- Email: [email protected]
Conversion from hollow fiber to spiral technology in large seawater RO systems — process design and economics
DESALINATION ELSEVIER
Desalination 156 (2003) 239-247 www.elsevier.com/locate/desal
Conversion from hollow fiber to spiral technology in large seawater RO systems - process design and economics Vashek Polasek”*, Santi Talob, Tamer ShariP “Hydranautics Europe, Zeppelinstrasse 4, 85399 Munchen-Hallbergmoos, Germany Tel. +49 (811) 99 8790; Fax +49 (811) 998 7979 “Hydranautics (Spain), CKonstitucio, 3er 5a, edtjkio Diagonal 2, 08960 Sant Just Desvern, Barcelona, Spain hTel. +34 (93) 473 1722; Fax +34 (93) 473 1485; email: [email protected] ‘Hydranautics Middle East, POB 5728, Jeddah 21431, Saudi Arabia Tel. + 966 (2) 699 2280; Fax f 966 (2) 660 6508; email: [email protected]
Received 14 February 2003; accepted 19 February 2003
Abstract
In the fast growing desalination market, Hydranautics has succeeded as the technology leader, meeting the continuous challenge of providing innovative products to the desalination industry. One of the recent challenges was to develop products, which allow recovering reject water streams from varying applications. Using Hydranautics’ products, water is reused by converting used “waste” water resources into high standard quality water to be used in various applications. For example, Hydranautics’ membranes successfully produce high quality water from municipal wastewater. Other challenges include the successful conversion of aging plant membrane populations such as hollow tine fiber (HFF) into the latest state of art spiral wound (SW) technology for the production of high quality potable water from seawater. This paper presents two case studies of membrane conversion from HFF aramid membrane to Hydranautics’ SW membrane technology. The first plant site discussed is in the Kingdom of Saudi Arabia and the second is located in Spain. Interchangeable reverse osmosis (RO) systems using SW membranes have emerged as a short/medium term objective for many RO plant owners and operators throughout the world. Conversion to SW RO technology is necessitated by the lack of availability of direct HFF membrane replacements from DuPont. In addition, the conversion from HFF membranes to SW technology helps reduce the operating costs of the plant, reduce of membrane replacements costs, and the total water cost. Keywords:
Hollow fiber; Spiral wound; Conversion; Seawater reverse osmosis; Design; Economics
*Corresponding
author.
Pre.sented at the European Conference on Desalination and the Environment: Fresh Waterfor All, Malta, 4-8 May 2003. European Desalination Societ;v International Water Association. 00 I I-91 64/031$- See front matter 0 2003 Elsevier Science B.V. All PTI: so01
l-9164(03)00346-l
rights reserved
340
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1. Introduction
2. Case studies
of water desalting using The economics reverse osmosis (RO) technology have been continuously improving. Due to the improvement of RO membrane technology, the present cost of desalting low and medium salinity brackish water using RO is competitive with conventional water supply. This is especially true if pumping over a long distance is required or if extensive water treatment has to be applied. Even though the prices of seawater membrane elements are only slightly higher than the prices of brackish membrane elements, the cost of RO desalting of seawater is significantly higher than the cost of the RO process applied to treat the brackish feed. This is primarily because seawater systems require significantly more expensive pumping equipment, and use more corrosion resistant piping materials made of expensive alloy steel. Also, due to the nature of an open-intake seawater feed and a relatively low recovery rate, the pretreatment system is significantly more expensive than that which is required in brackish RO systems. At the open-intake design, the cost of spiral wound (SW) membrane elements contributes to only 6%-8% of the total water cost produced in RO seawater system. The major water cost contribution results from the cost of process equipment and power consumption. The RO process parameter, which has the largest effect on investment and operating cost, is the permeate recovery rate. The feed flow is inversely proportional to the design recovery rate. Therefore, the recovery rate affects directly the size and cost of all process equipment and power consumption. However, in seawater RO systems, the recovery rate cannot be increased at will, as higher recovery results in higher average feed salinity, which results in higher osmotic pressure and increased, permeate salinity. In order to overcome the high osmotic pressure generated by high recovery, it is necessary to increase the feed pressure generating higher power requirements.
