Fouling and wetting in membrane distillation (MD) and MD-bioreactor (MDBR) for wastewater reclamation

Desalination 323 (2013) 39–47

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Fouling and wetting in membrane distillation (MD) and MD-bioreactor (MDBR) for wastewater reclamation Shuwen Goh, Jinsong Zhang, Yu Liu, Anthony G. Fane ⁎ Singapore Membrane Technology Centre, Nanyang Technological University, Singapore 639798, Singapore School of Civil and Environmental Engineering, Nanyang Technological University, Singapore 639798, Singapore

H I G H L I G H T S ► ► ► ►

Delayed membrane wetting in membrane distillation bioreactor (MDBR) Successful biological removal of organic and nutrients in MDBR Faster flux decline in MDBR Flux decline and wetting are dependent on fouling layer characteristics

a r t i c l e

i n f o

Article history: Received 30 April 2012 Received in revised form 31 October 2012 Accepted 3 December 2012 Available online 28 December 2012 Keywords: Membrane distillation Membrane distillation bioreactor Synthetic wastewater reclamation Wetting Fouling

a b s t r a c t The membrane distillation (MD) process is seldom employed in wastewater reclamation since the high organic and nutrient in wastewater promote wetting. The MD bioreactor (MDBR) can remediate this by biologically removing retentate carbohydrates and proteins. However, the inclusion of biomass in the MDBR can result in biofouling and flux decline. The objectives of this work are to determine the effectiveness of the bioprocess in delaying membrane wetting (by removing organics and nutrients) and the significance of the biofouling on flux decline. From this work, the MDBR flux can be maintained at more than 6.8 L/m 2 h (8% lower than the average MD flux) for at least 13 days. The faster flux decline in the MDBR is attributed to the thermal and mass transfer resistance of the biofilm but this can be controlled with periodic membrane cleaning and process optimization. Membrane fouling has been shown to compromise membrane hydrophobicity and accelerate wetting. By lowering the retentate organic and nutrient concentration, the MDBR has successfully delayed wetting by 1.7–3.6 times in this work, reducing the frequency of membrane cleaning and drying. With further process optimization, the MDBR could be a good option for reclamation of industrial wastewater with low volatile organic content and access to waste heat. © 2012 Elsevier B.V. All rights reserved.

1. Introduction Like the reverse osmosis (RO) process, membrane distillation (MD) is a high retention process that can retain recalcitrant, low-molecular weight organics (including chemicals of concern), non-volatiles solutes and viable-but-not-culturable (VBNC) bacteria [1–3]. The driving force in MD is the vapor pressure gradient. Operating on the principle of separation via phase equilibria, the hydrophobic membranes used in the MD process create the vapor–liquid interface at the membrane pore surface, which is essential for vaporization of volatiles, such as water. Heat and mass transfer mechanisms govern the permeation flux in the MD and have been well researched with numerous models developed for unfouled and fouled membranes [4–8]. The mass and heat

⁎ Corresponding author. Tel.: +65 67905272; fax: +65 67910756. E-mail address: [email protected] (A.G. Fane). 0011-9164/$ – see front matter © 2012 Elsevier B.V. All rights reserved. http://dx.doi.org/10.1016/j.desal.2012.12.001

balance across the fouled MD membrane can be summarized by the following two equations: Mass Balance : J ¼ ðP f −Pb Þ=Rf ¼ ðP b −P 1 Þ=Rb ¼ ðP 1 −P 2 Þ=Rm ¼ P 2 −P p =Rp

Heat Balance : hf ðTf −T b Þ ¼  kb =δb ðT b −T 1 Þ ¼ km =δðT 1 −T 2 Þ þ JΔHv ¼ hp T 2 −T p

ð1Þ

ð2Þ

Where (Pf − Pb) / Rf denotes mass transfer through the feed boundary layer, (Pb − P1) / Rb denotes mass transfer through the fouling layer, (P1 − P2) / Rm denotes mass transfer through the membrane, and (P2 − Pp) / Rp denotes mass transfer through the permeate (see Nomenclature). Eq. (1) assumes mass transfer can be described by an apparent vapor pressure difference and a mass transfer resistance. While this is physically realistic for transport through the membrane, it is also a convenient description for transport through liquid boundary layers. This

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S. Goh et al. / Desalination 323 (2013) 39–47

approach [4] allows comparison of the relative resistances in the process. In Eq. (2) hf(Tf − Tb) denotes heat transfer through the feed boundary layer, kb/δb (Tb − T1) denotes conductive heat transfer through the fouling layer, km/δ(T1 − T2) denotes conductive heat transfer through the membrane, JΔHv denotes heat transfer with the water vapor across the membrane (this is a ‘heat loss’), and hp(T2 − Tp) denotes heat transfer through the permeate boundary layer. The temperature profile across a fouled MD membrane can be presented in Fig. 1. The vapor pressure at the membrane surface is related exponentially to temperature (Antoine equation for dilute solutions) [9]. A fouling layer will confer an additional resistance which, like the membrane, is dependent on fouling layer characteristics such as porosity and thickness [10,11]. A non-porous protein fouling layer may result in thermal and hydraulic resistance while a porous calcium carbonate fouling layer is likely to only contribute to thermal resistance [8]. In addition to fouling, the MD faces another challenge: wetting. The Laplace equation shows the relationship between the membrane pore size and the breakthrough pressure [2]: ΔP interface ¼ P liquid –P vapor ¼ −2BγL cosθ=r

