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Reverse Osmosis To Concentrate Lisinopril Purified by Means of Liquid Chromatography: From Pilot-Plant to Industrial-Scale Unit David Senica,†,§ Janez Rzˇ en,† Matjazˇ Usˇ tar,† Pavel Drnovsˇ ek,† Stojan Kogej,† and Aleksander Pavko‡,* Lek Pharmaceuticals, Verovsˇ kova 57, 1526 Ljubljana, Slovenia, and Faculty of Chemistry and Chemical Technology, University of Ljubljana, Asˇ kercˇ eva 5, 1000 Ljubljana, Slovenia
Reverse osmosis to concentrate the lisinopril after a low-pressure liquid chromatography was tested under various process conditions on a pilot-scale unit. It was found that the permeate flux 20-30 L/hm2 can be kept over the whole concentration range with a feed pressure up to 35 bar and that fouling is negligible. The rejection of lisinopril is almost complete. These data were used for a conceptual design and installation of an industrial scale reverse osmosis unit with the membrane surface area 60 m2 and capacity of 1.5 m3/h permeate. The capital cost estimate is significantly lower compared to a single-stage vacuum evaporator. The operating cost estimate for evaporation is roughly $6.8/kg lisinopril, while the estimated cost for concentration of lisinopril solution with reverse osmosis taking into account the theoretical energy consumption by pumps is $0.23-0.33/kg lisinopril. Cost calculated from actual plant data is even lower, $0.16/kg lisinopril. The validation results of the full-scale unit and its trouble-free operation prove that the decision for reverse osmosis is economically justified and that the scale-up was technically successful. 1. Introduction Angiotensin-converting enzyme (ACE) inhibitors, such as enalapril and lisinopril, are effectively used for the control of hypertension and congestive hearth failure.1-3 Lisinopril dihydrate, chemically N-[N-[(1S)-1-carboxy3-phenylpropyl]-L-lysyl]-L-proline dihydrate, is produced by means of a stepwise chemical synthesis1-3 and purified by means of low-pressure liquid chromatography. A polar mobile phase (pH ) 10.5), which consists of water and small amounts of acetonitrile and ammonia, is used as an eluent. Purified lisinopril is precipitated, filtered, and dried. The concentration of lisinopril in the solution after chromatography is too low to be readily precipitated, and therefore the solution has to be concentrated from an initial lisinopril concentration of about 2 g/L to a final concentration of 90-100 g/L. During a pilot-scale production of lisinopril, the solution was concentrated to the specified concentration by evaporation of water in a simple single-stage vacuum evaporator.4 Full-scale production of lisinopril requires evaporation of large quantities of water, which is expensive, and for that reason reverse osmosis was chosen as a better option. A diagram of the process is shown in Figure 1. Reverse osmosis is a pressure-driven membrane process whereby the natural phenomenon of osmosis is reversed by the application of pressure to a concentrated solution in contact with a semipermeable membrane. If the applied pressure exceeds the solution’s natural * To whom correspondence should be addressed. Tel.: +386 1 2419506. Fax: +386 1 2419530. E-mail: saso.pavko@ uni-lj.si. † Lek Pharmaceuticals. ‡ University of Ljubljana. § Present address: Delft University of Technology, Faculty of Applied Sciences, Department of Biotechnology, Julianalaan 67, 2628 BC Delft, The Netherlands.
Figure 1. Schematic layout of the process.
