Evidence, Determination, and Implications of Membrane-Independent

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Evidence, Determination, and Implications of MembraneIndependent Limiting Flux in Forward Osmosis Systems Christopher Morrow, and Amy Childress Environ. Sci. Technol., Just Accepted Manuscript • DOI: 10.1021/acs.est.8b05925 • Publication Date (Web): 19 Mar 2019 Downloaded from http://pubs.acs.org on March 26, 2019

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Environmental Science & Technology

Evidence, Determination, and Implications of Membrane-Independent Limiting Flux in Forward Osmosis Systems

a

Christopher P. Morrow , *Amy E. Childress

a

a

University of Southern California, Los Angeles, CA, USA Sonny Astani Department of Civil and Environmental Engineering

*Corresponding author: email: [email protected] tel: +1 (213) 740-6304

A manuscript prepared for possible publication in Environmental Science and Technology

October 2018

1 ACS Paragon Plus Environment


1

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TOC/Abstract art

2

Membrane dependent: initial flux,RSF, limiting Δϖ

ea SF sing

cr

Membrane independent limiting flux

*Limiting Δϖ

*

SYSTEM and OPERATING improvements

A/B

R

Membrane

A/B

2) Constant Δϖ > limiting Δϖ Membrane

In

Water Flux (Lm-2h-1) Reverse Salt Flux (gm-2h-1)

1) Stepwise increases in Δϖ

Δϖ

*

Δt

3

Abstract

4

A stepwise method for determining limiting flux and limiting osmotic pressure and a constant osmotic

5

pressure method to validate the limiting flux were developed. First, five of the most commonly used FO

6

membranes were characterized for water permeability (A), solute permeability (B), and structural

7

parameter (S). During both stepwise and constant osmotic pressure fouling experiments, membrane

8

fouling constrained water flux to a singular, common upper limit, the limiting flux, for all membranes

9

despite very different A and A/B values for the membranes. Conversely, there was not an upper limit to

10

reverse salt flux. It was observed that reverse salt flux increases as S decreases; however, this does not

11

mean that higher S values are desirable. Higher S values (> ~600 μm) also increase dilutive internal

12

concentration polarization, which is recognized as the major impediment to achieving high FO water flux.

13

Thus, for osmotic processes where membrane fouling occurs, membrane transport parameters A and B

14

may not be useful performance indicators and the goal of improving water flux by developing highly

15

permeable, highly selective membranes may not be realistic. Instead, optimizing fouling mitigation

16

strategies, hydrodynamics at the membrane surface, and membrane module configuration may be more

17

promising alternatives for improving performance.

18 19


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20


Keywords: forward osmosis; membrane fouling; reverse salt flux; critical flux; limiting flux

21 22

Highlights:

23

•

A membrane-independent limiting flux concept was defined for forward osmosis processes

24

•

Methods were developed for determination of limiting flux and limiting osmotic pressure

25

•

Water flux declined over time and tended toward a singular limiting flux for five membranes

26

•

For fouled membranes, increasing osmotic pressure increased RSF without increasing water flux

27

•

Fouled membranes with higher A/B values had higher cake layer resistances

28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47



48 49

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1. Introduction Perhaps the greatest advantages of forward osmosis (FO) membrane processes are that they can 1-3

50

treat high fouling wastewaters with more flux stability

51

removed than those in pressure-driven membrane processes

52

hydraulic pressure, FO is driven by the osmotic pressure difference between a high osmotic pressure

53

(high salinity) draw solution and low osmotic pressure (low salinity) feed solution

54

promising applications of FO is as an integrated pretreatment process in membrane bioreactor (MBR)

55

systems for wastewater reuse

56

activated sludge using FO; an additional process (e.g., reverse osmosis (RO) or membrane distillation

57

(MD)) is then used to separate high-quality product water and reconcentrate the draw solution. Membrane

58

fouling causes some water flux decline in the FO process

59

reduction in the effective driving force due to dilution of the draw solution within the porous support layer

60

(dilutive internal concentration polarization (ICP)) and diffusion of draw solutes into the feed solution

61

(reverse salt flux (RSF))

62

8, 9, 19

10, 11

and membrane fouling layers can be more easily 4-6

. Unlike processes that are driven by

7-9

. One of the more

. In the osmotic MBR (OMBR) system, water is extracted from

7, 12-18

. Flux decline is also caused by a

.

