2008
Ind. Eng. Chem. Res. 2006, 45, 2008-2016
Prevention of Scaling of Reverse Osmosis Membranes by “Zeroing” the Elapsed Nucleation Time. Part I. Calcium Sulfate Natalie Pomerantz,† Yitzhak Ladizhansky,† Eli Korin,*,† Michael Waisman,‡ Naphtali Daltrophe,‡ and Jack Gilron*,‡ Department of Chemical Engineering and Department of Desalination and Water Treatment, Zuckerberg Institute for Water Research, Ben-Gurion UniVersity of the NegeV, P.O. Box 653, Beer-sheVa 84105, Israel
Precipitation of sparingly soluble salts is one of the main factors limiting the recovery in reverse osmosis (RO) of brackish water sources. Recoveries can be increased and antiscalant usage eliminated or reduced by applying flow reversal to RO process trains. Flow reversal works by changing the place of the entrance and exit of the pressurized feed before the induction time of the supersaturated solution along the membrane wall runs out and precipitation occurs. Reversing the flow before the induction time of the system is reached replaces the supersaturated brine at the exit with the unsaturated feed flow and thus “zeroes the elapsed nucleation time”, thereby resetting the induction clock. Laboratory experiments successfully demonstrated the technical feasibility of this concept by periodically exposing RO membranes to undersaturated solution after exposure to supersaturated calcium sulfate solution (bulk saturation index up to 3.0) formed from calcium chloride and sodium sulfate. The induction time to precipitation fouling without periodical switching of solutions was 150-270 min for stirred solutions. Periodic switching of solutions prevented precipitation for >480 min from stirred solutions. In follow-up experiments, a small scale pilot unit containing 2.5 in. diameter spiral RO elements was run continuously for 22 h with a calcium sulfate supersaturation index of 5.4 on the wall of the last element, when operated under reverse flow conditions. Without reverse flow the last RO element began to scale within 1 h. 1. Introduction 1.1. Background. There is increasing need for more water as populations grow, and water sources decline in quality and quantity. Desalination is an increasingly viable option for taking marginal resources, for example, brackish water and wastewater, and recovering them for agricultural, urban, and industrial use. At inland locations the application of desalination requires that adequate provision be made for disposal of concentrate. Glueckstern and Priel1 have pointed out that the costs of concentrate disposal can range from 5 to 30% of total desalinated water costs at inland locations. By increasing the fractional recovery of raw water as desalinated product water, significant reductions in product water cost can be achieved.2 As recovery of desalinated water from the feed stream is increased, the concentration of rejected ions in the concentrate becomes ever higher. In many conditions this concentration exceeds the solubility of sparingly soluble salts as expressed by saturation indices that are greater than 1. In addition, with increased recovery the bulk flow rate next to the membrane is also decreased. This leads to lower mass transfer and backmixing of rejected salts. As a result, even though the flux drops with increasing osmotic pressure of the more concentrated brine, concentration polarization will not drop as much, if at all. Thus, the saturation index at the membrane wall will be even higher than in the bulk concentrate solution. Membrane manufacturers usually specify minimum exit flow rates from spiral wound elements so that the concentration polarization (Cw/Cb) will not exceed 1.2.3 This requirement can often impose design restric* To whom correspondence should be addressed. E-mail:
[email protected] (J.G.),
[email protected] (E.K.). † Department of Chemical Engineering. ‡ Department of Desalination and Water Treatment, Zuckerberg Institute for Water Research.
tions that are difficult to deal with, especially for low-pressure reverse osmosis (RO) and NF membrane elements.4 To date, one or both of the following two strategies are used to prevent precipitation fouling and allow increased recovery:5 (1) Removal of scaling ions from the feed by base softening,6 nanofiltration softening,7 or ion exchange8 to remove hardness ions and acidification and air stripping to remove carbonate species is one strategy. A combination of these processes has recently been proposed to allow operation at high pH to avoid silica scaling.9 (2) Application of antiscalants10-12 which increase the induction time until precipitation (threshold inhibitors) or changing the crystal habits to prevent scale from adhering to the membrane (dispersants and crystal habit modifiers) is another strategy. While the first strategy is effective, the use of reagents to precipitate or remove scaling ions involves high chemical consumption and is often involved with large volumes of solids (lime softening sludges) or treatment brines (ion exchange). Indeed, a method has been proposed to soften water electrolytically13 which would eliminate the need for large amounts of chemicals. In the case of membrane softening by nanofiltration, the recovery limitations due to scaling may be transferred from the RO stage to the nanofiltration stage. As a result of the high chemical expense involved in chemical softening, dosing with antiscalants has become the method of choice for controlling precipitation fouling. Today’s antiscalants can allow a supersaturation index of 3 in calcium sulfate. However, even with antiscalants there is a limit to the supersaturation values that can be reached. What is more, precipitation of metal salts of the antiscalant and biofouling is often exacerbated by high levels of antiscalant because many are composed of polycarboxylic acids and phosphonate groups.10,14 The authors have recently proposed a method to allow increased recovery with reduced levels of antiscalant by using
10.1021/ie051040k CCC: $33.50 © 2006 American Chemical Society Published on Web 02/16/2006
Ind. Eng. Chem. Res., Vol. 45, No. 6, 2006 2009