ð3Þ

(Where γL = liquid surface tension, θ = liquid–solid contact angle, r = pore radius, B = geometric factor determined by pore structure). ΔPinterface denotes the pressure difference across the membrane pores. Membrane wetting will occur if ΔPinterface exceeds the breakthrough pressure of the membrane. From Eq. (3), an increase in membrane pore size, a decrease in liquid surface tension or contact angle would accelerate membrane wetting. MD is seldom applied in wastewater reclamation since wastewater usually contains organic, inorganic and amphiphilic components that can accelerate membrane wetting [8,12]. As MD is a highly retentive process [13], the accumulation of non-volatile solutes such as salts, carbohydrates and proteins in the MD retentate could exacerbate membrane wetting. Furthermore MD alone merely concentrates the waste and does not treat it. The membrane distillation bioreactor (MDBR) [1,14,15] has the potential to expand MD application to the reclamation of niche industrial wastewaters that are hot (or have access to waste heat) and have a low volatile solute content. By coupling the thermophilic bioprocess with the MD process, the inclusion of biomass can result in the biological removal of organics and nutrients. There have been studies on using MD/hybrid MD systems in wastewater containing recalcitrant organics (e.g., textiles) [16,17]. The MDBR may also be suitable for treating such wastewater since the decoupling of organic retention time (ORT) from hydraulic retention time (HRT) in the MDBR implies that there will be more time for the recalcitrant organics to be biologically degraded. The feasibility of the MDBR in synthetic wastewater and petrochemical wastewater reclamation has already been confirmed [1,14,15] but existing MDBR studies have not compared the effect of biomass inclusion on flux decline (due to fouling) and membrane wetting. While the biomass can result in biodegradation of wetting agents such as amphiphilic proteins, it can also lead to biofouling which will reduce flux. As such, this study aims to determine

Tf Tb T1

T2 Tp

Fig. 1. Schematic of temperature profile across a fouled MD membrane.

the effects of biomass inclusion on retentate organic carbon and nitrogen concentration, flux and membrane wetting by comparing the performance of the MDBR with MD in synthetic wastewater reclamation. 2. Materials and methods 2.1. Submerged MD and MDBR The MD and MDBR experiments were done using the same reactor and membrane set up and the same synthetic feed. The operating conditions were identical except that the MDBR included biomass. The submerged MD and MDBR systems (Fig. 2) consisted of a reactor (4.5 L aerated tank) with a submerged flat sheet membrane module. Polyvinylidene fluoride (PVDF) hydrophobic flatsheet membranes (MILLIPORE® Durapore GVHP) with nominal pore size of 0.22 μm and total effective area of 192 cm 2 were used for all experiments. Aeration rate and dissolved oxygen (DO) were maintained at 4 L/min (from each aeration tube located right below the membrane surface) and 2–4 mg DO/L, respectively. Permeate recirculation rate was fixed at 350 ml/min. The superficial velocity of the feed and permeate (estimated from the air flow rate and permeate recirculation rate) were 5.1 cm/s and 1.2 cm/s, respectively. The retentate pH in both systems was similar. Forty milliliters of samples were extracted from the MD and MDBR daily; this corresponds to approximately 0.9% of the reactor volume per day and translates to a sludge retention time (SRT) of approximately 110 days for the MDBR. The hydraulic retention time (HRT) varies with flux and ranged from 28 to 52 hours. The hydraulic diameter (dh) at the feed and permeate sides are 1.9 cm and 1.8 cm, respectively. Synthetic wastewater (0.67 g COD/L and 0.04 g TN/L) was prepared from the concentrated feed solution (4.27 g/L glucose, 0.85 g/L meat extract, 1.07 g/L peptone, 0.19 g/L KH2PO4, 0.19 g/L MgSO4, 0.16 g/L FeCl3, 3.2 g/L CH3COONa). The level sensors installed in the reactor allowed automatic pumping of feed into the MD and MDBR reactor. The MD and MDBR were maintained at reactor and permeate temperatures of 55.5±1 °C and 19.5±1 °C respectively and the experiments were terminated shortly after membrane wetting had occurred in each system. Feed organic carbon and nitrogen loading were maintained at 0.35±0.03 kg TOC/m3 day and 0.036±0.007 kg TN/m3 day, respectively, for both MD and MDBR. The dominant foulant in the MD treatment of tapwater solutions has been identified to be calcium carbonate in various studies [8,14]. To avoid this and to determine the effect of biofouling per se (instead of carbonate fouling) on flux decline, the feed solution in this study was prepared with Milli-Q to minimize carbonate fouling. As mentioned earlier, the setup and operating conditions for the MD and MDBR were identical except that 10% (v/v) of thermophilic seed sludge was added to the MDBR reactor at the start of its operation. Both experiments are recognized as unsteady state (the MDBR could need possibly 300 days (3 SRTs) for stabilization). However the main objective was to compare MD with and without biomass in terms of operation. 2.2. Analytical methods Samples were centrifuged at 13200 rpm using Hitachi CT15RE for 20 minutes and the supernatant were removed for Total Organic Carbon (TOC) and Total Nitrogen (TN) analysis by Shimadzu TOC Analyzer (Model: TOC-V CSH). Conductivity, DO and pH within the reactor were respectively measured with Mettler Toledo InPro 4010/120/PT1000 Sensor, InPro 6050/120 DO sensor and InPro 7108 Cond Sensor. Permeate flux was computed from permeate mass data (measured with Mettler Toledo PB8001-S/FACT Precision Balance). Biomass concentration was determined via the Total Suspended Solids (TSS) test [18]. Particle size was determined with a Malvern 2000S Mastersizer. The contact angle was measured with the OCA SCA (DataPhysics Instruments, Fiderstadt, Germany) goniometer using 9 μl sessile drops at 5 locations and the Laplace–Young fitting method. For confocal images, the fouled