For pressure-driven RO and nanofiltration (NF) membrane processes with high salt rejection, solutes

63

in the feed solution are rejected at the membrane surface and subsequently diffuse back into the bulk

64

feed solution. When membrane fouling occurs, the foulant cake layer hinders back-diffusion of solutes,

65

increasing the osmotic pressure that must be overcome and possibly causing significant flux decline; this

66

has been termed cake-enhanced concentration polarization (CECP)

67

typically operate at lower water fluxes than pressure-driven processes, the foulant layer can still hinder

68

back diffusion of solutes, but perhaps to a lesser extent. However, due to RSF, there can also be solute

69

accumulation in the foulant layer from diffusion of the draw solution across the FO membrane; this has

70

been termed cake enhanced osmotic pressure (CEOP)

71

osmotic pressure resistance increases and water flux declines. In FO, CEOP and the resulting water flux

72

decline trigger a self-correcting mechanism termed the ICP self-compensation effect

73

declines, draw solute concentration in the support layer increases, causing the driving force across the

74

selective layer to increase, thus inherently resulting in flux stability.

75

4, 22

20, 21

. Although FO processes

. CEOP has the same effect as CECP; the

2, 12, 23

. As water flux

At steady state, an upper limit to FO water flux has been observed for fouled membranes (S1.0) and


Page 5 of 22


76

this phenomenon has been compared to critical flux behavior in microfiltration (MF) and ultrafiltration (UF)

77

membrane processes

78

membranes

79

pressure (e.g., CECP/CEOP). Also, the critical flux value depends on membrane properties (e.g., pore

80

size)

81

relating it to the limiting flux concept, which was developed for nonporous RO and NF membranes

82

more appropriate.

83

37

34-36

2, 17, 24-33

. However, MF and UF critical flux describes processes with porous

, which pass dissolved solutes and are not affected by phenomena involving osmotic

. Thus, instead of relating the upper limit in water flux of FO membranes to the critical flux concept, 38, 39

, is

The limiting flux concept for RO/NF asserts that increasing hydraulic pressure causes greater 38-40

84

concertation polarization and greater foulant accumulation, which asymptotically limit water flux

85

RO/NF limiting flux concept is membrane-independent; after initial foulant deposition occurs and the

86

limiting flux is reached, foulant-foulant interactions (not membrane-foulant interactions) dominate. Limiting

87

flux behavior has also been reported in pressure-retarded osmosis (PRO) studies, where the support

88

layer faces the feed solution

89

pressure and initial flux, but strongly dependent on foulant/draw solute interactions and membrane

90

structural parameter

91

scaling that occur on and within the support layer in PRO membrane orientation.

92

41

30, 31, 41

. The

. PRO limiting flux was determined to be independent of applied

. The strong dependence on membrane structural parameter is due to fouling and

In FO membrane orientation, where the selective layer faces the feed solution, fouling would occur on

93

the selective layer. Therefore, FO limiting flux depends on the feedwater chemistry and fouling potential,

94

hydrodynamics of the membrane module, and operating conditions

95

for FO is defined as the maximum steady-state water flux obtainable in the presence of foulants; the

96

limiting flux cannot be overcome by increasing the osmotic pressure driving force. Limiting osmotic

97

pressure is defined as the minimum osmotic pressure required to achieve the limiting flux.

98 99

2, 17, 24-33

. In this paper, the limiting flux

The limiting flux concept in FO considers ICP self-compensation, RSF, and CEOP. As FO water flux declines due to membrane fouling, the ICP self-compensation effect determines the limiting osmotic

100

pressure (i.e., the driving force) and ultimately, the system achieves equilibrium with a limiting flux

101

independent of membrane water permeability (A), solute permeability (B), and structural parameter (S)

102

values. Increasing the draw solution osmotic pressure beyond the limiting osmotic pressure does not

103

increase steady-state water flux; however, it does increase draw solution concentration at the



104

selective/support layer interface, which may result in greater RSF and CEOP (S1.1, Fig. S1). The

105

dynamics that determine the limiting flux for FO are quite different than those for pressure-driven

106

membrane processes, although the core principle remains the same; greater driving force does not result

107

in greater water flux, but does result in greater concentration polarization and membrane fouling.

108

Research in FO membrane development has focused on attaining high FO water flux by improving

109

membrane transport and structural parameters (i.e., achieving higher A, lower B, and lower S), although

110

the dynamics between concentration polarization and membrane fouling at the limiting flux may diminish

111

the significance of membrane properties in higher fouling applications.

112

The objectives of this paper are to formally define the limiting flux concept for FO, to introduce two

113

methods developed to identify limiting flux, to experimentally demonstrate that limiting flux is independent

114

of membrane parameters, and to discuss implications of the limiting flux concept. First, the A, B, and S

115

values of five of the most commonly tested FO membranes were characterized. Next, two methods were

116

developed using either stepwise increases in osmotic pressure or constant osmotic pressure and both

117

methods were used to evaluate limiting flux and limiting osmotic pressure for the five membranes. Draw

118

solution osmotic pressure and RSF were evaluated under limiting flux conditions and a resistance-in-

119

series model was used to determine the relationship between membrane selectivity and hydraulic

120

resistance of the foulant layer. Alternative methods besides improved membrane properties for achieving

121

higher water flux were proposed.