osmotic pressure, the solvent will flow through the membrane from a more concentrated to a diluted solution.5 Ions and molecules dissolved in the solution are rejected. Reverse osmosis membranes are of a broad commercial interest and used in many industrial applications in various industry branches.5 Traditional technologies are substituted by reverse osmosis due to lower investment, operating and maintenance costs, and other advantages such as low operating temperatures, simple automation, and straightforward scale-up.6,7 Concentration polarization and fouling phenomena, which lower the process efficiency, can be substantially reduced with the use of the cross-flow technique, with the proper construction of the membrane elements and with the proper pretreatment of the feed. Concentration polarization and fouling are of a complex and transient nature, and they are very difficult to predict, especially fouling.6 Because of the complex nature of the phenomena associated with reverse osmosis, it is impossible to predict the process performance in the majority of industrial cases. Therefore, a design of a full-scale unit
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has to be based on the results of pilot-plant experiments. Fortunately, the scale-up from lab- or pilot-scale to fullscale is simple and reliable.6 During pilot-plant testing, the influences of transmembrane pressure, temperature, permeate recovery, volumetric concentration factor, and feed pretreatment methods on flux, fouling rate, and membrane rejection properties have to be carefully evaluated. The procedures of cleaning and sanitation and their intervals are also important.5,6 As a result of pilot-scale experiments, optimal transmembrane pressure and temperature ranges, and feed pretreatment methods, are determined. With known flow rates and required full-scale production capacity, the necessary membrane area can be calculated and a conceptual design of a full-scale unit together with capital and operating cost estimates can be made.6,8 The aim of the present work is to illuminate the approach by an industrial user to the specific design problem of a full-scale reverse osmosis process and solutions we have found. Furthermore, practical information about design, cost estimation, installation, and operation, which are given here, can be helpful to a wide range of potential users of membrane technology in industry. 2. Theory In general, reverse osmosis processes are evaluated in terms of three important parameters: permeate (water) flux, recovery, and solute rejection. The degree of solvent transport can be quantified in terms of solute rejection, which is defined as7,9
R)1-
Cp Cc
(1)
where Cp is the concentration of the solute in the permeate and Cc in the concentrate. The rejection defined as above is also referred to as apparent or observed rejection.7,9 The true rejection of solute is somewhat larger, because the concentration of the solute at the membrane surface is larger than that in the bulk concentrate. Reverse osmosis membranes reject ions, suspended solids, and certain uncharged dissolved substances (mainly organic compounds and silica compounds). Rejection of ions depends on their charge. The more charges ions have, the better they are rejected. Rejection of uncharged organic compounds generally depends on their size which is related to the molecular weight.10 Besides charge and size, the rejection of solutes is a function of pressure, temperature, concentration, and pH (for ionisable organic solutes). The organics rejection also depends on the membrane polymer type and on the structures and membrane/solute interactions.7 To evaluate rejection with a simple analysis of experimental data, eq 1 can be rearranged into the following form
Cp ) (1 - R)Cc
(2)
If the difference between transmembrane pressure and osmotic pressure is kept constant during concentration, then according to a solution-diffusion model,7,9 the true rejection and consequently the apparent rejection stay constant. Thus, a plot of instantaneous concentra-
