- Email: [email protected]
Recovery of brines from cheese making using membrane distillation at lab and pilot scale
Accepted Manuscript Recovery of brines from cheese making using membrane distillation at lab and pilot scale Lies Eykens, Kristien De Sitter, Charlotte Boeckaert, Jan Boeckx, Guy Borgmans PII:
S0260-8774(17)30406-5
DOI:
10.1016/j.jfoodeng.2017.09.016
Reference:
JFOE 9020
To appear in:
Journal of Food Engineering
Please cite this article as: Lies Eykens, Kristien De Sitter, Charlotte Boeckaert, Jan Boeckx, Guy Borgmans, Recovery of brines from cheese making using membrane distillation at lab and pilot scale, Journal of Food Engineering (2017), doi: 10.1016/j.jfoodeng.2017.09.016 This is a PDF file of an unedited manuscript that has been accepted for publication. As a service to our customers we are providing this early version of the manuscript. The manuscript will undergo copyediting, typesetting, and review of the resulting proof before it is published in its final form. Please note that during the production process errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain.
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Highlights -
The feasibility of brine treatment from cheese making was investigated using membrane distillation Microfiltration was successfully used as a pretreatment for the brines.
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Fouling occurred due to the deposition of proteins and calcium phosphates.
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Both at lab and pilot scale the technical feasibility of MD was shown.
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The treatment of cheese brines with MD was found to be economically viable
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Recovery of brines from cheese making using membrane distillation at lab and pilot scale
1 2
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Lies Eykens , Kristien De Sitter , Charlotte Boeckaert , Jan Boeckx , Guy Borgmans
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VITO - Flemish Institute for Technological Research, Boeretang 200, 2400 Mol, Belgium
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2
VLAKWA – Vlaams Kenniscentrum Water, Graaf Karel de Goedelaan 34, 8500 Kortrijk
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*Corresponding author – Email address: [email protected], [email protected]
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Abstract
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Brines are a problematic waste stream due to their high salinity. In cheese manufacturing,
9
a salinity up to 220 g/l is used for the pickling bath. This waste stream is currently
10
transported and treated externally at high costs. Membrane distillation is able to treat
11
much higher salinities compared to the traditional desalination techniques. In this article,
12
membrane distillation was therefore proposed as an alternative treatment technology,
13
simultaneously allowing the recovery of the salts and the production of pure water.
14
Microfiltration was successfully used as a pretreatment for the brines of two different
15
companies. The denaturation temperature of the proteins was found at 65.5°C and it was
16
shown that the temperatures in the MD process should be below this temperature.
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Different membranes were explored at lab scale and module performance tests confirm
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the feasibility of MD at pilot scale for this application. Additionally, it was shown that the
19
total water production costs are mainly affected by the availability of waste heat. Without
20
waste heat, the lowest cost of 35 €/m3 is achieved for the 24 m2 modules, with a total
21
thermal power demand of 34 kW. The costs for discharge of the brines are estimated at 50
22
– 150 €/m3, while the value of the salts recuperated equals 20 €/m3 distillate produced.
23
Therefore, even when there is a low availability of waste heat, application of membrane
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distillation for the recovery of brines from cheese making is an economically viable and
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technically feasible alternative for external treatment.
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Keywords: Economic evaluation, Fouling, Pretreatment, Technical feasibility
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1 Introduction
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During the cheese production process, cheese is pickled in a concentrated salt solution, to
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improve the preservability, consistency and the flavor of the cheese. The cheese is
31
releasing water, while taking up salts, which decreases the salt concentration of the
32
pickling bath. To keep the salinity constant, additional salt is added to the pickling bath and
33
regularly, part of the brine is drained. This brine is a problematic waste stream within the
34
production process, containing a sodium chloride concentration of 200 to 220 g/l with a
35
complex composition, including organic material, calcium and proteins. Nowadays, these
36
brines are transported and treated externally, which is a non-sustainable way of coping
37
with these brines at a high cost of 50 – 150 €/ton [1].
38
Within the Pekelmem project (Milieu- en energietechnologie Innovatie Platform (MIP) no.
39
150362), MD was proposed as a new technology to treat the excess brine and reuse the
40
salts in the process. This would severely reduce the amount of salts required in the process
41
and the volumes of waste water to be transported. Within the project, the brines of two
42
different cheese types are treated [2]. For the first type of cheese the pickle bath is not
43
actively clarified and currently a substantial amount of the brine is discharged and salt is
44
added to maintain the salinity as visualised in Figure 1 A. Alternatively, process scheme B in
45
Figure 1 is proposed, where the brine is pretreated with microfiltration (MF) to protect the
46
MD modules and the permeate of the MF is sent to the MD. The volume of concentrate of
47
the MF depends on the recovery that can be achieved and is only a small fraction of the
48
original brine volume containing both organics and salts.
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Figure 1
For a second type of cheese, the pickling bath is currently continiously clarified with a
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cartridge filter to actively remove suspended solids and excess organics. To maintain the
51
salinity, also a substantial volume of brine is discharged and salt is added (Figure 2 A).
52
Scheme B in Figure 2 proposes an alternative treatment, where microfiltration or
53
ultrafiltration replaces the cartridge filter to clarify the pickling bath. The resulting
54
permeate is a clarified salt solution containing small molecular weight organics. A part of
55
this volume is redirected towards the membrane distillation step for removal of excess
56
water and reuse of the concentrated salt. Since a large volume of the pickling bath is
57
continuously filtered, a relatively large volume of concentrate containing organics and salts
58
needs to be regenerated. The proposed regeneration makes use of diafiltration to remove
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contains organics, which can be treated by the standard biological waste water treatment
61
facility of the company. The permeate of the diafiltration is a saline stream with lower
62
amount of salts, depending on the number of diafiltrate volumes. This stream is
63
concentrated to the required salinity using membrane distillation. Figure 2
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To the best of our knowledge, the proposed treatment schemes are not tested before,
65
although membrane distillation itself was already proposed in the dairy industry as an
66
alternative for classical evaporation for the production of milk powder and the
67
concentration of whey proteins [3–5]. This was experimentally explored in the PhD of
68
Hausmann et al. [6–11]. Kezia et al. [12] focused on the concentration of salty whey, a
69
waste effluent from cheese making, containing up to 80 g/l NaCl as the major component.
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For all these applications, prevention and cleaning of protein fouling was reported as the
71
major challenge for the use of membrane distillation for these types of applications.
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Moreover, it was also found that milk fat partially wetted the membrane. Therefore, the
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major challenge in this article was to evaluate the applicability of membrane distillation for
74
the recovery of brines of the cheese making. This is done via a stepwise characterization
75
procedure (see Figure 3), including the feed water characteristics and pretreatment choice,
76
the fouling behavior and feasibility at lab scale, the validity of these labtests at pilot scale
77
and an economic evaluation (Figure 3). In a later phase of the project, on-site testing will
78
be performed at pilot scale. At lab scale, DCMD was used to evaluate the fouling potential
79
of the pickle brine. The DCMD configuration is ideal for this purpose, due to its simplicity in
80
design and sufficiently high production capacity at small scale. At full scale however, AGMD
81
is preferred as stated in a recent publication [13]. The first reason is that DCMD has a
82
higher installation cost due to the additional heat exchanger that is required to recuperate
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heat. Secondly, DCMD has lower fluxes compared to AGMD and therefore requires more
84
membrane surface for the same production capacity. Moreover, especially for longer
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modules, AGMD was also found to be more efficient compared to DCMD.
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2 Materials and methods
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2.1
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The samples of the brines obtained by company 1 are obtained from an actively clarified
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pickling bath (Figure 2). The samples of company two are obtained from a pickling bath
90
without active clarification (Figure 1). The composition of the samples as received is given
91
in Table 1.
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Feed water characteristics
Company 1 4.88 225 11 117 342
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pH Conductivity (mS/cm) Turbidity (FNU) Suspended solids (mg/L) Chloride (g/L) Phosphates (mg/L)
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Table 1: Chemical composition of the cheese brines of the two cheese companies. Company 2 5.05 215 65.3 58 104 288
The protein denaturation temperature was determined visually after the filtration step
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(discussed in section 2.2) on the permeate of company 1. In a first test, temperature was
94
raised every five minutes with 10°C under continuous stirring. In a second test, the
95
temperature is increased stepwise by 1°C every 10 minutes. The point where the sample
96
becomes turbid and flakes are observed is defined as the protein denaturation point.
97
2.2
98
For the pretreatment, an in house built filtration setup was used (Figure 4). Two different
99
types of spiral wound membranes were explored; a BN PVDF 50kDa UF membrane and a
100
v0.1 PVDF MF membrane from Synder Filtration. The membranes had an area of 2.14 m2
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and a spacer thickness of 1.2 mm. The experiments were carried out in crossflow mode
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with a cross flow velocity of 3 m3/h equal to a linear flow velocity of 0.3 m/s.
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MF/UF filtration
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Figure 4
2.3
Lab scale MD setup
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The properties of the membrane selected for the lab scale tests are given in Table 2. This
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membrane was produced by Lydall (Central Manchester, Connecticut, U.S.A) and selected
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based on its availability in pilot modules.
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Table 2: Properties of the PE membrane [14,15].