122 123

2. Materials and methods

124

2.1. Membranes and bench-scale FO system

125

Five commercially available FO membranes were tested; a TFC membrane from Porifera Inc.

126

(Hayward, CA), a TFC membrane from Toray Industries Inc. (Tokyo, Japan), a TFC membrane from

127

Oasys Water Inc. (Cambridge, MA), and a CTA and a TFC membrane from Hydration Technology

128

Innovations, LLC (Albany, OR). All membranes were tested with the selective layer facing the feed

129

solution. Details of the bench-scale FO system (Fig. S2) used in this study have been described

130

previously

131

area of 42 cm , and each membrane coupon had independent draw solution (1-L sidearm flasks) and

33

. Briefly, the membrane module contained three membrane coupons, each with an effective 2


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132

feed solution (2-L tanks) reservoirs. 31-mil (0.79 mm) diamond-style spacers (Sterlitech Corporation,

133

Kent, WA) were used on the draw solution side and no spacers were used on the feed solution side.

134

Membranes were tested at 24 ± 1 °C with a crossflow velocity of 10 cm/s. Analytical grade NaCl (VWR

135

International, Radnor, Pennsylvania) was used to prepare the draw solution. Conductivity probes (Cole-

136

Parmer, Vernon Hills, IL) in the draw solution reservoirs were used to monitor draw solution conductivity.

137

An automatic dosing system (Fig. S2) kept the draw solution reservoirs at constant concentration; when

138

conductivity fell below the set point, the dosing system was activated and 292-g/L (5-M) NaCl solution

139

was transferred to the appropriate draw solution reservoir until the set point was reached. The mass of

140

overflow from each draw solution reservoir was measured with an analytical balance (PA3102, OHAUS

141

Corporation, Parsippany, NJ) and used to calculate water flux in 5-min increments. Feed solution

142

conductivities were measured with a conductivity probe (EC215, Hanna Instruments, Woonsocket, RI)

143

and used to calculate RSF:

𝑅𝑆𝐹 =

150

𝐶& 𝑉& − 𝐶) 𝑉) (1) 𝐴+ 𝛥𝑡

144

where CF and Ci are final and initial NaCl feed concentrations (calculated from conductivity

145

measurements); VF and Vi are final and initial feed volumes; AM is membrane area; and Δt is elapsed

146

time. Data acquisition and control devices (USB-6009, NI 9208, NI ER-8) were connected to LabVIEW

147

software (National Instruments, Austin, TX) and used to record data and trigger the dosing system. Errors

148

in water flux and RSF values were calculated from the standard deviation between triplicate membrane

149

coupons.

151

2.2. Membrane characterization

152

Results from four experiments were used to determine A, B, and S values using an FO 42

153

characterization method

154

solution were conducted with draw solution concentrations (6.4, 12.9, 25.5, and 37.7 g NaCl/L)

155

corresponding to 0.5-, 1-, 2-, and 3-MPa osmotic pressure. Osmotic pressures and NaCl concentrations

156

were calculated using OLI Stream Analyzer (OLI Systems, Inc., Morris Plains, NJ). Average NaCl

157

concentrations, osmotic pressures, water flux, and RSF from the four experiments were substituted into

158

the governing solution-diffusion transport equations for water flux (Jw) and RSF (Js)

. Four 30-minute experiments using dilute NaCl (1.0 g/L; 0.1 MPa) feed


42, 43

:


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159

𝐽 𝐽6 𝑆 − 𝜋& 6 𝐷 𝑘 𝐽6 = 𝐴 𝐽6 𝐽 𝑆 𝐵 1+ exp − exp − 6 𝐷 𝐽6 𝑘

(2)

160

𝐽 𝐽6 𝑆 − 𝐶& 6 𝐷 𝑘 𝐽A = 𝐵 𝐽6 𝐽 𝑆 𝐵 1+ exp – exp − 6 𝐽6 𝐷 𝑘

(3)

𝜋8 exp −

𝐶8 exp −

161

where πD and πF are osmotic pressures of the bulk draw and bulk feed solutions; D is draw solute

162

diffusivity; k is mass transfer coefficient of the membrane module

163

concentrations of the bulk draw and bulk feed solutions. To determine A, B, and S, the system of eight

164

transport equations was fit to the experimental data using a least squares non-linear regression.

