SEPARATIONS

Separation of Ammonia/Water/Sodium Hydroxide Mixtures Using ReverseOsmosis Membranes for Low Temperature Driven Absorption Chillers

Simona Steiu,† Joan Carles Bruno,*,† Alberto Coronas,† Ma Fresnedo San Roman,‡ andInmaculada Ortiz‡

UniVersitat RoVira i Virgili, CREVER-Grup d’Enginyeria Termica Aplicada, AV. Paısos Catalans,26, 43007 Tarragona, Spain, and UniVersidad de Cantabria, ETSIIyT, Dpto de Ingenierıa Quımica y QI,AV. Castros s/n, 39005 Santander, Spain

The conventional working fluids used in absorption chillers (water/lithium bromide and ammonia/water) presentseveral disadvantages that limit their effective application. Recently, some works have reported the additionof NaOH to the ammonia/water working pair to improve the separation of ammonia in the generator, reducingthe chiller driving temperature by taking advantage of the common ion effect. However, the presence ofNaOH in the absorber has a negative impact on the absorption process. This study analyzes the technicalviability of separating NaOH from ammonia/water/NaOH mixtures by using reverse osmosis membranes toincorporate this separation method into future chiller designs that work with these mixtures. The concentrationrange analyzed covers the solution concentration values of interest for absorption chiller applications(approximate 0.02-0.05 mass fraction of NaOH and 0.3 mass fraction of NH3). The results obtained showthat, by using an in-series configuration of the modules, reverse osmosis technology is suitable for separatingNaOH from the ternary mixtures studied.

1. Introduction and Objectives

In recent years, the demand for refrigeration in buildings hasincreased significantly mainly due to higher standards of indoorcomfort conditions. This demand for refrigeration is usuallycovered by the electricity supply available, which produces ahigh peak load in the electricity distribution grid during thesummer. To reduce this demand, absorption chillers activatedwith waste heat or thermal solar energy are an interesting optionfor the near future. Conventional working fluids used inabsorption chillers (water/LiBr and ammonia/water) have somedrawbacks that limit the number of applications for absorptionchillers: for example, limited solubility and corrosion problemsin water/LiBr chillers and the need for rectification in the caseof the mixture ammonia/water. Moreover, the relatively highdriving temperatures for ammonia/water chillers are a problemfor the application of these type of chillers in a field that isattracting great interest, such as the solar assisted air-condition-ing of buildings.1 In this case, the possibility of using ammonia/water chillers to dissipate the absorption and condensation heatin the ambient air offers advantages in comparison with water/LiBr chillers. The addition of other substances, such ashydroxides, can help to reduce these problems and improve theabsorption cycle performance.2-5 In particular, the addition ofNaOH (eq 1) to the ammonia/water mixture causes a shift inthe solution equilibrium toward the ammonia gas form throughthe common ion effect, favoring a more effective separation ofammonia in the chiller generator with a lower energy require-ment.

NaOHfNa++OH- (1)

NH3 +H2Orf

NH4++OH- (2)

Existing concentration measurements of the vapor-liquidequilibrium of the mixtures ammonia/water/NaOH and am-monia/water/KOH confirm that the ammonia concentration inthe liquid phase decreases significantly as the addition of sodiumor potassium hydroxide increases.6-8 As is expected andconfirmed, comparing the equilibrium data from both ternarymixtures, the reduction of ammonia in the liquid phase is moresignificant in the case of sodium hydroxide so it has beendecided to conduct the study using this hydroxide. Other possiblefactors such as the corrosion potential have not been consideredat this stage.

The main benefit from the addition of NaOH to the workingfluid is the reduction in the chiller driving temperature andconsequently an increase in the cycle COP (ratio between thechiller refrigeration capacity and the required heat supplied).However, the presence of NaOH in the absorber produces anegative effect because it reduces the ammonia absorptioncapacity of the solution with lower ammonia content, the so-called weak solution. Thus, it is necessary to include some kindof separation system between the generator and absorber9,10 inthe refrigeration cycle working with ammonia/water/NaOHmixtures. Up to now, several separation systems have beenproposed for the separation of the hydroxyl ions from the ternarymixtures.5 In this study, we examine the analysis of the technicalviability of the reverse osmosis system.