Mean pore
Water contact
Hexadecane
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Thickness Porosity (µm)
(%)
size (µm)
angle (°)
contact angle (°)
99 ± 1
76 ± 1%
0.32 ± 0.02
120 ± 3
0
The feasibility of membrane distillation and the fouling potential of the brine were
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evaluated at lab scale with a DCMD setup (Figure 5). The DCMD configuration is ideal for
109
this purpose, due to its simplicity in design and sufficiently high production capacity at
110
small scale. The flat-sheet module has a feed and permeate channel with a width of 6 cm, a
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length of 18 cm and a height of 2 mm. The setup is described in more detail elsewhere
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[15,16]. The electrical conductivity at the feed and permeate side were logged every 15
113
minutes by portable conductivity meters (WTW GmbH, pH/Cond 340i, Weilheim,
114
Germany). The feed and permeate temperatures were varied, but the temperature
115
difference across the membrane was kept constant at 15°C. The flow velocity was fixed at
116
0.07 m/s. The experiments were carried out in isoconcentration mode, meaning that pure
117
water was regularly added to the feed tank at the same rate as the permeation occurs, to
118
maintain a constant feed volume. Each flux measurement was performed for at least 30
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minutes. The weight variations were automatically logged every 2 minutes and the
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averages and standard deviations were calculated based on this weight variation, both on
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feed and permeate side.
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Figure 5
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Fouled membranes have been analyzed with a Nexus FT-IR spectrometer from Thermo
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Nicolet with Smart Orbit module in attenuated total reflectance (ATR) mode, controlled
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with the OMNIC™ software. Samples were analyzed on the fouled side and compared with
125
a new sample. Field emission scanning electron microscopy and energy-dispersive X-ray
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spectroscopy were performed on the fouled membrane surface using a Nova Nanosem
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(FEI).
128
2.4
129
The pilot MD setup was delivered by Aquastill BV, The Netherlands (Figure 6). A P&ID of
130
the pilot setup was descibed by Hitsov et al. [13] and can also be found in the appendix. A
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standard AGMD module of Aquastill was used, with an area of 14.4 m2, 6 parallel channels,
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a height of 0.4 m, a spacer thickness of 2 mm and the PE membrane. At full scale, AGMD is
133
preferred over DCMD as stated in a recent publication [13]. The first reason is that DCMD
134
has a higher installation cost due to the additional heat exchanger that is required to
135
recuperate heat. Secondly, DCMD has lower fluxes compared to AGMD and therefore
136
requires more membrane surface for the same production capacity. Moreover, especially
137
for longer modules, AGMD was also found to be more efficient compared to DCMD. A
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volume of 200 l brine was added to the feed tank to start the experiment. After each day,
139
the volume in the feed tank was removed and fresh brine was added. A cooling
140
temperature of 10°C was aimed at, since it is equal to the pickling bath temperature.
141
Nevertheless, this temperature could not be achieved by the cooling circuit available and
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therefore the coolant inlet temperature was fixed at 15°C to ensure a constant
143
temperature. The feed temperature was fixed at 50°C and a flow velocity of 0.05 m/s was
144
used. Each flux measurement was performed for at least 60 minutes after reaching stable
145
temperature values. All temperatures, pressures, flow rates and fluxes where
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automatically logged every 2 minutes. The flux was calculated based on the height
147
variation of the feed and permeate tank. The average and standard deviation were
148
calculated based on these values over at least 60 minutes monitoring time. Figure 6
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Pilot scale MD setup
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2.5
Economic evaluation
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The model developed by Hitsov et al. was used to calculate flux, energy consumption and
151
production capacity [13,17,18]. The evaluation of energy consumption is performed using
152
the gained output ratio (GOR), which is a dimensionless ratio used for thermal desalination
153
processes. It is defined as the ratio of the total latent heat of evaporation of the produced
154
water to the input thermal energy [19].
155
The model was calibrated and validated against experiments with synthetic NaCl solutions
156
with different salinities, temperatures and flow rates using the modules of Aquastill. As 6
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158
Only the cold inlet temperature was 10°C in the simulations, which is equal to the pickle
159
bath temperature. The simulated module sizes were 4.8 m2, 7.2 m2, 14.4 m2 and 24 m2.
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Using these simulations, a prediction for the number of modules and thermal power
161
required to treat a certain volume can be estimated in the optimal case, i.e., without
162
severe flux reduction due to fouling. Based on this, a rough estimation of the cost for MD
163
for the different module dimensions was made to enable a choice of the optimal module
164
dimensions for pilot tests on location. This calculation includes the depreciation of the
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installation within 10 years and module lifetime of 5 years [20]. During an energy audit
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within the company it was found that waste heat is available at a price of 0.025 €/kWh and
167
sufficient free waste cooling is available within the factory. This estimation does not take
168
into account labor costs, pretreatment, cleaning and maintenance.
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3 Results and discussion
170
3.1
171
Table 1 indicates the presence of suspended solids in the feed water, which are removed
172
during a first prefiltration step. Figure 7 shows the stable flux and transmembrane pressure
173
(TMP) during the MF prefiltration experiment for over 2 hours using the brine of
174
company 1. This is a strong indication of the feasibility of the prefiltration of the brine
175
before the MD. Similar results were observed for the UF filtration and for the MF filtration
176
of the brine of company 2. The average flux, TMP and recovery are reported in Table 3. The
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recovery was not limited due to fouling, as the flux remains constant, but due to the
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minimum volume required in the filtration setup. This indicates that in real application,
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even higher recoveries can be achieved.
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Pretreatment: MF/UF filtration
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Figure 7 Company 1 – UF -1
Company 1 - MF
Company 2 - MF
Flux (kg.m .h )
347 ± 1
353 ± 1
358 ± 1
TMP (bar)
0.88 ± 0.01
0.89 ± 0.01
0.99 ± 0.01
Recovery (%)
99.3
99.5
99.02
Table 3: Average flux, TMP and recoveries achieved during the prefiltration of the cheese brines using UF or MF.
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3.2
181
3.2.1 Scaling potential
182
To be able to evaluate the scaling potential of the MF/UF permeates in the MD, the exact
183
salt composition of the prefiltered samples was determined together with the total organic
184
carbon (TOC) present in solution (Table 4). Based on this information, the saturation index
185
of the salts present in the brines was calculated using the open source software PHREEQC.
186
This calculation only indicates saturation of hydroxyapatite (Ca5(PO4)3OH) for all three
187
samples (Figure 8). The composition of the deposits are further investigated during the
188
labtests are evaluated in section 3.3.3.
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Scaling/fouling potential of MD feed (MF permeate)
Conductivity (mS/cm) Total salinity (g/l) -
Cl (mg/l) 3
o-PO4 (mg/l) 2-
SO4 (mg/l) 2+
Ca (mg/l) 2+
Mg (mg/l)
Company 1 – UF
Company 1 - MF
Company 2 - MF
2500
2900
7700
206
199,7
209
180
197
211
109000
103000
113000
282
255
517
282
280
284
2200
2330
1360
78
77.2
137
68100
91300
95400
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Organic carbon (mg C/l)
+
Figure 8
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Na (mg/l)
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Table 4: Chemical composition of the UF and MF permeates (=MD feed)
3.2.2 Protein denaturation
191
Besides scaling, also protein fouling is seen as one of the major issues while treating dairy
192
effluents. As stated by Hausmann et al., the high hydrophobicity of MD membranes can
193
result in the establishment of hydrophobic interactions between the membrane and
194
proteins [9]. In order to diminish these interactions, feed temperature of the MD has to be
195
chosen wisely to avoid denaturation of the proteins present in the brines.
196
The protein denaturation temperature was determined visually on the MF filtrate from
197
company 1. In a first test, temperature was raised every five minutes by 10°C, showing that
198
the denaturation temperature is between 60 and 70°C (Figure 9 A). Secondly, the
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temperature is increased stepwise by 1°C every 10 minutes. Figure 9 B shows that the
200
proteins denaturate at a temperature above 63.3°C. Above this temperature the solution
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becomes turbid and flocculation is observed. Similar results were observed for the two
202
other samples. Figure 9
3.3
Lab MD experiments
204
3.3.1 Maximum feed temperature
205
The effect of protein denaturation on the MD process was investigated by lab testing
206
below (45°C) and above the denaturation temperature (80°C). These tests were performed
207
on the UF permeate of company 1 (Table 4). Figure 10 A shows that the flux remains
208
constant at 4.5 kg.h-1.m-2 during 20 hours at a feed temperature of 45°C. At 80°C, a much
209
higher initial flux was achieved. However, due to deposits of the proteins on the
210
membrane, this flux quickly decreases to 3.5 kg.h-1.m-2 within the first 90 minutes, after
211
which it remains constant. The fast decrease is caused by fouling of the denaturated
212
proteins on the membrane. After this time period, all proteins precipitated and are
213
removed from the feed solution. Therefore, no further flux decrease is observed.
214
Figure 10 B shows a difference of 46% with the clean water flux on a clean membrane and
215
with the initial flux with the brine at 80°C. After 15 hours, flux decreased with 71%. In
216
contrast, at 45°C the decrease in flux from clean water to the brine is only 33%, whereas
217
after 15 hours a decrease of only 6% is observed. The deposits of the proteins of both tests
218
are visualized in Figure 11. A white deposit was observed on the heat exchanger after both
219
tests. This deposit was easily removed using a NaOH cleaning at pH 10. On the spacer, the
220
membrane and in the feed tank, white flakes were observed for the MD-tests at 80°C. At
221
45°C, no deposits were observed visually on the spacer and in the feed tank. Deposits were
222
still found on the heat exchanger and on the membrane, visually showing an increased
223
reflectance. These tests already indicate that even at 45°C, deposits are found on the heat
224
exchanger showing the importance of a thorough cleaning strategy. However, the salt
225
retention of 99.9% obtained at both temperatures during the entire testing time shows
226
that the deposited proteins do not wet the membrane.
