Impacts of silica on the sustainable productivity of reverse osmosis membranes treating low-salinity brackish groundwater

Desalination 279 (2011) 210–218

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Desalination j o u r n a l h o m e p a g e : w w w. e l s ev i e r. c o m / l o c a t e / d e s a l

Impacts of silica on the sustainable productivity of reverse osmosis membranes treating low-salinity brackish groundwater Mohammad Badruzzaman a,⁎, Arun Subramani a, James DeCarolis b, William Pearce c, Joseph G. Jacangelo d a

MWH, 618 Michillinda Ave, Suite 200, Arcadia, CA 91007, United States MWH, 9444 Farnham Street, Suite 300, San Diego, CA 92123, United States City of San Diego Water Department, 600 B Street, Suite 600, San Diego, CA 92101, United States d MWH, 40814 Stoneburner Mill Lane, Lovettsville, VA 20180, United States b c

a r t i c l e

i n f o

Article history: Received 3 November 2010 Received in revised form 3 June 2011 Accepted 6 June 2011 Available online 7 July 2011 Keywords: Brackish groundwater Reverse osmosis Antiscalant Chemical cleaning Silica Scaling Specific flux Rejection Feedwater recovery

a b s t r a c t Brackish water desalination is increasingly being considered as an alternative drinking water treatment processes. The impacts of high silica (63 mg/L) on the performance of brackish water reverse osmosis (BWRO) while treating a brackish groundwater (1315 mg/L of TDS) were investigated. This study demonstrated the importance of pilot-scale testing in understanding the long-term impacts of silica fouling in order to maintain the sustainable flux and recovery of new generation BWRO membranes. Two commercial membranes, selected from bench-scale coupon testing, were evaluated at pilot scale. The results suggested that due to the presence of high silica, operational flux and recovery were limited to 16.9 Lm − 2 h− 1 and 75%, respectively, for both RO membranes. The silica fouling predominantly occurred in the second stage of the RO systems, and membrane autopsy revealed that the foulant layer was primarily comprised of inorganic elements (10% organic, 90% inorganic) including silicon and a small amount of aluminium. The antiscalant and chemical cleaning were essential to minimize silica scaling and to restore flux; however, the flux decline rates were faster in the subsequent cycles indicating the detrimental effects of silica deposition. The annual chemical cost could be significantly high due to the increased frequency of chemical cleaning. © 2011 Elsevier B.V. All rights reserved.

1. Introduction Brackish groundwater constitutes more than half of the total available groundwater comprising the global water balance [1]. This water is expected to be considered more frequently as an alternative source of drinking water in the near future [2–4,6]. Based on a recent survey, about 19% of the desalination plants in the world are installed for brackish water treatment [5]. Brackish groundwater typically contains a high level of total dissolved solids (TDS) ranging from 1000–10,000 milligrams per liter (mg/L). As such, the treatment of this water requires a desalination process such as reverse osmosis, electrodialysis reversal, multiple effect distillation, or multistage flash evaporation. Reverse osmosis (RO) is one of the most widely considered treatment technologies for brackish groundwater desalination in the US and other parts of the world [7–10]. According to a recent survey, about 59% of the total desalination systems installed in the world (i.e. 13,869 desalination plants) are reverse osmosis

⁎ Corresponding author. Tel.: + 1 626 568 6013; fax: + 1 626 568 6015. E-mail addresses: [email protected] (M. Badruzzaman), [email protected] (A. Subramani), [email protected] (J. DeCarolis), [email protected] (J.G. Jacangelo). 0011-9164/$ – see front matter © 2011 Elsevier B.V. All rights reserved. doi:10.1016/j.desal.2011.06.013

systems [5]. Moreover, with the continual emergence of microcontaminants, water treatment trains are increasingly being designed to achieve multiple treatment objectives that can be effectively met by this technology [10]. High concentration of ions in the brackish groundwater source often increase membrane scaling through precipitation of sparingly soluble salts [11,12]. In brackish RO systems, the dissolved salts are normally concentrated 4 to10 times, often causing high concentrations of ions exceeding the solubility at the membrane surfaces [13,14]. Therefore, the inorganic composition and scaling potential of the brackish groundwater need to be carefully examined prior to designing RO systems for desalination. While seawater chemistry is dominated by sodium and chloride ions; divalent calcium and sulfate ions constitute a major fraction of the TDS content of the brackish groundwater [15,16]. Thus, mineral scaling primarily through crystallization and particle formation at the membrane surface by these ions are important factors determining the operational conditions of the RO process [17–20]. Calcium sulfate, calcium carbonate, and barium sulfate are some of the most common inorganic salts responsible for surface (i.e., heterogeneous) crystallization and scaling on the membrane surface [21,22]. In order to alleviate mineral precipitation and improve membrane productivity, strategies like pH adjustment, antiscalant addition, and chemical cleaning have been successfully used [23–26].

