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Separation and Purification Technology 75 (2010) 165–173

Contents lists available at ScienceDirect

Separation and Purification Technology

journa l homepage: www.e lsev ier .com/ locate /seppur

erformance of a ceramic ultrafiltration membrane system in pretreatment toeawater desalination

ia Xua,∗, Chia-Yuan Changb, Congjie Gaoa

Key Laboratory of Marine Chemistry Theory and Technology, Ministry of Education, College of Chemistry and Chemical Engineering, Ocean University of China, Songling road 238,ingdao, Shandong, 266100, ChinaDepartment of Environmental Engineering and Science, Chia Nan University of Pharmacy and Science, 60, Erh-Jen Road, Sec. 1, Tainan 71710, Taiwan

r t i c l e i n f o

rticle history:eceived 12 May 2010eceived in revised form 28 July 2010ccepted 30 July 2010

eywords:eramic membraneltrafiltrationesistanceretreatmenteawater desalination

a b s t r a c t

Pretreatment to reverse osmosis (RO) seawater desalination historically has been achieved using organic(polymer) ultra- and micro-filter (UF/MF) membranes. Newly developed inorganic (ceramic) membranesoffer unique advantages over the currently employed membranes, and were recently introduced for thepurpose. In this work, we investigate the performance of a zirconium dioxide ceramic membrane with0.05 �m pore diameter to clarify raw seawater under different operating conditions. The influences ofcross-flow velocity (2.7–4.9 m s−1), temperature (15–55 ◦C), transmembrane pressure (TMP) 0.1–0.2 MPa,and seawater pH (5.5–9.5) on permeate flux and rejection were assessed. The results show that thezirconium dioxide membrane at cross-flow velocity of 3.7–4.2 m s−1, TMP of 0.14–0.18 MPa, temperatureof 25–30 ◦C and seawater pH of 8.0–9.0 exhibited a high flux of 420–450 L m−2 h−1 with turbidity andCODMn rejection of 99.0–99.5% and 32–35%, respectively. This work also introduces a new experimentalprocedure to determine different filtration resistance (Rm, Rsads, Rpads, Rcp, Rpb and Rcr) components, whichhelps the analysis of performance characteristics and the better understanding of interactions between

the membrane and foulants. The present results show that the effect of concentration polarization wasdominant in the ultrafiltration of raw seawater, and various fouling components contribute to the totalin the order of Rcp > Rsr > Rm > Rsads > Rpads > Rirr. Further, reversible fouling was much more significantthan irreversible fouling. The experimental method also showed interrelation between the reversibleand irreversible resistances, whereby part of Rsr is transformed into Rirr within 20 min of filtration. The

ate thrmea

results of this study indiccan achieve consistent pe

. Introduction

Fresh water available to communities and factories in Chinand worldwide is typically reliant on capturing surface water andurrently inflows have been depleted below record levels. Strenu-us endeavors are under way to address this problem to produceufficient fresh water and improve water supplies such as wastewa-er reuse and brackish/seawater desalination in particular. Reversesmosis (RO), pursued for over 40 years, has been proven to ben efficient and reliable technology applied into seawater desali-ation allowing for more certain supply from the high abundantcean resource. The technologies for mature RO systems are how-

ver limited to high energy demand and membranes themselveshat are prone to fouling, degradation and delamination of the

embrane films [1–3]. Pretreatment using MF/UF and nanofiltra-ion (NF) technologies has been promoted as an effective means

∗ Corresponding author. Tel.: +86 532 66781872; fax: +86 532 66781872.E-mail addresses: gaocjie@mail.hz.zj.cn, qdxujia@sina.com.cn (J. Xu).

383-5866/$ – see front matter © 2010 Elsevier B.V. All rights reserved.oi:10.1016/j.seppur.2010.07.020

at zirconium dioxide ultrafiltration pretreatment before RO desalinationte quality and low fouling potential at high permeate fluxes.

© 2010 Elsevier B.V. All rights reserved.

to reduce the fouling potential and provide consistently high qual-ity feed to RO desalination systems [4,5]. An overview of seawaterdesalination pretreatment associated with the application of dif-ferent pressure driven membrane processes including MF, UF andNF has been reported by Van der Bruggen and Vandecasteele [6].In the article, the authors described the superiority of MF, UF andNF membrane over the traditional pretreatment processes for theseawater desalination pretreatment based on the considerations ofhigh feed water quality of RO, less energy consumption and highsystem compatibility.