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Figure 10 Figure 11
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Similar tests were performed using the MF-filtrate from company 1 and 2, with similar
228
observations. The flux decreases quickly when performing membrane distillation tests 9
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230
of company 2, the deposits of the proteins even formed an obstruction in the module inlet.
231
The feed liquid could not be circulated properly and the test was stopped after a few
232
hours.
233
3.3.2 Lab MD experiments on different cheese brines
234
Figure 12 A shows the flux as a function of time for the different samples using the PE
235
membrane. The concentration was kept constant during this test by regularly refilling with
236
water. This reduced the effect of the salinity on flux and enabled visualization of the effect
237
of fouling on the membrane flux. The largest decrease in flux was observed for the MF
238
sample from company 2 (31%). This decrease was less severe for the samples from
239
company 1. For the MF filtrate, no fouling was observed during the 25 hours MD-test.
240
Surprisingly the UF filtrate showed a stronger decrease of flux of 10%, despite the fact that
241
this pretreatment removes more organics from the sample. This might be caused by the
242
variability between the batch of MF and UF filtration.
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3.3.3 Fouling analysis
244
The difference in flux between the start and end of an isoconcentration experiment is a
245
first indication of fouling on the membrane. The strongest decline was observed for the
246
sample of company 2 after MF. Therefore, this sample was selected for further
247
investigation of the fouling. The clean water fluxes before and after the experiment is used
248
to evaluate the fouling. Figure 13 shows that the flux with brine at the start of the
249
experiment is already 32% lower compared to the original clean water flux. This lower flux
250
can have several causes:
251 252 253 254
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Figure 12
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pressure and hence driving force. The activity of the different samples was
measured and equals 0.856. After calculation of the bulk driving force (Δp = pbrine,45°C –pw,30°C = awpw,45°C – pw,30°C), it appears that the bulk driving force
255 256
The presence of salinity results in a decrease of the activity of water, vapor
decreases with 26 % due to the presence of salts. -
Concentration polarization results in a higher concentration at the membrane
257
surface compared to the bulk solution and therefore negatively affects the driving
258
force and the flux. 10
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-
Liquid properties including density and viscosity are different compared to pure
260
water, affecting the hydrodynamics in the channels and therefore also temperature
261
and concentration polarization.
263
-
The startup, where feed and permeate are directed towards the aimed temperatures, takes about 1 hour. Fouling can already take place during this time.
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During the experiment the flux decreases further by 31% (5.5 to 3.8 kg.h-1.m-2). The
265
concentration remains constant throughout the experiment and the start and end-flux
266
were determined at equal concentration. Hence, the decrease in flux is mainly caused by
267
fouling. The measurement of the clean water flux after the measurement shows that the
268
original flux is completely restored after rinsing with pure water only. This indicates that
269
the fouling, if still present after rinsing, at least does not affect the flux. After the rinsing
270
step and clean water flux, deposits are still observed on the membrane (Figure 11 H). This
271
is an indication that long term operation requires a cleaning strategy to reduce the effect
272
of fouling on the process performance. The TOC retention is 95%, while the salt retention
273
remains above 99.97%, indicating that no wetting occurred during the test. The lower
274
retention of the TOC can be caused by the transport of volatile organics present in the MD
275
feed stream or by surface diffusion.
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Figure 13
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Table 5: Membrane performance during an MD isoconcentration experiment. Process conditions: Tf = 45°C, Tp = 30°C, v = 0.07 m/s, membrane: PE1, sample: company 2 after MF. Average flux with -1 -2 brine (kg.h .m )
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4.9 ± 0.7
Flux decrease after 25 hours (%)
Salt retention (%)
TOC retention (%)
31%
99.97% ± 0.01%
95%
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The fouling after the experiment and after the rinsing step with clean water was
278
characterized using the infrared spectra of a clean and a fouled PE membrane, given in
279
Figure 14. The C-H and CH2 peaks are visible for both spectra and are assigned to the
280
chemical structure of PE. The peaks that are only visible for the fouled membrane are
281
caused by the presence of proteins on the membrane surface. Moreover, the peaks at
282
1030 and 557 cm-1 are typical for inorganic phosphates (PO43-), suggesting the presence of
283
hydroxyapatite crystals, as predicted by PHREEQC (Figure 8).
11
ACCEPTED MANUSCRIPT Figure 14
The SEM/EDX analysis in Figure 15 again shows that after the isoconcentration experiment
285
and after rinsing, the deposits contain phosphor, calcium and oxygen. This is again a strong
286
indication for the presence of hydroxyapatite on the membrane. Figure 15
287
3.4
288
A 14.4 m2 AGMD module was used to perform a test at pilot scale. The experiment started
289
with 1 batch of 200 l prefiltered (MF) brine from company 2. After 24 hours, this batch was
290
replaced by a new batch of 200 l in order to better simulate the real case, where
291
continuously fresh feed was treated by the MD unit. Figure 16 A shows the flux and the
292
permeate conductivity as function of time for the three different batches. On average a
293
flux of 0.76 ± 0.04 kg.h-1.m-2 was achieved, while a simulation under equal conditions
294
predicted a flux of 0.84 kg.h-1.m-2. The clean water flux before and after the experiment
295
shown in Figure 16 B equals 1.69 ± 0.02 kg.h-1.m-2 and 1.61 ± 0.02 kg.h-1.m-2 respectively,
296
indicating that the foulants exposed to the membrane during 3 days do not severely affect
297
the flux. As the feed is first heated on the aluminum foil side of the module and in the
298
external heat exchanger, fouling is probably mainly located there. This effect was also
299
observed during the pilot scale experiments to a minor extent. The evaporator and
300
condenser inlet pressures using the brine are increased compared to the pure water flux
301
experiments. However, these pressures do not increase over time, indicating that no
302
severe blocking occurs due to fouling in the module on the condenser foil or the
303
membrane.
304
During the 3 batches, on average a salt retention of 99.8% was obtained. This value is
305
lower compared to the lab tests described in Section 3.3.2, where a salt retention of
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99.97% was found for the PE1 membrane. As this salt retention is constant for all batches,
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gradual wetting is not expected to cause the lower salt retention. This lower retention is a
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known phenomenon and can be caused by defects in module sealing or the membrane,
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which are more probable at larger surface area [21]. Similarly to the labscale results in
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Section 3.3.2, the TOC retention was 95.6%, which is mainly caused by the volatile organic
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compounds present in the feed stream that can easily migrate through the vapor space in
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the membrane.
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3.5
Economic evaluation
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3.5.1 Module performance simulation
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A MATLAB model was used to simulate the pilot scale modules. The mass transfer through
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the membrane was calculated using the Dusty Gas Model [22]. The heat and mass transfer
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in the feed and permeate channels were calculated using Nusselt and Sherwood type
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equations, which are calibrated and validated as described in [23]. The membrane
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calibration was extensively discussed elsewhere [24]. More details on the construction of
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the model and calibration are described elsewhere [13,17,18,24].
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These simulations allow calculating flux, gained output ratio and production capacity of
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one module for modules with different sizes. This enabled a quick estimation of the effect
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of module size and a proper selection of a suitable module for module testing using the
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brine. Figure 17 A shows higher fluxes for smaller modules. The calculation of the GOR in
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Figure 17 B shows that small modules are less efficient. The internal energy recuperation
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becomes more efficient for larger modules with a lower temperature difference. Therefore
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a balance between flux and energy efficiency should be considered when selecting the
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length of the modules. Figure 17 C indicates that the production capacity of one module
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also increases with increasing module size, despite the decreasing fluxes.
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Figure 17
3.5.2 Economics of MD for brine treatment
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Using the simulated fluxes and required thermal energy, the water production cost (€/m3)
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was estimated. The number of modules was calculated based on the production capacity
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(Figure 17 C), assuming a distillate production requirement of 650 m3/year. This is
334
calculated based on the current discharge volumes of the companies. The installation of
335
4 x 24 m2 AGMD modules would be sufficient to cover the required production capacity. 5
336
modules of the 7.2 or 14.4 m2 AGMD type or 6 modules of the 4.8 m2 would be required to
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compensate for its lower production capacity per module.
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An estimation of the costs for MD was made depending on the membrane area and the
339
availability of the waste heat. This calculation includes a depreciation of the installation of
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10 years and a membrane lifetime of 5 years. The total water production costs are mainly
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affected by the availability of waste heat (Figure 18). The lowest cost without waste heat
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34 kW. As larger modules are more efficient (Figure 17 B), the energy costs decrease
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substantially with increasing membrane area. For the case with 50% waste heat, the 24 m2
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module again shows the best balance between energy and membrane costs. Four modules
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and a thermal power of 34 kW are required to cover the production capacity at a cost of 24
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€/m3 distillate. In the case 100% of the heat requirements are covered by waste heat, the
348
lowest cost is 12 €/m3 independent of the module size. The larger modules are more
349
productive so less, but slightly more expensive modules are required for the 24 m2
350
modules. The heat requirement increases from 34 up to 120 kWh with decreasing module
351
size. Similar calculations for DCMD were performed, where it was found that AGMD was 10
352
to 60% cheaper compared to the DCMD configuration, depending on the membrane area
353
installed in the modules.
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The costs for discharge of the brines are estimated at 50 – 150 euro/m3 and may increase
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due to the stricter discharge regulations in the future. Moreover, the salts in this
356
discharged brine have a value of 20 €/m3. The removal of 1 m3 pure water with MD, avoids
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a discharge of an equal amount of pickle water. Even without the availability of waste heat
358
at a cost of 35 €/m3, the high discharge costs make the application of membrane
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distillation for the recovery of brines from cheese making economically viable.
