Effect of draw solution concentration and operating conditions on forward osmosis and pressure retarded osmosis performance in a spiral wound module

Journal of Membrane Science 348 (2010) 298–309

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Effect of draw solution concentration and operating conditions on forward osmosis and pressure retarded osmosis performance in a spiral wound module Yuan Xu a,1 , Xiaoyu Peng a,1 , Chuyang Y. Tang b,c,∗ , Q. Shiang Fu d,∗∗ , Shengzhe Nie a a

Environmental and Clean Technology Laboratory, Suzhou Institute of Sichuan University, Suzhou, China School of Civil and Environmental Engineering, Nanyang Technological University, Singapore Singapore Membrane Technology Centre, Nanyang Technological University, Singapore d Nanjing Institute of Geography and Limnology, Chinese Academy of Sciences, Nanjing, China b c

a r t i c l e

i n f o

Article history: Received 22 July 2009 Received in revised form 8 November 2009 Accepted 9 November 2009 Available online 13 November 2009 Keywords: Forward osmosis (FO) Pressure retarded osmosis (PRO) Reverse osmosis (RO) Internal concentration polarization (ICP) Osmotic driving force

a b s t r a c t Forward osmosis (FO) and pressure retarded osmosis (PRO) are concentration-driven membrane processes. While they can be potentially used in water, wastewater, and energy applications, these processes suffer from the concentration polarization inside the porous membrane support resulting in severe flux decrease, a phenomenon known as internal concentration polarization (ICP). Researchers have investigated the effect of ICP both in theoretical and experimental studies. The current study extends the existing ICP model to include the effect of draw solution dilution by membrane permeate flow in a spiral wound FO module (SWFO). FO and PRO experiments were performed using a Hydrowell® SWFO under both submerged and cross-flow conditions. The effect of draw solution concentration, draw solution flow rate, feed water flow rate, and membrane orientation on FO and PRO water flux performance was systematically investigated. Permeate flow increased with greater draw solution concentration in both FO and PRO modes. ICP was found to drastically limit the available membrane flux in the concentration-driven membrane processes, and its adverse effect was more severe at greater draw solution concentration. Membrane flux was also affected by the dilution of draw solution when the permeate flow rate was comparable or greater than the draw solution flow rate. The submerged FO configuration performed nearly as good as the cross-flow configuration with feed water circulating outside of the membrane envelope (shorter flow path). In this case, the feed water flow rate only had limited effect on membrane flux likely due to its low mass transfer resistance. In contrary, the membrane flux can be adversely affected at low feed water flow rate when it was circulated inside of the membrane envelope (longer flow path). © 2009 Elsevier B.V. All rights reserved.

1. Introduction Membrane separation has gained increasing popularity in water, wastewater, and many other industrial applications. Traditionally, pressure-driven membrane processes, such as microfiltration (MF), ultrafiltration (UF), nanofiltration (NF), and reverse osmosis (RO), have received much attention. For example, lowpressure porous MF and UF membranes are widely used for surface water treatment [1,2], in membrane bioreactors (MBRs) [3,4], as well as for pretreatment of RO processes [5]. In parallel, the application of reverse osmosis (RO) is fuelled by the increasing

∗ Corresponding author at: Department of Civil and Environmental Engineering, Nanyang Technological University, Blk N1, Rm #1b-35, Singapore 639798. Tel.: +65 6790 5267; fax: +65 6791 0676. ∗∗ Corresponding author. E-mail addresses: [email protected] (C.Y. Tang), fu [email protected] (Q.S. Fu). 1 These author contributed equally to this study. 0376-7388/$ – see front matter © 2009 Elsevier B.V. All rights reserved. doi:10.1016/j.memsci.2009.11.013

needs for seawater desalination, water treatment, and wastewater reclamation [5,6], along with the significant improvements of RO membrane properties and module design [6,7]. On the other hand, although their full-scale implementation is still lacking, there is also increased interest in concentration-driven forward osmosis (FO) and pressure retarded osmosis (PRO) processes for water and energy applications due to their potential for energy production (or saving) [8–10]. During FO, a high concentration solution (draw solution) is separated from a low concentration solution by a semi-permeable (water permeable but not solute permeable) membrane (Fig. 1(a)). Due to the osmotic pressure difference across the membrane, water flows spontaneously through the membrane from the low concentration side to the draw solution side. This water permeating through the FO membrane is of high quality, with nearly complete retention of organic matter, particulates, and microorganisms as well as significant retention of dissolved salts [10–13]. Where a natural source of osmotic energy is available (e.g., seawater) [8,9,14], FO can be highly attractive since the process requires very little additional energy input (except a small amount of energy is

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Fig. 1. Illustration of FO, PRO, and RO processes. (a) FO process where no pressure is applied on the high concentration solution. Water flows from low concentration side to high concentration side. (b) PRO process where applied pressure on the high concentration solution is less than the osmotic pressure difference across the membrane. Water flows from low concentration side to high concentration side. (c) RO process where applied pressure on the high concentration solution is greater than the osmotic pressure difference across the membrane. Water flows from high concentration side to low concentration side. (d) Classification of FO, PRO, and RO in a flux versus pressure plot. Adapted from Refs. [8,10].

Fig. 2. Modeling internal concentration polarization. (a) Active layer is facing feed water. Draw solution concentration in the support layer is diluted by the water permeate flux, causing the available driving force (the concentration difference across the rejection layer) to be much lesser than the apparent driving force (concentration difference between the draw solution and the feed water). (b) Active layer is facing draw solution. The solute concentration at the rejection-support layer interface is much higher than the feed concentration, leading to reduced available driving force. Adapted from [8,10].