Typical MF/UF pretreatment technologies to RO desalination useorganic membranes made of polyvinyl chloride (PVC), polyvinyli-dene fluoride (PVDF), polyether sulfone (PES) and other polymers.Raw seawater is characterized by high salt content, presence oforganic foulants and biological activity, which result in the rapid

fouling of polymeric membranes [7,8]. It is generally known thatceramic membranes, such as alumina and silica, have superiormechanical properties, chemical inertia, long working life, andthermal stability [9–12] over polymeric membranes. Their abil-ity to provide consistent and high quality feed to RO membranes

166 J. Xu et al. / Separation and Purification Technology 75 (2010) 165–173

Table 1Typical composition of raw seawater.

Parameter Value Parameter Value

TSS (mg L−1) 4.24 × 104 pH 7.9–8.2Ferric/ferrous oxides (mg L−1) 9.14–10.23 Na+ (mg L−1) 1.36 × 104

Turbidity, NTU 0.82–15.7 Cl− (mg L−1) 1.70 × 104

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Alkalinity as CaCO3 (mg L−1) 2.01 × 102 SO42+ (mg L−1) 2.43 × 103

COD (mg L−1) 1.40–2.40 Ca2+ (mg L−1) 4.90 × 102

Silica (mg L−1) 0.36–0.90 Mg2+ (mg L−1) 1.38 × 103

t high permeate flux and low fouling potential, is still largelynexplored.

In this study we investigate the performance of a developedaboratory-scale zirconium dioxide ceramic UF membrane in treat-ng raw seawater. The objectives of this study are (i) to assesseramic membrane filtration as pretreatment technology; (ii) tolucidate the impact of different operating parameters on perme-te flux and foulants rejection; (iii) to evaluate fouling potentialnd membrane/foulant interactions by a full-scale assessment ofltration resistance using an improved experimental procedure.

. Materials and methods

.1. Source water

The experiments were conducted at Huangdao power plantJiaozhou Bay of Yellow Sea, China). The main characteristics of theeawater used in this study are listed in Table 1, noting that wateremperatures varied between 4 and 26 ◦C.

.2. Ceramic membrane

A commercial monolith ceramic UF membrane (XS-T-01) withmean pore size of 0.05 �m was used to filter raw seawater underifferent operating conditions. The specifications of this membraneupplied by the manufacture are given in Table 2, while the exper-mental setup is shown in Fig. 1.

As shown in Fig. 1(a), the raw seawater was used as collectedwith no pre-filtration). The temperature of raw seawater (20 L) inhe feed tank was kept constant at a desired temperature through-ut experiments, using a water bath. The pH of the feed wasdjusted between 5.5 and 9.5 with NaOH or HCl solutions. Theeed was circulated through the UF membrane module by a Grund-

os CHI-4 centrifugal pump. The valves upstream and downstreamV03–V06) of the membrane module, including of the bypass loopere adjusted to obtain the desired operating transmembraneressure (TMP) and cross-flow velocity.

able 2aterial characteristics and module details of the membrane used in this study.

Item Description

Manufacturer Research Center of MembraneScience and Technology, NanjingUniversity of Technology, China

Surface area (m2) 0.1Pure water permeability

(L m−2 h−1 bar−1)610

Active layer material Zirconium dioxideMembrane support material �-Alumina oxidePore size (�m) 0.05Membrane type TubularMWCO (kDa) 100Porosity (%) 37.5Length (mm) 550Number of channels 19Channel diameter (mm) 3

Fig. 1. Experimental setup of the ceramic UF membrane filtration system: (a)schematic diagram; (b) picture.