reverse flow to periodically switch the entrance and concentrate exit ports on the membrane pressure vessel at times less than the induction time for the precipitating salt.15 While the use of reverse flow in pressure driven membrane processes has been reported since 198016-18 and proposed for RO in particular,18 there are no known published reports on its use in RO, and no one has previously linked reverse flow frequency to induction times. In this article we present the methodology and report its demonstration with calcium sulfate dihydrate as the sparingly soluble salt. 1.2. Basic Parameters and Theoretical Background. The solute concentration of the brine retentate stream exiting the membrane pressure vessel is given by
1 - Y(1 - Rnom) Cr ) Cf 1-Y 1 ≈ Cf when Rnom ≈ 1 1-Y
(1)
where Y is the fraction of feed stream that leaves the membrane device as permeate and is referred to as the membrane recovery. Cr is the bulk concentration of the rejected species exiting in the concentrate stream, and Rnom refers to the overall membrane rejection of the solute relative to the feed concentration:
Rnom ≡ 1 -
Cp Cf
(2)
Once Cr exceeds the solubility, the precipitation of sparingly soluble salts on the membrane surface can cause serious flux loss. The supersaturation for a sparingly soluble salt Mν+Xν- is usually expressed in terms of a supersaturation index which expresses the thermodynamic driving force for precipitation:19
( )
[C+]ν+ [C-]ν- γ+/SI ) KSP γ+/-,eq
ν
(3)
Assuming a completely rejected sparingly soluble salt and that the activity coefficient ratio is relatively constant, combining eqs 1 and 3 shows that the supersaturation index can increase very rapidly with increasing recovery:
SIr ) SIf
1 (1 - Y)(ν
(4)
++ν-)
where the subscripts r and f refer to the saturation index of the concentrate and feed streams, respectively. For the example of a 1:1 salt, 80% recovery implies that SIr increases 25-fold over the SIf. In actuality, the increase will not be this great because the solubility product increases with increased ionic strength of the concentrate solution. When sparingly soluble salts are supersaturated, they do not precipitate immediately as expected from the classical theory of homogeneous nucleation. It takes time for some of the clusters to grow to embryos and to exceed a critical size above which they will continue to grow instead of redissolving. This time is called the induction time, and it is related to the supersaturation index by the following equation:20
τ ) AC exp
(
)
16πσ3υ2 3kB3T3(ln SI)2
(5)
This can be simplified to relate the induction time to the saturation index, SI:
ln(τ) ) ln A +
B (ln SI)2
(6)
The SI relevant for the induction time is that at the membrane wall because the supersaturation is highest there due to concentration polarization. For a 1:1 salt like calcium sulfate, this will be given by the following equation:11-12
SIW ) SIB‚CP2 where
CP ≡
()
CW Jv ) exp CB kd
(7)
for a completely rejected ion. Jv is the volumetric flux, and kd is the mass transfer coefficient which can usually be taken as the effective mass transfer coefficient for the entire salt. One could argue that the supersaturated solution does not stay in the membrane module long enough to exceed the induction time, if one related to the average residence time of the fluid in the membrane module (feed channel volume/ volumetric flow rate). This ignores the fact that the feed side mass transfer boundary layer is buried inside the hydrodynamic boundary layer and the layer of most concentrated solution next to the membrane is actually moving at a very low velocity. Essentially, the antiscalants are designed to extend the induction time beyond the residence time of the slowest moving fraction of fluid in the membrane module.12 If the supersaturated solution next to the membrane were to be replaced by an undersaturated solution, the collection of crystal nuclei with a developed size distribution would be swept away, and the next time supersaturated solution were placed next to the membrane another complete induction time would be necessary to again develop the same size distribution of nuclei. By rinsing the membrane periodically with undersaturated solution, at a time less than the induction time, completion of the nucleation process and crystal growth is never allowed to occur. In principle this could be done by lowering the pressure in the RO unit and fast flushing the RO train with feed, but this has the disadvantage of stopping production during the flushing procedure. If, on the other hand, the feed was inserted into the concentrate port of the membrane train and the concentrate was allowed to exit from the former feed port, the same effect could be achieved without releasing pressure and stopping production. For this to be practical, the induction time at the supersaturation conditions next to the concentrate exit must be long enough that the frequency of the reverse flow is not too high. For extremely high supersaturations, this will require the addition of a certain amount of antiscalant to extend the induction time to allow flow reversal. In the following we demonstrate these principles of the reverse flow methodology for the calcium sulfate system. 2. Experimental Section 2.1. Lab System Demonstrating Effect of Flow Reversal on Induction Times. The laboratory system used to simulate the effect of switching the direction of flow in a RO desalination system is illustrated schematically in Figure 1, and a description listing the system components is found in Table 1. The system setup had two 30 L plastic tanks containing aqueous solutions. The first tank (T1) contained a supersaturated solution of calcium sulfate to simulate the RO concentrate steam, and the second tank (T2) contains an unsaturated solution to simulate the RO
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was calculated from the permeate flow rate according to the equation
Lp )
Qp/Am ∆P - ∆π
(8)
where the difference in osmotic pressure was determined from measuring the conductivities at the feed and permeate. While the temperature was maintained with a coiled stainless steel heat exchanger connected to a cooler, the temperature would vary (1 °C. The value of the specific flux was corrected for these slight variations and normalized to its value at 25 °C using the relation
Figure 1. Lab setup for measuring induction times for scaling and demonstrating effects of reverse flow on delaying the onset of scaling.