As an example of the advantages of the addition of hydroxidesto ammonia, Table 1 shows the COP values for an absorption

* To whom correspondence should be addressed. E-mail:juancarlos.bruno@urv.cat.

† Universitat Rovira i Virgili.‡ Universidad de Cantabria.

Table 1. COP Values as a Function of the NaOH Mass Fraction10

NaOH (mass fraction) tgenerator (°C) COP

0 107.1 0.480.05 90.8 0.650.10 75.5 0.88

Ind. Eng. Chem. Res. 2008, 47, 10020–1002610020

10.1021/ie8004012 CCC: $40.75 2008 American Chemical SocietyPublished on Web 11/14/2008

chiller of 135 kW (chilled water capacity) with the correspond-ing NaOH mass fractions in the ammonia/water mixture.10

Table 1 shows that the driving temperature goes down andCOP values grow as more NaOH is added to the cycle. Theevaporation temperature of the cycle for the case is shown inTable 1.

The aim of this article is to study the technical viability ofthe separation of NaOH from ammonia/water/NaOH mixturesby using reverse osmosis technology in the concentration rangeuseful for the development of new absorption chillers based onthese mixtures to improve performance with respect to con-ventional ammonia/water chillers, that is, reducing the requireddriving temperature and increasing the cycle COP.

When two solutions at different concentrations are placed onboth sides of a semipermeable membrane, the solvent movesfrom the more diluted solution through the membrane to theother side, until the chemical potential on both sides is equal.This spontaneous phenomenon is known as direct osmosis. Theosmotic pressure exerted on the membrane is a measurementof this concentration difference. The reverse osmosis processproduces the opposite effect. An external pressure higher thanthe osmotic pressure is applied to a concentrated solution incontact with a semipermeable membrane, and the solventpermeates across the membrane. In this case, two streams aregenerated, one with the solvent that crosses the membrane(called permeate) and another one (rejection) more concentratedin solute than the feeding solution.11,12

Figure 1 shows a block diagram of the proposed absorptionrefrigeration cycle using a reverse osmosis unit to separate theammonia/water/NaOH mixture. The poor solution coming fromthe generator with large sodium hydroxide content enters thereverse osmosis unit that may consist of several membranemodules. In this unit, the feed is separated into two solutions:one of them more concentrated in hydroxide (stream 1) that isrecycled back to the generator together with the rich solutionfrom the absorber, and the second one with a reduced contentin hydroxide that is sent directed to the absorber (rich solution,stream 8).

COP values of the cycle shown in Figure 1 are a function ofthe temperatures of the cycle (evaporator, condenser, absorber,and generator) and also depend on the concentration of NaOHand the efficiency of the separation system.

2. Experimental Setup

The experimental setup used to carry out the experimentsconsists of a laboratory scale stainless steel (316SS) reverseosmosis cell SEPA CF, prepared to withstand pressures of upto 69 bar. The main characteristics of this cell are: max.

temperature, 177 °C; max. pressure, 69 bar; and floodingvolume, 70 mL. The membrane effective area is 155 m2.

The cell is designed to simulate the membrane fluid dynamicbehavior of industrial modules, testing their performance in awide range of operating conditions. The experimental setupconfiguration employed for the separation of NaOH from themixture of ammonia/water is shown in Figure 2.

As shown in Figure 2, the experimental setup consists of acell and cell body, a hydraulic pump, a positive displacementcirculation pump, a manometer, a concentrate control valve, andthe necessary storage tanks for the feed, permeate, andconcentrate when the plant is operated in a continuous mode.

The experimental procedure was as follows: once themembrane had been placed in the cell body, hydraulic pressurewas applied to produce the required hermetic sealing, thusensuring a good pressurization. The circulation pump ensuresthe desired feeding solution flow rate from the tank to the cellbody at the required operating pressure. The inlet solution isparallel to the membrane, and the filtration takes place tangen-tially producing the permeate flow. Cleaning of the system isnecessary as soon as a decrease of permeate flow is perceived.The concentrate flow, that is, the flow that does not cross themembrane, is controlled using a stainless steel valve thatproduces the required trans-membrane pressure on the mem-brane surface. The continuous flow operation is simulated usingthree different tanks for the feeding solution, permeate, andconcentrated solution. The objective pursued with the continuousoperation mode is to simulate the continuous flow of solutionin and out of the membrane and to avoid the increase in osmoticpressure if the system is operated in discontinuous mode. Thereverse osmosis membranes used in the experiments conductedwere BW30 and SW30HR LE, both supplied by Filmtec-Dow.These membranes come in tubular form, and it was necessaryto adapt them to the shape corresponding to the membrane cell.The characteristics of these membranes are given in Table 2.