2.3. Ultrafiltration process

The laboratory (bench) scale filtration experiments were con-ducted in total recycle mode, whereby both retentate and permeatewere returned to the feed tank. The pure water flux (Jw0) ofthe membrane was carefully measured and recorded before eachrun. Conventional one-factor-at-a-time analysis was employed in4 tasks to investigate the influence of cross-flow velocity, TMP,temperature and feed pH on flux decline, with testing rangesof 2.7–2.9 m s−1, 0.1–0.2 MPa, 15–55 ◦C and 5.5–9.5, respectively.Flow rate data of permeate and concentrate were collected at 5-minintervals. Preliminary runs found that steady-state operation wasestablished in less than 20 min, thus the experimental runs werelimited to maximum 80-min duration. Samples were collected foranalysis from both the feed tank and permeate stream at starts,and after 20-, 40-, 70-min operation. Seawater was drained off aftereach run and the membrane was rinsed three times using deionizedwater feed for 30 min at a TMP of 0.14 MPa and a cross-flow veloc-ity of 3.5 m s−1. Next, fresh deionized water was recirculated at0.01 MPa and 10.5 m s−1, in the system until the initial clean waterpermeability was completely restored.

2.4. Resistance analysis

Darcy’s law is commonly used to determine filtration resistancein permeate transport through porous membranes according to Eq.(1)

J = �P

�Rt(1)

where J is the permeate flux (L m−2 h−1), �P is the trans-membrane pressure (Pa); � is the dynamic viscosity of

cation Technology 75 (2010) 165–173 167

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J. Xu et al. / Separation and Purifi

he feed (Pa s); and Rt is the total filtration resistancem−1).

Resistance-in-series models [13,14] of membrane fouling oftenubdivide the total membrane resistance Rt into three components,amely intrinsic membrane resistance Rm external (concentra-ion polarization) resistance Rc; and internal resistance Rf, dueo fouling. In this study, however, Rt is divided into six compo-ents related to the membrane: intrinsic resistance Rm; surfacedhesion resistance Rsads; pore adhesion resistance Rpads; concen-ration polarization resistance Rcp; pore blockage resistance Rpb;nd cake layer resistance Rcr. Moreover, a new experimental proce-ure was established to measure the different filtration resistanceomponents, and thus to better understand membrane/foulantnteractions and factors affecting membrane performance.

The assumptions of resistance analysis are as follows:

Osmosis pressure by particles and organic matter in seawater isinsignificant.Rm, Rsads, Rpads and Rcp are unavoidable for ultrafiltration of sea-water.Rsads and Rpads are caused by weak physical intermolecular forces(adsorption) and thermodynamically unavoidable. These resis-tances can be targeted by hydraulic rinsing at a low pressure andhigh flow rate.Rsads is constant during ultrafiltration, and commences oncemembrane and seawater are in contact. This component is inde-pendent of the permeate flux and can be measured by means ofcirculation of the feed in the membrane module.Rpb and Rcr together provide the semi-reversible (Rsr) and irre-versible (Rirr) resistances. Unlike Rsr, Rirr cannot be eliminatedusing chemical cleaning.

The protocol to measure resistance analysis data is as follows:

The pure water flux of a new membrane ceramic UF membraneis measured to obtain Jwo.Adsorption experiments: the membrane is contacted with sea-water at given cross-flow velocity for 1 h in the absence ofpermeate flux (termed static mode), then the pure water flux ismeasured to obtain Jw1.Flux decline experiments: first seawater filtration is performedto obtain flux J, followed by the determination of the pure waterflux Jw2.Flushing experiments: strong shear force flushing is carried outthree times for 2 h using cross-flow velocity in the turbulent flowregime (10.0–10.5 m s−1) at an operating pressure of 0.01 MPa,followed by a pure water flux measurement to obtain Jw3.Chemical cleaning: 0.01% 40 ◦C NaClO solution is recirculated for30 min, then the pure water flux is determined to obtain Jw4.

.5. Analytical methods

Feed and permeate samples were collected regularly for wateruality analysis, to evaluate the suspended solids and organic mat-er removal efficiencies of the ceramic membrane. Turbidity andH were measured with a turbidity (LP2000-11, HANNA, Italy) andpH meter (HI8424, HANNA, Italy), respectively. CODMn was deter-ined according to the GB 17378.4-1998 (China) using KMnO4

xidant, which is the same as the method described in USEPA

1983) [15].

The permeate flux (J) was calculated as the permeate flow overnit surface area of the membrane. Normalized flux is calculated ashe ratio of permeate flux (J) to the initial flux at the beginning ofhe ultrafiltration (J0).

Fig. 2. Effect of cross-flow velocity on permeate flux (temperature at 28 ◦C, TMP at0.14 MPa, pH at 8.2).

