concentrated miscible form through the re-addition of carbon dioxide. The part of the process where the immiscible SPS separates from the aqueous phase can function as a solvent extraction process, as seen in the second vessel of Fig. 7, where “organic molecules” including any POP can be carried into the organic nonpolar SPS phase [4]. A consistent feature of POP is a water–octanol partition coefficient that favors the

organic phase (KOW = 3.0–8.2). Any POP that passed the FO membrane and entered the system would concentrate in the organic with each pass through the system. In the overall process, Fig. 8, a mechanical liquid separator could be placed between the gas contactor that produces the concentrated hydrophilic SPS and the FO cell. The separator would draw off any POP that became concentrated through repeated cycling in the closed loop draw system to the point of being immiscible with the hydrophilic SPS solution. Removal of model POP from the product water stream with SPS was demonstrated by dissolution of toluene (water solubility 470 μg/mL, KOW =2.69) into water and then mixing that water with concentrated miscible SPS to form a solution that was found to have a toluene concentration of 270 μg/mL. The SPS was switched to the immiscible form through the addition of NaOH to avoid the loss of solvent experienced in a gas purge. Samples were taken from each phase and analyzed with GC–FID where the aqueous phase contained 110 μg/mL and the organic phase contained 1200 μg/mL, yielding a partition co-efficient logK(organic/aqueous) of 1.1. The majority of POP have a higher KOW that would more dramatically favor the organic phase. The final separation of organics can be accomplished by switching the SPS back to its hydrophilic form, as shown in the third (right) vessel of Fig. 7. This was demonstrated by mixing toluene with the polar and nonpolar SPS forms. When toluene was mixed with N(Me)2Cy, it was completely miscible and only soluble in trace amounts (~0.3 wt.%) in a 7.6 mol/kg [HN(Me)2Cy HCO3] solution. Any organic that passes through the FO membrane will be removed through repeated cycling in the closed loop draw system combined with a SPS extraction process providing a system inherently effective at removing POP from a water feed. 3.4. Proposed SPS FO process The complete SPS FO system illustrated in Fig. 8 above is expected to require five major components: an FO draw cell, a polar to nonpolar switching cell, an RO cell, a finishing step (not presented in Fig. 8), and an SPS recycler. In addition to these major items, there will be various pumps, heat exchangers, holding tanks, and separators for the two phase solutions. The operating conditions for the FO cell are expected to be straight forward, but a more robust membrane is required. The RO cell and SPS recycler will require some engineering. The polar to nonpolar switching cell and a finishing step will both require development and optimization. Since this study is preliminary, additional work is required before this total process is demonstrated to a degree that performance, energy, and economic comparisons can be made to more developed next generation FO systems such as the ammonium carbonated system [17,22]. The SPS has the added benefit of higher osmotic strength for similar operating conditions and reduced back diffusion because of the greater molecular mass of SPS relative to ammonia.

CO2 Hydrophilic SPS

Gas Contactor

Feed Solution

Forward Concentrated Osmosis Material Membrane

Hydrophobic SPS

Mechanical Liquid Separator

CO 2 Separation

Fig. 8. Proposed SPS FO water purification process.

Fully Processed Water

Reverse Osmosis Membrane

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4. Conclusion SPS draw solutes have been found to enable an entirely new SPS FO process for purifying water and concentrating solutions. The osmotic pressure and FO water flux properties of SPS reported here are substantial, suggesting that high osmotic pressure feeds such as seawater, produced water, landfill leachate, industrially relevant waste streams, and brine solutions may be dewatered using this method. The novel phase switching from a high osmotic pressure aqueous solution to purified water and a nonpolar liquid, which can be physically separated, is a potentially revolutionary advance. The time, gas/vacuum, and heat required to achieve the switch and remove residual amine in the purified water has not yet been optimized, but initial work is encouraging. Future work will include refining the choice of tertiary amine and the effect this choice has on performing the other aspects of the scheme. The sensitivity of the commercially available FO membrane remains an issue that currently limits the use of this technology, but when the more advanced membranes currently under development become available, [31–37] the SPS FO system may have the capacity to re-concentrate diluted acids and bases as used in many industries, dewater solution mining extracts, and other systems in which the feeds are chemically aggressive. Acknowledgments This work was supported by the United States Department of Energy through contract DE-AC07-05ID14517. Funding was supplied by Idaho National Laboratory via the Laboratory Directed Research and Development Fund (LDRD) and the Battelle Memorial Institute through Intellectual Property Development Fund (IDF). A patent has been filed USPTO titled “Methods and Systems for Treating Liquids Using Switchable Solvents” application number 13/480,053. The authors thank Jeffrey R. McCutcheon for useful discussion and advice. The authors also acknowledge Hydration Technology Innovations for providing the membranes for this research. The authors thank Eastern Idaho Regional Medical Center for the use of their osmometer prior to the procurement of an osmometer at INL. The authors thank Chris Voxland at Katadyn North America for corresponds addressing their product. References [1] M. Elimelech, W.A. Phillip, The future of seawater desalination: energy, technology, and the environment, Science 333 (2011) 712–717. [2] M.A. Shannon, P.W. Bohn, M. Elimelech, J.G. Georgiadis, B.J. Mari|[ntilde]|as, A.M. Mayes, Science and technology for water purification in the coming decades, Nature 452 (2008) 301–310. [3] Q. Schiermeier, Water: purification with a pinch of salt, Nat. News 452 (2008) 260–261. [4] L. Ritter, K. Solomon, J. Forget, M. Stemeroff, C. O'leary, A review of selected persistent organic pollutants, International Programme on Chemical Safety (IPCS). PCS/95.39, World Health Organization, Geneva, 1995. [5] B. Van der Bruggen, L. Lejon, C. Vandecasteele, Reuse, treatment, and discharge of the concentrate of pressure-driven membrane processes, Environ. Sci. Technol. 37 (2003) 3733–3738. [6] Desalination of Seawater, American Water Works Association, 2011. [7] Hisham M. Ettouney, Hisham T. El-Dessouky, Ron S. Faibish, Peter J. Gowin, Evaluating the economics of desalination, Chem. Eng. Prog. (2002) 32–39.

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