M. Badruzzaman et al. / Desalination 279 (2011) 210–218

The presence of high levels of silica along with sparingly soluble ions may pose substantial challenge to the implementation of brackish groundwater desalination using RO membranes [27]. The complex chemistry of silica could cause scaling effects in various ways [28]. A number of mechanisms of silica scaling have been hypothesized including (i) colloidal submicron silica deposition on the membrane in an agglomerated form [29], (ii) polymerization of silicates in the bulk solution and on the membrane surface [30,31], and (iii) co-precipitation of silicates and deposition on the membrane surface [32]. Koo et al. (2001) showed that an increase in the concentration of calcium and magnesium enhanced the polymerization of silica on the membrane surface [33]. Many of these published studies investigated silica scaling on RO membranes through short-term bench-scale tests, some of which demonstrated that the scaling caused by silica could be different compared to mineral salt precipitation. For example, Braun et al. demonstrated that the scaling layer formed by divalent ions (e.g., Ca 2+) is significantly different from the scaling layer developed by silica [34]. Forensic evaluation of the exhausted membrane surface indicated that silica formed an amorphous solid layer rather than a crystalline solid typically caused by heterogeneous crystallization of divalent ions [29,33]. Therefore, due to the difference in scaling mechanisms, the effectiveness of mineral scaling inhibition strategies (e.g., antiscalant addition, chemical cleaning) in minimizing the effect of silica scaling on the membrane performance is still requires better understanding. For instance, there is still a need to understand the impact of silica on the RO membranes over an extended period of time in order to determine the impact of silica scaling on the chemical cleaning interval and consequently on annual chemical consumptions. Furthermore, bench-scale testing and computer modeling approaches for membrane performance prediction have inherent limitations to evaluate the effects of antiscalants and chemical cleaning on the sustainable flux and recovery. Such information could be critically important for widespread application of the RO for the brackish groundwater treatment containing high concentrations of silica. The specific objectives were to: (a) evaluate the long-term effect of high silica on the sustainable flux and recovery of two BWRO membranes, (b) understand the rate of flux decline for a number of subsequent cleaning cycles, (c) determine the effectiveness of antiscalants and chemical cleaning minimizing the impacts of silica scaling, and (d) estimate economic impacts of the silica scaling on the chemical consumption. 2. Materials and methods 2.1. Bench-scale experiments 2.1.1. Feed water Groundwater samples from the San Pasqual Basin of San Diego, CA were employed. Upon receipt, the water samples were stored at 4 °C to prevent any precipitation and microbial growth during storage. Historical water quality of the groundwater source is presented in Table 1. The average TDS and silica concentrations were about 1315 and 63 mg/L, respectively. 2.1.2. Membranes Four BWRO membranes were evaluated at bench scale: Membrane A (Saehan BLR), Membrane B (Toray TMG10), Membrane C (Hydranautics ESPA2) and Membrane D (Toray TM710). Brief descriptions of the membrane elements are presented in Table 2. The membranes were supplied as 4 inch spiral wound elements. Flat sheet membranes (20.3 cm × 10.1 cm) were cut from each module for the bench-scale experiments; the effective membrane surface area for filtration was 154.8 cm 2.

211

Table 1 Average feed water quality used for the bench- and pilot-scale evaluations. Analytes

Units

Alkalinity

mg CaCO3/ L Aluminum mg/L Ammonia mg-N/L Barium μg/L Boron μg/L Bromide μg/L Bicarbonate mg/L Chloride mg/L Chromium μg/L Copper μg/L Fluoride μg/L Ca Hardness mg CaCO3/ L Total mg CaCO3/ Hardness L Calcium mg/L Magnesium mg/L Iron μg/L Mercury μg/L Nickel μg/L Nitrate mg/L pH pH Units Ortho μg/L Phosphate Selenium μg/L Silica mg/L Sodium mg/L Sulfate mg/L TDS mg/L TOC mg/L Vanadium μg/L Zinc μg/L SDI −

Method detection level (MDL)

No of samples (n)

Values

3

19

238

10 0.03 40 5 0.1 0.5 1 2 0.1 2

1 1 7 9 20 11 20 1 5 19 20

5 4 173 84 697 289 305 333 7 440 449

2

19

511

2 2 2 0.2 2 1.95 0.01 0.2

11 10 3 21 1 20 19 5

175 16 592 1 3 37 7 254

3 2.5

6 11 13 21 19 13 10 6 19

4 63 236 307 1315 1 18 16 5

0.5 10 0.3 3 8 −

2.1.3. Experimental set-up and test procedure Bench-scale evaluations were conducted using a SEPA® cell obtained from Osmonics (Osmonics, Minnetonka, MN). The testing module which was constructed of stainless steel accommodated a 154.8-cm 2 flat sheet membrane. The experiments were conducted in a closed loop RO unit as shown in Fig. 1. Feed water was applied to the test cell using a high-pressure pump (Ryan Harco, USA). The volume

Table 2 Description of membranes employed in this study.