The foulant removal efficiency (R), which is useful for determin-ing the UF performance, is defined as:

R (%) = Cf − Cp

Cf× 100 (2)

where Cp and Cf are the foulant concentrations in the permeate andfeed solution, respectively. This study tested each run over time forinitial stability measurements.

3. Results and discussion

3.1. Membrane performance with cross-flow velocity

The transient behavior of the permeate flux for different cross-flow velocities at a TMP of 0.14 MPa is shown in Fig. 2. It is importantto note that the ceramic membrane achieves much bigger permeateflux than polymeric membranes having similar pore size.

There is a relatively sharp initial decrease of permeate flux dueto concentration polarization, adsorption, pore blockage and for-mation of fouling layer within 20-min operation, after which thepermeate flux reaches a relative steady-state. Cross-flow velocityhas significant influence both on the initial and stabilized permeateflux. The lowest cross-flow velocity (2.7 m s−1) leads to the lowestinitial permeate flux (410 L m−2 h−1), indicating inadequate turbu-lance to combat fouling effectively. The steady-state permeate fluxsignificantly increases at cross-flow velocity from 2.7 to 3.7 m s−1

but declines at 4.9 m s−1. The maximum steady-state permeate fluxof 443 L m−2 h−1 is observed at a cross-flow velocity of 3.7 m s−1.It is assumed that the adsorption layer, concentration polariza-tion layer and membrane surface fouling layers are disrupted bystrongly turbulent flow and hydrodynamic shear achieved at highercross-flow velocities. However, beyond a certain cross-flow veloc-ity (3.7 m s−1 in this study), a small (4.4%) drop in permeate fluxis observed at 4.9 m s−1, contrary to results reported by previousresearch [14,16]. This finding might be explained considering by thesystem specific ‘threshold’ cross-flow velocity that causes changesbetween surface foulant/membrane interactions. At higher cross-flow velocities, larger particle types present in the feed might beincreasingly removed due to vigorously turbulent flow, and thusonly smaller particle types can form deposits on the membrane

surface. This effect leads to the formation of different, more com-pact cake layer structure having increased resistance, which resultsin the reduction of the permeate flux.

The trends of normalized flux decline are illustrated in Fig. 3. Themost significant flux declines are observed at cross-flow velocities

168 J. Xu et al. / Separation and Purification Technology 75 (2010) 165–173

Ft

3stdspf

f4e

citcafsahrvselhca

Fig. 5 shows the decline in permeate flux of the ceramic mem-

ig. 3. Fouling of ceramic membrane with various cross-flow velocities (tempera-ure at 28 ◦C, TMP at 0.14 MPa, pH at 8.2).

.2, 2.7 and 4.9 m s−1, indicating that the membrane module is moreusceptible to fouling at a too low or too high cross-flow velocity. Athe optimum cross-flow velocity of 3.7 m s−1 the normalized fluxecreases to 96.7% of the initial flux after 20-min operation and thentabilizes in the 96.5–96.8% range. This finding indicates that theroper hydraulic conditions can control the extent of membraneouling but cannot eliminate it entirely.

Turbidity and CODMn were used as indicators of filtration per-ormance. The initial concentrations of turbidity and CODMn were.00–4.53 NTU and 1.32–1.40 mg L−1, respectively. Fig. 4 shows theffect of cross-flow velocity on permeate quality.

The ceramic membrane exhibits good turbidity removal effi-iency (over 97%) in each run (Fig. 4(a)). The removal efficiencys highest (99.8%) at the optimum cross-flow velocity of 3.7 m s−1,hough the trend shows a slight decline (to 99.5%) at the maximumross-flow rate of 4.9 m s−1. From many experimental observations,cake or a gel layer is usually formed on the membrane sur-

ace that influences the ultimate permeate flux. In addition to theolute concentration in the feed solution, the feed velocity wouldlso affect the extent of gel layer formation. Several researchersave reported the effects of cross-flow velocity on membraneejection performance. It is clearly expected that as cross-flowelocity increases, the mass and thickness of each fouling layerhould decrease, resulting in decreased filtration resistance. How-ver, results from Choi et al. [17] indicated that irreversible fouling

ayer formed regardless of filtration velocity and suggested that theigh density of fouling layer at high cross-flow velocities must beaused by effective fouling layer compaction due to high perme-tion drag force. In this study, for the cross-flow velocities from

Fig. 4. Effect of cross-flow velocity on permeate quality: (a) turbidity

Fig. 5. Effect of TMP on permeate flux (temperature at 28 ◦C, cross-flow velocity at3.2 m s−1, pH at 8.2).