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Figure 18
4 Conclusions
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Due to the high disposal costs, cheese factories are looking for a way of reusing their pickle
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brines. The brine composition is out of the scope of classical desalination techniques due
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to the high salinity and the presence of organic components, which denaturate above
364
60°C. Therefore, membrane distillation is an interesting alternative, avoiding the current
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costly external treatment and discharge of the salts. This article shows the promising
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potential and the technical and the economic viability of membrane distillation for the
367
treatment of brines in the cheese industry.
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Microfiltration was found sufficient to clarify the brine before directing the feed stream
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towards the MD unit. The labtests show that the choice of the process conditions is mainly
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limited by the protein denaturation temperature (63.3°C). Above this temperature severe
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fouling is observed, whereas a slightly declining flux was obtained below this temperature
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recovered to its initial value and basic and acid cleaning are proven to be suitable to clean
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the heat exchangers and the membrane. The availability of waste heat is an important
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parameter affecting the cost of MD. Without waste heat, the lowest cost of 35 €/m3 is
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achieved for the 24 m2 module with a thermal energy demand of 34 kW. In the case of
377
100% availability of the waste heat, the lowest cost of 12 €/m3 is estimated for all modules,
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with a heat requirement increasing with decreasing module size.
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Acknowledgements
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The authors thankfully acknowledge the funding provided by VLAIO (150362).
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APPENDIX
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Figure A.1: P&ID of the MD pilot in AGMD mode
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References
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[1]
K. De Sitter, MIP Pekelmem - Deliverable 1.1: Water- en energiebalansen, 2016.
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[2]
Pekelmem, (n.d.). February 21, 2017).
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[3]
S.N. Moejes, A.J.B. van Boxtel, Energy saving potential of emerging technologies in milk powder production, Trends Food Sci. Technol. 60 (2016). doi:http://dx.doi.org/10.1016/j.tifs.2016.10.023.
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[4]
M. Tomaszewska, L. Białończyk, Influence of proteins content in the feed on the course of membrane distillation, Desalin. Water Treat. 51 (2013) 2362–2367. doi:10.1080/19443994.2012.728052.
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[5]
K. Christensen, R. Andresen, I. Tandskov, B. Norddahl, J.H. du Preez, Using direct contact membrane distillation for whey protein concentration, Desalination. 200 (2006) 523–525. doi:10.1016/j.desal.2006.03.421.
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A. Hausmann, PhD Dissertation: Membrane distillation in diary processing, Victoria University, 2013.
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A. Hausmann, P. Sanciolo, T. Vasiljevic, U. Kulozik, M. Duke, Performance assessment of membrane distillation for skim milk and whey processing, J. Dairy Sci. 97 (2014) 56–71. doi:10.3168/jds.2013-7044.
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A. Hausmann, P. Sanciolo, T. Vasiljevic, E. Ponnampalam, N. Quispe-Chavez, M. Weeks, M. Duke, Direct contact membrane distillation of dairy process streams, Membranes (Basel). 1 (2011) 48–58. doi:10.3390/membranes1010048.
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A. Hausmann, P. Sanciolo, T. Vasiljevic, M. Weeks, K. Schroën, S. Gray, M. Duke, Fouling mechanisms of dairy streams during membrane distillation, J. Memb. Sci. 441 (2013) 102–111. doi:10.1016/j.memsci.2013.03.043.
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[10]
A. Hausmann, P. Sanciolo, T. Vasiljevic, M. Weeks, K. Schroën, S. Gray, M. Duke, Fouling of dairy components on hydrophobic polytetrafluoroethylene (PTFE) membranes for membrane distillation, J. Memb. Sci. 442 (2013) 149–159. doi:10.1016/j.memsci.2013.03.057.
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A. Hausmann, P. Sanciolo, T. Vasiljevic, M. Weeks, M. Duke, Integration of membrane distillation into heat paths of industrial processes, Chem. Eng. J. 211–212 (2012) 378–387. doi:10.1016/j.cej.2012.09.092.
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[12]
K. Kezia, J. Lee, M. Weeks, S. Kentish, Direct contact membrane distillation for the concentration of saline dairy effluent, Water Res. 81 (2015) 167–177. doi:10.1016/j.watres.2015.05.042.
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[13]
I. Hitsov, K. De Sitter, C. Dotremont, I. Nopens, Full-scale validated Air Gap Membrane Distillation ( AGMD ) model without calibration parameters, J. Memb. Sci. 533 (2017) 309–320. doi:10.1016/j.memsci.2017.04.002.
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L. Eykens, PhD Dissertation: A comprehensive study of membrane distillation:
(accessed
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membrane development, configuration assessment and applications, KU Leuven, 2017. [15]
L. Eykens, K. De Sitter, C. Dotremont, L. Pinoy, B. Van der Bruggen, K. De Sitter, C. Dotremont, L. Pinoy, B. Van Der Bruggen, Characterization and performance evaluation of commercially available hydrophobic membranes for direct contact membrane distillation, Desalination. 392 (2016) 63–73. doi:10.1016/j.desal.2016.04.006.
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[16]
L. Eykens, I. Hitsov, K. De Sitter, C. Dotremont, L. Pinoy, I. Nopens, B. Van der Bruggen, Influence of membrane thickness and process conditions on direct contact membrane distillation at different salinities, J. Memb. Sci. 498 (2016) 353–364. doi:10.1016/j.memsci.2015.07.037.
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[17]
I. Hitsov, PhD Dissertation: Model-based analysis and optimization of membrane distillation, Ghent University, 2017.
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[18]
I. Hitsov, L. Eykens, W. De Schepper, K. De Sitter, C. Dotremont, I. Nopens, Full-scale Direct Contact Membrane Distillation (DCMD) model including membrane compaction effects, J. Memb. Sci. 524 (2017) 245–256. doi:10.1016/j.memsci.2016.11.044.
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[19]
J. Koschikowski, M. Wieghaus, M. Rommel, V.S. Ortin, B.P. Suarez, J.R. Betancort Rodríguez, Experimental investigations on solar driven stand-alone membrane distillation systems for remote areas, Desalination. 248 (2009) 125–131. doi:10.1016/j.desal.2008.05.047.
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[20]
A.S. Alsaadi, N. Ghaffour, J.D. Li, S. Gray, L. Francis, H. Maab, G.L. Amy, Modeling of air-gap membrane distillation process: A theoretical and experimental study, J. Memb. Sci. 445 (2013) 53–65. doi:10.1016/j.memsci.2013.05.049.
447 448 449 450
[21]
H.C. Duong, A.R. Chivas, B. Nelemans, M. Duke, S. Gray, T.Y. Cath, L.D. Nghiem, Treatment of RO brine from CSG produced water by spiral-wound air gap membrane distillation - A pilot study, Desalination. 366 (2015) 121–129. doi:10.1016/j.desal.2014.10.026.
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[22]
R.W. Field, H.Y. Wu, J.J. Wu, Multiscale Modeling of Membrane Distillation: Some Theoretical Considerations, Ind. Eng. Chem. Res. 52 (2013) 8822–8828. doi:10.1021/ie302363e.
454 455 456
[23]
J. Phattaranawik, R. Jiraratananon, A.G. Fane, Heat transport and membrane distillation coefficients in direct contact membrane distillation, J. Memb. Sci. 212 (2003) 177–193. doi:10.1016/S0376-7388(02)00498-2.
457 458 459 460
[24]
I. Hitsov, L. Eykens, K. De Sitter, C. Dotremont, L. Pinoy, B. Van der Bruggen, I. Nopens, Calibration and analysis of a direct contact membrane distillation model using Monte Carlo filtering, J. Memb. Sci. 515 (2016) 63–78. doi:10.1016/j.memsci.2016.05.041.
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Figure 1: Pickling bath without active clarification: Process scheme for (A) the current situation and (B) the alternative proposed in this study with brine recuperation.
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Figure 2: Pickling bath with active clarification: Process scheme for (A) the current situation and (B) the alternative proposed in this study with brine recuperation.
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Figure 3: Evaluation process of the applicability of MD.
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Filtration membraan Permeate
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Figure 4: Filtration equipment and the process scheme of the filtration equipment.
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Figure 5: A) MD setup, B) Membrane module and spacer, C) Scheme of the MD setup.
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Figure 6: Pictures of the MD-pilot and the module in AMGD mode.
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Figure 7: Time course of the MF prefiltration experiment using the brine of company 1.
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Figure 8: Saturation index of salts present in the different samples.
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Figure 10: A) MD flux as function of time of the UF-filtrate of company 1, B) clean water flux initial flux and the flux after 15 hours. ΔT=15°C, v=0.07 m/s, Membrane: PE1.
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Figure 11: Depositions at the end of the MD test at 80°C (A,B,C,D) and 45°C (E,F,G,H) on the heat exchanger (A,E), spacer (B,F), feed tank (C,G) and the membrane (D,H)
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Figure 12: A) flux as function of time for the 3 prefiltered samples and B) the initial flux using the pickle brine and the flux after 25 hours. Process conditions: Tf = 45°C, Tp = 30°C, v = 0.07 m/s, membrane: PE1.
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Figure 13: Clean water fluxes before and after an experiment and the start and end process flux with pickling water. Process conditions: Tf = 45°C, Tp = 30°C, v = 0.07 m/s, membrane: PE1
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Figure 14: Infrared spectra of a clean (blue, bottom) and fouled (red,up) PE1 membrane (MD process conditions: Tf = 45°C, Tp = 30°C, v = 0.07 m/s, membrane: PE1, sample: MF permeate company 2).