feed stream. The solution concentrations ranged from 30 to 45 mM CaSO4 in tank T1 and 7-13 mM CaSO4 in tank T2. In some cases antiscalants (3-12 mg/L of PermeaTreat 191) were added to extend the induction time of the supersaturated solution. The calcium sulfate was formed from equimolar solutions of Na2SO4 (Frutarom, Ltd., 99% pure) and CaCl2‚2H2O (Frutarom, Ltd., 97% pure). In-line filters (5 µm polypropylene) were installed to ensure that there would be no precipitation in the bulk flow to the membrane, and two pumps were used to apply the desired amount of pressure across the test cell membrane. Because bulk precipitation could commence after a certain time in the supersaturated solution, it was replaced every 2-4 h during switch experiments. The membrane test cell was fabricated from polycarbonate and closed between two 4 mm thick brass plates. The rectangular flow channel cross section was 0.038 m wide by 0.001 m thick. The flow channel length was 0.09 m long so that the active membrane area was 34.2 × 10-4 m2. The membrane used was taken from a 2.5 in. × 21 in. spiral of ESPA1 from Hydranautics, Inc. The flow rate through the flow cell was maintained at 100 L/h (0.73 m/s, Re ≈ 1600). Initial fluxes in the various experiments ranged from 12 to 30 L/(m2 h). To monitor the flux through the membrane, the permeate was collected on a scale (WIT) connected to a computer which recorded the scale readings versus time. The turbidity of the feed was measured to see if bulk precipitation occurs during the course of the experiment. Feed and permeate conductivity, temperature, and pressure were also monitored during the course of the experiment. The effective membrane permeability, Lp,
LP,25 ) LP,T1.035(25-T)
(9)
where T is the process fluid temperature in °C. Induction times for supersaturated solutions were determined by monitoring the specific flux over time and determining when the specific flux began to decline. Turbidity was used as a secondary indicator because its increase over time reflected bulk precipitation and not in situ precipitation on the membrane. In a “switch” experiment, the supersaturated solution is run through the membrane test cell for a period of time less than the induction period previously measured for a supersaturated solution of the same composition. The unsaturated solution is then run across the membrane for 10 min to allow enough time for the boundary layer to rebuild itself at a composition corresponding to unsaturated solution. This switching of solutions is performed several times until the experiment is over. The switching from the supersaturated solution (T1) to the undersaturated solution (T2) is effected by turning the open position on the bottom three-way valve depicted in Figure 1 from the left to the right position, followed by switching the upper three-way valve from the left to the right position after a 45 s delay. The delay is designed to clear the lines and system of the supersaturated solution before starting to recycle solution into tank T2. The procedure is exactly reversed (right to left) when switching back to the supersaturated solution (tank T1). The switching of the three-way valves as well as data acquisition and processing was carried out using a program written in the Labview (National Instruments) environment and by connecting to the experimental system through two USB to RS232 converters, one interfacing the system through Advantech, Ltd., ADAM A/D and D/A cards (I.C.P.C., Israel) and the other directly interfacing the permeate weight scale.
Table 1. Description of Components of the Laboratory Test System Depicted in Figure 1 symbol
device
manufacturer
CIT F-1 F-2 HPS HTS LPS P-1 P-2 PI PIT R T1 T2 TIT TMIT V-1 V-2 WIT
conductivity indicator and transmitter filter filter high pressure switch high temperature switch low pressure switch booster filter pump high pressure pump pressure indicator pressure indicator and transmitter rotometer tank 1 tank 2 temperature indicator and transmitter turbidity indicator and transmitter three-way valve three-way valve weight indicator and transmitter
El Hamma Amiad Aqua Tal Beta Beta Beta Touton Pumps, Ltd. Cat Pumps Mego Afek Michshur S. Z. Ltd. Tecfluid Technologic Hach Baccara Geva Baccara Geva Gibertini
range 0-20
precision (0.01
60 40 0.2 0-1.00 0-40 25-250 30 30
units mS bar °C bar
(0.01 (1 (5%
0.01-10000
(0.05 (0.03
bar bar L/h L L °C NTU
0.01-2000
(0.03
g
Ind. Eng. Chem. Res., Vol. 45, No. 6, 2006 2011
Figure 2. Pilot flow test system for demonstrating reverse flow in spiral wound RO elements.