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required for recirculating the feedwater and draw solution) [10]. In other cases, a concentrated brine may be used as draw solution, but the diluted brine after the FO process has to be regenerated to complete a closed loop for the draw solution (e.g., via filtration or thermal processes) [15,16], which can be highly energy intensive. The PRO process is similar to the FO process, except an additional back pressure is applied on the draw solution (Fig. 1(b)). Water flows from low concentration side to the high concentration side as long as the applied pressure is lower than the osmotic pressure difference across the membrane [10]. When the applied pressure is greater than the osmotic pressure difference, however, the direction of water flux is reversed, which is the well known reverse osmosis condition (Fig. 1(c)). PRO, along with RO and FO, is summarized in a flux versus pressure plot as shown in Fig. 1(d). Unlike RO desalination where mechanical energy (pumping) is used to overcome the osmotic pressure of seawater, the PRO process converts the chemical potential (osmotic energy) of a concentrated draw solution into mechanical energy [8,17,18]. Loeb and coworkers [9,14,17,19] proposed the idea of PRO based osmotic power plant for electricity generation by running the pressurized permeate flow through a turbine generator. A Norwegian company in the clean energy sector, Statkraft, has built the world’s first prototype seawater osmotic power plant [20,21]. It is planning to commercialize osmotic power by year 2015 with a projected energy cost competitive against other renewable energy sources such as solar energy, offshore wind power, and hydroelectric power [20,21]. Despite the many potential applications [10], concentrationdriven FO and PRO processes are severely limited by internal concentration polarization (ICP) in the porous membrane support, i.e.: (i) the accumulation of rejected solute from low concentration feed when the dense active layer is oriented towards the draw solution (i.e., support facing the feed water) or (ii) the dilution of draw solution concentration when the active layer is facing the feed water (Fig. 2 and Refs. [8,22,23]). A direct consequence of ICP is a reduction of the membrane flux due to the significantly reduced concentration gradient across the membrane rejection film [23]. This was evident in view of the impractically low water flux when conventional reverse osmosis membranes were tested in FO and PRO modes, as a result of their thick membrane support [8,24,25]. Recently, Hydration Technology Inc. (HTI, Albany, Oregon) has developed a commercial cellulose triacetate membrane with a thin support (<50 ␮m in thickness), which showed superior flux performance and less ICP in the FO mode [12,13,25,26]. Many of the existing studies on the HTI membrane were performed for small flat sheet membrane coupons in the FO mode, while few studies are available in the PRO mode. To the authors’ best knowledge, systematic tests on spiral wound modules in both FO and PRO modes are not available. The purpose of the current study was to evaluate the effect of draw solution concentration and operational conditions on the permeate flux in FO and PRO processes. Filtration tests were performed with a spiral wound HTI membrane module using sodium chloride brine as a draw solution. The flux performance of the spiral wound module was also modeled via classical ICP models [8,24]. This study helps us to understand the role of various operational parameters on membrane flux performance in both FO and PRO modes.

of the dense rejection layer and the porous support layer (Csupport ) is significantly lower than the bulk draw solution concentration (Cdraw ) due to the convective transport of solute away from the support back to the draw solution (dilutive internal concentration polarization (DICP) [23]). Consequently, the available driving force (i.e., the concentration difference across the dense membrane rejection layer) is much smaller than the apparent concentration difference between the draw solution and the feed [8,10]. Similar ICP effect has also been reported when the draw solution is oriented towards the dense rejection layer (Fig. 2(b)), where solute from the feed solution and that transmitted through the membrane from the draw solution are accumulated and concentrated in the support layer [12,22,27]. In this case, the concentration in the support (Csupport ) is significantly higher than the feed water concentration (Cfeed ), which in turn reduces the available driving force (concentrative internal concentration polarization (CICP) [23]). The effect of internal concentration polarization can be modeled by adopting the classical solution-diffusion theory for the dense rejection layer coupled with convection and diffusion transport of the solute in the porous support layer [8,22,24]. Following Lee et al. [8] and Leob et al. [24] (refer to Appendix A), the water flux in FO can be expressed as



Jv = Km ln

Adraw + B Afeed + Jv + B



(FO, active layer facing feed water, DICP) and



Jv = Km ln

Adraw − Jv + B Afeed + B



(FO, active layer facing draw solution, CICP)

(2)

where Jv is the volumetric flux of water; A and B are the transport coefficients for water and solute, respectively; draw and feed are the osmotic pressure of the bulk draw solution and that of the bulk feed water, respectively. In Eqs. (1) and (2), Km is the mass transfer coefficient, which is given by the ratio of the solute diffusion coefficient (D) over the membrane structural parameter (S): Km =

D S

(3)

The S parameter is a property of the support structure. Its value is proportional to the thickness (t) and the tortuosity () of the support layer, but it is inversely proportional to the porosity (ε) of the support [20]. In a way, the S parameter in ICP, which provides a length scale of the concentration polarization in the support layer, is analogous to the boundary layer thickness in external concentration polarization. A larger structural parameter (e.g., thicker membrane support) leads to a lower mass transfer coefficient, which can significantly reduce permeate flux due to the exponential dependence of concentration polarization on Km . Eqs. (1) and (2) may be extended to include the effect of the applied pressure (P) in the PRO mode [8,28] (refer to the brief derivation in Appendix A as well):



Jv = Km ln 2. Theory

Adraw + B((AP/Jv ) + 1) (Afeed + Jv + AP) + B((AP/Jv ) + 1)



(PRO, active layer facing feed water) A unique feature in osmotically driven membrane processes (both FO and PRO) is that the direction of water flux is opposite to the direction of solute flux. This causes internal concentration polarization in the porous support layer. When the active rejection layer is facing the feed water (i.e., the porous support facing the draw solution, Fig. 2(a)), the solute concentration at the interface

(1)

 Jv = Km ln

(Adraw − AP − Jv ) + B((AP/Jv ) + 1) Afeed + B((AP/Jv ) + 1)

(PRO, active layer facing draw solution)

(4)



(5)

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Fig. 3. Experimental setup. (a) An illustration of an open Hydrowell® spiral wound module. (b) A cross-flow filtration test setup where the spiral wound membrane module was housed inside a pressure vessel. Cross-flows can be maintained both inside and outside of the membrane envelope. (c) Different test configurations adopted in the current study.