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Figure 15: SEM images of a clean (A) and fouled (B) PE-membrane. C) EDX spectrum of the fouled PE1 membrane.
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Figure 16: A) Flux and B) Evaporator and condenser inlet pressures and flow rate as function of time using the 7.2 m2 module from Aquastill. Process conditions: Te,in = 50 °C, Tc,in = 15°C, v = 900 l/h
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Figure 17: Simulated flux (A), GOR (B) and production capacity per module (C) for AGMD modules with varying membrane area. Process conditions: Tf = 50°C, Tp = 10°C, v = 900 l/h
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Figure 18: Treatment cost estimation in €/m3 destillate devided in installation costs, membrane costs and energy costs, without waste heat, B) with 50% waste heat and C) with 100% waste heat availability.
S0260-8774(17)30406-5
DOI:
10.1016/j.jfoodeng.2017.09.016
Reference:
JFOE 9020
To appear in:
Journal of Food Engineering
Please cite this article as: Lies Eykens, Kristien De Sitter, Charlotte Boeckaert, Jan Boeckx, Guy Borgmans, Recovery of brines from cheese making using membrane distillation at lab and pilot scale, Journal of Food Engineering (2017), doi: 10.1016/j.jfoodeng.2017.09.016 This is a PDF file of an unedited manuscript that has been accepted for publication. As a service to our customers we are providing this early version of the manuscript. The manuscript will undergo copyediting, typesetting, and review of the resulting proof before it is published in its final form. Please note that during the production process errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain.
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Highlights -
The feasibility of brine treatment from cheese making was investigated using membrane distillation Microfiltration was successfully used as a pretreatment for the brines.
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Fouling occurred due to the deposition of proteins and calcium phosphates.
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Both at lab and pilot scale the technical feasibility of MD was shown.
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The treatment of cheese brines with MD was found to be economically viable
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Recovery of brines from cheese making using membrane distillation at lab and pilot scale
1 2
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1,2
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Lies Eykens , Kristien De Sitter , Charlotte Boeckaert , Jan Boeckx , Guy Borgmans
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1
VITO - Flemish Institute for Technological Research, Boeretang 200, 2400 Mol, Belgium
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2
VLAKWA – Vlaams Kenniscentrum Water, Graaf Karel de Goedelaan 34, 8500 Kortrijk
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*Corresponding author – Email address: [email protected], [email protected]
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Abstract
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Brines are a problematic waste stream due to their high salinity. In cheese manufacturing,
9
a salinity up to 220 g/l is used for the pickling bath. This waste stream is currently
10
transported and treated externally at high costs. Membrane distillation is able to treat
11
much higher salinities compared to the traditional desalination techniques. In this article,
12
membrane distillation was therefore proposed as an alternative treatment technology,
13
simultaneously allowing the recovery of the salts and the production of pure water.
14
Microfiltration was successfully used as a pretreatment for the brines of two different
15
companies. The denaturation temperature of the proteins was found at 65.5°C and it was
16
shown that the temperatures in the MD process should be below this temperature.
17
Different membranes were explored at lab scale and module performance tests confirm
18
the feasibility of MD at pilot scale for this application. Additionally, it was shown that the
19
total water production costs are mainly affected by the availability of waste heat. Without
20
waste heat, the lowest cost of 35 €/m3 is achieved for the 24 m2 modules, with a total
21
thermal power demand of 34 kW. The costs for discharge of the brines are estimated at 50
22
– 150 €/m3, while the value of the salts recuperated equals 20 €/m3 distillate produced.
23
Therefore, even when there is a low availability of waste heat, application of membrane
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distillation for the recovery of brines from cheese making is an economically viable and
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technically feasible alternative for external treatment.
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Keywords: Economic evaluation, Fouling, Pretreatment, Technical feasibility
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1 Introduction
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During the cheese production process, cheese is pickled in a concentrated salt solution, to
30
improve the preservability, consistency and the flavor of the cheese. The cheese is
31
releasing water, while taking up salts, which decreases the salt concentration of the
32
pickling bath. To keep the salinity constant, additional salt is added to the pickling bath and
33
regularly, part of the brine is drained. This brine is a problematic waste stream within the
34
production process, containing a sodium chloride concentration of 200 to 220 g/l with a
35
complex composition, including organic material, calcium and proteins. Nowadays, these
36
brines are transported and treated externally, which is a non-sustainable way of coping
37
with these brines at a high cost of 50 – 150 €/ton [1].
38
Within the Pekelmem project (Milieu- en energietechnologie Innovatie Platform (MIP) no.
39
150362), MD was proposed as a new technology to treat the excess brine and reuse the
40
salts in the process. This would severely reduce the amount of salts required in the process
41
and the volumes of waste water to be transported. Within the project, the brines of two
42
different cheese types are treated [2]. For the first type of cheese the pickle bath is not
43
actively clarified and currently a substantial amount of the brine is discharged and salt is
44
added to maintain the salinity as visualised in Figure 1 A. Alternatively, process scheme B in
45
Figure 1 is proposed, where the brine is pretreated with microfiltration (MF) to protect the
46
MD modules and the permeate of the MF is sent to the MD. The volume of concentrate of
47
the MF depends on the recovery that can be achieved and is only a small fraction of the
48
original brine volume containing both organics and salts.
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Figure 1
For a second type of cheese, the pickling bath is currently continiously clarified with a
50
cartridge filter to actively remove suspended solids and excess organics. To maintain the
51
salinity, also a substantial volume of brine is discharged and salt is added (Figure 2 A).
52
Scheme B in Figure 2 proposes an alternative treatment, where microfiltration or
53
ultrafiltration replaces the cartridge filter to clarify the pickling bath. The resulting
54
permeate is a clarified salt solution containing small molecular weight organics. A part of
55
this volume is redirected towards the membrane distillation step for removal of excess
56
water and reuse of the concentrated salt. Since a large volume of the pickling bath is
57
continuously filtered, a relatively large volume of concentrate containing organics and salts
58
needs to be regenerated. The proposed regeneration makes use of diafiltration to remove
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60
contains organics, which can be treated by the standard biological waste water treatment
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facility of the company. The permeate of the diafiltration is a saline stream with lower
62
amount of salts, depending on the number of diafiltrate volumes. This stream is
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concentrated to the required salinity using membrane distillation. Figure 2
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To the best of our knowledge, the proposed treatment schemes are not tested before,
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although membrane distillation itself was already proposed in the dairy industry as an
66
alternative for classical evaporation for the production of milk powder and the
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concentration of whey proteins [3–5]. This was experimentally explored in the PhD of
68
Hausmann et al. [6–11]. Kezia et al. [12] focused on the concentration of salty whey, a
69
waste effluent from cheese making, containing up to 80 g/l NaCl as the major component.
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For all these applications, prevention and cleaning of protein fouling was reported as the
71
major challenge for the use of membrane distillation for these types of applications.
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Moreover, it was also found that milk fat partially wetted the membrane. Therefore, the
73
major challenge in this article was to evaluate the applicability of membrane distillation for
74
the recovery of brines of the cheese making. This is done via a stepwise characterization
75
procedure (see Figure 3), including the feed water characteristics and pretreatment choice,
76
the fouling behavior and feasibility at lab scale, the validity of these labtests at pilot scale
77
and an economic evaluation (Figure 3). In a later phase of the project, on-site testing will
78
be performed at pilot scale. At lab scale, DCMD was used to evaluate the fouling potential
79
of the pickle brine. The DCMD configuration is ideal for this purpose, due to its simplicity in
80
design and sufficiently high production capacity at small scale. At full scale however, AGMD
81
is preferred as stated in a recent publication [13]. The first reason is that DCMD has a
82
higher installation cost due to the additional heat exchanger that is required to recuperate
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heat. Secondly, DCMD has lower fluxes compared to AGMD and therefore requires more
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membrane surface for the same production capacity. Moreover, especially for longer
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modules, AGMD was also found to be more efficient compared to DCMD.
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2 Materials and methods
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2.1
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The samples of the brines obtained by company 1 are obtained from an actively clarified
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pickling bath (Figure 2). The samples of company two are obtained from a pickling bath
90
without active clarification (Figure 1). The composition of the samples as received is given
91
in Table 1.
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Company 1 4.88 225 11 117 342
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Table 1: Chemical composition of the cheese brines of the two cheese companies. Company 2 5.05 215 65.3 58 104 288
The protein denaturation temperature was determined visually after the filtration step
93
(discussed in section 2.2) on the permeate of company 1. In a first test, temperature was
94
raised every five minutes with 10°C under continuous stirring. In a second test, the
95
temperature is increased stepwise by 1°C every 10 minutes. The point where the sample
96
becomes turbid and flakes are observed is defined as the protein denaturation point.
97
2.2
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For the pretreatment, an in house built filtration setup was used (Figure 4). Two different
99
types of spiral wound membranes were explored; a BN PVDF 50kDa UF membrane and a
100
v0.1 PVDF MF membrane from Synder Filtration. The membranes had an area of 2.14 m2
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and a spacer thickness of 1.2 mm. The experiments were carried out in crossflow mode
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Lab scale MD setup
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The properties of the membrane selected for the lab scale tests are given in Table 2. This
105
membrane was produced by Lydall (Central Manchester, Connecticut, U.S.A) and selected
106
based on its availability in pilot modules.
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Table 2: Properties of the PE membrane [14,15].