S. Goh et al. / Desalination 323 (2013) 39–47

41

Flowmeter

T

FM T

Fridge

pH

DO

C

T

C

Heater

Feed Pump

Pump

Cooling Tank

Feed Tank

Mass Balance C pH

Air pump

Permeate Tank

Airflow meter

DO T

Legend for probes Conductivity pH Dissolved Oxygen Temperature Level Sensor

Fig. 2. Schematic diagram of submerged MD and MDBR system.

membranes were stained with molecular probes (Invitrogen, Baclite L7012) consisting of fluorescent dyes SYTO 9 (to stain live cells green) and propidium iodide (PI) (to stain dead cells red) and observed with the Carl Zeiss LSM 516 at 488/561 nm. For scanning electron microscope (SEM) images, the fouled MDBR membrane was pre-treated with glutaraldehyde/cacodylate buffer and dehydrated in increasing concentration of ethanol and dried. The fouled MD membrane was dried without further treatment. The dried samples were sputter-coated with EMITECH gold sputter coater SC762 and scanned with Carl Zeiss EVO 50 SEM. For FTIR, the virgin and fouled membrane surfaces were dried overnight in a 50 °C vacuum oven and scans were performed at a resolution of 4 cm−1 by Fourier Transform Infra Red spectrometer (FTIR, Shimadzu IR Prestige-21) using the Attenuated Total Reflection (ATR) equipped with ZnSe crystal method at ambient temperature. Nitrate, nitrite and ammonium ions analysis were conducted using Hach test kits (nitrate TNT836, nitrite TNT839 and Nitrogen-Ammonia Reagent Set #2606945). The extraction of loose and bound extracellular polymeric substances (EPS) was conducted as per reported by Zhang et al. [19] and the analysis of carbohydrates and proteins were conducted as per the phenol-sulfuric method [20] and Bradford test [21]. The bound EPS refers to the EPS on the biomass surface extracted via the centrifugation-NaOH method [22] while the loose EPS refers to the EPS in the supernatant.

the biological removal efficiency of the MDBR, a mass balance can be conducted on the substrate (which can be TOC or TN): V ðdC=dt Þ ¼ Q f C f –Q p C p –Q w C−r su V

ð6Þ

Where V(dC/dt) denotes substrate accumulation rate within the MDBR retentate, QfCf denotes mass flow rate of substrate into the retentate, QpCp denotes mass flow rate of substrate out of the retentate via the permeate stream, QwC denotes mass flow rate of substrate out of the retentate via the sludge wasting stream and rsuV denotes rate of biological substrate degradation within the MDBR retentate. The equation assumes negligible substrate loss via evaporation from the reactor. Rearranging Eq. (6) and dividing it by reactor volume (V) throughout yields the following equation:   Biological substrate utilization rate; r su¼ ðQ f C f Þ=V− Q p C p =V −ðQ w C Þ=V–dC=dt Biological substrate removal efficiencyð% Þ biological substrate utilization rateðrsuÞ  100 ¼ substrate influent mass flow rateððQ f C f Þ=V Þ

2.3. Determination of overall and biological TOC and TN removal efficiency Fig. 3 shows the schematic of a submerged MDBR where Q denotes flow rate (L/day), V denotes volume (L), C denotes TOC concentration (g TOC/L), N denotes TN concentration (g TN/L) and X denotes biomass concentration (g TSS/L). The subscripts f, p and w denote the feed stream, permeate stream and sludge wasting streams respectively. The overall MDBR removal efficiency of the system and can be presented as follows: Overall TN removal efficiency : Noverall ¼ 1−Np =Nf

ð4Þ

Overall TOC removal efficiency : C overall ¼ 1−C p =C f

ð5Þ

Within the MDBR retentate, the thermophilic biological process removes TOC and TN from the retentate to reduce excessive accumulation of retentate TN and TOC within the MDBR retentate. To determine

ð7Þ

Fig. 3. Schematic diagram depicting mass balance in a submerged MDBR.