Manufacturer Product ID Element Size (cm) Configuration Membrane area (m2) Nominal rejection (%) Permeate flow (L/min) Membrane Type Membrane material

Membrane Aa

Membrane Bb

Membrane Cc

Membrane Dd

Saehan BLR 10 × 101

Toray TMG 10 10 × 101

Hydranautics ESPA 2 10 × 101

Toray TM 710 10 × 101

Spiral wound 7.9

Spiral wound 7.9

Spiral wound 7.9

Spiral wound 7.9

99.5

99.5

99.6

99.6

5

4.2

5.0

5.8

TFC

TFC

TFC

TFC

Composite Polyamide

Composite Polyamide

Composite Polyamide

Composite Polyamide

a Test conditions: 1500 mg/L NaCl, 150 psig pressure, 770 F, recovery rate 15%, feed pH 7. b Test conditions: 1500 mg/L NaCl, 150 psig pressure, 770 F, recovery rate 15%, feed pH 7. c Test conditions: 500 mg/L NaCl, 110 psig pressure, 770 F, recovery rate 15%, feed pH 7. d Test conditions: 2000 mg/L NaCl, 225 psig pressure, 770 F, recovery rate 15%, feed pH 7.

and and and and

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Chiller

PI

PI

Fig. 1. Schematic of the bench-scale RO experiment setup.

of water in the feed tank used during the test was 5 L. Flow rates were controlled by varying the pump speed and pressure was controlled using a back pressure regulator (Swagelok, USA). Temperature was maintained by circulating cooling water through a stainless steel coil immersed in the feed tank. A constant temperature of 25 °C was maintained throughout the experiment. Membranes sheets were cut and placed in the test cell and equilibrated with deionized water spiked with NaCl at 1722 kPa until there was no decrease in specific flux. Since polyamide RO membranes compact when pressure is initially applied, all flat sheet RO coupons were compacted first with NaCl spiked distilled water in order to differentiate between compaction and fouling. Equilibration time typically ranged from 6 to 8 h. Appropriate quantities of NaCl were added during membrane compaction to represent similar TDS of the groundwater being tested. At the end of this period, membrane fouling experiments were conducted using actual brackish groundwater (Table 1) collected from the pilot test site. An initial flux of 33.9 Lm − 2 h − 1 and constant feed flow of 850 mL/min (0.5 m/s) was maintained in all the experiments. Both permeate and concentrate were recycled back (total recycle) to the feed tank. Since silt density index (SDI) measurements for the raw groundwater were high (5.0), pretreatment with a cartridge filter was used. 2.2. Pilot-scale evaluation 2.2.1. Pilot site Pilot evaluations were conducted in Escondido, CA. Source water was pumped from the San Pasqual Aquifer using an existing well that was 15.8 m deep and consisted of 10-cm diameter schedule 40 PVC casing. The well was screened with 0.05-cm slotted PVC screen from 1 to 6 m below the ground surface. Average water quality of the water tested is presented in Table 1.

for Stage 2. Each vessel housed three RO elements with nominal dimensions of 10 cm by 101 cm. Fig. 2 shows a schematic of the RO pilot system with the flow scenarios. All equipment related to the pilot equipment such as pressure gauges, flow meters, and safety switches were calibrated on-site throughout the pilot-test period. The system rotameters, digital flow meters, system pressure and vacuum gauges were verified at the beginning, mid and end of the testing. Feed, permeate, and concentrate stream pressures at different stages of the operation are presented in Table 3. Also, the net osmotic pressures and net operating pressures at both stages of the membrane systems are listed in Table 3. It should be noted that the permeate pressure in stage 1 was high because the permeate flow was throttled to provide sufficient cross-flow velocity to the second stage elements. 2.2.3. Materials Piloting was conducted with Membrane A and Membrane B. A bag filter (Hayward, USA) and a cartridge filter (1 micron, Ryan Herco, USA) were used to remove particles. Chemical cleaning was performed using three different chemicals: Cleaning Chemical A (30 min cleaning and 30 min flush), Cleaning Chemical B (30 min cleaning, 60 min soak, 30 min cleaning and 20 min flush) and Cleaning Chemical C (20 min cleaning and 20 min flush). All chemicals used in the pilot evaluations including antiscalants and cleaning chemicals are listed in Table 4. 2.3. Analytical methods All analyses were performed using the standard methods [35] and EPA recommended methods [36]. A list of analytical methods used for this study is presented in Table 5. 2.4. Calculation of operational parameters

2.2.2. Description of the RO pilot unit The RO pilot system was configured in a 2–1 array. Stage 1 consisted of four pressure vessels arranged as two parallel arrays and Stage 2 contained two pressure vessels. The stages were arranged in series, which allowed concentrate from Stage 1 to serve as feed water

A number of calculated parameters were used to establish the performance of the RO membrane systems. These calculated parameters are defined below. The equations presented here were developed from Schippers et al., AwwaRF, and Hydranautics [37–39].

Antiscalant Feed Tank

Cleaning Tank Groundwater (Well)

Fig. 2. Schematic of the pilot-scale membrane system.

M. Badruzzaman et al. / Desalination 279 (2011) 210–218 Table 3 Operating pressures at different points of the membrane process.

Table 5 A list of analytical methods used in this study.