2. I. C’ase .vtuc+ 1: King Fuisal 5x I.515 rd/dSWROplant
Nuvul
Base
The plant was built and commissioned by Metito Riyadh in 1995 and put into operation in 1996. It is located about 20 km south of Jeddah, Saudi Arabia. The naval base chose to use RO technology after using multi-stage flush (MSF) plants (2x0.67 MGD or 2,536 m3/d) for more than 15 years. The two MSF plants were built in 1978 by the US Corps of Engineers (subcontractor is Enregenecs). This plant produces potable water for the consumption at the naval base. Both the MSF and RO plants are presently in operation. The seawater source is an open water intake from approximately I .5 km inside the Red Sea and 6 m deep. Salinity is 40,000 ppm - 41,000 mg/l, and the feed water temperature range is between 28°C and 32°C. The pretreatment consists offilter feed pumps (7 on-line with 2 on standby, the capacity is 2 IO m’/h each at 4.2 bar). There are 6 epoxy coated horizontal media filters (5 on-line and I on standby), with a capacity of 210 m’/h each. Air scouring is available with 3 blowers (each with 360 n-+/h at 0.55 bar). Coagulant used is FeCI: with a dosing rate of -3 ppm. Acid is also dosed to reduce RO feed pH down to 6.9. There were 2 12 cartridge filters/vessel. There is no continuous chlorination in place, however a provision is made for shock dosing with I-1.5 mg/l residual chorine for about IO h (once every 30 d). High-pressure pumps are lngersol Rand (2 10 m3/h at 74.8 bar each). There is no energy recovery system installed. The plant consists of 5 RO trains with a 400,000 GPD (I ,5 15 m3/d) capacity each. The plant has been built with Du Pont B- 10, 68807 Twin Permeators. Each permeator has a surface area of 9,3 18 ft2. The design was to produce permeate quality of <500 mg/l at 35% overall plant recovery. In the year 2000, two RO trains were refurbished with Hydranautics’ SWCI seawater RO
I! Polasek et al. / Desalination
membrane elements installed in Codeline pressure vessels. Water and Environmental Services Co. (W ESCO) in Jeddah refurbished the 2 RO trains. WESCO has also provided extensive engineering and pretreatment optimization support to the end user to ensure successful performance ofthe plant after the refurbishment. These two Hydranautics trains have been operating successfully since November 2000. In the two years of operation, and despite poor raw water feed quality, Hydranautics’ Train I has only been cleaned every 6 months and Train 2 every 9 months. The cleaning cycle consisted of Step 1: Sod TPP + DDBS (pH=lO); Step 2: Citric acid 2% at pH= 4 solution. ??
??
DuPont trains were operated for 5 years. A third train was partially replaced (65% of the elements). Hydranautics’ RO trains, after more than 2 years of operation, did not require any membrane replacement. After replacing the hollow fiber membranes with Hydranautics’ SW elements, the pretreatment has been optimized. The coagulant was replaced (cationic polyelectrolyte to ferric chloride) and the SD1 improved, contributing significantly to the reduction of the cleaning frequency and improvement of the overall plant utility factor. Membrane comparison and test conditions are presented in Tables I and 2. 2.2. Case stuc$2: Marbella, Spain 8 x 7,050 mi/d SWROplant The desalination plant of Marbella, built between 1995 and 1996 (Inima-Endesa), was conceived initially to produce potable water starting from seawater beach wells. It was considered in the design that the salinity of the seawater extracted from the wells would be 36.747 ppm. The design capacity of this plant is 56,400 m3/d, produced by 8 trains or 162 (144 connected) permeators, containing the Permasep model BIO, 6882TM manufactured by DuPont. Each train’s contractual production is 293,75 mj/h at conversion of 45%.