The NaOH/water and ammonia/water/NaOH solutions usedas feed were prepared by dissolving NaOH pellets and ammonia,with a purity of 28-30% and a density of 0.9 kg/L; bothreagents were supplied by MERCK.

Figure 1. Absorption refrigeration cycle using ammonia/water/NaOHmixtures.

Figure 2. Reverse osmosis experimental setup configuration.

Table 2. Technical Data of the FILMTEC-DOW Membranes

membrane model BW30SW30HR

LE

membrane area 155 cm2 155 cm2

material polyamide polyamidemax. operating pressure 41 bar 83 baroperating pressure 40 bar 40 barpH operating range 2-11 2-11max. operating temperature 45 °C 45 °CNaCl rejection (%) 99 99.6max. permeate flow (L/m2 ·h) 120 80permeability coefficient (L m2 h bar) 3 2

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The experimental results were monitored by measuring thesolution conductivity using a CRISON BASIC 30 conductimeterin the range of 0.01-199.9 mS/cm. For the analysis of theammonium ion, a K-355 distillation unit from BUCHI was usedwith a detection limit of nitrogen g0.1 mg. The analyticalmethod of this unit is based on the Kjeldahl method.

Additionally, permeate samples were collected every 10 min,and NaOH concentration was determined. In the case of theexperiments with solutions that also contained ammonia, onlythe composition of the final product was analyzed because ofthe difficulty in manipulating solutions with high ammoniacontent. The composition of NaOH and ammonia of the finalproduct was determined with the conductimeter BASIC 30 forNaOH and the distillation unit K-355 for ammonia.

Once the experiment was completed, the final volume of thepermeate tank was used to calculate the permeate flux.

3. Results and Discussion

The parameters commonly used to determine the separationefficiency factor in reverse osmosis membranes are the permeateflux through the membrane and the rejection coefficient. Thepermeate flux or solvent flux (Jw) is given by the followingequation:

Jw )K(∆P-∆π) (3)

where K is the permeability coefficient characteristic of eachmembrane, ∆P is the transmembrane differential pressure, and∆π is the osmotic pressure difference. The term (∆P - ∆π)represents the net driving force that is responsible for the solventpermeation through the membrane. To calculate the osmoticpressure developed in the experiments, the following equationbased on thermodynamic considerations13 was used:

Π) 0.08308Φ(T+ 273.15)∑ mi (4)

where Φ is the osmotic coefficient (usually 0.93 for brackishwater and 0.902 for seawater), T is temperature in °C, and mi

is molality of all of the solution, ionic, or nonionic componentsin mol/L that can be calculated by:

mi )Ci

1000MWi106 -TDS

106

(5)

where Ci is the concentration of component “i” in the solution,in mg/L; MWi is the molecular weight of component “i”; andTDS is the total dissolved salt content in the solution, in mg/L.

The retention or retention coefficient (R) is defined as:

R) 1-Cpermeate

Cfeed(6)

This parameter expresses the extent to which a solute isretained by the membrane. In an ideal case, the retentioncoefficient could be 1 or 100%, that is, the membrane does notallow the solute to cross the permeate stream (Cpermeate ) 0). Inpractice, it is not possible to achieve a retention of 100%, butvalues between 95% and 99% are possible.14

Besides the permeate flux and the retention coefficient, thevariables that influence the reverse osmosis performance aretemperature (influences the flux through the membrane and thetendency to compaction), net pressure, flux density, soluteconcentration in the permeate and concentrate, and the mem-brane lifetime.