Net osmotic Net operating Permeate Conc. Time Feed pressure (kPa) pressure pressure (kPa) (h) pressure pressure (kPa) (kPa) (kPa) Stage 1 Stage 2 Stage 1 Stage 2 Stage 1 Stage 2 Membrane A 0 1061 258 1068 0 689 133 758 135 717 460 1034 478 799 593 930

214 276 193 234 207 496 276 393

17 10 0 0 0 69 69 10

792 792 537 634 579 903 675 792

147 147 149 149 149 150 149 150

203 202 205 205 205 206 205 205

641 577 312 340 310 353 343 353

641 671 373 455 393 659 432 473

Membrane B 0 758 652 744 0 827 165 1350 0 717 160 785 0 703 165 1054

262 276 358 889 269 351 262 551

79 54 55 52 53 51 55 52

634 668 703 1233 572 655 572 965

150 153 174 159 153 153 154 181

206 204 274 255 209 210 209 251

295 277 260 268 260 246 255 298

359 369 401 950 347 423 342 720

Qp  e−0:0239ðT−25Þ S

ð2Þ

where, Pnet = net operating pressure (kPa), Pi = pressure at the inlet of the membrane module (kPa), Po = pressure at the outlet of the membrane module (kPa), P p = permeate pressure (kPa), and Δπ = net osmotic pressure of the feed and permeate (kPa). Specific flux: specific flux was calculated according to the following Eq. (3). Jsp =

Jtm ðat 25-C Þ Pnet

Reference

EPA Method 300.1 2340C 4500-CO2 D 2320B 2550B 4500H+ 2510B 200.7

US EPA (2010) APHA (2005) APHA (2005) APHA (2005) APHA (2005) APHA (2005) APHA (2005) USEPA (2010)

ð4Þ

where, ΔP = differential pressure (kPa), Pf = pressure measured in RO feed (kPa), and Pc = pressure measured in RO concentrate (kPa). Feedwater recovery: feed water recovery (FWR) represents net water production of the RO systems which was calculated according to Eq. (5):

ð1Þ

ðPi + Po Þ −Pp −Δπ 2

Method Number

Chloride, nitrate, sulfate Total hardness Bicarbonate Total alkalinity Temperature pH Conductivity Calcium, sodium, silica and other trace metals

ΔP = Pi −Pc

where, Jtm = instantaneous flux (Lm − 2 h − 1), Qp = permeate flow, (m 3h − 1), T = temperature, (°C), and S = membrane surface area, (m 2). Net operating pressure: average net operating pressure for the RO membrane system was calculated according to Eq. (2): Pnet =

Parameter

Differential pressure: differential pressure of the RO membranes is the difference between the feed pressure and concentrate pressure, calculated according to Eq. (4):

Temperature adjustment for flux: temperature correction to 25 °C for transmembrane flux of the RO membranes was made according to Eq. (1) which is based on the variation of water viscosity with temperature: Jtm ðat 25-C Þ =

213

ð3Þ


 Qp FWR ¼  100% Qf

ð5Þ

where, Qf = feed flow (m 3h − 1), and Qp = product flow (m 3h − 1). Solute rejection: rejection of ions by the RO process was calculated according to Eq. (6): R=

! Cp  1− 100% Cf

ð6Þ

where, R = rejection (%), Cp = product water concentration (mg/L), and Cf = feedwater concentration (mg/L). Net osmotic pressure: the net osmotic pressure is the difference between feed and permeate osmotic pressures as calculated by the following Eq. (7) Δπ = IAF  πf 

Cp 1− Cf

!

where, πf = Water osmotic pressure = 0.0115* TDS*6.89 (kPa); Cp = product water concentration (mg/L), and Cf = feedwater concentration (mg/L), IAF = Integrating average factor calculated based on the following Eq. (8)  1  ln 1−FWR FWR

where, Jsp = specific flux (Lm − 2 h − 1/kPa); Jtm(at 25° C) = temperature corrected flux Lm − 2 h − 1), and Pnet = net operating pressure (kPa).

IAF =

Table 4 Chemicals used in the pilot evaluations.

3. Results and discussion

Chemical name

Suppliers

Antiscalant A (Pretreat King Lee Plus Y2K) Technologies, USA Antiscalant B (VITEC 3000) Avista Technologies, USA Cleaning Chemical A King Lee (Diamite LpH) Technologies, USA Cleaning Chemical B (High King Lee Flux A) Technologies, USA Cleaning Chemical C King Lee (Diamite AFT) Technologies, USA

Application type

ð7Þ

ð8Þ

where, FWR = feed water recovery expressed in decimal, calculated according to Eq. (5).

3.1. Selection of membranes for piloting

Antiscalant Antiscalant Low pH chemical cleaning (pH = 2.5) Chemical cleaning (pH = 3.8) High pH chemical cleaning

A large number of RO membranes are available from different manufacturers for brackish groundwater desalination. The evaluation of a wide variety of membranes through pilot-scale testing is often time consuming and economically prohibitive. Therefore, flat-sheet bench-scale evaluations were utilized in this study to select two BWRO membranes. Specific fluxes of four tested membranes, a measure of membrane resistance and productivity, were determined. The results suggested that Membranes A (0.82 Lm− 2 d− 1/kPa) and B