ð8Þ

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S. Goh et al. / Desalination 323 (2013) 39–47

(a)

2.4. Determination of heat and mass transfer resistance From Eq. (2), the following equations can be derived:

T2 ¼

ð9Þ

hm T 1 þ hp T p þ JΔHv hp þ hm

h T þ hb T 1 Tb ¼ f f hb þ hf

ð10Þ

ð11Þ

(Where hf = heat transfer coefficient of the feed boundary layer, hb = heat transfer through the fouling layer, hm = heat transfer coefficient through the membrane, hp = heat transfer through the permeate boundary layer). Solving Eqs. (9) to (11) will result in the following equation [4,8]:

T1 ¼

Retentate TOC (mg/L)

hb T b þ hm T 2 −JΔHv H b þ hm

2000 1500 1000 500 0

1

2

ð12Þ The heat transfer coefficients for the boundary layers in this study were calculated via the Graetz–Leveque empirical model for laminar flow (Re b 2100) in a flat sheet module: Nu = 1.86(Re Pr(dh/L)) 0.33 [4]. Salt accumulated in the MD and MDBR system to around 8.3 g NaCl/L but the effect on vapor pressure depression was calculated to be insignificant (less than 0.3% without taking into account effect of concentration polarization for such ‘dilute’ solutions) [6]. The mass transfer resistances of the membrane, boundary and fouling layers were calculated using Eq. (1) and experimentally determined flux data. 3. Results and discussion

4 Time (days)

5

6

7

5

6

7

MD MDBR Theoretical

160

1=hb þ 1=hf þ 1=hm þ 1=hp

3

(b)

  1=hp þ 1=hm T f þ ð1=hf þ 1=hb ÞT p −JΔhvðð1=hb Þð1=hm Þ þ ð1=hf Þð1=hm ÞÞ

Retentate TN (mg/L)

T1

MD MDBR Theoretical

2500

140 120 100 80 60 40 20 0

1

2

3

4 Time (days)

Fig. 4. Variation of retentate (a) TOC and (b) TN with time, where the black triangles denote MD and the gray squares denote MDBR.

3.1. Performance of thermophilic bioproces in the MDBR system By combining the bioprocess with the MD process, the MDBR system can simultaneously produce water and treat the wastewater by biologically removing organic and nutrients from the retentate. Fig. 4 shows that the retentate TOC and TN concentration in the MDBR were significantly lower than that in the MD, implying that the biomass had successfully resulted in the removal of organics and nitrogenous nutrients from the retentate. The retentate results were compiled from duplicate runs of the MD and MDBR process under similar operating conditions. The theoretical TOC and TN concentrations for the reactors, assuming no bioconversion, were calculated based on the TOC and TN loading rate and solutes wasting rate. The TOC and TN concentration in the MD system were similar to the theoretical values for the first few days of operation; the slight deviation from the theoretical values may be due to the sampling timing (for example, if the samples were extracted just after feed has been pumped into the reactor, the MD value will be similar to the theoretical value). A dip in TOC and TN in the MD was observed from the third day onwards. Since it was not possible to operate the MD under sterile conditions, some microbial contamination may have occurred. The MD results were verified with batch tests (results not shown here) operated by fed-batch process where feed were added daily to nutrient feed bottles heated at 55 °C. The fed-batch results show that before feeding, the retentate TOC and TN values were lower than the theoretical values but right after feeding, the TOC and TN coincided with the theoretical values. Similarly, the TOC and TN results (after feeding) deviated from the theoretical values after 5 days of operation, indicating possible microbial contamination.

Using Eq. (8), the thermophilic biological TOC and TN removal efficiency within the MDBR retentate was calculated to be 88±3% and 67±8% respectively. Organic carbon and nitrogen were biologically removed from the MDBR supernatant for biomass growth (as shown in Fig. 5 by the increase in biomass concentration with time). In the conventional mesophilic bioreactor, nitrogen can also be biologically transformed via the ammonification-nitrification-denitrification (NH4+ →NO3− →N2) pathway. The exact mechanism for biological nitrogen transformation in the thermophilic system is not known but it is generally agreed that organic nitrogen is degraded to ammonium via ammonification [15,23]. The occurrence of thermophilic nitrifiers is rare but not unheard of [24,25]. Analysis of the MDBR retentate for ammonium, nitrite and nitrate via Hach tests showed the presence of these three components, indicating that ammonification and nitrification may have occurred in the system to some extent. Further study is required to verify this. In addition to ammonification, other nitrogen transformation (for example, conversion of organic nitrogen to nitrogen gas, ammonia and nitrous oxide (N2O)) [23] may have occurred but no tests have been conducted to verify this. From the preliminary results, the two possible means of nitrogen removal from the MDBR supernatant are ammonification (conversion of organic nitrogen to ammonium by biomass and removal from solution via membrane sparging) and uptake by biomass for growth. The free ammonia (FA) concentration in the retentate is estimated from the ammonium analysis using the equation: FA (mg l − 1) = ([NH4-N]× 10pH) /{[exp(6334/(273+ T)]+ 10pH} [26]. [NH4-N] refers to the ammonium-nitrogen concentration and T refers to the temperature in C. The results show that the FA concentration in the retentate did not exceed 6 mg/L NH3-N throughout the experiments, hence,