2.2. Spiral Element Test Flow System. Figure 2 shows a portion of the RO test flow system (supplied by Tambour, Ltd., Akko) which can accommodate both 2.5 in. and 4 in. diameter spiral wound elements (PV). The high-pressure pump (HPP, Gould multistage centrifugal) feeds the process solution from the feed tank (FT) to a flow manifold containing two threeway valves (V1 and V2), one two-way valve (V4), and a backpressure valve (V3 not shown in figure) for developing/ maintaining pressure in the flow loop. The feed tank (FT) contained 100 L of process fluid, which was fed at a rate of 950-1120 L/h to the RO pressure vessels. When operating with six 2.5 in. diameter by 40 in. long (2540) spiral wound membrane elements in series, permeate recoveries were varied between 65 and 82% by varying the operating pressure between 15 and 21 bar. Both permeate and concentrate were returned to the feed tank in total internal recycle mode, so the temperature was maintained with an in-line heat exchanger (HE) supplied with chilled water. All process parameters (permeate and concentrate flowrate, feed and permeate conductivity, and feed pressure and axial pressure drop) were transmitted to a data logger with which the RO system was equipped, and the data was transferred to a computer for later processing. The unit was equipped with an in-line 5 µm cartridge filter (FH) to remove bulk solids that might otherwise serve as nuclei. The feed calcium sulfate solutions were prepared by mixing equimolar amounts of sodium sulfate and calcium chloride solutions. The feed solutions used were 10 mM in calcium and sulfate ion and 20 mM in chloride and sodium ion. Because the feed solution was always undersaturated, the experiment could be conducted indefinitely without danger of bulk precipitation. Highly supersaturated conditions were only obtained in the last pressure vessel in the series. The extent of supersaturation depended on the extent of recovery. To determine the supersaturation at the membrane wall at the concentrate exit from the last pressure vessel, the flow conditions and flux values were fed into the membrane process design program provided
Figure 3. Flow scheme for reversing flow through a train of three pressure vessels in series.
Hydranautics, Inc. (IMSdesign), which calculates the concentration polarization at the membrane wall at the end of each membrane element in series. Analyses for the calcium ion were carried out on samples from the feed, concentrate, and permeate streams by colorimetric titration. Overall salt rejection was determined using the conductivity values of the feed, concentrate, and permeate streams. In reverse flow experiments with supersaturated solutions, the system was loaded with six spiral wound LPRO elements (ESPA1 2540, Hydranautics, Inc.) divided between three twoelement pressure vessels arranged in series (PV1, PV2, and PV3 in Figure 3). As shown in Figure 3, the feed flow is initially delivered to the left end of pressure vessel PV1, and the concentrate flow exits from the right end of the pressure vessel PV3. Under these conditions, valve V-4 is closed and valves V-1 and V-2 are positioned toward the up position. In switching to reverse flow (Q1 enters PV3 from the right and Q4 exits the PV1 from the left), the following sequence of actions is taken: valve V-4 is opened followed by moving valve V-2 to the down
2012
Ind. Eng. Chem. Res., Vol. 45, No. 6, 2006 Table 2. Induction Times Measured in RO Lab Experiments with CaSO4 (from CaCl2 and Na2SO4) at a 100 L/h Recycle Ratea induction time experiments bulk supersaturation ratio
pressure (bar)
4.37 3.07 3.07 3.07 3.07 3.07 1.86 1.86
10 10 10 10 10 10 10 15
“switch” experiments
Q (L/h)
antiscalant (ppm)
τ (h)
100 100 100 100 50 100 100 100
12 0 3 9 9 12 0 0
0.7 ∼0 1 2.4 1.3 9.6 2.5 2
run time for supersaturated soln between switches (h)
total time of experiment (h)
NA
NA
0.5 1 0.75 NA 1 NA
5.92 7.50 7.67 NA 8.00 NA
a In the switch experiments, the undersaturated solution was 0.00750.010 M CaSO4.
Figure 4. Induction time experiment in laboratory the RO setup: 30 mM CaSO4 and 60 mM NaCl formed from calcium choride and sodium sulfate. The feed tank is unstirred.
position followed by moving valve V-1 to the down position. This is done in this sequence to prevent causing a “water hammer” effect which could damage the membrane element if the flow direction were reversed too suddenly. The unit is returned to standard the flow direction by conducting the sequence in exact reverse of the first sequence. In experiments in which direction switching was performed, the flow reversal took place every 0.5-1 h depending on the recovery to be achieved in the experiment. To check the effect of repeated switching of the system on the element integrity, a single element was installed in a one-element pressure vessel (the thinner PV shown in Figure 2). The flow direction was switched every 10 min for 10 days. There was no drop in the rejection (>99.5%) or change in the membrane flux during this time. For a reverse flow operation with switching every 2 h, this would be the equivalent of 4 months of operation. The permeate flow rate from pressure vessel PV3 was monitored throughout the course of each experiment. In experiments designed to determine induction times, the flow reversal was not performed at all. Induction times were then determined by examining changes in the PV3 permeate flow rate, Qp3, and the level of calcium in the feed tank. 3. Results and Discussion 3.1. Laboratory System. A sample run with induction time is given in Figure 4. The results are reported as specific flux, Lp (flux/net driving pressure), versus time. As can be seen from the graph, there is a sudden drop at 100 min due to scaling of dead regions in the test cell between the walls of the flow channel and the high-pressure O-ring seal. Afterward, the flux remains constant until 170 min have passed, when it starts to drop steadily. So 170 min would be taken as the induction time for this run. Induction time measurements with 30 mM calcium sulfate formed from calcium chloride and sodium sulfate (supersaturation index ∼ 1.8) gave 2.9 h when operated at 10 bar and 3.3 h when operated at 5 bar. This compares with an induction time of ∼2.5 h in the work of Hasson et al.11 [4] operating in a similar RO experiment with a bulk supersaturation index of 1.9. In these experiments, the turbidity of the feed never exceeded 1 NTU, and the conductivity was constant. This shows that most of the nucleation was heterogeneous and on the membrane. These induction times are significantly shorter than the bulk induction time (>10 days)21 and show that nucleation in the RO experiment is due to heterogeneous nucleation on the membrane and not homogeneous bulk nucleation.