214

M. Badruzzaman et al. / Desalination 279 (2011) 210–218

(0.93 Lm − 2 d − 1/kPa) had higher initial specific fluxes (at t = 0) when compared to Membranes C (0.72 Lm − 2 d − 1/kPa) and D (0.62 Lm− 2 d− 1/kPa) under similar hydrodynamic conditions. Since initial specific fluxes were different for the membranes evaluated, values observed at the different time points were normalized with respect to the initial specific flux (at t=0) for each membrane. A comparison of the normalized specific flux for four membranes is shown in Fig. 3a. The comparison of normalized specific flux results indicate that Membrane D exhibited the highest degree of fouling potential when compared to the other three membranes. Normalized specific fluxes were not substantially different for Membranes A, B and C; therefore, other membrane properties such as salt rejection was considered. The observed TDS rejection by four membranes is shown in Fig. 3b. All four membranes exhibited stable rejection (N97%) of TDS. The high rejection was expected since all the membranes were constructed of cross-linked aromatic polyamine (PA) chemistry which typically shows very high removal efficiencies of dissolved solutes [40]. However, Membrane C (97.7%) showed the lowest rejection of the four evaluated. It should be noted that based on such a small difference in rejection, it was not possible to eliminate Membrane C as a potential choice for pilot testing since the pilot testing may have demonstrate that this membrane has less fouling potential, higher specific flux, or other properties that are desirable for long-term performance. But, due to logistical concerns, the investigators opted to select only two membranes. As such, Membranes A and B were selected for pilot testing.

CSMPro3 Software, the maximum saturation limit of silica and the antiscalant type and dosage selected. Temperature of the feed water remained within the range of 20 to 24 °C. The temperature corrected flux results are shown in Fig. 4a. The specific flux remained fairly constant in the first stage, but decreased in the second. Although antiscalant (4 mg/L, Antiscalant A as described in Table 4) was added to the feed stream, the net operating pressure of the overall system increased (641 to 647 kPa, Table 3) and the specific flux decreased more than 20% compared to the start-up condition in that period. The specific flux decline rate was about 1.1 × 10 − 3 Lm − 2 d − 1/kPa/h. These data indicated a need for chemical cleaning at 250 h of operation (i.e., 10 days) which was significantly lower than the economical cleaning frequency [15]. The high frequency of chemical cleaning would increase the annual chemical consumption cost and could also reduce the life-time of the membrane. Therefore, the system flux was reduced to 16.9 Lm − 2 h − 1 and operated at the same recovery of 75%. The net operating pressure of the overall system was reduced to 312 kPa for this lower flux operation. The performance was monitored for three cleaning cycles. The specific fluxes for both first and second stages are shown in Fig. 4b. The specific flux for the first stage remained stable for 650 h of operation, whereas it decreased steadily in the second stage. The first chemical cleaning was performed at 133 h of operation. During that period specific flux at 25 °C was reduced about 14% of the initial value at a rate of 1.2 × 10 − 3 Lm − 2 d − 1/kPa/h, which was similar to the rate observed at 25.5 Lm − 2 h − 1 operation. Specific flux in the second stage was recovered completely after the first chemical cleaning cycle. After restarting the operation, the flux in

3.2. Pilot-scale evaluation of membrane performance

Normalized Specific flux @ 25°C

(a) 1.0

0.8

Membrane A Membrane B Membrane C Membrane D

0.6

(a) Specific Flux at 25°C (Lm-2d-1/kPa)

3.2.1. Performance of membrane at different flux scenarios The operation of Membrane A was started at 25.5 Lm − 2 h − 1 and 75% feedwater recovery based on the modeling using Saehan

2.0

First Stage Second Stage 1.5

Slope: -8.8x10-4 Lm-2d-1/kPa/hr

1.0 Slope: -1.1x10-3 Lm-2d-1/kPa/hr

0.5

0.0 0

50

1

2

3

4

5

6

7

8

9

150

200

250

300

10

(b)

Test Duration, hr

(b) 100

98.87

98.89

99

98.24 97.76

98

97

96

Specific Flux at 25°C (Lm-2d-1/kPa)

0

TDS Rejection, %

100

Time of Operation, hrs

0.4

2.0

First Stage Second Stage Slope: -7.7x10-5 Lm-2d-1/kPa/hr

1.5

1.0 Slope: -1.2 x10-3 Lm-2d-1/kPa/hr

0.5

Slope: -1.54x10-3 Lm-2d-1/kPa/hr

Slope: -2.9x10-3 Lm-2d-1/kPa/hr

0.0 0

100

200

300

400

500

600

700

Time of Operation, hrs

95 Mem A

Mem B

Mem C

Mem D

Fig. 3. Bench-scale experimental results with BWRO membranes (a) normalized specific fluxes at 25 °C, (b) TDS rejections observed.

Fig. 4. Specific flux profiles adjusted at 25 °C with Membrane A at different operating conditions (a) 4 mg/L of antiscalant A at 15 gfd and 75% recovery, (b) 4 mg/L of antiscalant A at 10 gfd and 75% recovery.