Mean pore
Water contact
Hexadecane
SC
Thickness Porosity (µm)
(%)
size (µm)
angle (°)
contact angle (°)
99 ± 1
76 ± 1%
0.32 ± 0.02
120 ± 3
0
The feasibility of membrane distillation and the fouling potential of the brine were
108
evaluated at lab scale with a DCMD setup (Figure 5). The DCMD configuration is ideal for
109
this purpose, due to its simplicity in design and sufficiently high production capacity at
110
small scale. The flat-sheet module has a feed and permeate channel with a width of 6 cm, a
111
length of 18 cm and a height of 2 mm. The setup is described in more detail elsewhere
112
[15,16]. The electrical conductivity at the feed and permeate side were logged every 15
113
minutes by portable conductivity meters (WTW GmbH, pH/Cond 340i, Weilheim,
114
Germany). The feed and permeate temperatures were varied, but the temperature
115
difference across the membrane was kept constant at 15°C. The flow velocity was fixed at
116
0.07 m/s. The experiments were carried out in isoconcentration mode, meaning that pure
117
water was regularly added to the feed tank at the same rate as the permeation occurs, to
118
maintain a constant feed volume. Each flux measurement was performed for at least 30
119
minutes. The weight variations were automatically logged every 2 minutes and the
120
averages and standard deviations were calculated based on this weight variation, both on
121
feed and permeate side.
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Figure 5
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Fouled membranes have been analyzed with a Nexus FT-IR spectrometer from Thermo
123
Nicolet with Smart Orbit module in attenuated total reflectance (ATR) mode, controlled
124
with the OMNIC™ software. Samples were analyzed on the fouled side and compared with
125
a new sample. Field emission scanning electron microscopy and energy-dispersive X-ray
5
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spectroscopy were performed on the fouled membrane surface using a Nova Nanosem
127
(FEI).
128
2.4
129
The pilot MD setup was delivered by Aquastill BV, The Netherlands (Figure 6). A P&ID of
130
the pilot setup was descibed by Hitsov et al. [13] and can also be found in the appendix. A
131
standard AGMD module of Aquastill was used, with an area of 14.4 m2, 6 parallel channels,
132
a height of 0.4 m, a spacer thickness of 2 mm and the PE membrane. At full scale, AGMD is
133
preferred over DCMD as stated in a recent publication [13]. The first reason is that DCMD
134
has a higher installation cost due to the additional heat exchanger that is required to
135
recuperate heat. Secondly, DCMD has lower fluxes compared to AGMD and therefore
136
requires more membrane surface for the same production capacity. Moreover, especially
137
for longer modules, AGMD was also found to be more efficient compared to DCMD. A
138
volume of 200 l brine was added to the feed tank to start the experiment. After each day,
139
the volume in the feed tank was removed and fresh brine was added. A cooling
140
temperature of 10°C was aimed at, since it is equal to the pickling bath temperature.
141
Nevertheless, this temperature could not be achieved by the cooling circuit available and
142
therefore the coolant inlet temperature was fixed at 15°C to ensure a constant
143
temperature. The feed temperature was fixed at 50°C and a flow velocity of 0.05 m/s was
144
used. Each flux measurement was performed for at least 60 minutes after reaching stable
145
temperature values. All temperatures, pressures, flow rates and fluxes where
146
automatically logged every 2 minutes. The flux was calculated based on the height
147
variation of the feed and permeate tank. The average and standard deviation were
148
calculated based on these values over at least 60 minutes monitoring time. Figure 6
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Pilot scale MD setup
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2.5
Economic evaluation
150
The model developed by Hitsov et al. was used to calculate flux, energy consumption and
151
production capacity [13,17,18]. The evaluation of energy consumption is performed using
152
the gained output ratio (GOR), which is a dimensionless ratio used for thermal desalination
153
processes. It is defined as the ratio of the total latent heat of evaporation of the produced
154
water to the input thermal energy [19].
155
The model was calibrated and validated against experiments with synthetic NaCl solutions
156
with different salinities, temperatures and flow rates using the modules of Aquastill. As 6
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158
Only the cold inlet temperature was 10°C in the simulations, which is equal to the pickle
159
bath temperature. The simulated module sizes were 4.8 m2, 7.2 m2, 14.4 m2 and 24 m2.
160
Using these simulations, a prediction for the number of modules and thermal power
161
required to treat a certain volume can be estimated in the optimal case, i.e., without
162
severe flux reduction due to fouling. Based on this, a rough estimation of the cost for MD
163
for the different module dimensions was made to enable a choice of the optimal module
164
dimensions for pilot tests on location. This calculation includes the depreciation of the
165
installation within 10 years and module lifetime of 5 years [20]. During an energy audit
166
within the company it was found that waste heat is available at a price of 0.025 €/kWh and
167
sufficient free waste cooling is available within the factory. This estimation does not take
168
into account labor costs, pretreatment, cleaning and maintenance.
169
3 Results and discussion
170
3.1
171
Table 1 indicates the presence of suspended solids in the feed water, which are removed
172
during a first prefiltration step. Figure 7 shows the stable flux and transmembrane pressure
173
(TMP) during the MF prefiltration experiment for over 2 hours using the brine of
174
company 1. This is a strong indication of the feasibility of the prefiltration of the brine
175
before the MD. Similar results were observed for the UF filtration and for the MF filtration
176
of the brine of company 2. The average flux, TMP and recovery are reported in Table 3. The
177
recovery was not limited due to fouling, as the flux remains constant, but due to the
178
minimum volume required in the filtration setup. This indicates that in real application,
179
even higher recoveries can be achieved.
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Figure 7 Company 1 – UF -1
Company 1 - MF
Company 2 - MF
Flux (kg.m .h )
347 ± 1
353 ± 1
358 ± 1
TMP (bar)
0.88 ± 0.01
0.89 ± 0.01
0.99 ± 0.01
Recovery (%)
99.3
99.5
99.02
Table 3: Average flux, TMP and recoveries achieved during the prefiltration of the cheese brines using UF or MF.
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3.2
181
3.2.1 Scaling potential
182
To be able to evaluate the scaling potential of the MF/UF permeates in the MD, the exact
183
salt composition of the prefiltered samples was determined together with the total organic
184
carbon (TOC) present in solution (Table 4). Based on this information, the saturation index
185
of the salts present in the brines was calculated using the open source software PHREEQC.
186
This calculation only indicates saturation of hydroxyapatite (Ca5(PO4)3OH) for all three
187
samples (Figure 8). The composition of the deposits are further investigated during the
188
labtests are evaluated in section 3.3.3.
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Scaling/fouling potential of MD feed (MF permeate)
Conductivity (mS/cm) Total salinity (g/l) -
Cl (mg/l) 3
o-PO4 (mg/l) 2-
SO4 (mg/l) 2+
Ca (mg/l) 2+
Mg (mg/l)
Company 1 – UF
Company 1 - MF
Company 2 - MF
2500
2900
7700
206
199,7
209
180
197
211
109000
103000
113000
282
255
517
282
280
284
2200
2330
1360
78
77.2
137
68100
91300
95400
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+
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Table 4: Chemical composition of the UF and MF permeates (=MD feed)
3.2.2 Protein denaturation
191
Besides scaling, also protein fouling is seen as one of the major issues while treating dairy
192
effluents. As stated by Hausmann et al., the high hydrophobicity of MD membranes can
193
result in the establishment of hydrophobic interactions between the membrane and
194
proteins [9]. In order to diminish these interactions, feed temperature of the MD has to be
195
chosen wisely to avoid denaturation of the proteins present in the brines.
196
The protein denaturation temperature was determined visually on the MF filtrate from
197
company 1. In a first test, temperature was raised every five minutes by 10°C, showing that
198
the denaturation temperature is between 60 and 70°C (Figure 9 A). Secondly, the
199
temperature is increased stepwise by 1°C every 10 minutes. Figure 9 B shows that the
200
proteins denaturate at a temperature above 63.3°C. Above this temperature the solution
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becomes turbid and flocculation is observed. Similar results were observed for the two
202
other samples. Figure 9
3.3
Lab MD experiments
204
3.3.1 Maximum feed temperature
205
The effect of protein denaturation on the MD process was investigated by lab testing
206
below (45°C) and above the denaturation temperature (80°C). These tests were performed
207
on the UF permeate of company 1 (Table 4). Figure 10 A shows that the flux remains
208
constant at 4.5 kg.h-1.m-2 during 20 hours at a feed temperature of 45°C. At 80°C, a much
209
higher initial flux was achieved. However, due to deposits of the proteins on the
210
membrane, this flux quickly decreases to 3.5 kg.h-1.m-2 within the first 90 minutes, after
211
which it remains constant. The fast decrease is caused by fouling of the denaturated
212
proteins on the membrane. After this time period, all proteins precipitated and are
213
removed from the feed solution. Therefore, no further flux decrease is observed.
214
Figure 10 B shows a difference of 46% with the clean water flux on a clean membrane and
215
with the initial flux with the brine at 80°C. After 15 hours, flux decreased with 71%. In
216
contrast, at 45°C the decrease in flux from clean water to the brine is only 33%, whereas
217
after 15 hours a decrease of only 6% is observed. The deposits of the proteins of both tests
218
are visualized in Figure 11. A white deposit was observed on the heat exchanger after both
219
tests. This deposit was easily removed using a NaOH cleaning at pH 10. On the spacer, the
220
membrane and in the feed tank, white flakes were observed for the MD-tests at 80°C. At
221
45°C, no deposits were observed visually on the spacer and in the feed tank. Deposits were
222
still found on the heat exchanger and on the membrane, visually showing an increased
223
reflectance. These tests already indicate that even at 45°C, deposits are found on the heat
224
exchanger showing the importance of a thorough cleaning strategy. However, the salt
225
retention of 99.9% obtained at both temperatures during the entire testing time shows
226
that the deposited proteins do not wet the membrane.