Figure 5. Switching experiment with 30 mM CaSO4 and 60 mM NaCl formed from calcium chloride and sodium sulfate. The undersaturated solution was 7.5 mM CaSO4 and 15 mM NaCl. The switching points can be seen as periodically scattered points on the horizontal data line. The applied pressure was 10 bar, and the recycle flow rate was 100 L/h.
Table 2 summarizes the induction times obtained with different operating conditions. The decrease in induction time with increased pressure arises from the increased flux giving higher concentration polarization and higher supersaturation ratios at the membrane wall than in the bulk solution confirming the trends found by Hasson et al.11 As can be seen in the left block of the table, the induction time drops with supersaturation, and above a supersaturation ratio of 3, it becomes almost instantaneous. Only by adding antiscalant was it possible to extend the induction time. Figure 5 displays a trace of a switch experiment in which the supersaturated solution (bulk supersaturation ratio ) 1.8) is replaced each hour for 10 min with a solution of 7.5 mM calcium sulfate solution and 15 mM NaCl solution. As can be seen the specific flux is maintained for ∼480 min until the end of the experiment. This is almost three times the normal induction time found without flow reversal. If there was no “zeroing” of the induction clock then one would expect the induction time to occur around 170 min of accumulated run time on the supersaturated solution. The switch experiments are summarized in the right block of Table 2. As can be seen, by briefly switching to undersaturated solution, the system could be run for hours without flux decline even when the induction time in the parallel nonswitched experiment was only ∼80 min. Because no flux decline was observed between switches, it is more reasonable to attribute the effect to zeroing the induction clock rather than dissolving any pre-existing scale even though the undersaturated solutions had saturation indices of 0.320.45. Essentially the undersaturated solution sweeps away any
Ind. Eng. Chem. Res., Vol. 45, No. 6, 2006 2013 Table 3. Operating Conditions of Pilot Experiments To Measure Induction Times with an Initial Feed Composition of 10 mM CaSO4 and 20 mM NaCl flow rate time in expt
recovery (%)
feed (L/h)
permeate (L/h)
concentrate (L/h)
pressure (bar)
average flux [L/(m2 h)]
polarization at concentrate exit, Cw/Cb
start finish start finish
65.8 70.8 82.2 79.8
961 1118 1061 1090
632 792 872 870
329 326 189 220
15.9 20.3 20.2 21
40.5 50.8 55.9 55.8
1.23 1.28 1.47 1.43
of the pre-existing supersaturated solution with its distribution of nuclei. Once the supersaturated solution returns to the membrane wall, the nuclei distribution must develop again resulting in another induction time being required before any precipitation can begin. A further demonstration of this affect was made when in one experiment a solution of 0.02 M CaSO4 was used as the undersaturated solution. This solution is nearly at saturation. When the switch was carried out twice after each of the first 2 h of running (140 total min) and then discontinued, the overall induction time was 410 min (390 min net run time) as opposed to 170 min without any switching at all. 3.2. Pilot Test Unit. 3.2.1. Induction Times. Table 3 summarizes the operating conditions for the induction time experiments. Results from induction time experiments in the pilot unit are shown in Figures 6-9. As can be seen for Figures 6 and 7, the onset of precipitation can be distinguished by the drop in all three parameters: effective membrane permeability (Lp) of the elements in the third pressure vessel, calcium content in the concentrate, and calcium content in the feed tank. This is summarized in Table 4. According to Figure 6, which shows the experiment with 6572% recovery, the most sensitive indicators were change in calcium content in the concentrate stream (Figure 6a) and the
effective membrane permeability, Lp (Figure 6b). The calcium content of the concentrate had the greatest change (from 25 to 18 mM), and the Lp dropped by almost the same factor (from 3.9 to 2.8 L/(m2 h bar)). The steady-state Lp value first begins to drop as early as 17 h into the experiment, and it begins an uninterrupted decline from 20 h onward. Therefore, one can take the induction time as occurring sometime within that span. One can also see in Table 4 that, at the time of the greatest drop in Lp in both experiments (starred entry for each experiment), the ratio of the calcium value in the concentrate to its value in the feed, CFexpt, is less than the volume concentration factor (VCF). This can be explained by there being a significant rate of precipitation at this point in time so that not all of the feed calcium reaches the concentrate exit. On examining Table 4 and Figure 6, one can see that the steady-state feed concentration of calcium sulfate for 65-72%
Figure 7. Pilot unit with six elements in series, 82% recovery, and 10 mM CaSO4 and 20 mM NaCl in the feed tank. Shown are the calcium concentration in concentrate (2) and effective membrane permeability, Lp, in the last pressure vessel PV3 (9).