M. Badruzzaman et al. / Desalination 279 (2011) 210–218

the second stage continued to decline until the next chemical cleaning was performed at 460 h of operation (Fig. 4b). During that period, the rate of specific flux decline was about 1.54 × 10 − 3 Lm − 2 d − 1/kPa/h, which was about 1.3 times of the flux decline rate observed in the previous cycle. After chemical cleaning of the membrane at 478 h of operation, the recovery of the flux was again established. However, a steeper flux decline continued to be observed (2.9 × 10 − 3 Lm − 2 d − 1/kPa/h) which was about three times of that observed at 25.5 Lm − 2 h − 1 operation. This continual increase in the flux decline indicated that a potential residual of foulants/scalants were not properly removed by chemical cleaning and consequently caused accelerated fouling propensity. Similar results were observed by Koo et al. who showed that chemical cleaning solutions are not effective to restore the flux decline caused by the colloidal deposition of silica [33].

3.2.2. Effects of membrane type At the end of Membrane A operation, Membrane B was operated at 16.9 Lm − 2 h − 1 flux and 75% recovery in order to compare its performance with Membrane A. An antiscalant A dose of 4 mg/L was added to the system. A net operating pressure of about 323 kPa was required to achieve the flux; this pressure was similar to that required for Membrane A. Initial system specific flux for the Membrane B was determined to be 1.36 Lm − 2 d − 1/kPa, slightly higher than Membrane A which was 1.12 Lm − 2 d − 1/kPa. The specific fluxes of the both stages of Membrane B are in Fig. 5a. The specific flux results obtained in the first 650 h of operation suggest that the specific flux and net operating pressure (359 to 369 kPa, Table 3) did not change significantly (less than 7%). The rate of specific flux decline in the second stage was about 1.4 × 10 − 4 Lm − 2 d − 1/kPa/h, which was about an order of magnitude lower than the rate observed with Membrane A at similar operating conditions suggesting the impacts of scaling was dependent on the membrane utilized. Although the properties of the membranes were not investigated in this study, it has been established by several researchers that surface properties

(a)

such as membrane roughness could significantly influence adhesiveness of the scalants and consequently contribute to flux decline [41,42].

3.2.3. Effects of increased recovery Since Membrane B performance was stable during the initial run, the recovery was increased to 80% at the same flux (16.9 Lm − 2 h − 1) and antiscalant dose (4 mg/L). During the operation at 80% recovery the specific flux in the first stage remained stable whereas it decreased drastically in the second stage (Fig. 5b). The specific flux declined about 58% within 165 h of operation at the rate of 4 × 10 − 3 Lm − 2 d − 1/kPa/h which was an order of magnitude higher than the rate observed at the operation at 75% recovery. Due to the increase in recovery to 80%, the silica concentration in the end element of the second stage was 280 mg/L compared to 235 mg/L for 75% recovery (based on Hydranautics IMS Model results). Hence, a drastic decrease in specific flux was observed in the second stage, probably due to the scaling of high level of silica.

3.2.4. Evaluation of antiscalant type and dose Additional tests were performed with Membrane B to assess the performance of different antiscalant addition strategies including varying antiscalant type and dosage. Antiscalant B at a dose of 4 mg/L was used in order to compare performance with antiscalant A applied at similar dose and test conditions (16.9 Lm − 2 h − 1 and 75% recovery). The specific fluxes for the first and second stage for Membrane B using antiscalant B is shown in Fig. 5c. The specific flux in the first stage was stable. The specific flux in the second stage decreased steadily within 160 h of operation. The flux decline rate was 1.7 × 10 − 3 Lm − 2 d − 1/kPa/hr which was an order of magnitude higher than the flux decline rate observed using antiscalant A at similar test conditions (16.9 Lm − 2 h − 1 and 75% recovery). Hence, these results illustrate that the impacts of scaling and the performance of the membranes is dependent on the type of antiscalant used.

(b)

3

215

3 First Stage

Specific Flux at 25°C (Lm-2d-1/kPa)

Specific Flux at 25°C (Lm-2d-1/kPa)

First Stage Second Stage

Slope: 7.0x10 -4 Lm-2d-1/kPa/hr

2

1 Slope: -1.4x10 -4 Lm-2d-1/kPa/hr

Second Stage

Slope: -6.9x10-4Lm-2d-1/kPa/hr

2

1

Slope: -4x10-3Lm-2d-1/kPa/hr

0

0 0

100

200

300

400

500

600

700

0

800

Time of Operation, hrs

(c)

40

80

120

160

(d)

3

3

Specific Flux at 25°C (Lm-2d-1/kPa)

Specific Flux at 25°C (Lm-2d-1/kPa)

First Stage Second Stage

Slope: 3.7 x10-4 Lm-2d-1/kPa/hr

2

200

Time of Operation, hrs

1

Slope: -1.7 x10-3 Lm-2d-1/kPa/hr

First Stage Second Stage

Slope: -1.3 x10-3 Lm-2d-1/kPa/hr

2

1

Slope: -5.8 x10-3 Lm-2d-1/kPa/hr 0

0 0

40

80

120

160

Time of Operation, hrs

200

0

40

80

120

160

200

Time of Operation, hrs

Fig. 5. Specific flux profiles adjusted at 25 °C with Membrane B at different operating conditions (a) 4 mg/L of antiscalant A at 10 gfd and 75% recovery, (b) 4 mg/L of antiscalant A at 10 gfd and 80% recovery, (c) 4 mg/L antiscalant B at 10 gfd and 75% recovery; (d) 8 mg/L of antiscalant A at 10 gfd and 80% recovery.