20 18 16 14 12 10 8 6 4 2 0

43

300 250 200 150 100 50

Loose EPS (mg/L)

BIomass concentration (gTSS/L), pH, Bound EPS concentration (mg/gTSS)

S. Goh et al. / Desalination 323 (2013) 39–47

0 1

3

5

7

9

11

13

15

17

19

21

23

Time (days) Biomass

Retentate pH

Total Bound EPS

Total Loose EPS

Fig. 5. Variation of biomass concentration, pH, bound and loose EPS concentration with time. Fig. 6. Change in flux in MD and MDBR with time.

nitrogen removal via ammonia gas release during membrane sparging is assumed to be negligible. Some of the organic and nutrient utilized by the biomass are metabolized, forming by-products such as extracellular polymeric substances (EPS) which either form a protective layer around cells (bound EPS) or are released into the MDBR retentate (loose EPS) [27–29]. The retentive nature of the MD system allows the retention and accumulation of residual feed and EPS, resulting in the increasing concentration of loose EPS observed in the MDBR (Fig. 5). In the MD and MDBR systems, retentate pH was observed to increase to 10 in as short as 5 days. This pH increase phenomenon is not uncommon and has been observed in several thermophilic studies [30–32]. The pH increase could be due to carbon dioxide stripping at elevated temperature (which resulted in the removal of carbonic acid from solution) [33]. 3.2. Effect of biomass inclusion on fouling and flux decline Fig. 6 shows the normalized and actual fluxes as observed in the MD and MDBR systems. The initial fluxes observed in the MD and MDBR were 8.08 + 0.13 L/m 2 h and 8.37 + 0.09 L/m 2 h respectively. The initial flux (J0) is defined as the average flux in the MD and MDBR systems on day 1. Only 7 days of flux results are shown for the MD experiment because the membrane was wetted after 7 days of operation. While the initial flux observed in the MDBR system was slightly higher (around 4.6% greater) than that in the MD system, the MDBR flux decreased to 7.88 + 0.04 L/m2 h after 3 days of operation (similar to the MD flux of 7.68+ 0.13 L/m 2 h). From day 3 onwards, the MDBR flux remained lower than the MD flux and reached a low of 4.05+ 0.59 L/m 2 h after 23 days of operation. The MDBR flux results are comparable to the flux observed in the study of petrochemical wastewater treatment using the MDBR[14]. A comparison of the normalized flux yielded similar observations: MDBR and MD flux decreased at similar rate from days 1 to 3 before the MDBR flux decreased more rapidly to 86% of initial flux on the 4th day and continued to decline for the next 20 days (Fig. 6). The average flux in the MD and MDBR systems were relatively similar with less than 10% difference in the first 6 days of operation. Despite the inclusion of biomass, the average MDBR flux at the 5th to 13th days was maintained at 6.79 ± 0.34 L/m 2 h, which is around 81% of its initial flux and only 8% lower than the average flux of 7.4 L/m 2 h observed in the MD system on days 5 and 6. However the MDBR flux continued to decrease with time to around 4.63 LMH at the 22nd day (55% of initial flux). As mentioned in Section 1, the fouling layer confers additional mass and heat transfer resistance. From Fig. 6, the flux in the MDBR system decreased at a much faster rate than the MD system (despite

the lower retentate TOC, TN and conductivity observed in the MDBR) after 3 days of operation. The faster flux decline in the MDBR may be due to biofouling and this assumption is supported by the SEM image of the PVDF membrane after 3 days of operation in the submerged MDBR (Fig. 7). Non-uniform biofouling had occurred, with an initial biofilm development (layers of overlapping single cells coated with EPS) shown on the left side of the image and single biofloc attachment observed on the centre right side. Single cells attachment were observed on the bottom left corner of the image and 5000× magnification of the same spot show cells coated in EPS (section e). A cross-sectional view of the fouling layer shows a thickness of 1.2 to 5.6 μm in just 3 days (section d). While some degree of biofouling had occurred on some parts of the membrane, certain portions of the membrane remained unfouled. A fouling layer thickness of 2 to 8 μm was observed from confocal images of the fouled MDBR membrane (Fig. 8) after 7 days of operation; this is similar to the fouling layer thickness observed from the cross-sectional view of the same fouled membrane under SEM (results not shown here). On the other hand, the fouling layer thickness of the fouled MD membrane (after 7 days of operation) was similar at 3.2–5.5 μm. With a longer operational period of 22 days, the fouling layer thickness of the fouled MDBR membrane was observed to be as high as 20 μm. The total overall transfer resistance (R = J/(Pf − Pp)) in the MDBR system (computed based on the vapor pressure, derived from Tf, Tb, T1, T2 and Tp using Antoine equation) increased from 1637 Pa m2 h/kg on day 1 to 2988 Pa m 2 h/kg on day 22 (an 82% increase). The term “overall” is used because mass flux is affected by a number of factors, including thermal and hydraulic resistances and other factors (such as vapor pressure depression due to reduced membrane hydrophilicity after fouling). The fouling layer may have created an additional resistance for heat and water transfer to the membrane pore but the question is how significant is this barrier? The fouling layer likely impeded convective heat transfer so heat transfer to the membrane surface would have occurred via diffusion and conduction [34,35]. The thermal conductivity of biofilms have been estimated to be 0.57–0.71 W/m K [36], which is almost 75% lower than the thermal conductivity of inorganic deposits such as CaCO3 and CaSO4 (2.6 and 2.3 W/m K respectively) [37]. Despite all these agruments, a 20 μm fouling layer on the 22nd day of MDBR operation with an assumed effective thermal conductivity (the overall thermal conductivity of the biofilm, including its water-filled pores) of 0.57 W/m K will result in a heat transfer coefficient of 28500 W/m2 k, which is of 2 magnitudes greater than heat transfer coefficients of the membrane, feed and permeate boundray layers. From the experimental