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Figure 10 Figure 11
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Similar tests were performed using the MF-filtrate from company 1 and 2, with similar
228
observations. The flux decreases quickly when performing membrane distillation tests 9
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230
of company 2, the deposits of the proteins even formed an obstruction in the module inlet.
231
The feed liquid could not be circulated properly and the test was stopped after a few
232
hours.
233
3.3.2 Lab MD experiments on different cheese brines
234
Figure 12 A shows the flux as a function of time for the different samples using the PE
235
membrane. The concentration was kept constant during this test by regularly refilling with
236
water. This reduced the effect of the salinity on flux and enabled visualization of the effect
237
of fouling on the membrane flux. The largest decrease in flux was observed for the MF
238
sample from company 2 (31%). This decrease was less severe for the samples from
239
company 1. For the MF filtrate, no fouling was observed during the 25 hours MD-test.
240
Surprisingly the UF filtrate showed a stronger decrease of flux of 10%, despite the fact that
241
this pretreatment removes more organics from the sample. This might be caused by the
242
variability between the batch of MF and UF filtration.
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3.3.3 Fouling analysis
244
The difference in flux between the start and end of an isoconcentration experiment is a
245
first indication of fouling on the membrane. The strongest decline was observed for the
246
sample of company 2 after MF. Therefore, this sample was selected for further
247
investigation of the fouling. The clean water fluxes before and after the experiment is used
248
to evaluate the fouling. Figure 13 shows that the flux with brine at the start of the
249
experiment is already 32% lower compared to the original clean water flux. This lower flux
250
can have several causes:
251 252 253 254
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Figure 12
-
pressure and hence driving force. The activity of the different samples was
measured and equals 0.856. After calculation of the bulk driving force (Δp = pbrine,45°C –pw,30°C = awpw,45°C – pw,30°C), it appears that the bulk driving force
255 256
The presence of salinity results in a decrease of the activity of water, vapor
decreases with 26 % due to the presence of salts. -
Concentration polarization results in a higher concentration at the membrane
257
surface compared to the bulk solution and therefore negatively affects the driving
258
force and the flux. 10
ACCEPTED MANUSCRIPT 259
-
Liquid properties including density and viscosity are different compared to pure
260
water, affecting the hydrodynamics in the channels and therefore also temperature
261
and concentration polarization.
263
-
The startup, where feed and permeate are directed towards the aimed temperatures, takes about 1 hour. Fouling can already take place during this time.
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During the experiment the flux decreases further by 31% (5.5 to 3.8 kg.h-1.m-2). The
265
concentration remains constant throughout the experiment and the start and end-flux
266
were determined at equal concentration. Hence, the decrease in flux is mainly caused by
267
fouling. The measurement of the clean water flux after the measurement shows that the
268
original flux is completely restored after rinsing with pure water only. This indicates that
269
the fouling, if still present after rinsing, at least does not affect the flux. After the rinsing
270
step and clean water flux, deposits are still observed on the membrane (Figure 11 H). This
271
is an indication that long term operation requires a cleaning strategy to reduce the effect
272
of fouling on the process performance. The TOC retention is 95%, while the salt retention
273
remains above 99.97%, indicating that no wetting occurred during the test. The lower
274
retention of the TOC can be caused by the transport of volatile organics present in the MD
275
feed stream or by surface diffusion.
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Table 5: Membrane performance during an MD isoconcentration experiment. Process conditions: Tf = 45°C, Tp = 30°C, v = 0.07 m/s, membrane: PE1, sample: company 2 after MF. Average flux with -1 -2 brine (kg.h .m )
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4.9 ± 0.7
Flux decrease after 25 hours (%)
Salt retention (%)
TOC retention (%)
31%
99.97% ± 0.01%
95%
277
The fouling after the experiment and after the rinsing step with clean water was
278
characterized using the infrared spectra of a clean and a fouled PE membrane, given in
279
Figure 14. The C-H and CH2 peaks are visible for both spectra and are assigned to the
280
chemical structure of PE. The peaks that are only visible for the fouled membrane are
281
caused by the presence of proteins on the membrane surface. Moreover, the peaks at
282
1030 and 557 cm-1 are typical for inorganic phosphates (PO43-), suggesting the presence of
283
hydroxyapatite crystals, as predicted by PHREEQC (Figure 8).
11
ACCEPTED MANUSCRIPT Figure 14
The SEM/EDX analysis in Figure 15 again shows that after the isoconcentration experiment
285
and after rinsing, the deposits contain phosphor, calcium and oxygen. This is again a strong
286
indication for the presence of hydroxyapatite on the membrane. Figure 15
287
3.4
288
A 14.4 m2 AGMD module was used to perform a test at pilot scale. The experiment started
289
with 1 batch of 200 l prefiltered (MF) brine from company 2. After 24 hours, this batch was
290
replaced by a new batch of 200 l in order to better simulate the real case, where
291
continuously fresh feed was treated by the MD unit. Figure 16 A shows the flux and the
292
permeate conductivity as function of time for the three different batches. On average a
293
flux of 0.76 ± 0.04 kg.h-1.m-2 was achieved, while a simulation under equal conditions
294
predicted a flux of 0.84 kg.h-1.m-2. The clean water flux before and after the experiment
295
shown in Figure 16 B equals 1.69 ± 0.02 kg.h-1.m-2 and 1.61 ± 0.02 kg.h-1.m-2 respectively,
296
indicating that the foulants exposed to the membrane during 3 days do not severely affect
297
the flux. As the feed is first heated on the aluminum foil side of the module and in the
298
external heat exchanger, fouling is probably mainly located there. This effect was also
299
observed during the pilot scale experiments to a minor extent. The evaporator and
300
condenser inlet pressures using the brine are increased compared to the pure water flux
301
experiments. However, these pressures do not increase over time, indicating that no
302
severe blocking occurs due to fouling in the module on the condenser foil or the
303
membrane.
304
During the 3 batches, on average a salt retention of 99.8% was obtained. This value is
305
lower compared to the lab tests described in Section 3.3.2, where a salt retention of
306
99.97% was found for the PE1 membrane. As this salt retention is constant for all batches,
307
gradual wetting is not expected to cause the lower salt retention. This lower retention is a
308
known phenomenon and can be caused by defects in module sealing or the membrane,
309
which are more probable at larger surface area [21]. Similarly to the labscale results in
310
Section 3.3.2, the TOC retention was 95.6%, which is mainly caused by the volatile organic
311
compounds present in the feed stream that can easily migrate through the vapor space in
312
the membrane.
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3.5
Economic evaluation
314
3.5.1 Module performance simulation
315
A MATLAB model was used to simulate the pilot scale modules. The mass transfer through
316
the membrane was calculated using the Dusty Gas Model [22]. The heat and mass transfer
317
in the feed and permeate channels were calculated using Nusselt and Sherwood type
318
equations, which are calibrated and validated as described in [23]. The membrane
319
calibration was extensively discussed elsewhere [24]. More details on the construction of
320
the model and calibration are described elsewhere [13,17,18,24].
321
These simulations allow calculating flux, gained output ratio and production capacity of
322
one module for modules with different sizes. This enabled a quick estimation of the effect
323
of module size and a proper selection of a suitable module for module testing using the
324
brine. Figure 17 A shows higher fluxes for smaller modules. The calculation of the GOR in
325
Figure 17 B shows that small modules are less efficient. The internal energy recuperation
326
becomes more efficient for larger modules with a lower temperature difference. Therefore
327
a balance between flux and energy efficiency should be considered when selecting the
328
length of the modules. Figure 17 C indicates that the production capacity of one module
329
also increases with increasing module size, despite the decreasing fluxes.
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Figure 17
3.5.2 Economics of MD for brine treatment
331
Using the simulated fluxes and required thermal energy, the water production cost (€/m3)
332
was estimated. The number of modules was calculated based on the production capacity
333
(Figure 17 C), assuming a distillate production requirement of 650 m3/year. This is
334
calculated based on the current discharge volumes of the companies. The installation of
335
4 x 24 m2 AGMD modules would be sufficient to cover the required production capacity. 5
336
modules of the 7.2 or 14.4 m2 AGMD type or 6 modules of the 4.8 m2 would be required to
337
compensate for its lower production capacity per module.
338
An estimation of the costs for MD was made depending on the membrane area and the
339
availability of the waste heat. This calculation includes a depreciation of the installation of
340
10 years and a membrane lifetime of 5 years. The total water production costs are mainly
341
affected by the availability of waste heat (Figure 18). The lowest cost without waste heat
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ACCEPTED MANUSCRIPT of 35 €/m3 is achieved for the 24 m2 modules, with a total thermal power demand of
343
34 kW. As larger modules are more efficient (Figure 17 B), the energy costs decrease
344
substantially with increasing membrane area. For the case with 50% waste heat, the 24 m2
345
module again shows the best balance between energy and membrane costs. Four modules
346
and a thermal power of 34 kW are required to cover the production capacity at a cost of 24
347
€/m3 distillate. In the case 100% of the heat requirements are covered by waste heat, the
348
lowest cost is 12 €/m3 independent of the module size. The larger modules are more
349
productive so less, but slightly more expensive modules are required for the 24 m2
350
modules. The heat requirement increases from 34 up to 120 kWh with decreasing module
351
size. Similar calculations for DCMD were performed, where it was found that AGMD was 10
352
to 60% cheaper compared to the DCMD configuration, depending on the membrane area
353
installed in the modules.