For the special case where P = 0, Eqs. (4) and (5) become identical to Eqs. (1) and (2), respectively. Eqs. (1), (2), (4), and (5) are valid for small membrane coupons where the membrane permeate flow (Qp ) is small compared to the flow rate of draw solution (i.e., the cross-flow of draw solution, Qcf ), e.g., in batch experiments with low recoveries. In a spiral wound module, however, the permeate flow can significantly dilute the bulk draw solution concentration, which results in reduced flux performance compared to small coupon tests. Such dilution effect needs to be explicitly accounted for in a spiral wound module. Consider a module with a draw solution concentration Cdraw,0 and a cross-flow of Qcf,0 at the module inlet. The bulk draw solution concentration Cdraw,x anywhere inside the module can be determined by Cdraw,x = Cdraw,0

Qcf,0 Qcf,0 +

x

(6)

J dAm 0 v

x

where Am is the membrane area, and the term 0 Jv dAm represents the accumulative permeate flow rate from the module inlet to the location of interest (x). Thus, the diluted bulk concentration Cdraw,x is simply related to the original draw solution concentration Cdraw,0

x

via a volumetric dilution factor (Qcf,0 /(Qcf,0 + 0 Jv dAm )). At the outlet of the module, the draw solution is diluted to a concentration

Cdraw,exit : Cdraw,exit = Cdraw,0

Qcf,0 Qcf,0 + Qp

(7)

where Qp is the total permeate flow rate through the entire membrane module, i.e.:

 Qp =

exit

Jv dAm

(8)

0

The average permeate flux in the module Jv is given by Jv =

Qp Am

(9)

Eq. (6) can be used to determine the bulk draw solution concentration Cdraw anywhere in the module, and Cdraw can then be used to determine the local membrane flux Jv via the FO and/or PRO flux models (Eqs. (1), (2), (4), or (5)). Thus, Eq. (6), coupled with the ICP flux models, can be used for flux performance modeling of spiral wound FO/PRO modules. A numerical example is provided in Appendix A. For PRO applications, the power output is of interest. The total power available from the module (˚PRO ) can be determined as a

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function of membrane permeate flow rate [8]: ˚PRO = PQp

(10)

The power density, that is, the total power normalized by the membrane area in the module, is given by PRO = PJv

(11)

cross-flow rate at the feed solution side. A separate pump was used to circulate the draw solution from a 40-L stainless steel pressure tank. The pressure inside the draw solution tank was controlled by adjusting a needle valve located at the top of the tank, and it was measured by a mechanical pressure gauge (Fig. 3(b)). Where no back pressure was required in the FO mode, the needle valve was fully open so that the pressure inside the draw solution tank reached atmospheric pressure.

3. Materials and methods 3.1. Chemicals Unless otherwise specified, all reagents and chemicals were of analytical grade. Deionized water was used throughout the experiments. Analytical grade sodium chloride was used to prepare draw solutions of various concentrations. 3.2. Membrane material and module configuration The commercially available Hydrowell® membrane elements were purchased from HTI [29]. This membrane was specially developed for FO applications with an ultrathin polyester screen mesh support (comprising of polyester fibers ∼20 ␮m in diameter arranged orthogonally at a fiber-to-fiber spacing of 120 ␮m, overall thickness < 50 ␮m) to reduce internal concentration polarization [10,12]. According to the membrane supplier, the active rejection layer of the HTI membrane is based on cellulose triacetate, with a sodium chloride rejection of 93–95%. The water permeability of HTI membrane is 2.2 × 10−12 m/s Pa [12], which is significantly lower than that for typical brackish water RO membranes (∼1–2 × 10−11 m/s Pa with NaCl rejection > 95%) but comparable to typical seawater RO membranes (e.g., 3 × 10−12 m/s Pa for SWC4) [30,31]. Despite of the relatively poorer separation properties of the HTI membrane in comparison to commercial RO membranes, the HTI membrane demonstrated much superior flux performance in the FO mode, which was attributed to the drastically lower ICP due its special thin support layer design (Section 2 and Refs. [12,25]). The HTI membrane elements were supplied in the spiral wound configuration (Fig. 3(a)). Each element comprises of two pieces of flat sheet membranes (∼225 cm in length and 30 cm in width). These membrane sheets are to be glued face to face on three edges to form a membrane envelope (i.e., a membrane leaf). The remaining edge of the envelope is connected to a central tube with holes for water collection. The active layers of the membranes are on the outer sides of the envelope. The membrane leaf is rolled into a spiral wound configuration, and spacers (diamond pattern, ∼0.8 mm in thickness and ∼2.5 mm spacing) are used both inside and outside of the membrane envelope for maintaining flow channels. Unlike a typical spiral wound RO element, the Hydrowell® element has additional glue lines along the central line of the membranes (Fig. 3(a)), and the center of the water collection tube is also plugged. This design forces two independent cross-flows: (1) one within the membrane envelope (along the length of the membrane leaf), and (2) the other one outside the envelope (along the width of the membrane leaf). Each Hydrowell® element has a total active membrane area of ∼0.94 m2 (excluding glued membrane area). 3.3. FO and PRO experiments 3.3.1. Test setup To allow the application of hydraulic pressure as required for PRO tests, the Hydrowell® element was housed in a pressure vessel self-manufactured with a maximum pressure rating of 30 bar (Fig. 3(b)). The pressure vessel was connected to a feed water tank (unpressurized) and a circulation pump was used for controlling the