165

2.3. Synthetic activated sludge feed solution

166

44

; and CD and CF are NaCl

A synthetic activated sludge feed solution (Section S2.0) was prepared to simulate the fouling

167

potential of activated sludge in an OMBR system treating domestic wastewater. Sodium alginate, bovine

168

serum albumin, and humic acid (Sigma Aldrich, St. Louis, MO) were used as model foulants representing

169

polysaccharides, protein, and NOM, respectively

170

layers that simulate the structure of biologically formed foulant cake layers

171

environment was simulated using ACS reagent-grade ionic salts (Sigma Aldrich, St. Louis, MO); NH4Cl,

172

Ca(NO3)24H2O, KH2PO4, and MgSO4 were sources of ammonium, nitrate, phosphate, sulfate, hardness,

173

and salinity; NaHCO3 and NaCl were additional sources of hardness and salinity. The feed solution was

174

prepared to final NH4 , NO3 , PO4 , SO4 , and HCO3 concentrations representative of OMBR systems

175

treating domestic wastewater

176

Cl were also representative of commonly reported values in OMBRs

+

-

3-

2-

10, 17, 49-51

45, 46

. CaCl2 was used to form cross-linked alginate gel . The chemical

-

2+

+

2+

+

. Ionic concentrations of the counterions, Ca , K , Mg , Na , and

-

177

47, 48

50-52

.

Total solids and volatile solids were analyzed in duplicate using Standard Methods 2540B and 2540E

178

53

179

Conductivity and pH were measured using a handheld probe (SevenGo Duo Pro, Mettler Toledo Inc.,

180

Hightstown, NJ). The conductivity of the synthetic foulant solution was 1.7 mS/cm and pH was 7.2.

. Total solids concentration was 5.15 ± 0.04 g/L and volatile solids concentration was 2.49 ± 0.07 g/L.

181 182


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183

2.4. Experimental procedures

184

2.4.1. Stepwise draw solution osmotic pressure experiments

185

Based on experiments for MBR critical flux determination 42

35, 36, 54, 55

, FO fouling observation

28

, and FO

186

membrane characterization

187

fouled FO membranes. Baseline water flux and RSF experiments were conducted using an NaCl feed

188

solution prepared to the same conductivity (1.7 mS/cm) as the synthetic foulant solution. NaCl draw

189

solution osmotic pressure was increased in a stepwise manner corresponding to 0.5, 1, 2, 3, 5, and 10

190

MPa (6.4, 12.9, 25.5, 37.7, 60.9, and 112 g NaCl/L). Each step was carried out for two hours and water

191

flux values were determined as in Section 2.1. Initial and final feed solution conductivity and mass values

192

were recorded and used to calculate RSF during each step (Eqn. 1).

193

, a stepwise FO experiment was developed to identify limiting flux values for

Experiments were repeated with the synthetic activated sludge feed solution and results were used to

194

calculate water flux and RSF under fouling conditions. Feed solution volumes decreased during each

195

step. This would have caused the foulant concentration of the feed solution to increase over time;

196

however, this was prevented by refilling the feed solution reservoirs to the initial (2-L) volume with

197

deionized water between each step. This ensured that each step began with the same foulant

198

concentration in the feed solution.

199

2.4.2. Limiting flux and limiting osmotic pressure with stepwise draw solution osmotic pressure

200

experiments

201

For stepwise experiments, membranes were considered fouled when water flux declined by at least

202

10% during a given step. Limiting flux values were calculated from a minimum of three steps with

203

membrane fouling. To observe three membrane fouling steps for CTA membranes, it was necessary to

204

use higher draw solution osmotic pressures (15 and 20 MPa; 153 and 184 g/L NaCl). Final water flux was

205

calculated over the last 15 minutes of each step. Limiting flux was taken as the average of the final water

206

flux values from steps where membrane fouling was observed. For each membrane, final water flux

207

values in each step were plotted as a function of draw solution osmotic pressure and a fourth order

208

quadratic equation with an R value of 0.99 or greater was fit to experimental data. Limiting flux values for

209

each membrane were used in the corresponding quadratic equations to solve for the minimum (limiting)

2



210

osmotic pressure required to achieve the limiting flux (see example in Fig. S3). The methodology for

211

calculating membrane and foulant cake layer resistances is detailed in Section S4.0.

212

2.4.3. Constant draw solution osmotic pressure experiments

213

A second set of membrane fouling experiments was conducted to validate limiting flux values

214

obtained from the stepwise experiments. These experiments were carried out for 12 hours with a constant

215

draw solution osmotic pressure that was approximately twice that of the calculated limiting osmotic

216

pressure from the stepwise experiments. Draw solution osmotic pressure was kept constant by the

217

automatic dosing system (Section 2.1.). Feed solution reservoirs were replenished with deionized water

218

every four hours to prevent increased foulant concentration in the feed solution over the duration of the

219

experiments.