Figure 6. Pilot run on 10 mM CaSO4 and 20 mM NaCl, with 65-72% recovery, no switching, and an average flux of 40.5 L/(m2 h). (a) Calcium concentrations of concentrate (2) and feed (9) solutions. (b) Lp of the elements in last pressure vessel.
Figure 8. Online record from the pilot experiment at 82% recovery with a 10 mM CaSO4 feed: permeate flow rate from pressure vessel PV3 (Qp, left axis, [) and applied pressure in bar (right axis, 9).
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Ind. Eng. Chem. Res., Vol. 45, No. 6, 2006
Figure 9. Reverse flow experiment at 82% recovery, with a 10 mM CaSO4 and 20 mM NaCl feed, applied pressure of 21.5 bar, and feed flow rate of 1100 L/h. Table 4. Summary of Results of Induction Time Experiments on Pilot Unit, 10 mM CaSO4 and 20 mM NaCl in the Original Feed [Ca] (mM) [L/(m2
recovery (%)
expt duration (h)
time
feed
concentrate
bulk supersaturation
VCFcalc
CFexpt
τind (h)
65-72
45
start end
4.09 2.93
9.1 6.8
25.4 18.2
1.8
2.9 3.4a
2.79 2.68a
17
82
24
start end
4.08 2.85
7.2 4.4
37.2 20.2
2.57
5.6 5.4a
5.24 4.9a
1
a
Lp
h bar)]
This was calculated at the point of maximum flux drop.
recovery before onset of precipitation was only 9 mM. This is the case even though the initial feed composition contained 10 mM calcium sulfate. The discrepancy is caused by the relatively low total volume of feed (100 L) and the relatively high holdup volume in feed channels of the membrane elements and exit line of the pressure vessels (10-20 L). The average calcium concentration in the feed channel of the elements and the pressure vessel exit line is higher than in the feed tank because of the permeate recovery. This holdup volume thus has a higher proportion of the total calcium in the system, and the steadystate concentration in the feed tank must be lower than the initial value by mass balance. Clearly, as the volume recovery increases, the calcium concentration in each membrane element increases, and a higher proportion of the total calcium of the system is in the element and pressure vessel holdup volume leaving less in the feed tank. This was indeed seen when the recovery was increased to over 80% as shown in the bottom block of Table 4. Here the steady feed calcium concentration was initially 7.2 mM. At 82% recovery, the onset of precipitation appears to take place within 1 h, as can be seen by the change in concentrate calcium concentration in Figure 7. (The initial concentrate calcium value of 39 mM was measured before attaining the steady-state distribution of the salts between feed and concentrate as discussed above.) According to Figure 7, the effective membrane permeability of the third pressure vessel (Lp) appears to drop almost immediately. This is seen even more clearly in Figure 8 in the time trace from the online flow transducer monitoring permeate flow from the third pressure vessel. The momentary stabilization of the permeate flow around 4 h of elapsed time occurred when the applied pressure increased from
20 to 20.8 bar. The effective membrane permeability was actually dropping even then. It should be noted that the induction times for the pilot system were significantly longer than for the laboratory system at comparable bulk supersaturations. This may reflect that the effective mass transfer in the lab system is actually poorer because the test cell has a larger portion of the membrane close to the wall regions where the rate of fluid flow next to the membrane is much slower. The feed spacer in the spiral element may do a better job of mixing the boundary layer with the bulk fluid and thus give better mass transfer and shorter effective residence time distribution than the flat laminar test cell. Alternatively, it may be the case that, despite the presence of an in-line 5 µm filter in the lab system, there are nanocrystallites in the supersaturated bulk solution that do reach the membrane and effectively cut short the induction time. In the pilot flow system no such crystallites can form in the feed solution because it is always undersaturated. 3.2.2. Effect of Reverse Flow. On the basis of the baseline experiment at 82% recovery with 10 mM CaSO4 and 20 mM NaCl original feed solution, a reverse flow experiment was conducted. It should be noted that the feed was approximately 1000 L/h, so the concentrate was only 180 L/h. This is significantly below the manufacturers recommended minimum flow rate (300-350 L/h) for 2.5 in. spiral elements. Also the flux was very highs40-50 L/(m2 h)seven in the last pressure vessel. As a result the concentration polarization on the membrane element at the pressure vessel exit was very highs Cw/Cb ∼ 1.45. As a result the supersaturation index at the membrane wall was 2.57 × 1.45 ) 5.4! Because the induction
Ind. Eng. Chem. Res., Vol. 45, No. 6, 2006 2015 Table 5. Summary of Reverse Flow Experiments at Two Different Recoveries, 10 mM CaSO4 and 20 mM NaCl in the Original Feed [Ca] (mM) recovery (%) 71-85 80-82
reversal frequency (h-1)
elapsed time (h)
Lp [L/(m2 h bar)]
feed