3.3.2. Modes of testing Three modes of testing were compared in this study (Fig. 3(c)): (1) submerged FO mode (s-FO), (2) FO mode with cross-flows for both feed water and draw solution (x-FO), and (3) PRO mode with cross-flows for both feed water and pressurized draw solution (xPRO). In the submerged FO mode (s-FO-in), the pressure vessel was not used and the Hydrowell® module was submerged directly and completely in the feed water tank. The high-concentration draw solution was pumped into the central collection tube (Fig. 3(a)) and was circulated inside the membrane envelope. In this configuration, the draw solution was in direct contact with the porous support layer, while the feed water reached the active sides of the membranes via the spacer outside of the membrane envelope. Feed water was drawn through the HTI membrane under the concentration gradient across membrane. This configuration was somewhat similar to that in a submerged membrane bioreactor in that no circulation pump was used on the feed water side. Aeration was not used due to the thin membrane spacer. The s-FO configuration has been used for water production for emergency uses (military applications, disaster relief, etc.) [10], as originally advertised for Hydrowell® elements by HTI [29]. In the cross-flow FO mode (x-FO in Fig. 3(c)), the membrane module was housed inside the pressure vessel although no pressure was applied in the module by fully open the pressure control valve of the draw solution tank. Both feed water and draw solution were circulated via two circulation pumps (Fig. 3(b)), with the draw solution either inside the membrane envelope in direct contact with the porous membrane support (x-FO-in configuration) or outside of the envelope (x-FO-out configuration) in direct contact with the active membrane layer. The cross-flow PRO mode (x-PRO) had a similar configuration to that of x-FO, except a back pressure was exerted on the draw solution side. The x-FO mode may find potential applications for water treatment and desalination, while the x-PRO mode has been suggested for osmotic power harvesting [10,32,33]. 3.3.3. Experimental conditions Most of the experimental work was performed using a single SWFO element. However, several SWFO elements were used for x-PRO-in mode due to the resulting damage of glue lines (refer to Section 4.3) under the high pressure in this test mode. All filtration tests were performed at room temperature (22–24 ◦ C). Before each experiment, the SWFO module was thoroughly rinsed and flushed with deionized water on both sides of the membrane envelope. The volumetric flow rate through the Hydrowell® module was determined by measuring the time needed to fill up a volumetric cylinder (20 ml or 50 ml). The flux was calculated by normalizing the volumetric flow rate by the effective membrane area according to Eq. (9). Combination of different testing modes and membrane orientation (s-FO-in, x-FO-in, x-FO-out, xPRO-in, and x-PRO-out, refer to Fig. 3(c)) were evaluated. For each testing mode, different draw solution concentrations and the cross-flow velocities were evaluated. In PRO tests, the effect of applied pressure on membrane flux performance was also monitored.

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Fig. 4. Effect of draw solution concentration on permeate flow in s-FO-in configuration. The error bars shown in the figure represent one standard deviation based on at least three different tests. Note: 1 L/m2 h = 2.78 × 10−7 m/s.

4. Results and discussions 4.1. s-FO configuration Fig. 4 presents the permeate flow rate of the Hydrowell® element in the submerged FO mode (s-FO-in) as a function of draw solution concentration at different circulation rate of the draw solution (50–190 ml/min, corresponding to a cross-flow velocity of 1–4 cm/s at the SWFO inlet). Clearly, greater permeate flow rate was observed at higher draw solution concentration. This was due to the increased driving force (osmotic pressure difference in this case) across the membrane. However, unlike the classical solution-diffusion model which predicts that the water flux Jv in the absence of applied pressure shall be directly proportional to the osmotic pressure difference, i.e., Jv = A(draw − feed ), the experimental flux was highly non-linear with respect to the draw solution concentration (Cdraw,0 ). For relatively dilute draw solutions (Cdraw,0 < 0.5 M), significant flux enhancement was achieved at increased concentration. A plateau was reached at higher Cdraw,0 (>0.6 M). The maximum permeate flow rate was ∼100 ml/min, corresponding to a permeate flux of ∼6 L/m2 h. Similar non-linear flux behavior has been reported for FO by several research groups [8,12,13,34], which was attributed to internal concentration polarization in the porous support layer. As explained in Section 2 and Fig. 2(a), the solute concentration in the support layer (Csupport ) was significantly lower than the bulk concentration in the draw solution, which led to a drastic reduction of the available driving force. This simple ICP model [8,10,12] also predicts, in consistency with our experimental observation, that: (1) Flux behavior is less affected by ICP at lower membrane flux and (2) At high permeate flux, ICP plays an increasingly dominant role due to its exponential dependence on flux (Eq. (1)). Under this condition, any further increase in the draw solution concentration is offset by a much more severe ICP, resulting in less effective flux enhancement. In addition to the ICP effect, external concentration polarization (ECP) may also play an important role at higher draw solution concentrations in s-FO-in mode. This is due to the lack of forced recirculation at the feed solution side in this mode. The ECP in s-FO-in mode can be potentially controlled by bubbling (not

Fig. 5. Comparison of flux performances in FO and RO modes. (a) Permeate flux of RO, FO coupon, and FO module (s-FO-in configuration). Results of RO and FO coupon were obtained from reference [12]. The s-FO-in experimental results were obtained at a draw solution cross-flow rate of 100 ml/min. The s-FO-in simulation results were obtained from Eq. (6) coupled with ICP models (Section 2 and Appendix A). (b) FO over RO flux ratio as a function of draw solution concentration. Note: 1 L/m2 h = 2.78 × 10−7 m/s.

studied in this work), similar to that for submerged membrane bioreactors. It is worthwhile to note that the flux achieved in s-FO-in mode was relatively low (5–7 L/m2 h for a 0.5 M NaCl draw solution (corresponding to an osmotic pressure difference of ∼22 bar), depending on the cross-flow rate). In contrast, significantly higher membrane flux can be achieved if the same HTI membrane were tested in the RO mode (Fig. 5(a) and Ref. [12]). She et al. [12] reported a water permeability of 2.2 × 10−12 m/s Pa for the HTI membrane. This corresponds to a flux of 35.6 L/m2 h in the RO mode at an applied pressure of 45 bar, an order of magnitude larger than a flux of 6.6 L/m2 h observed in the s-FO-in mode under the same apparent driving force (1 M draw solution, osmotic pressure = 45 bar). Such comparison clearly confirms that the FO process performed less efficiently compared to RO under the same driving force, as a result of internal concentration polarization in FO. The flux efficiency of FO may be define by the ratio of FO flux over RO flux at identical apparent driving force [34]. This ratio is shown in Fig. 5(b) as a function of draw solution concentration. The FO flux efficiency was drastically reduced at greater draw solution concentration, which is a clear indication that ICP was more severe under larger draw solution concentration. This observation is consistent