220

2.4.4. Limiting flux with constant draw solution osmotic pressure experiments

221

During experiments with constant draw solution osmotic pressure, flux decline could occur due to

222

mass transfer resistance from foulant deposition and/or decreased driving force from RSF into the feed

223

solution. Flux decline from foulant deposition would likely occur until the cake layer is fully formed.

224

However, flux decline from RSF would occur until the system reaches equilibrium. At the true equilibrium

225

state for the closed-loop testing system, water and solute diffusion would occur in both directions at the

226

same rate (water flux and RSF equal zero). For this reason, the system was not operated until the true

227

equilibrium state was reached. Instead, a t-test (n = 3 (triplicate coupons), α = 0.05, μ = 0, and t* = 2.92)

228

comparing the rate of flux decline between two successive 2-hr periods was performed. Water flux for

229

each membrane coupon was plotted as a function of time and the rate of flux decline over discrete 2-hr

230

periods was obtained from linear regressions. Membrane fouling was considered fully developed when

231

flux decline between the periods was no longer statistically significant (P ≥ 0.05, see example in S2.1).

232

Water flux at the beginning of the first two-hour period where flux decline was found to be statistically

233

insignificant was taken as the limiting flux. Additionally, decreased driving force due to RSF during the

234

fouling experiments was calculated using a mass balance (S2.2).

235 236

3. Result and discussion

237

3.1. Membrane transport and structural parameters


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238


A, B, and S values determined using the FO characterization method are given in Table 1 for the five

239

membranes. The A/B ratio is also shown to indicate selectivity of the membranes, with higher A/B

240

indicating higher selectivity and lower A/B indicating lower selectivity

241

exhibited the lowest A/B, and A/B increased in the order of: HTI TFC, Oasys TFC, Porifera TFC, and

242

Toray TFC membranes. Discrepancies between A, B, and S values determined here (with the FO

243

method) and literature values obtained with the conventional RO-FO method are likely due to the different

244

driving forces in FO (osmotic pressure) and RO (hydraulic pressure)

245

assumed to be non-porous, it has been suggested that defects in the selective layer result in a

246

combination of pore flow and diffusion when hydraulic pressure is applied (e.g., with the RO-FO

247

characterization method). On the other hand, pore flow is unlikely in the absence of hydraulic pressure

248

(e.g., with the FO characterization method)

249

passage, or higher B values. Furthermore, compaction of the porous support structure under hydraulic

250

pressure in the RO-FO method may reduce water permeability, or the A value, by increasing membrane

251

density. It is noted that lower S values (between 175 and 344 μm) were obtained for Porifera TFC

252

membranes using the RO-FO method

253

be due to differences in membranes provided by the manufacturers or damage during shipping or use.

58

29, 59, 60

56, 57

42, 58

. The HTI CTA membrane

. Although FO membranes are

. Thus, the RO-FO method may lead to greater salt

. Difference with values reported in the literature could also

Table 1. Water permeability (A), solute permeability (B), and structural parameter (S) determined using the FO characterization method. -2 -1 -1 -2 -1 -1 Membrane A (L m h bar ) B (L m h ) S (μm) A/B (bar ) HTI CTA

0.58 ± 0.05

0.49 ± 0.06

385 ± 18

1.18

HTI TFC

2.78 ± 0.21

1.36 ± 0.14

513 ± 36

2.04

Oasys TFC

5.26 ± 0.88

2.44 ± 0.68

554 ± 50

2.16

Porifera TFC

3.35 ± 0.51

1.18 ± 0.18

398 ± 35

2.84

Toray TFC

2.29 ± 0.34

0.30 ± 0.08

198 ± 44

7.63

254 255

3.2. Limiting flux and limiting osmotic pressure from stepwise experiments

256

Water flux as a function of time is shown in Figs. 1a-e for experiments with stepwise increases in

257

draw solution osmotic pressure (solid line). Baseline water flux (dashed line) is with sodium chloride feed

258

solution and water flux from membrane fouling experiments (open circles) is with synthetic activated

259

sludge feed solution. For each membrane, baseline water flux increased with each osmotic pressure step.

260

Baseline water flux was higher for membranes with successively higher A/B values (Toray TFC > Porifera



261

TFC > Oasys TFC > HTI TFC > HTI CTA) and was relatively constant during each step. Water flux with

262

activated sludge was constant during steps with lower draw solution osmotic pressures (0.5 and 1.0 MPa)

263

but decreased during steps with higher draw solution osmotic pressures (> 10 MPa for the CTA

264

membrane and > 2 MPa for TFC membranes). This transition between water flux remaining constant over

265

a step duration and water flux declining over a step duration is the key indicator that the limiting flux has

266

been exceeded. In other words, flux decline due to membrane fouling only occurs when the limiting flux,

267

or the maximum steady-state water flux obtainable in the presence of foulants, is exceeded.