concentrate
VCFcalc
1 1 2 2
0 24 0 22
3.76 3.62 3.87 3.77
8.4 7.8 7.5 7.2
26.2 32.0 34.2 34.2
3.6 3.5 4.69 4.94
time in the baseline experiment was between 0.5 and 1 h, the flow was reversed every 1/2 h. The data logger printout results of permeate flow from the third (last) pressure vessel are given in Figure 9. One can clearly see the changes in feed flow direction from the graph. When the third pressure vessel was last, it gave the lower levels of permeate flow, and when it was first during reverse flow, it gave the higher levels of permeate flow. This followed directly from the fact that when the third pressure vessel was last it saw the lowest net driving pressure as both the osmotic pressure of the solution was higher and the applied pressure was lowest after the axial pressure drop. When it was first it was exposed to the highest net driving pressure. As can also be seen in Figure 9, the initial permeate flow from PV3 is higher than its steadystate value (∼250 LPM) when the PV3 switches from being first to last in the treatment train. This is to be expected because initially the boundary layer osmotic pressure still corresponds to the feed solution and not to the concentrate. Similarly the initial permeate flow from PV3 is lower than its steady-state value (∼340 LPM) when PV3 is switched from being last to first in the treatment train. As the mixing dynamics are the same for both switches, the transients in the first and last PVs will cancel each other out and not affect the overall output. Most importantly, it can clearly be seen that the permeate flow is stable both when the third pressure vessel is first and when it is last. This demonstrates the use of reverse flow to stabilize flux even when the conditions are extremely supersaturated and the unit is operating outside the manufacturer’s recommended limits. Additional proof for the stability obtained with reverse flow is found in comparing Table 5 (reverse flow summary) to Table 4. Unlike the base case, the calcium level of both the feed and the concentrate are stable over the length of the experiment within 10% of the initial steady-state value. In addition, the actual concentration factor (CFexpt ≡ Cr/Cf) does not decrease over the course of the experiment relative to the volume concentration factor. This shows that almost no calcium sulfate precipitated. In practical application, the benefits of flow reversal must be weighed against the possible costs. Aside from the modest outlay in additional equipment and controls to effect the reversal, one might consider potential loss of production during the switching time. However, the only loss due to flow reversal is during the very short time that the average linear feed flow velocity next to the membrane is somewhat lower during the switching procedure. During that time the osmotic pressure will increase as a result of a momentary increase in the concentration polarization. This effect of the reduction in production rate during the transient time of reversing the flow on the total production rate depends on the ratio between the period that the system operates at steady state and the transient switch period. The steady-state period is determined by nucleation induction time, which will depend on the feed composition and other operating conditions and which could be increased by using a relatively small amount (compared to conventional doses) of antiscalant. We consider the minimum practical steadystate period to be no less than 0.5 h. Considering that switching
CFexpt 3.13 4.1 4.56 4.75
τind (h) not reached not reached
takes no more than 1 min and giving a very conservative flux estimate of 50% of the steady-state value during the switching procedure, we see that the maximum water loss would be only 1.6% of the overall production rate. If the value of the water is $0.40 (USD)/m3, then this results in a loss of 0.64 cents/m3 of product water. On the other hand if this allows recovery to be increased from 75 to 85% and the brine disposal costs are $0.50/ m3 brine, then the savings in brine disposal costs will be 0.5 × [(25/75) - (15/85)] or 7.8 cents/m3 of product water. This simple example gives a feel for the potential relative benefits and costs of the technique. 4. Conclusions The concept to exploit flow reversal to prevent scaling of RO membranes was presented. It was hypothesized that this is achieved by removing supersaturated solutions from the vicinity of the RO membrane before heterogeneous nucleation induction times were reached. The use of undersaturated solutions to “zero the induction clock” was demonstrated for calcium sulfate with a laboratory simulation system. For a saturation index of 1.86 the induction time was ∼170 min. By periodically switching the membrane to the undersaturated solution once an hour for 10 min, we prevented scaling for 480 min of continuous running. The feasibility shown in the laboratory was confirmed in actual reverse flow pilot experiments. In the pilot system with six spiral wound elements