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Fig. 6. Effect of draw solution cross-flow on permeate flow in s-FO-in configuration. Both experimental (solid symbols) and simulation results (continuous lines) were shown. The s-FO-in simulation results were obtained from Eq. (6) coupled with ICP models (Section 2 and Appendix A). The error bars shown in the figure represent one standard deviation based on at least three different tests. Note: 1 L/m2 h = 2.78 × 10−7 m/s.

with the previous findings by McGinnis and Elimelech [35], who reported lower flux efficiency at greater ammonia carbonate draw solution concentrations. It is also worthwhile to compare the flux performance of the spiral wound FO module (in s-FO-in configuration) to that of a flat FO coupon (Fig. 5(a)). The latter was measured in a cross-flow FO testing cell using a small HTI membrane coupon (∼60 cm2 ) cut from a Hydrowell® module [12]. In both cases, the porous support was orientated towards the draw solution. Interestingly, the spiral wound module (s-FO-in configuration) had much lower membrane flux at comparable draw solution concentrations. This might be explained by the major difference between testing in a spiral wound configuration versus testing a small-area flat coupon—the bulk concentration of the draw solution Cdraw was significantly diluted by a sizable permeate flow in the spiral wound module (Eq. (6)). Consider a cross-flow rate of 50 ml/min at the inlet to the module and a permeate flow of ∼100 ml/min, the total cross-flow rate was increased to 150 ml/min with Cdraw diluted by a factor of 3 at the module outlet. This is somewhat analogous to testing an RO spiral wound module where brine concentration increases from module inlet to outlet as pure water permeate through the membrane, which tends to reduce the flux performance of the module at high recovery in comparison to that of a small coupon (recovery ∼0). In the present case, the low flux in the s-FO-in configuration (Fig. 4 and Fig. 5(a)) was likely due to ICP as well as dilution of the bulk draw solution concentration. Model simulation was performed in accordance to Appendix A, where the dilution effect by the permeate flux was simulated by Eq. (6), and the ICP effect was determined by Eq. (1). The experimental results agreed very well with the simulation results at lower draw solution concentrations (Fig. 5(a)). On the other hand, the experimental flux was slightly lower than the theoretical prediction for Cdraw,0 above 0.8 M. This discrepancy at high draw solution concentration might be due to the mass transfer resistance and ECP on the feed water side which became increasingly important at high permeate flow (i.e., higher Cdraw,0 ) as forced feed water circulation was not employed in the submerged configuration (refer to Section 4.2 for additional discussion). The effect of draw solution cross-flow rate on submerged FO flux performance at a fixed draw solution concentration is shown in Fig. 6. At low draw solution flow rate, increasing cross-flow resulted in substantial permeate flux enhancement. Such enhancement

effect became less prominent at relatively high cross-flow rate (>100 ml/min). Flux enhancement at greater cross-flow velocity has been well documented for pressure-driven membrane processes (e.g., RO) thanks to the reduced external concentration polarization and membrane fouling as a result of the increased mass transfer coefficient [36,37]. However, the internal concentration polarization is less likely to be affected by cross-flow velocity as the salt accumulation (or dilution) occurred in the porous membrane support which acts like an unstirred layer [38,39]. For example, She et al. [12] tested flat sheet HTI membrane coupons in the FO mode, and they reported only a marginal increase (<5%) in FO flux when the cross-flow rate was doubled. Therefore, the increased FO flux at greater cross-flow rate in the current study cannot be solely attributed to the enhancement of mass transfer at the liquid–membrane interface on the draw solution side. A more plausible explanation is the dilution of bulk concentration of the draw solution (Eq. (6)). As discussed in Section 2, the draw solution concentration in the spiral wound module became significantly diluted by the permeate flow. The dilution factor, i.e., the ratio of the inlet draw solution concentration to the outlet concentration, is given by (Qp + Qcf,0 )/Qcf,0 , where Qp and Qcf,0 are the total permeate flow rate and the inlet cross-flow rate, respectively (refer to Eq. (7)). Thus, excessive dilution of draw solution (large dilution factor) may occur at low draw solution cross-flow rate, which was likely responsible for the reduced FO flux. The dilution factor decreased at greater cross-flow, corresponding to enhanced flux performance (Fig. 6). However, at relatively high cross-flow rate (draw solution flow rate  permeate flow rate), further increase in cross-flow became less effective in reducing the dilution factor, which is consistent with the experimental observation of reduced effectiveness of cross-flow. Fig. 6 also presents the simulation results for various draw solution cross-flows and concentrations based on Eq. (1) (ICP effect) and Eq. (6) (dilution effect). The simulated results agreed well with the experimental ones, which confirms that the dependence of the permeate flow of FO module on draw solution cross-flow was mainly caused by the dilution of bulk draw solution concentration. Based on the above discussion, the cross-flow in a spiral wound FO module needs to be carefully optimized. Permeate flux can be adversely affected due to dilution effect if the draw solution flow rate is relatively low in comparison to the permeate flow rate. On the other hand, high cross-flow of draw solution demands greater pumping energy, yet its effect on flux enhancement may be limited. A balance of pumping energy and flux performance shall be achieved at the optimal cross-flow. In Fig. 6, the lines corresponding to dilution factors of 3, 2 and 1.5, respectively, are indicated. For example, a dilution factor of 2 is represented by a line with 1:1 slope in a permeate flow rate versus cross-flow plot. It seems that substantial flux reduction occurred when the dilution factor was greater than 2. Such conditions shall be avoided to maintain an acceptable FO flux. 4.2. x-FO configurations Two cross-flow FO configurations (x-FO-in and x-FO-out) were tested in the current study, and their flux performance is compared to the submerged FO configuration in Fig. 7. For all the three configurations, permeate flow increased at higher draw solution concentration as a result of increased driving force. In addition, the slope of each curve reduced as concentration increased. Once again, this is consistent with Eqs. (1) and (2) that ICP is more severe at increased flux level and greater draw solution concentration due to its exponential dependence on permeate flux. In Fig. 7, the flux performance of x-FO-in was very similar to that of s-FO-in, except slightly higher flux was achieved in the cross-flow configuration at higher draw solution concentrations

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Fig. 7. Comparison between s-FO-in, x-FO-in, and x-FO-out configurations. Draw solution flow rate at 75 ml/min, and feed water flow rate at 140 ml/min (where applicable). Simulation results were obtained from Eq. (6) coupled with ICP models (Section 2 and Appendix A). The error bars shown in the figure represent one standard deviation based on at least three different tests. Note: 1 L/m2 h = 2.78 × 10−7 m/s.