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Fig. 7. SEM images of (a) clean membrane, (b) fouled MD membrane (after 15 hours of operation), (c) fouled MDBR membrane (after 72 hours of operation), (d) cross-sectional view of image c — fouling thickness around 1.2 to 5.6 μm, (e) 5000× magnification of image c.

data and equations from Section 2.4, the fouling layer only reduced the temperature by 0.24 °C (Tf − Tb). This corresponds to a vapor pressure resistance of 28 Pa m 2h/kg, indicating that heat transfer resistance in the fouling layer is rather insignificant. On the other hand, the heat transfer coefficient of the membrane layer (calculated as Hm + JΔHv/(T1 − T2)) decreased by 42% at the end of 22 days of operation. This is similar to the flux decline pattern observed in Fig. 6 where the normalized flux at the 22nd day of operation was 55% of initial flux. T1 and T2 are related to factors such as to Hm, Tp, which are dependent on temperature, solution and membrane properties. Since feed and permeate are maintained at fixed temperatures

throughout the experiments, and heat conductivity of solution and membranes are assumed to remain relatively constant, Eqs. (10) and (12) can be simplified to show that T1 is mainly dependent on flux, and T2 is predominatly affected by the flux and T1. When flux decreases, the term JΔHv will decrease and T1 will increase. In physical terms, when flux decreases, less water is brought to the membrane surface to be vaporized. Less heat is removed from surrounding (e.g,. membrane surface) to vaporize the water, resulting in a higher T1. When flux decreases and T1 increases, the decrease in JΔHv due to flux decline is of one magnitude greater than the increase in hmT1, resulting in an overall decrease in T2 when flux J decrease. In physical terms, when

Fig. 8. Confocal images of different section of the same fouled MDBR membrane after 7 days of operation.

S. Goh et al. / Desalination 323 (2013) 39–47

The use of hydrophobic membranes in the MD and MDBR implies that prior to membrane wetting, only volatiles can pass through the MD membrane. Therefore, large macromolecular organic carbon (represented by TOC) would be retained in the retentate. Increase in permeate TOC would indicate that membrane wetting has occurred. As shown in Fig. 9, permeate TOC started to increase significantly from the 7th day in the MD and the 22nd day in the MDBR, implying that membrane wetting had been delayed in the MDBR. The overall TOC and TN removal efficiency of the MDBR and MD system were computed using Eqs. (4) and (5). Prior to membrane wetting, the overall organic removal efficiencies for both the MD and MDBR were similar at around 99.9% respectively in the first 7 days. Wetting in the MD system increased permeate TOC from 0.4 mg/L to 6.1 mg/L (a 15 fold increase in TOC transmission), reducing organic removal efficiency from 99.9% to 99.1%. The overall organic removal efficiency in the MDBR remained at 99.9% with permeate TOC lower than 0.8 mg/L for 22 days of operation (before membrane wetting). Delaying membrane wetting enables the MDBR system to maintain a high overall organic removal efficiency of 99.9% for a longer operation period (compared to MD). Repeating the experiment for both MD and MDBR under similar conditions yielded similar results where

100.0

14.0 12.0

99.0

10.0

98.0

8.0

97.0

6.0

96.0

4.0

95.0

2.0 0.0

94.0 0

5

10 15 Time (days)

MD Permeate TOC MDBR Removal Efficiency

20

25

Overall TOC removal efficiency(%)