The experimental study is described below. First (task 1),the separation of binary solutions water/NaOH was tested todetermine the technical viability of the separation of NaOH usingreverse osmosis technology. In the second experimental step(task 2), the separation of ternary mixtures ammonia/water/NaOH was analyzed to determine the influence of the presenceof ammonia in the NaOH separation and the permeate flow ratewith respect to the results of task 1. For the purpose of thiswork, the concentrations of NH3 and NaOH used in theexperiment are close to the ones that could be found in arefrigeration absorption cycle (approximate 0.01-0.05 massfraction NaOH and 0.1-0.3 mass fraction NH3). Each experi-ment was conducted twice and with both membranes until steadystate was reached. Table 3 shows the operating conditions ofexperiments 1-8 using the water/NaOH mixtures, and Table 4provides similar information for experiments 9-20, using theternary mixtures. In both cases, the transmembrane pressure wasset to 39 bar, the operating maximum pressure for the BW30membrane. This pressure is high enough to overcome theosmotic pressure of the most concentrated analyzed solutionsand corresponds to the difference between the pressure of theconcentrate set by the control valve (40 bar) and the pressureon the permeate stream side, atmospheric pressure.

The theoretical osmotic pressure for each solution wascalculated using eqs 4 and 5. The results are shown in Table 5.As seen in this table, as the NaOH concentration increases, theosmotic pressure also increases. That means that working withthe same membrane and the same transmembrane pressure, thepermeate flux decreases as the NaOH mass fraction increases.Thus, for the most concentrated NaOH solutions (0.05 massfraction), it is necessary to pressurize the feed solution atpressures higher than 24 bar to obtain a significant permeateflux.

3.1. NaOH/Water Binary Mixture. Table 6 gives the valuesof permeation flux and retention coefficients corresponding toexperiments 1-8.

Table 3. Experimental Planning with Water/NaOH Solutions (Task1)

membrane BW30 membrane SW30HR LE

experiment feed (mass fraction) experiment feed (mass fraction)

EXP 1 0.007 EXP 5 0.008EXP 2 0.013 EXP 6 0.013EXP 3 0.019 EXP 7 0.021EXP 4 0.042 EXP 8 0.039

Table 4. Experimental Planning with Ammonia/Water/NaOHTernary Solutions (Task 2)

membrane BW30 membrane SW30HR LE

feed (mass fraction) feed (mass fraction)

experiment NH3 NaOH experiment NH3 NaOH

EXP 9 0.0786 0.0172 EXP 15 0.0755 0.0159EXP 10 0.0632 0.0321 EXP 16 0.0659 0.0290EXP 11 0.1959 0.0146 EXP 17 0.1984 0.0144EXP 12 0.1922 0.0432 EXP 18 0.1928 0.0400EXP 13 0.3093 0.0154 EXP 19 0.3028 0.0148EXP 14 0.3028 0.0449 EXP 20 0.3024 0.0418

Table 5. Calculated Osmotic Pressure as a Function of the NaOHConcentration

feed solution (mass fraction) osmotic pressure (bar)

0.012 4.60.020 8.40.028 11.80.050 24.0

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According to the results presented in Table 6, both thepermeate flux and the retention coefficients decrease workingwith both membranes as the NaOH mass fraction increases, asis to be expected from the increase in the osmotic pressure.The permeate flux obtained at the lowest NaOH mass fraction(approximate 0.012) is between 2.5 (BW30) and 5 (SW30HRLE) times lower than the maximum flux shown in Table 3 whenthe membranes are used to separate water and NaCl. Regardingthe results of the retention coefficient, working with bothmembranes, a better separation is achieved with the SW30HRLE than using the BW30 membrane, although the permeate fluxobtained is lower for the former membrane (SW30HR LE).Figures 3 and 4 present the NaOH mass fraction in the permeateflux as a function of time for EXP 1-8.

As shown in Figures 3 and 4, membrane SW30HR LE leadsto a better separation for the same operating conditions thandoes the membrane BW30, in accordance with the results ofthe retention coefficient shown in Table 6.

As can be see in Table 6, with one separation step, theretention values are lower than required for the future applicationof the separation process (it is necessary to maximize theseparation of NaOH and to avoid the presence of hydroxyl ionsin the absorber of the refrigeration cycle). To increase efficiencyin the retention of NaOH, a separation methodology is proposedto simulate the operation of a membrane process in-series. Thismethodology consists of performing an experiment with highconcentration of NaOH and to use the permeate obtained asfeed in a subsequent step, and so on, in such a way that amembrane configuration in-series is simulated. With this

procedure it is expected to achieve two objectives: (i) retentionvalues of 99% and (ii) higher permeate fluxes. This is also asuitable methodology for overcoming the osmotic pressure ofhighly concentrated solutions. The initial feed solution concen-trations for both membranes (BW30 and SW30HRLE) werethose corresponding to EXP 4 and EXP 8. With each membrane,a complete cycle of experiments was conducted, that is,successive separation steps until a retention value close to 99%was achieved. In the case of membrane BW30, the cycle ofexperiments was named SERIES_A, and for membraneSW30HRLE it was named SERIES_B. As a typical exampleof the results obtained, Figures 5 and 6 show the time evolutionof the NaOH mass fraction in the permeate at each step of bothcycles of experiments.