354
The costs for discharge of the brines are estimated at 50 – 150 euro/m3 and may increase
355
due to the stricter discharge regulations in the future. Moreover, the salts in this
356
discharged brine have a value of 20 €/m3. The removal of 1 m3 pure water with MD, avoids
357
a discharge of an equal amount of pickle water. Even without the availability of waste heat
358
at a cost of 35 €/m3, the high discharge costs make the application of membrane
359
distillation for the recovery of brines from cheese making economically viable.
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Figure 18
4 Conclusions
361
Due to the high disposal costs, cheese factories are looking for a way of reusing their pickle
362
brines. The brine composition is out of the scope of classical desalination techniques due
363
to the high salinity and the presence of organic components, which denaturate above
364
60°C. Therefore, membrane distillation is an interesting alternative, avoiding the current
365
costly external treatment and discharge of the salts. This article shows the promising
366
potential and the technical and the economic viability of membrane distillation for the
367
treatment of brines in the cheese industry.
368
Microfiltration was found sufficient to clarify the brine before directing the feed stream
369
towards the MD unit. The labtests show that the choice of the process conditions is mainly
370
limited by the protein denaturation temperature (63.3°C). Above this temperature severe
371
fouling is observed, whereas a slightly declining flux was obtained below this temperature
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373
recovered to its initial value and basic and acid cleaning are proven to be suitable to clean
374
the heat exchangers and the membrane. The availability of waste heat is an important
375
parameter affecting the cost of MD. Without waste heat, the lowest cost of 35 €/m3 is
376
achieved for the 24 m2 module with a thermal energy demand of 34 kW. In the case of
377
100% availability of the waste heat, the lowest cost of 12 €/m3 is estimated for all modules,
378
with a heat requirement increasing with decreasing module size.
379
Acknowledgements
380
The authors thankfully acknowledge the funding provided by VLAIO (150362).
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APPENDIX
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Figure A.1: P&ID of the MD pilot in AGMD mode
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References
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K. De Sitter, MIP Pekelmem - Deliverable 1.1: Water- en energiebalansen, 2016.
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S.N. Moejes, A.J.B. van Boxtel, Energy saving potential of emerging technologies in milk powder production, Trends Food Sci. Technol. 60 (2016). doi:http://dx.doi.org/10.1016/j.tifs.2016.10.023.
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M. Tomaszewska, L. Białończyk, Influence of proteins content in the feed on the course of membrane distillation, Desalin. Water Treat. 51 (2013) 2362–2367. doi:10.1080/19443994.2012.728052.
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K. Christensen, R. Andresen, I. Tandskov, B. Norddahl, J.H. du Preez, Using direct contact membrane distillation for whey protein concentration, Desalination. 200 (2006) 523–525. doi:10.1016/j.desal.2006.03.421.
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A. Hausmann, P. Sanciolo, T. Vasiljevic, U. Kulozik, M. Duke, Performance assessment of membrane distillation for skim milk and whey processing, J. Dairy Sci. 97 (2014) 56–71. doi:10.3168/jds.2013-7044.
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A. Hausmann, P. Sanciolo, T. Vasiljevic, M. Weeks, K. Schroën, S. Gray, M. Duke, Fouling mechanisms of dairy streams during membrane distillation, J. Memb. Sci. 441 (2013) 102–111. doi:10.1016/j.memsci.2013.03.043.
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A. Hausmann, P. Sanciolo, T. Vasiljevic, M. Weeks, K. Schroën, S. Gray, M. Duke, Fouling of dairy components on hydrophobic polytetrafluoroethylene (PTFE) membranes for membrane distillation, J. Memb. Sci. 442 (2013) 149–159. doi:10.1016/j.memsci.2013.03.057.
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A. Hausmann, P. Sanciolo, T. Vasiljevic, M. Weeks, M. Duke, Integration of membrane distillation into heat paths of industrial processes, Chem. Eng. J. 211–212 (2012) 378–387. doi:10.1016/j.cej.2012.09.092.
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I. Hitsov, K. De Sitter, C. Dotremont, I. Nopens, Full-scale validated Air Gap Membrane Distillation ( AGMD ) model without calibration parameters, J. Memb. Sci. 533 (2017) 309–320. doi:10.1016/j.memsci.2017.04.002.
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membrane development, configuration assessment and applications, KU Leuven, 2017. [15]
L. Eykens, K. De Sitter, C. Dotremont, L. Pinoy, B. Van der Bruggen, K. De Sitter, C. Dotremont, L. Pinoy, B. Van Der Bruggen, Characterization and performance evaluation of commercially available hydrophobic membranes for direct contact membrane distillation, Desalination. 392 (2016) 63–73. doi:10.1016/j.desal.2016.04.006.
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L. Eykens, I. Hitsov, K. De Sitter, C. Dotremont, L. Pinoy, I. Nopens, B. Van der Bruggen, Influence of membrane thickness and process conditions on direct contact membrane distillation at different salinities, J. Memb. Sci. 498 (2016) 353–364. doi:10.1016/j.memsci.2015.07.037.
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[17]
I. Hitsov, PhD Dissertation: Model-based analysis and optimization of membrane distillation, Ghent University, 2017.
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I. Hitsov, L. Eykens, W. De Schepper, K. De Sitter, C. Dotremont, I. Nopens, Full-scale Direct Contact Membrane Distillation (DCMD) model including membrane compaction effects, J. Memb. Sci. 524 (2017) 245–256. doi:10.1016/j.memsci.2016.11.044.
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J. Koschikowski, M. Wieghaus, M. Rommel, V.S. Ortin, B.P. Suarez, J.R. Betancort Rodríguez, Experimental investigations on solar driven stand-alone membrane distillation systems for remote areas, Desalination. 248 (2009) 125–131. doi:10.1016/j.desal.2008.05.047.
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H.C. Duong, A.R. Chivas, B. Nelemans, M. Duke, S. Gray, T.Y. Cath, L.D. Nghiem, Treatment of RO brine from CSG produced water by spiral-wound air gap membrane distillation - A pilot study, Desalination. 366 (2015) 121–129. doi:10.1016/j.desal.2014.10.026.
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R.W. Field, H.Y. Wu, J.J. Wu, Multiscale Modeling of Membrane Distillation: Some Theoretical Considerations, Ind. Eng. Chem. Res. 52 (2013) 8822–8828. doi:10.1021/ie302363e.
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[24]
I. Hitsov, L. Eykens, K. De Sitter, C. Dotremont, L. Pinoy, B. Van der Bruggen, I. Nopens, Calibration and analysis of a direct contact membrane distillation model using Monte Carlo filtering, J. Memb. Sci. 515 (2016) 63–78. doi:10.1016/j.memsci.2016.05.041.
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Figure 1: Pickling bath without active clarification: Process scheme for (A) the current situation and (B) the alternative proposed in this study with brine recuperation.
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Figure 2: Pickling bath with active clarification: Process scheme for (A) the current situation and (B) the alternative proposed in this study with brine recuperation.
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Figure 3: Evaluation process of the applicability of MD.
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Figure 4: Filtration equipment and the process scheme of the filtration equipment.
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Figure 5: A) MD setup, B) Membrane module and spacer, C) Scheme of the MD setup.
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Figure 6: Pictures of the MD-pilot and the module in AMGD mode.
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Figure 7: Time course of the MF prefiltration experiment using the brine of company 1.
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Figure 8: Saturation index of salts present in the different samples.
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Figure 9: Protein denaturation tests at different temperatures A) test 1, B) test 2.
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Figure 10: A) MD flux as function of time of the UF-filtrate of company 1, B) clean water flux initial flux and the flux after 15 hours. ΔT=15°C, v=0.07 m/s, Membrane: PE1.
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Figure 11: Depositions at the end of the MD test at 80°C (A,B,C,D) and 45°C (E,F,G,H) on the heat exchanger (A,E), spacer (B,F), feed tank (C,G) and the membrane (D,H)
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Figure 12: A) flux as function of time for the 3 prefiltered samples and B) the initial flux using the pickle brine and the flux after 25 hours. Process conditions: Tf = 45°C, Tp = 30°C, v = 0.07 m/s, membrane: PE1.
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Figure 13: Clean water fluxes before and after an experiment and the start and end process flux with pickling water. Process conditions: Tf = 45°C, Tp = 30°C, v = 0.07 m/s, membrane: PE1
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Figure 14: Infrared spectra of a clean (blue, bottom) and fouled (red,up) PE1 membrane (MD process conditions: Tf = 45°C, Tp = 30°C, v = 0.07 m/s, membrane: PE1, sample: MF permeate company 2).
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Figure 15: SEM images of a clean (A) and fouled (B) PE-membrane. C) EDX spectrum of the fouled PE1 membrane.
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Figure 16: A) Flux and B) Evaporator and condenser inlet pressures and flow rate as function of time using the 7.2 m2 module from Aquastill. Process conditions: Te,in = 50 °C, Tc,in = 15°C, v = 900 l/h
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Figure 17: Simulated flux (A), GOR (B) and production capacity per module (C) for AGMD modules with varying membrane area. Process conditions: Tf = 50°C, Tp = 10°C, v = 900 l/h
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Figure 18: Treatment cost estimation in €/m3 destillate devided in installation costs, membrane costs and energy costs, without waste heat, B) with 50% waste heat and C) with 100% waste heat availability.