3.3. Effect of biomass inclusion on membrane wetting

the time taken to wet the membranes in the MD and MDBR systems were 9 + 2 and 22 + 3 days, respectively. The results show that the inclusion of biomass in the MD system was able to delay wetting by 8 to 18 days in this work. What has resulted in this difference in wetting rate? Comparing the MD and MDBR retentate, the TOC and TN concentration in the MD reactor was around 3–5 times greater than that in the MDBR (Fig. 4). A higher MD retentate TOC and TN implies a higher concentration of foulants on the membrane surface and faster adsorption to the membrane surface, as shown by the thin coating on the MD membrane after 15 hours of operation (Fig. 7). The coating of the membrane pores indicates that the organic foulant in the feed solution had adhered to the membrane surface and changed the membrane surface property (e.g., the hydrophobicity). An examination of the contact angles of dried, fouled MD and MDBR membrane shows that the contact angles were as low as 32.9° and 59.1° respectively. Some parts of the MD and MDBR membranes remained relatively unfouled with contact angle as high as 68.6° and 65.9°, indicating non-uniform fouling on the membranes. The contact angle of the virgin PVDF hydrophobic membrane was 126.33 + 6.15°. Membrane fouling significantly reduced the hydrophobicity of both the MD and MDBR PVDF membranes. The contact angle results show that there are areas on the MD fouled membrane where the hydrophobicity of the membrane is more severely compromised and this may account for the faster wetting observed in the MD membrane. Why is the fouled MD membrane more hydrophilic than the fouled MDBR membrane? The answer may lie in the fouling layer characteristics and the types of foulants found on the membrane surface. From the SEM images (Fig. 7), the fouling layer on the MDBR and MD is visually different. The fouling layer in the MDBR consists of biomass and extracellular polymeric substances which are chemically and structurally different from the feed protein and carbohydrate that had adhered to the MD membrane surface. Fig. 10 shows the FTIR spectrum of the virgin PVDF membrane as well as the fouled MD and MDBR membranes. The transmittance dips observed at 1100 to 1300 cm − 1 for the virgin membrane correspond to CF2 and CF3 [41]. The signal is dampened/non-existent in the MDBR and MD fouled membrane, indicating that membrane fouling had effectively muted the hydrophobic functionality of the virgin membrane. The transmittance dips at 1030 to 1040 cm−1 indicate carbohydrate fouling [42] while the transmittance dips at the Amide I (1600–1700 cm−1) [43] and Amide II (1500–1550 cm−1) [43] bands indicate the presence of protein fouling on both MD and MDBR membranes. There are some differences between the MDBR and MD fouled membrane, for example, the C O stretch (indicative of ester and fatty acids) observed on MDBR fouled membrane at 1735 cm − 1 [44–46] and the transmission peak at 1242 cm − 1, which may be due to PO2− antisymmetric stretching vibration from nucleic acid [45]. The CH2 asymmetric stretch was also

Permeate TOC (mg/L)

flux decreases, less water condenses on membrane surface and less heat of condensation is liberated, leading to a lower T2. In short, the thin fouling layer probably did not confer significant heat transfer resistance but it probably did result in significant mass transfer resistance, thus accounting for the lower flux observed. A closer examination of the SEM image of the pre-treated, fouled MDBR membrane (after 22 days of operation, image not shown here) shows that there were pores smaller than 50 nm. Convective transport may take place within the more porous parts of the biofilm [38] but since the biofilm is predominantly made up of extracellular polymeric substance (EPS), the main mass transport mode for water through the hydrated EPS matrix would be diffusion. The self-diffusion coefficient of water within a biofilm has been estimated to be 15% lower than that in bulk water [39] so water diffusion through the hydrated EPS matrix would have conferred a degree of mass transfer resistance. On the other hand, the SEM image of a PVDF membrane after 15 hours of operation in the MD system shows a thin, non-uniform coating on the PVDF membrane pores (Fig. 7b). After 15 hours of operation, the thin fouling layer evidently did not result in significant mass or heat transfer resistance, as shown by the similar MD flux observed on days 1 and 2 of operation (Fig. 6). A cross-sectional examination of the fouled MD membrane after 7 days of operation shows a non-uniform fouling layer with thickness as high as 5.5 μm. While the observed maximum MD fouling layer thickness (after 7 days of operation) was similar to that of the MDBR after 7 days of operation, the MD flux was 19% higher than the MDBR flux, indicating that the flux in the MD process is influenced by not just the fouling layer thickness but also the characteristics of the fouling layer (e.g., structure, composition, coverage, pore blocking). Pore blocking by scalants and proteins crystallization have been known to reduce the effective membrane area in MD processes, resulting in flux decline [40]. A lower fouling layer coverage on the MD membrane may account for the lower flux decline. Another possible scenerio is that the bacteria in the MDBR may have covered some of the membrane pores (as shown in Fig. 7c). As the EPS accumulate around the bacteria, a “crust” of low permeability may have developed, resulting in a localized resistance effect on the membrane surface and hence, the lower flux observed in the MDBR (compared to the MD) after 7 days of operation. From the SEM images and fluxes of the MDBR and MD experiments, it is likely that the bulk of the biofilm may have played a minor role in contributing to the hydraulic resistance while the pore coverage may have accounted for the significant decrease in flux for the MDBR.

45

MDBR Permeate TOC MD Removal Efficiency

Fig. 9. Change in permeate TOC and overall TOC removal efficiency with time in MD and MDBR.