216

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To determine the influence of antiscalant dose on productivity, a higher dose of antiscalant A (8 mg/L) was used and the recovery of the RO system was increased to 80%. Specific fluxes of the first and second stage are plotted in Fig. 5d. Even by doubling the dosage of antiscalant from 4 to 8 mg/L, the decrease in specific flux in the second stage was not prevented; instead a flux decline rate 5.8 × 10 − 3 Lm − 2 d − 1/kPa/h was observed. This rate was higher than the flux decline observed at 80% recovery using 4 mg/L antiscalant A. These data suggest that the dose of antiscalant is important but a higher level of antiscalant dose might not be effective in reducing scaling effects. It should be noted here that a control without any antiscalant was not conducted. Although, data without antiscalant would have demonstrated a fouling trend under extremely aggressive conditions, it was not practical to run these tests at pilot-scale. Such tests would probably have damaged the membrane to an extent that no cleaning scheme would have been capable of recovering flux.

(a)

3.3. Evidence of silica scaling To determine the nature of a limiting foulant in the second stage, the end element from the second stage of Membrane B was removed and an autopsy was performed. The autopsy was conducted in two ways: (a) microscopic investigation of the fouled surface of the membrane and (b) extraction of foulant layer from the membrane to determine the composition of the foulants. The morphological images of the membrane surface using scanning electron microscopy (SEM) at different magnifications (Fig. 6) indicate that a complex and thick foulant material was present. Energy Dispersive Spectrum (EDS) of the foulant layer along with atomic composition of the foulant layer is shown in Fig. 7. The spectrum revealed the presence of predominantly silicon (23.2%) and aluminum (0.5%) as the chemical nature of the foulant material. It should be noted that the concentration of carbon was high due to the polyamide RO membrane, which contained a substantial amount of carbon and oxygen functionalities (COO-groups on membrane, rather than carbon due to carbonate formation on the membrane. The representative foulant from the surface of the element was harvested through washing several times with RO permeate by mixing and decantation similar to the techniques applied by other researchers [43]. After drying at 120–130 °C, the dried foulant was analyzed quantitatively for absolute percent by weight of carbon, hydrogen, nitrogen and ash. The sample was scanned using the SEM with EDS. The results as presented in Table 6 show that the foulant was composed of 10% organic and 90% inorganic materials. From elemental analysis (Table 6), it was reconfirmed that silica was the most significant foulant. The presence of other co-occurring ions (i.e., aluminum) on the exhausted surface may have increased silica fouling. Several other researchers have indicated that dissolved aluminum and iron influence silica polymerization and impact the performance of RO [30,32,44,45]. The efficacy of chemical cleaning employed in this study also provided evidence of silica fouling. The recovery of specific flux after chemical cleaning with respect to the specific flux of the fouled membrane was evaluated. After the first cleaning using Chemical A, the specific flux in the second stage remained unchanged relative to the fouled membrane. The second cleaning involved the use of Chemical B which is exclusive for removing silica precipitation. The specific flux recovery in the second stage increased 6% (comparing flux before and after chemical cleaning). After the third cleaning with Chemical B, a complete recovery was observed and the specific flux was close to the initial value observed with the virgin membrane. Since Chemical B and Chemical C (i.e., high pH cleaners as listed in Table 4) were effective in recovering the specific flux to initial value while low pH cleaner was not, it could also be inferred that major foulant in the RO surface could be silica. The calculated osmotic pressures in the second stage of both membranes as shown in Table 3

(b)

Fig. 6. SEM images of the end element of the second stage of Membrane B system at different magnifications: (a) 7000× and (b) 28,000×.

were significantly lower than the net operating pressure suggesting that the specific flux decline was not caused the hydraulic conditions (e.g., flow) at the second stage.

Elements Carbon

Atomic Composition (%) 36.2

Oxygen

40.2

Sodium

<0.1

Magnesium

<0.1

Aluminum

0.5

Silicon

23.2

Phosphorous

<0.1

Sulfur

<0.1

Chlorine

<0.1

Calcium

<0.1

Iron

<0.1

Potassium

<0.1

1.00 1.90 2.80 3.70 4.60 5.50 6.40 7.30 8.20 9.10 Fig. 7. EDS spectrum of the fouled surface of the end element of the second stage of Membrane B (inset — elemental composition).

M. Badruzzaman et al. / Desalination 279 (2011) 210–218

• The pilot results showed the necessary cleaning interval was different for Stage 1 and Stage 2 elements. The cost estimation was performed with an assumption that the Stage 1 elements would require cleaning once per year. • The stage 2 elements assumed to be required much more frequent cleaning; the cost estimations were conducted for 10, 20 and 30 day cleaning intervals. • The plant assumed to be operational at 16.9 Lm − 2 h − 1 flux and 75% recovery continuously for 365 days.