268 269 270 271 272 273 274

Figure 1. Baseline water flux, water flux with synthetic activated sludge feed solution, and final water flux values for fouled membranes with stepwise increases in draw solution osmotic pressure for (a) HTI CTA, (b) HTI TFC, (c) Oasys TFC, (d) Porifera TFC, and (e) Toray TFC membranes. Also, (f) limiting flux and limiting osmotic pressure for each membrane from stepwise experiments. Error bars represent standard deviation between triplicate membrane coupons.

275

The solid circles in Figs. 1a-e indicate final water flux values when water flux declined over a step

276

duration due to membrane fouling. For each individual membrane, the final water flux values for each

277

step varied by less than 6% of each other. For the HTI CTA, HTI TFC, Oasys TFC and Porifera TFC


Page 12 of 22

Page 13 of 22


278

membranes together, the average of all the final water flux values also varied by less than 6% from the

279

individual values, indicating that the limiting flux for the system was independent of the membrane used.

280

As can be seen by the data in Fig. 1f, limiting flux values were within standard deviations of the HTI CTA,

281

HTI TFC, Oasys TFC and Porifera TFC membranes and the average value for these four membranes

282

was 13.9 ± 0.5 L m h . The Toray TFC membrane had a higher limiting flux (17.7 ± 0.7 L m h ) likely

283

due to its higher initial water flux during each step with activated sludge feed solution compared to all

284

other membranes. Higher initial water flux causes greater flux decline

285

possible that foulant deposition, and resulting flux decline, were not complete when the experiment was

286

terminated. It should be noted that the higher flux of the Toray TFC membrane was not due to a higher A

287

value, in fact it had the lowest A among the TFC membranes (Table 1). Instead, it is likely due to Toray

288

TFC’s low S value that contributes to a lower propensity for dilutive ICP

289

effective driving force for water flux.

290

-2

-1

-2

12, 45, 61

62, 63

-1

. And in this case, it is

, and hence, a higher

The observation that similar limiting flux values were observed for HTI CTA, HTI TFC, Oasys TFC

291

and Porifera TFC membranes is consistent with a mechanistic understanding of membrane fouling; after

292

initial foulant deposition (driven by membrane-foulant interactions

293

toward the membrane surface (drag forces), foulant-foulant interactions at the membrane surface

294

(adhesion forces), and operational parameters (crossflow velocity and pressure) that dictate additional

295

foulant deposition and corresponding hydraulic resistance

296

and NF membranes

297

morphology; eventually water flux is independent of membrane properties and becomes a function of only

298

the operating conditions and fouling potential of the feed solution

299

38, 39

39, 65

64

), it is only convective water flux

. Similar to the limiting flux concept for RO

, only initial foulant deposition is dependent on membrane parameters and

64

.

The circles in Fig. 1f represent limiting osmotic pressures for the five membranes. For all TFC

300

membranes, limiting osmotic pressures were between 2.2 and 2.5 MPa; for the HTI CTA membrane, the

301

limiting osmotic pressure was 8.6 MPa. Thus, the TFC and CTA membranes required significantly

302

different draw solution concentrations (30.3 ± 2.4 vs. 111 ± 0.09 g/L NaCl) to achieve the limiting flux,

303

indicating that limiting osmotic pressure is membrane-dependent. This is not unexpected because higher

304

permeability TFC membranes with higher A/B values require less osmotic pressure to achieve similar

305

water fluxes as lower permeability CTA membranes

66, 67

. The lower limiting osmotic pressure of the TFC



Page 14 of 22

306

membranes is advantageous as less draw solute and draw solute replenishment are required, although

307

the need to replenish the draw solute also depends on losses due to RSF .

308

3.3. Effect of osmotic pressure increases on RSF

309

9

Water flux and RSF were also evaluated as a function of increasing draw solution osmotic pressure

310

(Fig. 2). For the non-fouling feed solution, baseline water flux (solid circles) and baseline RSF (solid

311

squares) both increase with draw solution osmotic pressure

312

activated sludge feed solution

313

osmotic pressure but water flux (open circles) only increases up to a maximum value (the limiting flux).