tions of the solute in the permeate as a function of concentrations of the solute in the retentate results in a straight line, passing through zero. The rejection of the solute is calculated from the slope. 3. Materials and Methods 3.1. Pilot-Scale Reverse Osmosis Unit. The pilotscale unit includes two feed tanks, two high-pressure centrifugal pumps, a composite material pressure vessel (Phoenix Vessels Ltd., United Kingdom) for one 4 in. spiral-wound membrane element, a heat exchanger, measuring sensors, regulating and manual valves, a switchboard with programmable logic controller, and piping. The whole unit is skid-mounted for easy transportation. The feed is pumped from one of the stainless steel feed tanks through a prefilter to the membrane element by two high-pressure centrifugal pumps (Grundfos, Denmark). The flow rate of the feed is regulated manually by ball valves. The concentrate flows from the membrane element through the shell-and-tube heat exchanger and can be directed either back to one of the feed tanks or to the concentrate outlet. The concentrate pressure/flow rate is regulated by a pneumatic regulating valve (Samson, Germany). The concentrate and permeate flow rates are measured by rotameters (Rota, Germany) and pressures by piezoelectric sensors (Wika, Austria). The temperature of the concentrate is regulated by a self-operated temperature regulator (Samson, Germany) coupled to a regulating valve (Samson, Germany) which controls the outflow of the cooling water from the heat exchanger. The pilot plant unit operation is controlled by an industrial programmable logic controller (Omron, Japan), which is programmed to support four different control modes (feedback control loops): constant feed pressure, constant permeate flow rate, constant concentrate flow rate, and constant volumetric concentration factor. The pilot-scale reverse osmosis has been designed for testing various applications and for pilot-scale production. The unit can be operated either in a batch or in a continuous (single pass) mode. A more detailed description of the pilot plant unit is given elsewhere.4 A four in. Filmtec BW30-4040 spiral-wound reverse osmosis membrane element with 7.6 m2 of the surface area was used (The Dow Chemical Company, Midland, MI). A rated product water flow rate of the membrane element is approximately 350 L/h or 45 L/hm2, and the minimum salt rejection is 98%. (The permeate flow rate and salt rejection are based on the standard test conditions: 2000 ppm NaCl, 16 bar, 25 °C, pH 8, permeate recovery 15%.) The maximum permissible operating pressure is 41 bar and the maximum operating temperature 45 °C. The operating pH range is 2-11. For continuous operation at pH g 11, the maximum temperature is 35 °C.11 3.2. Pilot-Scale Experiments. The pilot-scale experiments were performed within the lisinopril concentration range from 2 g/L to 90-100 g/L. Preliminary experiments were run with constant feed pressure, and all other experiments, both in the batch and in the continuous operation mode, were run with constant permeate flux, which is usually the way of the operation of full-scale industrial units. Considering the relatively low osmotic pressure (18.9 bar calculated with the van’t Hoff9 equation for a solution containing 100 g/L lisinopril, 12 g/L acetonitrile and 4 g/L NH3 at 25 °C) and membrane element specifications, it was predicted that
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Figure 2. Simplified scheme of a full-scale reverse osmosis unit. (Symbols denote: P1, P2, P3, feed pumps; PT1, PT2, PT3, pressure transmitters; FT1, FT2, flow rate transmitters; TC, self-operated temperature controller; RV1, RV2, regulating valves; LT, level transmitter; pHT, pH meter; UV, ultraviolet detector.)
the permeate flux 20-35 L/hm2 could be achieved with a moderate feed pressure of 25-35 bar. Therefore the experiments were performed within the corresponding flux and pressure limits. The flow rate was regulated by means of a programmable controller. The temperature in the system was maintained below 30 °C. The limit of 30 °C was chosen for two reasons: to avoid any degradation of lisinopril and to operate the system below the maximum permissible operating temperature. The permeate recovery was kept at 15-25%. The fouling was studied in a continuous, single pass operation mode at a constant concentration of the feed. To keep the feed concentration constant, the concentrate and the permeate were recycled back to the feed tank. 3.3. Full-Scale Reverse Osmosis Unit. The fullscale reverse osmosis has been designed according to the user requirements specification, and it consists of a feed tank, three high-pressure pumps, four pressure vessels for spiral-wound membrane elements, a heat exchanger, measuring sensors, regulating valves, a switchboard with programmable logic controller, piping, and manual valves. A simplified scheme of the unit is shown in Figure 2. The feed is pumped from the stainless steel feed tanks through the prefilter in the pressure vessels by three high-pressure centrifugal pumps, P1, P2, and P3 (Grundfos, Denmark). The pump P1 runs at a constant speed while the speed of the pumps P2 and P3 can be changed by means of a frequency converter. The prefilter protects the