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S. Goh et al. / Desalination 323 (2013) 39–47

observed at 2920 cm − 1 for the MDBR fouled membrane, indicating the presence of fatty acid [44]. Hydrophilic PVDF membranes also sport similar C O bonds at 1735 cm − 1 [46,47], implying that the C O bond would have conferred some degree of hydrophilicity on the MDBR fouled membrane. Presumably it is the chemical nature of the foulant (not the thickness) that will dictate the rate of wetting. A thin layer of amphiphilic coating can reduce the contact angle of the membrane and result in wetting. The contact angle results are supported by the FTIR results which show greater obscuration of the hydrophobic CF2 and CF3 functional sites on the MD membrane. Though the difference is slight, the transmission dip at the Amide I and II groups deposited on the MD and MDBR fouling layer occur at different wavelengths, indicating that the type and structure of the protein foulant are slightly different as well. Some foulants (especially those with amphiphilic structures) can accelerate wetting. While copious amounts of EPS have been observed on the MDBR membrane after 3 days of operation (Fig. 7), wetting was only observed after 22 days of operation. In comparison, the MD membrane (with a light coating of feed foulants) was wetted in 7 days, implying that the feed foulants probably promote wetting more than EPS. The higher feed TOC and TN in the MD retentate had likely resulted in the faster adsorption of critical foulant(s) onto the MD membrane, leading to faster membrane wetting. The variation in concentration and composition of the foulants in the MD and MDBR systems, as well as the observed differences in the flux decline rates, may account for varying degrees of hydrophilicity observed on the fouled membranes and the different rates of wetting. 4. Conclusions The biological removal of carbohydrates and proteins in the MDBR has been demonstrated. In this work, the MDBR flux was maintained at more than 6.8 L/m 2 h with less than 20% loss in flux for at least 13 days. This was only 8% lower than the average MD flux of 7.3 L/m 2 h (on day 6 of operation). The faster flux decline in the MDBR is likely due to the increased thermal and hydraulic resistance of the fouling layer. This can probably be controlled with periodic membrane cleaning [14,48] and/or optimization of operating parameters (e.g., sludge retention time, aeration rate). Mass transfer resistance in the fouling layer is likely to be more significant than heat transfer resistance. The significant decrease in flux in the MDBR may be due to pore coverage which resulted in a decrease in effective membrane area. It is noted that severe

temperature polarization (possibly due to bad fluid dynamics in the submerged MD/MDBR system) may have accounted for the initial flux of 8 L/m2 h, which is lower than that of traditional MBR and DCMD. Future work would look into improving the hydrodynamics of the system and limiting fouling to ensure that the MDBR flux is comparable with that of the traditional MBR and DCMD systems. Membrane fouling has been shown to compromise membrane hydrophobicity and accelerate wetting. The difference in concentration and composition of foulants in the retentate may have resulted in different rates and types of foulant adsorption and accounts for the different rates of wetting. Once the membrane has been wetted, it has to be cleaned and dried before it can be reused. Delaying membrane wetting thus reduces the frequency of membrane cleaning and drying. Despite the slightly lower flux (a difference of 8%), the MDBR was able to delay wetting by 1.7–3.6 times in this work. The MDBR would be a better option than MD in the reclamation of certain industrial wastewaters that are hot, or have access to waste heat, with low volatile organic content. Nomenclature C TOC concentration in MDBR retentate (gTOC/L) TOC concentration in feed (gTOC/L) Cf Overall TOC removal efficiency in the MDBR Coverall TOC concentration in permeate (gTOC/L) Cp Hydraulic diameter (m) dh Heat transfer coefficient in biofilm or fouling layer hb (W/m 2 K) Heat transfer coefficient in feed boundary (W/m 2K) hf Heat transfer coefficient in membrane (W/m 2 K) hm Heat transfer coefficient in permeate boundary (W/m 2 K) hp J Permeation flux (kg/m 2 s) Thermal conductivity of fouling layer (W/m K) kb Thermal conductivity of membrane (W/m K) km L Chamber length (m) N TN concentration in MDBR retentate (g TN/L) TN concentration in feed (g TN/L) Nf Noverall Overall TN removal efficiency in the MDBR TN concentration in permeate (g TN/L) Np Nu Nusselt number Vapor pressure at membrane surface on the feed side (Pa) P1 Vapor pressure at membrane surface on the permeate side P2 (Pa) Vapor pressure at fouling layer (Pa) Pb Vapor pressure at bulk feed (Pa) Pf

MD fouled membrane

MDBR fouled membrane

Virgin PVDF membrane

Fig. 10. FTIR spectrum of virgin PVDF membrane, fouled membranes from MDBR and MD setup.

S. Goh et al. / Desalination 323 (2013) 39–47

Pp Pr Qf Qp Qw Rb Re Rf Rm Rp rsu Rt T1 T2 Tb Tf Tp TSS V X δ δb ΔHv

Vapor pressure at bulk permeate (Pa) Prandtl number Feed flow rate into MDBR (L/day) Permeate flow rate out of MDBR (L/day) Sludge wasting rate out of MDBR (L/day) Resistance in biofilm or fouling layer (Pa m 2 s/kg) Reynolds number (dhρν/μ) Resistance in Feed boundary (Pa m 2 s/kg) Resistance in membrane (Pa m 2 s/kg) Resistance in permeate boundary (Pa m 2 s/kg) Rate of substrate removal by biomass (g/L day). Substrate can refer to TOC or TN. Total resistance (Pa m 2 s/kg) Temperature at membrane surface on the feed side (K) Temperature at membrane surface on the permeate side (K) Temperature at on fouling layer surface (K) Temperature in bulk feed (K) Temperature in bulk permeate (K) Total suspended solids Reactor Volume (L) Biomass concentration (g TSS/L) Membrane thickness (m) Fouling layer thickness (m) Heat of vaporization (J/kg)

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