As shown in Figures 5 and 6, to achieve a good separationof NaOH working with membrane BW30, four separation stepswere necessary, whereas only two steps were needed withmembrane SW30HRLE. Table 7 shows the values of permeateflux and retention coefficient obtained in experiments “SE-RIES_A” and “SERIES_B”. It is worth noting that the permeateflux values are higher for the BW30 membrane (28.2 L/m2 ·h)than for the SW30HRLE, as was also the case in the one-stepexperiments performed previously. The retention parametercalculated in the last step ranges between 86% (membraneBW30) and 94% (membrane SW30HRLE).

In the case of the membranes connected in series, the totalretention parameter was calculated as the ratio between theconcentration in the permeate flow in the last separation stepand the feed concentration in the first step (Figure 7):

Table 6. Permeate Flux and Retention Percentage in Experiments EXP 1-8

flux (L/m2 ·h) retention (%)

experiment BW30 (EXP 1-4) SW30HR LE (EXP 5-8) BW30 (EXP 1-4) SW30HR LE (EXP 5-8)

EXP 1/EXP 5 49.3 ( 0.6 16.0 ( 1.1 56.5 91.5EXP 2/EXP 6 35.9 ( 0.7 11.9 ( 0.3 48.7 90.2EXP 3/EXP 7 24.7 ( 0.8 9.6 ( 0.7 47.5 87.0EXP 4/EXP 8 24.3 ( 0.6 8.4 ( 0.3 31.7 87.6

Figure 3. Time evolution of the NaOH mass fraction in the permeatesolution for membrane BW30.

Figure 4. Time evolution of the NaOH mass fraction in the permeatesolution for membrane SW30HR LE.

Figure 5. Time evolution of the NaOH mass fraction in the permeate.Experiments SERIES_A.

Figure 6. Time evolution of the NaOH mass fraction in the permeate.Experiments SERIES_B.

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Rtotal ) total retention) 1-Cpermeate n+1

Cfeed(7)

For both membranes, the calculated total retention after foursteps (membrane BW30) and two steps (membrane SW30HRLE) was 99%.

3.2. Ternary Mixture Ammonia/Water/NaOH. The resultspresented above show the technical viability of the separationof NaOH from aqueous solutions in the range of mass fractionsbetween approximately 0.01 and 0.05 NaOH. As stated in Table4, in the experiments corresponding to task 2 EXP 9-20,ammonia was added to the studied solutions at several concen-trations (approximately 0.1, 0.2, and 0.3 ammonia mass frac-tion). The aim of this task was to analyze the effect of ammoniaon the separation of NaOH from the solution and to determinethe viability of the NaOH separation from the ammonia/water/NaOH ternary mixture. As with task 1, the results were analyzedby measuring the final permeate flux, the retention, and the massfraction of OH- and NH3, in the permeate as well as in thefinal concentrate stream (Figure 8, Figure 9, Table 8).

The experimental results presented in Figures 8 and 9 showthat the addition of ammonia leads to a slight decrease of thepermeate flux in comparison with the values obtained with theNaOH aqueous solutions. This reduction in the permeate fluxis justified by a slight increase in the osmotic pressure retentatesolutions that also reduces the separation of NaOH (Table 8).

Tables 9 and 10 show the influence of the ammoniaconcentration in the permeate concentration. The results pre-sented in Table 9 represent the NaOH concentration in thepermeate solution for the last separation step corresponding tothe water/NaOH binary mixtures and to the ammonia/water/NaOH ternary mixtures with a concentration of NaOH in thefeed stream of approximately 0.02. Likewise, Table 10 showsthe NaOH approximate concentration in the permeate solutionfor the binary mixture water/NaOH and for the ternary mixturesammonia/water/NaOH. In this case, the NaOH mass fractionin the feed solution is approximate 0.05. To calculate the mass

fraction of each component, the parameters densities presentedin Salavera et al.6 were used.