membrane elements against plugging with coarser solid particles. Eight 4 in. spiral-wound membrane elements Filmtec BW30-4040 (The Dow Chemical Company, Midland, MI) are placed in the four stainless steel pressure vessels, which are connected in parallel. The total membrane surface area is 60.8 m2. The concentrate flows from the pressure vessels through the shell-andtube heat exchanger back in the feed tank. The concentrate pressure is regulated by a pneumatic regulating valve RV1 (Ka¨mer Ventile, Germany). The flow rates are measured by mass flow meters FT1 and FT2 (Endress+Hauser, Germany), and pressures by piezoelectric sensors PT1, PT2, and PT3 (Endress+Hauser, Germany). The temperature in the system is regulated by a self-operated temperature controller TC. A temperature sensor is coupled to a thermostat and to a regulating valve RV2 (ABB, Germany), which controls the outflow of the cooling water from the heat ex-
changer. The feed tank is equipped with a stirrer, piezoelectric level sensor LT (Endress+Hauser, Germany), and standard pH electrode (Mettler Toledo, Switzerland). The pH value is measured by a pH meter pHT (Endress+Hauser, Germany). An ultraviolet detector (Optek, Germany) is mounted at the permeate outlet to detect lisinopril in case of membrane leakage. If lisinopril is detected, the permeate will automatically be redirected to the feed tank to prevent any loss of the product. The operation of the reverse osmosis is controlled by an industrial programmable logic controller (Siemens, Germany). The controller is programmed with a constant permeate flow rate feedback control loop and includes all necessary start-up and shutdown safety logics for automatic operation. The reverse osmosis is also connected to a plant-wide control system (GE Fanuc Intellution Inc., Foxborough, MA), which enables remote supervision and control of the unit and storage of all important information, such us measured process variables and events. 3.4. Analytical Procedures. The samples of the concentrate (retentate) and permeate collected during the pilot-scale experiments and the full-scale unit validation were analyzed for lisinopril, acetonitrile, and ammonia concentration by using the analytical methods described bellow. The concentrations determined were used to calculate the rejection of lisinopril. Rejections of acetonitrile and ammonia were determined later on the samples collected during the start-up and validation of the full-scale unit. The concentration of lisinopril in samples was determined by means of high-pressure liquid chromatography (HPLC). An HP 1100 HPLC system (Agilent Technologies, Palo Alto, CA) with a degasser, gradient pump, autosampler, column thermostat, and ultraviolet detector was used. An HPLC column with dimensions of 250 × 4.6 mm was packed with 5 µm Hypersil ODS stationary phase (ThermoHypersil Ltd., UK). A gradient was used for chromatographic separation. The mobile phase A contained 95% of a phosphate buffer solution and 5% of acetonitrile (Merck, Germany). The mobile phase B contained 50% of a phosphate buffer solution and 50% of acetonitrile. The pH of solution was adjusted at 4.0 with phosphoric acid (Merck, Germany). The flow rate of the mobile phase was 2.5 mL/min and the temperature of the column 60 °C. Lisinopril was detected at 210 nm. The retention time was 3.3 min. Acetonitrile was determined by gas chromatography (GC). A HP 5890A gas chromatography system (Agilent Technologies, Palo Alto, CA), equipped with a HP 7694 headspace autosampler (Agilent Technologies, Palo Alto, CA) was used. Chromatographic separation was achieved with the use of a HP FFAP column (Agilent Technologies, Palo Alto, CA), 50 m long and 0.32 mm in diameter with 0.52 µm film thickness. Helium was used as a carrier gas. The oven temperature was 200 °C. Samples were injected in splitless mode. A retention time of acetonitrile was 6.3 min. Ammonia was determined either with colorimetric or potentiometric titration. Samples were titrated with a solution of hydrochloric acid (Merck, Germany). Methylorange was used as indicator in the colorimetric titration. The potentiometric titration was performed with an automatic titrator Rondolino DL 50 (Mettler Toledo, Switzerland). The potential was measured by DG111SC pH electrode (Mettler Toledo, Switzerland).
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Figure 3. Influence of the lisinopril concentration, the temperature, and the permeate recovery on the permeate flux and the feed pressure during batch concentration of lisinopril. Data for two runs at different concentration ranges. (Filmtec BW30-4040 4 in. spiral-wound RO membrane, pH 10.5.)