(Cdraw,0 > 0.5 M). In both configurations, the draw solution was circulated at a rate of 75 ml/min. While a cross-flow of 140 ml/min pure water was circulated on the feed water side for the x-FO-in configuration, no forced circulation was available for submerged one. Thus, the slightly better flux in x-FO-in configuration at greater Cdraw,0 was likely due to the reduced mass transfer resistance and ECP on the feed water side, which is consistent with our earlier discussion in Section 4.1. For comparison purpose, simulation was performed based on Eqs. (1) and (6). It is worthwhile to note that the simple theoretical model in the current study does not take into consideration the feed water cross-flow by assuming that the membrane flux is not limited by the mass transfer and ECP on the feed water side. Therefore, identical simulation results were obtained for both x-FO-in and s-FO-in configurations. A comparison between the simulation results and the experimental data shows (1) that the model prediction agreed better with the experimental results at lower draw solution concentration; and (2) that the model simulation worked better for the x-FO-in configuration. Both trends can be explained by the lower mass transfer resistance and ECP on the feed water side under these conditions. The permeate flow of the x-FO-out configuration is also presented in Fig. 7. This configuration (where the dense active layer was facing the draw solution) had higher flux compared to the s-FO-in and x-FO-in configurations where active layer was facing the feed water. Similar results have been reported by various research groups [12,13,22]. Such difference was likely due to the more severe ICP in the s-FO-in and x-FO-in configurations, where a dilutive concentration polarization occurred in the support (i.e., the draw solution was diluted inside the porous support layer to cause a drastic loss of driving force, Fig. 2(a)). This explanation was also confirmed by the simulation results based on Eqs. (2) and (6)—the simulated permeate flow for x-FO-out is clearly superior to those for x-FO-in and s-FO-in (Fig. 7). Surprisingly, the experimental results in the x-FO-out configuration were significantly lower compared to the simulated results (Fig. 7). It is hypothesized that the mass transfer resistance on the feed water side was much greater in this configuration due to its much longer flow path. The flow path of feed water in the x-FOout configuration was ∼4.5 m (twice the length of the membrane leaf (Fig. 3(a)) as compared to a 0.3 m flow path length for x-FO-in where feed water flowed outside of the membrane envelope. The x-

305

Fig. 8. Effect of feed water flow rate on permeate flow in the x-FO-out configuration. Simulation results were obtained from Eq. (6) coupled with ICP models (Section 2 and Appendix A). Note: 1 L/m2 h = 2.78 × 10−7 m/s.

FO-out mode could result in more severe salt accumulation and ECP on the feed water side inside the membrane envelope. The effect of feed water flow rate on module performance in x-FO-out configuration is better illustrated in Fig. 8 for different feed and draw flow rates. At a draw solution cross-flow of 75 ml/min, the permeate flow at a feed flow rate of 60 ml/min was drastically lower than that at 140 ml/min feed water flow. Indeed, the maximum permeate flow rate (60 ml/min) was identical to the feed flow rate at the spiral wound module inlet. When this maximum permeate flow was achieved (which corresponds to 0.5 and 0.8 M draw solutions), the feed water flow at the module outlet was almost zero. Under this condition, nearly all the available feed water passed through the membrane resulting in a limitation of the maximum permeate flow to this amount. The same was true for the combination of a 140 ml/min draw solution and a 140 ml/min feed water. Fig. 8 clearly suggests that the permeate flow in the x-FO-out configuration was constrained by the feed water flow rate in addition to the ICP effect and the draw solution dilution effect. In contrast, the effect of feed water flow rate in the x-FO-in configuration was much less important. The flux performance of s-FO-in (where there was no forced circulation of feed water) was nearly identical to that of x-FO-in, as a result of significantly shorter flow path of feed water. 4.3. x-PRO configurations The PRO performance of the Hydrowell® spiral wound module was evaluated in both x-PRO-in and x-PRO-out configurations (Fig. 9). While the permeate flow in the x-PRO-out configuration seems to be more stable, the flux performance in x-PRO-in deteriorated drastically upon applying a back pressure. In the x-PRO-in configuration, the draw solution that flowed inside the membrane envelope was pressurized. Major leakage of draw solution was identified at a pressure of ∼3 bar which forced the experiment to be stopped. The leakage was due to the failure of the glue between the two pieces of membranes (Fig. 3(a)) under the applied pressure. Similar failures were observed when additional Hydrowell® modules were tested (data not reported here). It is concluded that the Hydrowell® module is not suitable for the x-PRO-in configuration, except where applied pressure is relatively low. Where it is desirable to have the dense rejection layer facing the feed water (e.g., for fouling control [12,13]), the module may need to be redesigned such that the rejection layers are inside of the membrane envelope and pressure is applied outside of the envelope to ensure the integrity of the module.

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Fig. 9. Comparison of x-PRO-out and x-PRO-in configurations. Draw solution concentration at 0.5 M. The open symbols represent the respective fluxes obtained in the x-FO-out and x-FO-in configurations. Note: 1 L/m2 h = 2.78 × 10−7 m/s.