The presence of ammonia in the solution reduces theefficiency of the NaOH separation using both membranes.Ac-cording to the results presented in Tables 9 and 10, the NaOHconcentration in the permeate is lower for the binary mixtures(EXP 2, 6, 4, and 8) with respect to the ternary mixtures (EXP9-20), so the retention is greater for the binary mixtures.Furthermore, it is interesting to note that the ammonia concen-tration is almost the same in the permeate and in the concentratesolutions; that is, the ammonia is not rejected by either of themembranes and goes through them freely.

The error standards in the measurements of the results,calculated for ammonia and NaOH mass fraction (in the feed,

Figure 7. Schematic diagram of a reverse osmosis system connected in series.

Figure 8. Influence of the NH3 concentration in the permeate flux (NaOHapproximate feed concentration ) 0.02).

Figure 9. Influence of the NH3 concentration in the permeate flux (NaOHapproximate feed concentration ) 0.05).

Table 7. Flux Permeate and Retention Coefficient for the Cycle ofExperiments SERIES_A and SERIES_Ba

membrane BW30(SERIES_A)

membrane SW30HRLE (SERIES_B)

step flux (L/m2 ·h) retention (%) flux (L/m2 ·h) retention (%)

1 24.3 ( 0.6 31.7 8.4 ( 0.3 87.62 25.6 ( 0.6 57.7 13.8 ( 0.6 94.13 26.6 ( 0.7 69.34 28.2 ( 0.8 86.3

a The retention was calculated with molar concentration of NaOH.

Table 8. Retention Percentage in Experiments EXP 9-20

retention (%) (calculated with mass fraction)

experiment BW30 (EXP 9-14) SW30HR LE (EXP 15-20)

EXP 9/EXP 15 39.2 73.4EXP 10/EXP 16 20.1 31.3EXP 11/EXP 17 38.9 71.2EXP 12/EXP 18 21.5 31.2EXP 13/EXP 19 37.5 70.2EXP 14/EXP 20 21.1 31.8

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permeate, and retentate streams), are 0.05% and between 0.001%and 0.02%, respectively, for the experiments with both typesof membranes.

Following the same methodology employed in task 1, twonew sets of experiments were performed to simulate theconfiguration of membranes connected in series: SERIES_C(membrane BW30) and SERIES_D (membrane SW30HR LE).The initial feed solutions concentrations for both membraneswere those corresponding to EXP 14 and EXP 20 of task 2(Table 4). In this methodology, the permeate collected fromeach step was used as feed solution in the following step. Figure10 shows the NaOH mass fraction in the permeate solutioncorresponding to experiments SERIES_C and SERIES_D.

The first step indicated in Figure 10 corresponds to the feedconcentration (for the BW30 membrane is 0.0449 mass fractionand 0.0418 for the other membrane) and not to the first permeateconcentration. To maximize the separation of the NaOH fromthe studied ternary mixture, seven separation steps are necessaryusing membrane BW30, whereas only three steps are needed

in the case of membrane SW30HR LE. In both cases, thenumber of necessary steps is higher than the required numberfor the binary mixtures that started with a feed of NaOH massfraction approximately of 0.05. The permeate flux and retentionvalues for both series of experiments are presented in Table11.

As shown in Table 11, membrane BW30 allows higherpermeate flux than does membrane SW30HR LE. However, theretention percentages are higher for the latter membrane,following behavior similar to that observed in the experimentswith the binary mixture NaOH/water. The total retentioncoefficient was 99% for both membranes.

The values of permeate flow rate obtained in this work arestill lower than the values required in the operation of anabsorption chiller.10 This difference can be explained byconsidering the small membrane surface (0.155 m2). Membraneswith a larger surface or modules in parallel could be used toincrease the permeate flow rates.

The experimental results will be used to perform simulationsof the refrigeration cycle with ammonia using Aspen Plussimulation software to evaluate the separation process ascompared to a conventional ammonia-water cycle.