4. Results and Discussion 4.1. Results of Pilot-Scale Experiments. A few preliminary experiments were performed to check the expected flux and feed pressure range. The results of two batch concentration experiments, performed within different concentration ranges, are shown in Figure 3. It can be seen that the permeate flux was indeed reasonably high over the whole concentration range, already at a moderate feed pressure. In the first experiment (A) the lisinopril solution was concentrated from 1.7 to 12.5 g/L. With the feed pressure between 15.5 and 12.5 bar, the corresponding normalized permeate flux (T ) 25 °C) was between 25 and 21 L/hm2. In the second experiment (B) the solution was concentrated from 25.7 g/L to 86.2 g/L with the feed pressure kept between 23 and 25 bar. The normalized permeate flux dropped from 23 L/hm2 to 15.5 L/hm2. Fouling was tested by an experiment in a continuous, single-pass mode. The concentration of lisinopril in the feed was 95.8 g/L and it was kept constant by recycling the concentrate and the permeate back to the feed tank. The permeate recovery was maintained between 17 and 20% and the temperature between 20 and 30 °C. The permeate flux and the feed pressure were nearly constant for more than 6 h (Figure 4). According to these results, no problems with fouling were expected at fullscale. Lisinopril rejection is shown in Figure 5. Negligible concentration of the lisinopril found in the permeate throughout the concentration range, indicates a very high, almost complete rejection of the molecule. The rejection, calculated according to eq 2 from the slope of the regression line, is 99.8%. A complete rejection of lisinopril was expected because of its relatively high molecular weight (423.5 g/mol). The pilot-scale experiments actually showed that the permeate flux 20-30 L/hm2 can be kept over the whole concentration range with a feed pressure up to 35 bar. 4.2. Conceptual Design of Full-Scale Reverse Osmosis Process. Based on the results of the pilotscale tests and process requirements, a conceptual design of a full-scale process was made, and operating and capital costs were estimated. The necessary capacity and time schedule of the full-scale reverse osmosis process are dependent on the chromatographic production characteristics. For that reason, both processes
Figure 4. Fouling experiment. (Continuous, single-pass operation at the constant feed concentration of 95.8 g lisinopril/L. Filmtec BW30-4040 4 in. spiral-wound RO membrane, lisinopril concentration 95.8 g/L, permeate recovery 17-20%, 20-30 °C, pH 10.5.)
have to be well synchronized to ensure minimum process times and maximum utilization of equipment. A batch (4 m3) of lisinopril solution has to be concentrated from the initial concentration of 2 g/L to the final concentration of 90-100 g/L with resulting concentrate volume smaller than 100 L. The maximum batch cycle time (including preparation and cleaning), estimated from the overall process time schedule is limited to 5 h. Conceptually, the concentration process was planned to run as follows (the time schedule and the time course of different variables are shown in Figure 6): The feed solution coming from the chromatography is first collected in the feed tank until a certain volume is reached (I), and then the process of reverse osmosis is started. In this part of the process (II) the volume of the concentrate in the feed tank has to be kept constant. When the chromatography is finished, the solution in the feed tank is further concentrated (III) until the final volume is reached (∼80-90 L). The permeate flow rate throughout the process is 1500 L/h which is equal to the flow from the chromatography. As shown in Figure 6, the concentration of lisinopril first changes only gradually (II), and therefore a rather low feed pressure is sufficient to keep the permeate flowing at the required
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energy consumption per batch are calculated for the following operating conditions: permeate flow rate 1500 L/h, feed pressure 35 bar and permeate recovery 15%. The feed pressure is intentionally at the upper limit, to get a more conservative energy estimate. First, the flow rate of the feed is calculated according to eq 3.12
Qp × 100% 1.5 m3/h × 100% ) ) 10 m3/h (3) PR 15%
Qf )
Next, the power input and total energy consumption are estimated at corresponding feed flow rate and pressure drop according to the following equations6,8,13
Qfpf 10 m3/h × 35 × 105 Pa ) ) 13.9 kW η 0.70 × 3600 s/h
(4)
ERO ) Pt ) 13.9 kW × 2.7 h ) 37.5 kWh
(5)
P) Figure 5. Concentration of lisinopril in the permeate as a function of the concentration of the lisinopril in the concentratessummarized results of pilot- and full-scale runs. (Filmtec BW30-4040 4 in. spiral-wound RO membrane, permeate flux 15-30 L/hm2, permeate recovery 15-20%, 20-30 °C, pH 10.5.)