The flux performance at various draw solution concentrations and applied pressures in the x-PRO-out configuration was shown in Fig. 10. As predicted correctly by Eq. (5), membrane flux reduced at higher pressure and/or lower draws solution concentration due to the reduced driving force (the osmotic pressure difference minus the applied pressure, refer to Fig. 1(b) and (d)). At a draw solution concentration of 0.5 M, the permeate flow was reduced by approximately 50% when the applied pressure increased from 0 to 5 bar. However, the experimentally measured permeate flow for 0.3 and 0.5 M draw solutions was significantly lower than the corresponding simulated results. As discussed in Section 4.2, this difference was likely due to the mass transfer resistance and ECP on the feed water side. Interestingly, it became much more difficult to pump feed water into the membrane module (i.e., inside the membrane envelope) upon pressurizing on the draw solution. In the Hydrowell® module, identical spacers were used both inside and outside the membrane envelope (Fig. 3(a)). These spacers were formed by plastic strings arranged in orthogonal manner, with spacing between adjacent strings ∼2.5 mm. This means that membranes that were unsupported over a 2.5 mm span had to withstand the high pressure applied on the draw solution side, which is not

Fig. 10. Effect of applied pressure on permeate flow in the x-PRO-out configuration at different draw solution concentrations. Cross-flow rate of draw solution at 75 ml/min. Simulation results were obtained from Eq. (6) coupled with Eq. (5) (Section 2 and Appendix A). Note: 1 L/m2 h = 2.78 × 10−7 m/s.

Fig. 11. Available power density from the Hydrowell® spiral wound module in xPRO-out configuration. A 0.5 M NaCl was used as the draw solution.

an ideal design for high pressure applications such as PRO. The deformation of the membranes under the high pressure and the corresponding narrowing of the flow channel inside the membrane envelope can significantly increase the mass transfer resistance on the feed water side. A spacer design that is similar to an RO permeate collector may work better in this regard. At a draw solution of 0.1 M, our simulated results agreed reasonably well with the observed permeate flow. Consistent with the discussion in Section 4.2, the membrane flux in this case was sufficiently low that it was not constrained by the feed water flow. A potential application of PRO is for harvesting osmotic power from seawater [14,17,20,32,40]. The power density, or the power output per membrane area from the PRO process, can be determined from Eq. (11). Fig. 11 presents the available power density using a 0.5 M NaCl draw solution. The dotted line represent the simulated power density assuming a large flow rate of draw solution such that dilution of the bulk draw solution by the permeate flow is negligible. The maximum power density is 1.5 W/m2 membrane area, which occurs at an applied pressure of 10–11 bar. This agrees well with earlier studies that the maximum power density occurs at an applied pressure of ∼50% of the draw solution osmotic pressure [8,18]. The power density can be significantly lower if dilution of the draw solution becomes significant (refer to the solid line in Fig. 11). For a draw solution flow rate of 75 ml/min, the maximum power density is only 1 W/m2 . The corresponding optimal pressure is also lower (∼9 bar) as a reflection of the dilution effect. Compared to the simulated data, the actual power performance of the Hydrowell® module was much worse due to the combination of dilution effect as well as the feed water mass transfer resistance. Improved module/spacer design and proper management of draw solution and feed water cross-flows are therefore required for optimized power performance from a PRO module. The maximum available PRO power density is presented as a function of draw solution concentration in Fig. 12. The simulated results without the dilution effect (dotted line) shows that the power density is approximately proportional to (Cdraw,0 )2 . Similar results have been reported by Ludwig et al. [41]. In contrast, the power density is proportional to (Cdraw,0 )1.7 when the dilution effect is considered for a draw solution of 75 ml/min (solid line), while the experimental results showed a nearly linear dependence on the draw solution concentration. Our results suggest that improved module design and cross-flow management are more critical for greater draw solution concentrations (higher flux levels).

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Acknowledgements This research received financial support from Suzhou Industrial Park Technology Innovation Zone and Suzhou Institute of Sichuan University. Xu and Peng were supported by Suzhou Institute of Sichuan University. Appendix A. A.1. Derivation of ICP models Based on the classical solution-diffusion model for the rejection layer, we have Jv = A(draw − support − P)

(A1)

Note: P = 0 for FO and P > 0 for PRO Js = B(Cdraw − Csupport ) Fig. 12. Effect of draw solution on maximum power density of PRO.

The current study did not perform any direct measurements on the salt flux through the FO membrane and the pressure drops cross the FO module (for both the feed water side and the draw solution side). Such measurements may provide additional insights into the mass transfer and flux behavior of the FO process. Therefore, it is recommended that future studies may incorporate these measurements as well. In addition, the cascading of multiple FO modules in a membrane array will also be of great interest, though beyond the scope of the current study. 5. Conclusions The effect of draw solution concentration and operating conditions on permeate flow rate of forward osmosis and pressure retarded osmosis was investigated in the current study using a spiral wound Hydrowell® module. The following conclusions can be made from the current study: • In both FO and PRO modes, the permeate flow increased at greater draw solution concentration. However, the flux behavior was highly non-linear with respect to Cdraw,0 due to internal concentration polarization as well as the dilution of draw solution by the permeate flow. • ICP, the concentration polarization in the porous membrane support, played an important role in membrane flux reduction in FO and PRO modes, and it became increasingly dominant at greater draw solution concentration. • Likewise, the dilution effect was more severe at greater draw solution concentration. The cross-flow of the draw solution shall be sufficiently high (dilution factor < 2) to limit the adverse dilution effect. • The feed water flow circulation had limited effect on membrane flux when it was circulated outside of the membrane envelope, and the submerged FO configuration performed nearly as good as the cross-flow x-FO-in configuration. However, the feed flow imposed an upper limit for the membrane permeate when it was circulated inside of the membrane envelope. The effect of feed flow may be explained by the mass flow resistance and ECP on the feed solution side. • A simple and useful model was developed to account for the ICP as well as the dilution effects. The model prediction agreed well with the experimental results when mass transfer resistance on the feed water side was insignificant. However, the model overestimates membrane flux when ECP on the feed solution side is significant.