2.7 to 3.7 m s−1, the formation of looser cake layers (dynamic ‘pre-filters’) might be the main result at higher cross-flow velocities,which lead to the increased particle leakage to the membraneto cause permeate quality degradation; for the cross-flow veloc-ities from 3.7 to 4.9 m s−1, the compact cake layers formed fromsmaller deposited particles at higher cross-flow velocities havean increased particle removal (pre-filtration) efficiency. Accord-ing to the speculation mentioned above, the transition from loosercake layer to compact cake layer should happened to the periodof 2.7–3.7 m s−1 and resulted in the highest turbidity removal atcross-flow velocities of 3.7 m s−1.

As shown in Fig. 4(b), an increase in cross-flow velocity also leadsto higher CODMn removal, CODMn removal efficiency increases withcross-flow velocity, from 16.2% to 37.1% at 2.7 and 4.9 m s−1 cross-flow velocities, respectively. One thing should be noted is that therewas only 2% difference of CODMn removal between the velocitiesof 3.7 and 4.9 m s−1. However, the highest turbidity removal andflux were found under the condition of 3.7 m s−1. It is clear thatthe condition of 3.7 m s−1 seems to be the optimal velocity of thesystem.

3.2. Membrane performance with TMP

brane at different TMP values. The permeate flux monotonouslyincreases with TMP to represent a 16.9% improvement in steady-state permeate flux when the TMP is increased from 0.1 to0.18 MPa. However, the increase of permeate flux is minimal

(b) CODMn (temperature at 28 ◦C, TMP at 0.14 MPa, pH at 8.2).

J. Xu et al. / Separation and Purification Technology 75 (2010) 165–173 169

Ffl

(toecaTflaitds4rsdFt0m

cw0its

ig. 6. Fouling of ceramic membrane with various TMP (temperature at 28 ◦C, cross-ow velocity at 3.2 m s−1, pH at 8.3).

453–457 L m−2 h−1) at a corresponding TMP increase from 0.18o 0.20 MPa. This may be attributed to the increasing compactionf the surface deposit on the membrane surface, to counteract theffect of higher TMP values. This observation is also related to theoncept of “critical flux” which is generally used to understandnd improve the operation of membrane filtration systems [8,18].he critical flux operationally defined as the “threshold” permeateux below which fouling does not increase with permeate volume,nd is fully reversible. More specifically in cross-flow filtration,t refers to the flux corresponding to the condition under whichhe drag force imparted to particles or the momentum causedue to cross-flow is greater than that due to permeation. In thistudy, the determined critical flux of the ceramic UF membrane is53 L m−2 h−1, which is significantly higher than relevant resultseported in the literatures [19–21]. While the details will be pre-ented in another publication, this finding would suggest to use aesign flux of less than 453 L m−2 h−1 for this membrane and feed.ig. 6 compares the observed normalized fluxes during ultrafiltra-ion at TMP values between 0.10 and 0.20 MPa. It can be seen that.18 MPa represent an optimum value, above which a decrease inembrane performance is observed.The effect of TMP on permeate quality is presented in Fig. 7. It

an be seen that the turbidity of permeate monotonously decreasesith TMP, while the CODMn of permeate decreases from 1.2 to

.85 mg L−1 with the increase of TMP from 0.10 to 0.18 MPa butncreases to 0.11 mg L−1 at a higher TMP. These two different pat-erns probably result from the different transport mechanisms ofuspended solids and dissolved organic matter in filtration of sea-

Fig. 7. Effect of TMP on permeate quality: (a) turbidity, (b) CODMn (te

Fig. 8. Effect of seawater temperature on permeate flux (TMP at 0.14 MPa, pH at8.2).

water. Concerning the suspended solids whose size is larger thanthat of dissolved organic matter, the formation of cake layer onthe membrane surface might contribute to ‘pre-filter’ (second-barrier) in ultrafiltration. The density and thickness of the cake layerincreases with TMP, yielding a lowest turbidity of 0.01 NTU in per-meate at TMP 0.20 MPa. Concerning the dissolved organic matter,concentration polarization and gel layer formation increases withTMP that affects mass transfer in ultrafiltration to yielding a higherCODMn in permeate. However, the coexistence of suspended solidsand organic matter, as well as of cake and gel layers leads to theircomplex interaction, which will be presented in a later publication.