in series, the calcium sulfate solution reached a saturation index of 5.4 at the wall based on mass balance and mass transfer correlations, and the flux dropped after 0.5-1 h. By switching the direction every 0.5 h, scaling was not observed even after 22 h. These results indicate that this method has the potential to increase recovery in desalination of feedwaters with high concentrations of sparingly soluble minerals. Ongoing work with calcium carbonate and silica shows that the concept is applicable in these systems as well. This will be reported in future publications. Acknowledgment Hanan Hayout installed and programmed the data acquisition and process control for the laboratory system. The Israel Ministry of Trade and Industry and Tambour Ecology, Ltd., provided financial support. Notation AC ) pre-exponential coefficient in eq 5, s A, B ) coefficients defined in eq 6 Am ) membrane area, m2 C ) concentration, mol/L CP ) concentration polarization modulus Jv ) permeate flux, L/(m2 h) kB ) Boltzman constant, J/(molecule K) kd ) mass transfer coefficient, m/s or L/(m2 h) Ksp ) solubility product
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Ind. Eng. Chem. Res., Vol. 45, No. 6, 2006
Lp ) effective membrane permeability, L/(m2 h bar) P ) pressure, bar Q ) volumetric flow rate, L/h Rnom ) membrane rejection SI ) supersaturation index T ) temperature, K Y ) recovery ratio Greek Symbols γ ) activity coefficient ν ) molecular volume, m3/molecule ν ) stoichiometric coefficient π ) osmotic pressure, bar σ ) surface energy, J/m2 τ ) induction period, s or h Subscripts b ) of the bulk eq ) at equilibrium f ) feed p ) permeate r ) retentate w ) on the membrane wall + ) of the anion - ) of the cation ( ) average of anion and cation Literature Cited (1) Glueckstern, P.; Priel, M. Optimized brackish water desalination plant with minimal impact on the environment. Desalination 1996, 108, 19. (2) Glueckstern P.; Priel, M. Upgrading of old brackish water reverse osmosis (BWRO) desalination plants. Proceedings of 6th Annual Conference of the Israel Desalination Society, Beer Sheva, Dec 2003; p 59. (3) IMSdesign Design Guidelines. Hydranautics, Inc., 2003. (4) Nemeth, J. Innovative system designs to optimize performance of ultra-low pressure reverse osmosis membranes. Desalination 1998, 118, 63. (5) Taylor, J. In Water Treatment Membrane Processes; Mallevialle, J., Odendaal, P., Wiesner, M., Eds.; McGraw-Hill: New York, 1996; Section 9.4. (6) Gilron, J.; Daltrophe, N.; Waissman, M.; Oren, Y. Comparison between compact accelerated precipitation softening (CAPS) and conven-
tional pretreatment in operation of brackish water reverse osmosis (BWRO). Ind. Eng. Chem. Res. 2005, 44, 5465. (7) Hassan, A. M.; Farooque, A. M.; Jamaluddin, A. T. M.; A1-Amoudi, A. S.; A1-Sofi, M. A. K.; A1-Rubaian, A. F.; Kither, N. M.; A1-Tisan, I. A. R.; Rowaili, A. A demonstration plant based on the new NF-SWRO process. Desalination 2000, 131, 157. (8) Van Hoek, C.; Kaakinen, J. W. Ion exchange pre-treatment using desalting plant concentrate for regeneration. Desalination 1976, 10, 471. (9) Mukhopadhyay, D. U.S. Patent 5,925,255, July 20, 1999. (10) Gill, J. A novel inhibitor for scale control in water desalination. Desalination 1999, 124, 43. (11) Hasson, D.; Drak, A.; Semiat, R. Inception of CaSO4 scaling on RO membranes at various water recovery levels. Desalination 2001, 139, 73. (12) Hasson, D.; Drak, A.; Semiat, R. Induction times induced in an RO system by antiscalants delaying CaSO4, precipitation. Desalination 2003, 157, 193. (13) Rinat, J.; Korin, E.; Soifer, L.; Bettelheim, A. Electrocrystallization of calcium carbonate on carbon-based electrodes. J. Electroanal. Chem. 2005, 575, 195. (14) Butt, F. H; Rahman, F.; Baduruthamal, U. Identification of scale deposits through membrane autopsy. Desalination 1995, 101, 219. (15) Gilron, J.; Korin, E. Method and system for increasing recovery and preventing precipitation fouling in pressure-driven membrane processes; No. PCT/IL2004/001110; December 7, 2004. (16) Breslau, B. R.; Testa, A. J.; Milnes, B. A.; Medjanis, G. Advances in Hollow Fiber Ultrafiltration Technology. In Ultrafiltration Membranes and Applications, Polymer Science and Technology; Cooper, A., Ed.; Plenum Press: New York, 1980; Vol. 13, p 109. (17) Hargrove, S. C.; Ilias, S. Flux Enhancement Using Flow Reversal in Ultrafiltration. Sep. Sci. Technol. 1999, 34, 1319. (18) Lauer, G. Conditioning process and device for producing pure water. U.S. Patent 5,690,829, 1997. (19) Meijer, J. A. M.; van Rosemalen, G. M. Solubilities and supersaturations of calcium sulfate and its hydrates in seawater. Desalination 1984, 51, 255-305. (20) Mullin, J. W. Crystallization; Butterworth-Heinemann: Oxford, 2001. (21) Korngold, E. Optimization of membrane methods for treating wastewater; Annual Report 2003, Contract No. 01-01-01486; Israel Ministry of Science.
ReceiVed for reView September 15, 2005 ReVised manuscript receiVed December 16, 2005 Accepted December 28, 2005 IE051040K