33, 49, 51

8, 9

. However, this is not the case for the

where RSF (open squares) increases continually with draw solution

314 315 316 317 318 319 320 321

Figure 2. Baseline water flux, baseline RSF, water flux with synthetic activated sludge, and RSF with synthetic activated sludge as a function of stepwise increases in draw solution osmotic pressure for (a) HTI CTA, (b) HTI TFC, (c) Oasys TFC, (d) Porifera TFC, and (e) Toray TFC membranes. Error bars represent standard deviations between triplicate membrane coupons. Unexpectedly, the membranes with the lowest B values (HTI CTA and Toray TFC membranes) had significantly higher RSF at higher osmotic pressures with activated sludge feed solution than membranes


Page 15 of 22


322

with higher B values (HTI TFC, Oasys TFC, and Porifera TFC membranes). For the HTI CTA membrane,

323

the observation of high RSF with low B may be attributed to a higher osmotic pressure to achieve similar

324

water flux as TFC membranes. The HTI CTA and Toray TFC membranes also have the lowest S values

325

(Table 1). Lower S indicates a shorter diffusive path length through the support structure and a lower

326

propensity for dilutive ICP

62

according to the dilutive ICP modulus EF,H

327

EF

= exp −

19

IJ K 8

(4)

328

where πD,i is the solute concentration at the selective layer/support layer interface. At the limiting flux,

329

increasing πD does not increase Jw. For a fixed Jw, πD,i increases exponentially as S decreases and linearly

330

as πD increases

331

Increasing πD,i increases RSF because rejection depends on solute concentration at the selective layer

332

surface. Consequently, RSF increases more severely with πD for membranes with small S than for

333

membranes with large S. The observation that RSF increases as S decreases should not be interpreted

334

to mean that higher S values are desirable. Higher S values (> ~600 μm) do result in lower RSF,

335

however, they also increase dilutive ICP, which is recognized as the major impediment to achieving high

336

FO water flux

337

low water flux and low RSF, a moderate S value (~300 to 600 μm) may be desired, particularly for TFC

338

membranes when membrane fouling is expected.

56

. Thus, at or above the limiting osmotic pressure, πD,i is more sensitive to S than πD.

44, 63, 67

. Given that low S values result in high water flux and high RSF and high S results in

339

For TFC membranes with higher B values and moderate S values (HTI TFC, Oasys TFC, and

340

Porifera TFC membranes), RSF with activated sludge feed solution was less than or equal to baseline

341

RSF at lower draw solution osmotic pressures (0.5 to 5 MPa). Lower RSF with membrane fouling may be

342

attributed to higher ICP for membranes with moderate S values and also to “fouling-reduced

343

concentration polarization”

344

improves NaCl rejection and results in lower RSF. However, even for these membranes, RSF with

345

activated sludge feed solution increases continually with draw solution osmotic pressure. Thus, results for

346

all membranes demonstrate that membrane fouling restricts water flux to the limiting flux whereas salt flux

347

(RSF) is not subject to an upper limit over the range of osmotic pressures tested here. For draw solutes

348

that may interact more specifically with foulants, RSF may lead to solute/foulant interactions and influence

12, 68

, a phenomenon where hydraulic resistance from the foulant cake layer



349

limiting flux in ways that were not observed under the conditions tested here. In those cases, limiting flux

350

may not be independent of membrane parameters.

351

3.4. Limiting flux from constant osmotic pressure experiments

352

To validate the limiting flux values obtained from the stepwise experiments, a second set of

353

experiments was conducted, also with activated sludge feed solution. These experiments were run with

354

constant draw solution osmotic pressures that were approximately twice that of the limiting osmotic

355

pressures calculated from stepwise experiments to ensure that flux decline from membrane fouling would

356

occur. 15 MPa was used with the HTI CTA membrane and 5 MPa was used with the TFC membranes. P

357

values (Fig. 3 inset) represent the significance of flux decline between each successive 2-hour period. P

358

values less than 0.05 (red font) indicate flux decline was significant and P values greater than 0.05 (black

359

font) indicate flux decline was not significant during that period. Only the periods with significant flux

360

decline for each membrane are shown by the graphed data in Fig. 3 (i.e., 0 to 8 hours for the Toray TFC

361

membrane, 0 to 6 hours for the HTI CTA and HTI TFC membranes, and 0 to 4 hours for the remaining

362

TFC membranes). Over time, water flux declined and tended toward a singular, limiting flux value for all

363

membranes. The limiting flux calculated by averaging the final flux values for all membranes was 13.80 ±

364

0.35 L m h (Table 2). Remarkably, this average limiting flux value was less than 3.0% different from the

365

final fluxes of each individual membrane.