3.3. Membrane performance with temperature

In this study, the temperatures of raw seawater feed were con-trolled in the range from 15 to 55 ◦C using a water bath. Fig. 8shows the variation of the final steady-state permeate flux after80 min of filtration at different feed temperatures. The perme-ate flux represents increases of 57.1% and 47.6% with increasesof feed temperature from 20 to 47 ◦C at cross-flow velocities of3.2 and 3.5 m s−1, respectively, indicating a significant effect. Theobtained data shows good linear regressions fits with coefficientsof determination (R2) of 0.9898 and 0.9876 at cross-flow rates 3.2and 3.5 m s−1, respectively, to suggest that permeate flux increases

linearly with feed temperature. Higher feed temperature leads tolower viscosity of feed and also to higher solubility of some feedconstituents. The same also reduces concentration polarization andthe transport of solvent through the membrane intensifies, yielding

mperature at 28 ◦C, cross-flow velocity at 3.2 m s−1, pH at 8.2).

170 J. Xu et al. / Separation and Purification Technology 75 (2010) 165–173

ity, (b

attsccfl

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3

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p

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tions. The reason might rely on the interaction between the positivecharged membrane and the negative charged NOM at a lower pH(5.5), leading to a more compact concentration polarization or gellayer, which increases NOM concentration close to membrane sur-

Fig. 9. Effect of seawater temperature on permeate quality: (a) turbid

higher permeate flux. It should be noted that the elevation of feedemperature normally increases the energy (operational) cost andhe potential of scaling, and reduces the durability of membraneystem in spite of superior thermal stability of ceramic membraneompared to polymeric membrane. It would, therefore, be practi-al to operate membranes at ambient temperature unless a greaterux is required at the expense of operational costs.

The effect of feed temperature on permeate quality in termsf turbidity and CODMn is shown in Fig. 9(a) and (b). It can beeen that both turbidity and CODMn in permeate monotonouslyecrease with feed temperature, exhibiting from 0.11 to 0.03 NTUnd 1.16–0.87 mg L−1 with feed temperature from 14 to 49 ◦C,espectively. And, the corresponding removal efficiency in per-eate turbidity and CODMn increases from 97.25 to 99.25% and

2.12 to 34.09% in the range of feed temperature, respectively.ith the increase of feed temperature, lower feed viscosity and

igher diffusion of foulants (suspended solids and organic matter)ead to stronger convection force during filtration, which hindershe passage of foulants so as to decrease turbidity and CODMn inermeate. Moreover, compared to organic membrane, the ceramicembrane yields a higher thermal stability and thus the enlarge-ent of membrane pore size is less obvious with the increase in

eed temperature, which indicates that the effect of the changes ofore size on permeate quality is insignificant.

.4. Membrane performance with seawater pH

The effect of seawater pHs in the range between 5.5 and 9.3n final steady-state permeate flux was investigated and shown inig. 10.

The change of permeate flux shows an exponential increase withH, as also indicated by the good fit (R2 of 0.9985) of Eq. (3).

= 395.57 + 0.17 exp(

pH1.66

)(3)

This finding may be explained considering the interactionsetween membrane and foulants at different solution pHs. Zirco-ium oxide, which is an amphoteric material with an isoelectricoint (IEP) slightly less than 7 [22,23], is the active layer of theeramic membrane used. Therefore, the surface of the membranehows different ionization states or zeta potentials, depending onhe feed pH [24,25]. The surface of zirconium oxide is occupiedy amphoteric MOH groups which are a result of an exposure toater in solution (hydroxylation). The amphoteric MOH groups

end to dissociate and become charged in contact with polar feeds,nd electrostatic repulsion or attraction results the interactionsetween the particles present in the solution and the charged mem-rane surface. Natural organic matter (NOM) has an IEP at around[26], and carries a negative charge in the typically range of pH of