4. Conclusions

The addition of NaOH to ammonia/water absorption refrig-eration cycles reduces the temperature required to drive the cycleand increases its efficiency. To obtain a significant improvementin the process performance by addition of NaOH to the cycle,and to avoid its presence in the absorption section of the cycle,it is necessary to incorporate a system that separates the hydroxylions from the solution sent to the absorber. The use of reverseosmosis technology is a suitable system for achieving thisseparation. The following conclusions can be drawn from theexperimental results obtained from this work.

The separation of NaOH from their aqueous solutions with amaximum mass fraction of 0.05 (5% weight) is viable usingreverse osmosis with the tested membranes, BW30 and SW30HRLE, providing enough contact time of the solution with theexperimental data of the membrane in series configuration. Inthis case, maximum retention values between 86% and 94%and average permeate fluxes between 14 and 28 (L/m2 ·h) wereobtained.

The experiments carried out with the ternary mixturesammonia/water/hydroxide reveal a slight influence of theammonia on the separation of NaOH with respect to the binarymixture NaOH/water. Comparing the results obtained withbinary and ternary mixtures, it was observed that the permeateflux and the percentage of retention decrease slightly with the

Table 9. Influence of the NH3 Concentration in NaOH Concentration in the Permeate (NaOH Feed Approximate Concentration ) 0.02)

feed NH3 concentration (mass fraction) permeate NaOH concentration (Mass fraction)

experiment membrane BW30 membrane SW30HR LE membrane BW30 membrane SW30HR LE

EXP 2/EXP 6 0 0 0.0066 0.0012EXP 9/EXP 15 0.0786 0.0755 0.0104 0.0042EXP 11/EXP 17 0.1959 0.1984 0.0089 0.0041EXP 13/EXP 19 0.3093 0.3028 0.0097 0.0044

Table 10. Influence of the NH3 Concentration in NaOH Concentration in the Permeate (NaOH Feed Approximate Concentration ) 0.05)

feed NH3 concentration (mass fraction) permeate NaOH concentration (Mass fraction)

experiment membrane BW30 membrane SW30HR LE membrane BW30 membrane SW30HR LE

EXP 4/EXP 8 0 0 0.0286 0.005EXP 10/EXP 16 0.0632 0.0659 0.0256 0.0219EXP 12/EXP 18 0.1922 0.1928 0.0339 0.0272EXP 14/EXP 20 0.3028 0.3024 0.0355 0.0290

Table 11. Permeate Flux and Retention Values for Experiments ofTernary Mixtures SERIES_C and SERIES_Da

membrane BW30(SERIES_C)

membrane SW30HRLE (SERIES_D)

step flux (L/m2 ·h) retention (%) flux (L/m2 ·h) retention (%)

1 22.6 ( 0.6 21.7 7.3 ( 0.8 30.92 23.9 ( 0.8 30.2 9.2 ( 0.8 48.83 24.2 ( 0.6 33.1 12.3 ( 0.8 61.24 25.4 ( 0.6 38.2 15.6 ( 0.7 78.35 30.0 ( 0.7 47.66 32.3 ( 0.8 56.17 34.6 ( 2.0 65.18 37.2 ( 1.0 79.2

a The retention was calculated with molar concentration of NaOH.

Figure 10. Permeate NaOH mass fraction of experiments SERIES_C andSERIES_D.

Ind. Eng. Chem. Res., Vol. 47, No. 24, 2008 10025

ternary mixture. In the range of solution concentrations studied(approximate 0.1-0.3 ammonia mass fraction, 10-30% weight),the retention parameter reached values of 78-79% and thepermeate flux comprised between 16-37 (L/m2 ·h) dependingon the NaOH concentration.

The ammonium ion is not retained; it permeates through bothtypes of membranes, as almost the same concentration wasmeasured in the permeate and concentrate solutions. Theseresults favor the use of these membranes in absorption refrigera-tion cycles because they will not modify substantially theammonia concentration at the absorber inlet.

Acknowledgment

We appreciate the funding support from the Spanish Scienceand Education Ministry through research projects DPI2003-04752 and CTQ2006-14360/PPQ. S. Steiu thanks Departamentd’Universitats, Recerca i Societat de la Informacio de laGeneralitat de Catalunya i del Fons Social Europeu, for ascholarship.

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ReceiVed for reView March 11, 2008ReVised manuscript receiVed October 6, 2008

Accepted October 7, 2008

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10026 Ind. Eng. Chem. Res., Vol. 47, No. 24, 2008