-2

-1

366


Page 16 of 22

Page 17 of 22


367 368 369 370 371 372 373

Figure 3. Water flux as a function of time for HTI CTA, HTI TFC, Oasys TFC, Porifera TFC, and Toray TFC membranes with synthetic activated sludge feed solution. Continuous draw solution osmotic pressures of 15 MPa for HTI CTA membrane and 5 MPa for all TFC membranes were used. P-values represent significance of flux decline between successive 2-hr periods. P ≥ 0.05 indicates no further significant flux decline was observed; only periods with significant flux decline are shown by the graphed data. Table 2. Limiting flux values (Jw,L) calculated from stepwise and constant draw solution osmotic pressure (πD) experiments. -2 -1 -2 -1 Jw,L (L m h ) Jw,L (L m h ) % Difference Membrane Stepwise osmotic pressure Constant osmotic pressure between methods HTI CTA

14.30

±

0.77

14.22

±

0.14

0.54

HTI TFC

14.40

±

0.45

13.70

±

0.37

4.99

Oasys

13.31

±

0.80

13.57

±

0.16

1.92

Porifera

13.59

±

0.45

13.41

±

0.27

1.31

Toray

17.73

±

0.66

14.12

±

0.27

22.68

*13.90

±

0.54

13.80

±

0.35

0.68

Average

*Does not include Toray TFC membrane from stepwise method.

374 375

Table 2 shows less than 5% difference in limiting flux values between the constant osmotic pressure

376

and stepwise osmotic pressure methods for the HTI CTA, HTI TFC, Oasys TFC, and Porifera TFC

377

membranes. This similarity indicates the stepwise method was validated for these membranes. For the

378

Toray TFC membrane, the limiting flux from the constant osmotic pressure method was significantly lower

379

than that with the stepwise osmotic pressure method. As can be seen in Fig. 3 inset data, fouling of the



380

Toray TFC membrane was not fully developed until after 8 hrs of operation with the constant osmotic

381

pressure method whereas all other membranes required 6 hrs or less. This supports the earlier

382

hypothesis (Section 3.2) that flux decline from fouling for the Toray TFC membrane did not reach

383

completion in the stepwise method; this invalidates the stepwise result for Toray TFC membrane and

384

suggests a longer step duration may be required for membranes with high A/B. The constant osmotic

385

pressure method can be used as a standalone procedure for determination of limiting flux because unlike

386

the stepwise method, step duration is not an experimental parameter. However, limiting osmotic pressure

387

cannot be obtained with the constant osmotic pressure method, only with the stepwise method.

388

Therefore, a combined approach using the stepwise method for determination of limiting flux and limiting

389

osmotic pressure followed by validation of the limiting flux and a subsequent recalculation of the limiting

390

osmotic pressure using the constant osmotic pressure method may be necessary.

391

A singular limiting flux value obtained with the constant osmotic pressure method for all membranes

392

despite very different A and A/B values indicates that A and A/B are poor predictors of water flux once

393

membrane fouling occurs. The lack of flux dependence on membrane properties is rationalized by two

394

key observations obtained from the resistance-in-series model (S4.1): 1) membranes with higher A/B

395

have higher foulant cake layer resistance (Fig. S5), and 2) foulant cake layer resistance continually

396

increases with draw solution osmotic pressure. Thus, for osmotic processes where membrane fouling

397

occurs, membrane transport parameters A and B may not be useful performance indicators and the goal

398

of improving water flux by developing highly permeable, highly selective membranes may not be realistic.

399

Instead, optimizing operational parameters such as fouling mitigation strategies (e.g., osmotic

400

backwashing, hydraulic scour, and air scour), hydrodynamics at the membrane surface (e.g., operating

401

conditions and/or spacer and module design), and membrane module configuration (e.g., flat-sheet,

402

tubular, or other) may be more promising alternatives for improving performance for applications where

403

limiting flux is set by the system, the operating conditions, and the foulant/draw solution interactions

404

rather than the membrane properties.

405

Acknowledgements

406

This work was supported by the Strategic Environmental Research and Development Program (SERDP

407

ER-2237), the National Science Foundation Graduate Research Fellowship Program (nsf13584), the


Page 18 of 22

Page 19 of 22


408

University of Southern California’s Viterbi School of Engineering Fellowship, and the 2016 American

409

Membrane Technology Association/U.S. Bureau of Reclamation Fellowship Program. The authors would

410

like to thank Porifera Inc., Toray Industries Inc., Oasys Water Inc., and Hydration Technologies LLC for

411

providing the FO membranes.

412

Supporting Information

413

Critical flux observations in FO literature; additional details and figures describing limiting flux concept in

414

FO; ancillary information on FO testing system and preparation of synthetic activated sludge; example of

415

limiting osmotic pressure calculation; example of limiting flux calculation with constant osmotic pressure,

416

resistance-in-series model and experimental results.

417 418

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419 420 421 422 423 424 425 426 427 428 429 430 431 432 433 434 435 436 437 438 439 440 441 442 443 444 445 446 447 448 449 450 451

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Evidence, Determination, and Implications of Membrane-Independent

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