) CODMn (TMP at 0.14 MPa, cross-flow velocity at 3.2 m s−1, pH at 8.4).

seawater. At the acidic pH of 5.5–6.0, the ceramic membrane has aslight positive charge (ZrOH2+) to attract the negatively chargedNOM. This attraction between the membrane and NOM causesadsorption inside the pores and on the membrane surface, whichis responsible for the lower permeate flux of 400 L m−2 h−1 at apH 5.5. In contract, when the pH of the feed is increased over theIEP of the membrane, the membrane surface becomes negativelycharged. The same also increases the negative charge of NOM [27],leading to increased particle repulsion, and the increase of perme-ate flux. The expected increase of precipitation of divalent metalspecies, as well as their complexation with humic substances withthe increase of experimental pH range appears to be insignificanton fouling.

Fig. 11 shows the change of permeate quality with feed pH.Higher turbidity and CODMn removals are achieved with theincrease in feed pH. For example, turbidity and CODMn removalsare up to 99.8% and 39.1% at feed pH 9.2, respectively. And, theincreasing trend is more pronounced for CODMn. This, in general,can be explained considering the effect of feed pH on the chargeson the membrane surface (as in the previous section). Surprisingly,the CODMn in permeate was higher than that in feed at pH 5.5and thus the removal efficiency of CODMn is negative, which is notreported in the previous literature and also contradicts expecta-

Fig. 10. Permeate flux variation with respect to seawater pH (TMP at 0.14 MPa,cross-flow velocity at 3.2 m s−1, temperature at 28 ◦C).

J. Xu et al. / Separation and Purification Technology 75 (2010) 165–173 171

F0

fpgh

3

Fisbcrc

ig. 11. Effect of feed pH on permeate quality: (a) turbidity, (b) CODMn (TMP at.14 MPa, cross-flow velocity at 3.2 m s−1, temperature at 28 ◦C).

ace far beyond that in bulk feed. Therefore, although CODMn in theermeate is lower than that in the concentration polarization orel layer due to the exclusion of the ceramic membrane, it is stilligher than that in feed bulk, as reported in above phenomenon.

.5. Resistance analysis of membrane fouling

The relation of various resistance components are illustrated inig. 12 and Eqs. (4)–(9), noting that the resistance analysis protocols described in Section 2.4 previously. As shown in Fig. 12, (1) repre-ents the system stablizing process (also to remove entrapped air

ubbles initially); (2) reflects Rsads after adsorption experiments,orresponding to Jw1; (3) reflects Rcp, corresponding to Jw2; (4)eflects Rt, corresponding to J; (5) reflects the sum of Rsads and Rpad,orresponding to Jw3; (6) reflects Rsr which is eliminated only be

Fig. 12. Schematic diagram of filtration resistance analysis.

Fig. 13. Evolution of total filtration resistance with filtration time (TMP at 0.14 MPa,temperature at 28 ◦C, cross-flow velocity at 3.2 m s−1, pH at 8.2).

chemical cleaning, corresponding to Jw4; while (7) reflects Rirr.

Jw0 = TMP(�wRm)

(4)

Jw1 = TMP[�w(Rm + Rsads)]

(5)

Jw2 = TMP[�w(Rm + Rsads + Rpads + Rsr + Rirr)]

(6)

Jw3 = TMP[�w(Rm + Rsr + Rirr)]

(7)

Jw4 = TMP[�w(Rm + Rirr)]

(8)

J = TMP[�sw(Rm + Rsads + Rpads + Rsr + Rirr + Rcp)]

(9)

where �w and �sw are the viscosities of pure water and seawater,respectively. At 28 ◦C, �w equals to 0.8360 MPa s and �sw equals to0.8910 MPa s for a seawater salinity of 32‰.

In this study, the variation of the total resistance (Rt) with theoperational time under the conditions of TMP 0.14 MPa, temper-ature 28 ◦C, cross-flow velocity 3.2 m s−1, and pH 8.2 is analyzedwith results shown in Fig. 13. The initial Rt increases quicklyfrom 1.26 × 109 to 1.34 × 109 m−1 in the beginning of operation(0–15 min) and remains relatively stable (1.35–1.36 × 109 m−1)after 20 min. The exponentially declining trend is confirmed withthe very good fit of Eq. (10):

Rt = 1.35 − 0.11 exp(−0.16t), R2 = 0.9966 (10)

It should be noted that the total resistance of (1.34 × 109 m−1),and thus the fouling potential of the ceramic membrane is signifi-cantly lower than that of polymeric membrane (1011 m−1) [28].