Table 6 Elemental composition of the foulant layer of the second stage RO elements (Membrane B). Element

% of ash

Silicon Calcium Aluminum Iron Sodium Total

91.6 3.1 2.3 1.9 1.1 100.0

3.4. Membrane rejection performance Rejection efficiency of Membranes A and B for major ionic composition such as TDS, alkalinity, hardness, sodium, chloride, magnesium, nitrate, and sulfate shown in Fig. 8. Both membranes achieved a high rejection of TDS with average values of 98.2 and 96.8% for Membrane A and Membrane B, respectively. The rejection observed in the bench-scale experiments was representative of the pilot-scale observation for membrane A, but lower for membrane B. Except for the rejection of total hardness, Membrane A exhibited a higher rejection of ions than Membrane B. Silica rejection was 96% by Membrane A and 94% by Membrane B. Considering the fact that Membrane A has a lower specific flux (1.22 Lm − 2 d − 1/kPa) compared to Membrane B (1.35 Lm − 2 d − 1/kPa), Membrane A can be considered as tighter than Membrane B and hence exhibited higher rejection efficiency. It appeared that membrane fouling did not have significant impact on the rejection properties of membrane. For example, the permeate conductivity was stable (b60 μS/cm) throughout the study period for Membrane A. The permeate conductivity for Membrane B was less than 60 μS/cm for the initial 600 h whereas it increased to 100 μS/cm for the remaining period of the study. Metals such as iron, copper, cadmium, and aluminum were removed to below detection limits by both membranes. Total organic carbon (TOC) was rejected below detection limits by both the membranes. Nitrate rejection was slightly higher for Membrane A (85%) as compared to Membrane B (71%). 3.5. Economic implications One of the major economic impacts of silica scaling in the BWRO membrane processes could be the cleaning chemical consumptions. Conceptual chemical cleaning costs were estimated based on the following assumptions: • The chemicals used and the cleaning intervals established from this pilot testing were considered for this estimation.

Membrane A

110

Membrane B

100

Rejection (%)

217

90 80 70 60

Fig. 8. Average rejection of ions by Membrane A and Membrane B.

Silica

Sulfate

Nitrate

Magnesium

Calcium

Chloride

sodium

Hardness

Alkalinity

TDS

Ions

The estimated annual chemical costs for treating per cubic meter of treated water for 10, 20 and 30 day cleaning intervals were $0.3/m 3, 0.15/m 3 and 0.1/m 3 indicating that the cleaning chemical cost can be a significant part of the operational and maintenance (O&M) costs. Based on a comprehensive survey conducted on reverse osmosis/ nanofiltration full-scale plants, Burbano et al. stated that more than 90% of the plants performed chemical cleaning with median chemical cleaning frequency of approximately twice a year (6 months) [46]. So the chemical cleaning frequency observed in this study was significantly higher than the survey results suggesting that silica scaling can severely impact cleaning frequency. It should be also noted that a higher cleaning interval increases the frequency of membrane replacement and consequently increases the O&M cost. Additionally, frequent membrane cleaning will also increase energy consumption due to more frequent operation of the cleaning pumps and will also generate significant volume of cleaning waste that will need to be disposed. Silica fouling on RO membrane is sometimes minimized through by adjusting feed water pH to above 10; however, in such application, hardness ions need to be removed prior to pH adjustment [47]. This approach requires the use of more complex pretreatment processes such as ion exchange before the RO. In addition, such a treatment train has patent rights (e.g., HERO process), is costly and has been utilized primarily for industrial water treatment. An alternate method involves lowering the pH of the feed water to less than 4 [48]. This method requires considerable addition of acid (based on the alkalinity of the groundwater) and substantially increases overall chemical costs. Moreover, the RO permeate needs to be post stabilized as the pH of RO permeate will be less than 3.5 when such a pH adjustment method is employed. Although these approaches may offer site specific solutions for minimizing silica fouling, the economic evaluation in this study was only focused on the utilization of antiscalant type and dose.

4. Conclusions Silica scaling and its impacts on RO membrane performance has been a major concern for water utilities. This study conducted with two new generation BWRO membranes showed that the presence of 63 mg/L of silica limited the operational flux and recovery to 16.9 Lm − 2 h − 1 and 75%, respectively for a brackish groundwater with a TDS content of about 1315 mg/L. Silica scaling predominantly occurred in the second stage of the RO system while the first stage specific flux was fairly consistent. At that operating condition, the cleaning frequency of the second stage membranes could be as frequent as 10 days depending on the membrane type that might cause a significant increase annual chemical consumption. To recover the specific flux of the second stage, this study examined the effectiveness of the combined applications of cleaning chemicals (low pH, silica cleaner and high pH) and antiscalant type and dose on the flux and recovery of BWRO membranes. The study showed that the cleaning chemicals were capable of restoring flux, but the flux decline rates were faster in the subsequent cycles indicating the detrimental effects of silica deposition for an extended period of operation.

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Acknowledgements The authors would like to express their gratitude to the following individuals and organizations for their contributions to the successful completion of the project: Dr. Samer Adham with Conoco Philips, and Dale Rohe, Karla Kinser, Zakir Hirani and Eric Bruce of MWH for their valuable advice. Susan Brannian and lab staffs at the City of San Diego Alvarado Water Quality Laboratory are also acknowledged for analytical support. This research was funded through California Proposition 13 under the Desalination Research and Innovation Partnership (DRIP) program. The comments and conclusions detailed in this paper do not reflect the views of the DRIP program, its officers, affiliates or agents.

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