In the estimation a 70% pump efficiency is assumed, which is common efficiency of centrifugal pumps at their rated capacity.6 The installed power has to be somewhat higher because of the electric losses. Electricity costs for EU-industrial users are on average $0.05/kWh. The total estimated energy costs for the concentration of one batch of the lisinopril solution with the reverse osmosis are therefore $1.9 ($0.23/kg lisinopril). The evaporation of an equal amount of water at 30 °C in a single-stage vacuum evaporator (p ) 42,4 mbars, ∆Hev ) 2431 kJ/ kg water14) requires
Eev ) m∆Hev )
Figure 6. Planned time schedule of the concentration process. (The Roman numbers denote: I, filling of the feed tank; II, concentration of the solution at a constant volume; III, final concentration of the solution; IV, remaining time used for preparation and cleaning.)
flowrate. When the concentration starts to increase rapidly (III), a higher pressure has to be applied. For calculations of the membrane area, a flux of 25 L/hm2 was assumed. Consequently, the necessary area for the first part of the concentration (II) is 60 m2. In the second part (III), 750 L of the solution is concentrated further to 90 L in 0.5 h. Taking into account the same flux, the calculation gives 52.8 m2. Therefore 60 m2 of the membrane area was considered as sufficient in further calculations. With membrane elements equal to those used for pilot-scale experiments, eight elements are needed (total area of 60.8 m2). The RO operation time during one concentration cycle is approximately 3.3 h (see Figure 6). The remaining time is used for preparation and cleaning (IV). If necessary, it can also be utilized to increase the capacity since a cycle of 5 h is acceptable. The specification for the full-scale RO are hence as follows: a batch system with an open loop recycle; a 60 m2 membrane area; feed flow rate 10 m3/h; four pressure vessels connected in parallel equipped with high-pressure multistage centrifugal pumps; a cleaningin-place (CIP) unit and a process controller. 4.3. Estimation of Operating and Capital Costs. Feed flow rate, power input of the feed pumps and
3910 kg × 2431 kJ/kg ) 2.64 MWh 3600 s/h (6)
or 3.7 tons of superheated steam (2 bar, 120 °C), condensed and cooled to 30 °C (Σ∆H ) 2581 kJ/kg steam).14 The steam costs are $15/ton ($0.021/kWh) 8,14 which gives $55.5/batch ($6.8/kg lisinopril) for the evaporation. Since the estimation does not include thermal losses and energy requirements of a vacuum system and evaporator condenser (cost of cooling medium) the actual costs are considerably higher. Considering the total number of batches per year, savings on energy consumption are substantial if reverse osmosis is used. Membrane replacement costs estimates based on current 4 in. membrane module prices ($45-50/m2) and expected membrane lifetime (2-3 years) are $900-1400 per year. Capital costs of RO units installed in LEK Pharmaceuticals (similar configurations, various sizes) were used to estimate the cost of the 60 m2 RO unit and the estimation was $70 000-80 000 (Figure 7). Cost of a single-stage stage vacuum evaporator (see Table 1) of same capacity was estimated to $150 000-200 000. The final purchasing price of the full-scale RO was $115 000 (including membranes, the on-site delivery costs and all the qualifications but excluding the installation costs and the costs of auxiliaries). Compared to the purchasing price, the estimate given above is by 40% lower, which is still within reasonable limits for such estimates.6 All costs are summarized in the Table 2. 4.4. Operation of the Full-Scale Reverse Osmosis Unit. After the installation, the full-scale unit was first validated and the validation showed that reverse osmosis works in accordance with the requirements,
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Figure 7. Logaritmic plot of capital cost of RO unit and capital cost per unit membrane area as a function of membrane area. The slope of the solid line is equal to the scale factor. A similar scale factor (0.52) is reported in the literature13 for small ready-made RO systems (0.02-380 m3 permeate/day). For large RO systems (>380 m3 permeate/day) scale factors usually range between 0.75 and 1.0.13
Figure 8. Typical plant data.
Table 1. Single-Stage Vacuum Evaporator Dataa evaporator capacity (kg water/h) system pressure (mbar) temperature of heating media (steam, 3 bar) (°C) temperature of cooling media (Tin/Tout) (°C) reboiler surface area (m2) condenser surface area (m2)
1500 100 150 2/12 52 25
a Designed by means of a process simulator Chemcad 5.2 (Chemstations Inc., Houston, TX).