(A2)

For the solute transport in the support layer Jv C + Js =

D dC ε dx

(A3)

where C is the concentration in the support layer at a distance x away from the interface of the support layer and the active layer, and D/ε is the effective diffusion coefficient of the solute in the porous support with a porosity of ε. Consider the active layer facing the draw solution orientation, the boundary conditions of Eq (A3) are C = Csupport at x = 0

(A4)

and C = Cfeed at x = t

(A5)

By combining Eqs. (A2) and (A3), one can derive:



Csupport + (B(Cdraw − Csupport )/Jv ) ln Cfeed + (B(Cdraw − Csupport )/Jv )



=

Jv Km

(A6)

where Km = S=

εD D = t S

(A7)

t ε

(A8)

If we assume that the osmotic pressure of a solution is proportional to the solution concentration, Eq. (A6) becomes



support Jv + B(draw − support ) ln feed Jv + B(draw − support )



=

Jv Km

(A9)

By substituting Eq. (A1) into Eq. (A9) to illuminate support , we can obtain the ICP equation for the active layer facing the draw solution orientation in the PRO mode (Eq. (5) in the main paper):



Jv = Km ln

(Adraw − AP − Jv ) + B((AP/Jv ) + 1) Afeed + B((AP/Jv ) + 1)



(5)

Eq. (2) can be obtained by setting P = 0 in Eq. (5). Eqs. (1) and (4) can be derived in a similar fashion. A.2. An example for modeling FO/PRO permeate flow in a spiral wound module The local membrane flux anywhere inside the spiral wound module can be modeled by ICP flux models presented in Section 2 (Eqs. (1)–(5)). Based on the flat coupon FO test results and additional independent RO tests by She [12], the HTI membrane used in the current studies had the following properties: A = 2.2 × 10−12 m/s Pa

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Table A1 An example for simulating spiral wound FO permeate fluxa . Stepb

Local draw solution concentration, Cdraw,x (M)c

Local membrane flux, Jv,x (L/m2 h)d

Incremental membrane area, Am (m2 )b

Incremental permeate flow, Qp,x (ml/min)e

Accumulative permeate flow, Qp,x (ml/min)f

1 2 3 4 – 100

0.500 0.492 0.484 0.477 – 0.237

7.90 7.79 7.68 7.58 – 4.04

0.0094 0.0094 0.0094 0.0094 – 0.0094

1.24 1.22 1.20 1.19 – 0.63

1.24 2.46 3.66 4.85 – 83.87

Notes: a Simulation conditions: s-FO-in/x-FO-in configurations; total effective membrane area in the module of 0.94 m2 ; a 0.5 M draw solution with a flow rate of 75 ml/min at the module inlet. b The total membrane area was divided into 100 equal small areas to compute the local incremental permeate flow. c The local draw solution was determined from the inlet draw solution concentration according to Eq. (6). d The local flux was obtained as a function of local draw solution concentration from Fig. A.1. Unit conversion factor: 1 L/m2 h = 2.78 × 10−7 m/s. e The incremental permeate flow was determined from Qp,x = Jv,x Am . f

The accumulative permeate flow was obtained from Qp,x = ˙(Qp,x ) or Qp,x =

and B = 1.7 × 10−7 m/s. The mass transfer coefficient Km inside the membrane support was 3.3 × 10−6 m/s when the dense rejection layer faced the draw solution, and it was 4.2 × 10−6 m/s in the alternative membrane orientation. These values agreed reasonably well with Gray et al. [22]. Based on Eqs. (1) and (2), the local FO flux can be determined as a function of local draw solution concentration (Fig. A.1). Similarly, flux in PRO can be determined by Eqs. (4) and (5) (results not shown here). The FO/PRO permeate flow rate of a spiral wound module can be modeled using both the ICP equations (e.g., the flux-concentration equations in Fig. A.1 for FO mode) together with Eq. (6) which accounts for the dilution effect. A worked example is shown in Table A1. With a 0.5 M NaCl draw solution at a flow rate of 75 ml/min at the module inlet, the local membrane flux is as high as 7.9 L/m2 h according to Fig. A.1. As the draw solution flows through the module, its concentration becomes increasingly diluted by the permeate flow. A total of 84 ml/min permeate flow is available from the entire membrane module. Consequently, the concentration at the module outlet is determined by 0.5 × 75/(75 + 84) = 0.24 M, which is much lower than the inlet concentration. The corresponding local flux at the outlet is only 4.0 L/m2 h. Clearly, the flux

x 0

Jv dAm .

performance is adversely affected due to the draw solution dilution effect. Nomenclature 2 A Am B Cdraw Cdraw,0 Cdraw,exit Cdraw,x Cfeed Csupport D Jv Jv Km P Qcf Qcf,0 Qp S t ε draw feed  ˚PRO PRO

transport coefficient for water (i.e., water permeability) (m/s Pa) membrane area (m2 ) transport coefficient for solute (m/s) draw solution concentration (M) draw solution concentration at module inlet (M) draw solution concentration at module outlet (M) draw solution concentration at location x (M) feed solution concentration (M) solute concentration at the interface of the dense rejection layer and the porous support layer (M) solute diffusion coefficient (m2 /s) permeate flux of water (m/s or L/m2 h) average permeate flux in a spiral wound module (m/s or L/m2 h) mass transfer coefficient of the membrane support (m/s) applied pressure on the draw solution (Pa) cross-flow rate of the draw solution (m3 /s or ml/min) cross-flow rate of the draw solution at module inlet (m3 /s or ml/min) membrane permeate flow (m3 /s or ml/min) membrane structural parameter (m) membrane support layer thickness (m) porosity of the membrane support layer osmotic pressure of the bulk draw solution (Pa) osmotic pressure of the bulk feed water (Pa) tortuosity of the membrane support layer total power available from the module (W) power density (W/m2 )

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Unit conversion factor: 1 L/m2 h = 2.78 × 10−7 m/s.

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