The various filtration resistance components are summarized inFig. 14. The contributions to the total resistance are presented in thefollowing order: Rcp > Rsr > Rm > Rsads > Rpads > Rirr. (Note: Rcp con-centration polarization resistance; Rsr semi-reversible resistance;Rm intrinsic resistance; Rsads surface adhesion resistance; Rpads poreadhesion resistance and Rirr irreversible resistance.)

Rm is related to only membrane properties and is consideredto be a constant value of 0.24 × 109 m−1. Hence, the Rm/Rt ratio

decreases slightly from 18.90% to 17.78% with the increase ofother resistance components with time. The initial Rcp value of0.44 × 109 m−1 contributes most to the total resistance with theratio of 35%, and increases significantly from 33.86% to 37.04%with filtration time. This finding reveals that Rcp is dominant in

172 J. Xu et al. / Separation and Purification

tiart(st0mcwsptatRttansi

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[

[

Fig. 14. Evolution of fractional resistances with filtration time.

hose conditions, and concentration polarization is the main foul-ng mechanism to explain membrane performance in terms of fluxnd permeate quality. Fig. 14 also shows that Rsads (0.22 × 109 m−1)emains nearly constant, with a slightly declining ratio from 17.45%o 16.35%. In contrast, Rpads increases from 0.12 to 0.15 × 109 m−1

ratio from 9.4% to 11.72%), revealing that adsorption on membraneurface with filtration volume is more pronounced than inside ofhe membrane. The increase of Rsr (maybe Rpb) by 6.38% in the–10-min operational period indicates pore blockage and the accu-ulation foulants on membrane surface. It is well understood that

oncentration polarization and adsorption take place first once sea-ater and membrane are in contact, and other forms of fouling,

uch as pore blockage and cake layer emerge with the increase ofermeate volume. However, the fact is not as expected that the con-ributions of pore blockage and cake layer are relatively high (6.38%)t the beginning (0–10-min operation) and present a decliningrend after 10-min filtration. The irreversible fouling componentirr is nearly zero during the first 10 min of operation, increaseso 0.015 × 109 m−1 after 15-min filtration and then remains rela-

9 −1

ively constant. In contrast, Rsr reduces from 0.28 to 0.24 × 10 mfter an initial increase, and thus the Rsr/Rt ratio also shows a sig-ificant decrease from 22.21% to 15.41%. This finding suggests thatome part of the initially reversible fouling Rsr is transformed intorreversible form (Rirr) with the increase of permeate volume/time.

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Technology 75 (2010) 165–173

4. Conclusion

The performance of a ceramic UF membrane with pore size of0.05 �m was investigated in treating raw seawater under differentoperating conditions, leading to the outcomes listed below:

(1) The ceramic membrane assessed in this study compares favor-able with organic counterpart in terms of permeate flux,permeate quality and fouling potential, while also offers supe-rior durability and chemical resistance in the filtration ofseawater.

(2) In this study, the superior performances with 99.0–99.5%turbidity and 32–35% CODMn removals were achieved atcross-flow velocities of 3.7–4.2 m s−1, TMPs of 0.14–0.18 MPa,temperatures of 25–30 ◦C and seawater pHs of 8.0–9.0.Under the conditions, a permeate flux ranged from 420 to450 L m−2 h−1 was obtained.

(3) The analysis of permeate quality under different operatingconditions showed different removal efficiencies of suspendedsolids and organic matter, which indicates complex interactionsbetween the membrane and foulants.

(4) A new experimental procedure was conducted to realize theperformance further of a commercial monolith ceramic UFmembrane (XS-T-01) and its fouling mechanisms. The presentresults show that the effect of concentration polarization wasdominant in the ultrafiltration of raw seawater, and vari-ous fouling components contribute to the total in the orderof Rcp > Rsr > Rm > Rsads > Rpads > Rirr. Further, reversible foulingwas much more significant than irreversible fouling. Theexperimental method also showed interrelation between thereversible and irreversible resistances, whereby part of Rsr istransformed into Rirr within 20 min of filtration.

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