Table 2. Summarized Cost Data Reverse Osmosis cost of concentration pilot-scale RO ($/kg) full-scale RO based on installed power ($/kg) based on recorded plant data ($/kg) other costs estimated membrane replacement costs ($/year) estimated full-scale RO capital costs (k$) full-scale RO purchasing price (k$) Evaporation estimated cost of evaporation ($/kg) estimated evaporator capital costsa (k$) a
0.23 0.33 0.16 900-1500 70-80 115 6.8 150-200
Excluding the vacuum system.
regarding the permeate flow-rate, cycle process time and rejections of solutes. The full-scale unit is operated in a batch, constantflux mode with a feed pressure of 15-30 bar and a permeate flux of 20-30 L/hm2. The process runs according to the planned time schedule shown in Figure 6. An example of a typical time course of different variables during a full-scale operation is shown in Figure 8. Because the permeate is returned back to the chromatography where it is used again for the preparation of a mobile phase (Figure 1), it is important to know the rejections of acetonitrile and ammonia. The rejections determined during validation runs are ∼10% for acetonitrile and ∼48% for ammonia (Figures 9 and 10). Acetonitrile (CH3CN) is a small (41 g/mol), uncharged molecule and as a consequence it is not rejected by the membrane.10 Incomplete ammonia rejection by RO membranes is also reported in the literature15-20 and explained with a pH dependent equilibrium between
Figure 9. Concentration of acetonitrile in the permeate as a function of the concentration of the acetonitrile in the concentrate. Summarized results of two full-scale runs. (Filmtec BW30-4040 4 in. spiral-wound RO membrane, permeate flux 15-30 L/hm2, permeate recovery 15-20%, 20-30 °C, pH 10.5.)
ammonia (NH3) and ammonium ions (NH4+). At high pH the equilibrium is shifted toward uncharged ammonia, which is difficult to reject. Rejection coefficients for all the three solutes are shown in Table 3. For acetonitrile and ammonia, the values from literature, determined with different membranes and under similar conditions, are also reported for comparison. Reuse of the permeate results in a considerable saving in the amounts of fresh technological water, since approximately 90% of the water is recycled. Significant is also the environmental effect, because more than 97% of toxic acetonitrile is recycled and thus kept in a closed cycle. The remaining 3% are stripped from the mother liquor, which is left as a waste after lisinopril precipitation and filtration. The full-scale RO has three high-pressure feed pumps with total installed power 16.5 kW. If we assume the maximum power use over the whole 3 h batch cycle, then the energy consumption is 49.5 kWh. Consequently, costs are $2.7 per batch ($0.33/kg lisinopril). Average energy consumption calculated from a recorded plant data, taking into account pump efficiency of 55%, is 25.1 kWh or $0.77 per batch ($0.16/kg lisinopril). Energy consumption and related costs are therefore in
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Figure 10. Concentration of ammonia in the permeate as a function of the concentration of the ammonia in the concentrate. Summarized results of four full-scale experiments. (Filmtec BW304040 4 in. spiral- wound RO membrane, permeate flux 15-30 L/hm2, permeate recovery 15-20%, 20-30 °C, pH 10.5.) Table 3. Determined Rejection Coefficients and Comparison with Literature Data solute lisinopril acetonitrile ammonia
reference
rejection (%)
worka
this this worka Riley20 b this worka Lee et al.15 c Koyuncu et al.16 d Bo´dalo-Santoyo et al.17 Naaktgeboren et al.18 f
e
>99 10 >25 48 82 5-60 9.4 40
a Filmtec BW30-4040*, pH ) 10.5, 1-4 g/L ammonia, 1-12 g/L acetonitrile, 2-100 g/L lisinopril**. b PA-300*, pH ) 6.3, 425 ppm acetonitrile**. c Hydranautics ESPA*, pH ) 9.5, 3400 ppm (NH4)2CO3**. d Filmtec BW30-2540*, pH ) 9.5-10.5, 1-186 ppm NH4+ **. e Filmtec HR95PP*, pH ) 9.8, 700 ppm (NH4)2SO4**. f TNO/Wavin/Stork-Wafilin WFC × 006*, pH ) 10.5, 3000 ppm NH4Cl**. *Thin-film composite membrane. **Concentration of solute in concentrate.
good agreement with the predictions based on the pilotscale results (Table 2). After 24 months of operation, plant data do not show any significant change in the performance of the membranes. The permeate flux and the rejection of lisinopril under given operating conditions have been within the same limits since the start-up. Therefore the expected membrane lifetime of 2-3 years is considered to be realistic. So far, the operation of the RO has been trouble-free. The unit demands only minimal operator attention because of a high level of automation and furthermore due to the integration into the plant-wide control system. The required maintenance is low, involving mostly only regular inspections of components and calibration of instrumentation. 5. Conclusions The purpose of this work is to show the important stages in a design of an industrial scale reverse osmosis unit for a 50-fold concentration of aqueous solution of lisinopril. The pilot plant testing showed that reverse osmosis is suitable for concentrating lisinopril. With a moderate feed pressure up to 35 bar the permeate flux 20-30 L/hm2 can be kept over the whole concentration
range. Fouling is negligible. The rejection of lisinopril is almost complete (>99%), while the rejections of acetonitrile and ammonia are ∼10% and ∼48%, respectively. According to the production process requirements and experimental data from the pilot scale unit, the required membrane surface area of the full scale unit is 60 m2. Based on the cost data of similar units which are used in LEK Pharmaceuticals, the capital cost estimate was $70 000 to $80 000. A single-stage batch vacuum evaporator, together with auxiliary equipment, costs $150 000 to $200 000. However, the price of a purchased full-scale RO was $115 000 which is 40% more than estimated, but still within reasonable limits for such estimates. Beside the investment, costs of energy and membranes replacement were also estimated. Based on the installed pump characteristics, the cost estimate for concentration with the reverse osmosis is $0.33/kg lisinopril. The actual energy consumption for concentration was also calculated, and the corresponding cost is $0.16/kg lisinopril, which implies that the installed pumps have some power reserve. On the other hand, the cost of concentration using evaporation is $6.8/kg lisinopril. Considering the lisinopril yearly production, the savings on lower energy consumption using reverse osmosis are substantial, which confirms the economical advantage of a membrane concentration process compared to a thermal. According to present prices of membrane modules, replacement cost was estimated at a maximum of $900-1500 per year, depending on the actual replacement frequency. The expected membrane lifetime in this application is 2-3 years or more. After 24 months of operation, the permeate flux and the rejection of lisinopril under demanded operating conditions have been within the same limits since the start-up. The validation results of the full-scale unit, its trouble-free operation and low associated operating costs prove that the decision on reverse osmosis is economically justified and that the scale-up based on classical engineering principles was technically successful. The described example of the use of reverse osmosis and the approach can therefore serve as a reference when similar applications are to be introduced in pharmaceutical or other industries. Acknowledgment The present work was produced at Lek Pharmaceuticals, Ljubljana, Slovenia. One part of the work was supported by the Ministry of Education, Science and Sport of the Republic of Slovenia. The authors also sincerely thank Mr. Janez Nosan of the company ABN doo, Celje, Slovenia, for his valuable comments and suggestions during the work. List of Symbols C ) concentration, kg/m3 E ) consumed energy (work done), Wh m ) mass of evaporated water, kg p ) pressure, Pa P ) power, W PR ) permeate recovery Q ) flow rate, m3/s R ) solute rejection t ) time, h ∆H ) enthalpy of evaporation, J/kg
Ind. Eng. Chem. Res., Vol. 44, No. 6, 2005 1867 Greek Letters η ) pump efficiency Subscripts ev ) evaporation f ) feed c ) concentrate p ) permeate RO ) reverse osmosis
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Received for review September 21, 2004 Revised manuscript received December 8, 2004 Accepted January 3, 2005 IE049075S