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Table 1 Membrane comparison Membrane
Hydranautics
DuPont
Model
SWCl
Surface, f? Permeate flow, GPD Rejection, % Diameter, inch
315 5,000 99.6 8
6880T twin bundle 9.318 16,000 99.55 8.5
After the initial start-up in September and November of 1996 (only train No. 1) and April 1997 (for all the trains) it was observed that all the trains worked under the design conditions, fulfilling the contractual flow. The feed conductivity was 30.000 and 35.000 pS. Since the flow ofthe wells was not enough to feed the whole plant, Decosol decided to build a seawater open intake. The salinity of the open intake is 38.750 mg/l. This resulted in use of all the permeators installed to meet the contractual flow as per original design. The change from RO feed from sea wells to open-intake followed a series of recommendations. Namely, the chlorine dosing point must be located near the open-intake of seawater (not used at present), installation of a detection system for free chlorine to eliminate potential hazard of permeator chlorination, thus destruction due to oxidation. All necessary design modifications have been made to obtain RO feed SD1 of less or equal to 3 .O. Dosing of polyelectrolyte and coagulant has been inserted into the pretreatment section of the plant. The configuration of the desalination plant after the plant update has resulted in the following design: the reception of seawater is some 360 m off the coast, and the water arrives by a collector piping of 2 m diameter until a break tank. From there, the seawater is pumped with 4 pumps (3-t 1 for reserve) at a flow velocity of 1956 m3/h. In the aspiration of these pumps there is dosing of an oxidant as sterilizer. After 2.5 km of piping, the seawater arrives at an intermediate tank. From this tank, 9 pumps with a capacity of 652 m3/h
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Table 2 Test conditions Hydranautics
SWCl
32,000 mg/l NaCl
800 psi, 56.3 bar, applied pressure 77”F, 25”C, operating temperature 10% permeate recovery 6.5-7.0 pH range
DuPont
6880T
Twin bundle
35,000 mgil NaCl
I .OOOpsi, 70.4 bar, applied pressure
77’F, 25°C operating temperature 75% permeate recovery 6.556.9 pH range
KFNB plant design
Technical summary
Membrane source Membrane model Plant capacity, m’/d Membrane surface, ft’ Total membranes surface, t? Average brine flow Average product flow Maximum feed pressure, bar Feed pH guideline Feed SDI. 15 min
DuPont, USA 6880TR 7,575 9,318 3,727,200 1 GFDil.7 Imh 0.54 GFDi0.92 Imh 82.7 ~6.9 ~4.0 recommended
Hydranautics, Japan/USA SWCl 7,575 320 256,000 14.67 GFD/24.9 Imh 7.1 GFD02.7 Imh 82.7 no guideline ~4.0 recommended
_--.
Start-up data Feed salinity, ppm Feed pressure, bar Permeate salinity, mg/l
42.295 62 415
42.295 61.5 275
Total recovery, %
35
35
(8+1 for reserve) transfer the seawater to the 16 horizontal closed sand filters. In the aspiration of these pumps the sodium metabisulphite is dosed to eliminate any oxidizer traces. The filtering speed is 11 m3/m2/h in operation and I 5 mj/m2/h during backwash cycle of one filter. The water goes out of the sand filters with a SDI of 2. After acid dosage to pH 6.9, the seawater goes to a common collector directly to the 12 vertical cartridge filters. In each cartridge filter are 69 cartridges of Pall’s High Flow type. The antiscalant dosage is injected before the cartridge filters. From the cartridge filters the seawater enters in another common collector that feeds the 9 high-pressure pumps with capacity of 652 m’/h
for reserve). Starting from this point the plant can work in a common collector or with independent lines. Each high-pressure pump feeds one RO train. The trains of RO present in the superior part a deposit of osmotic balance of 14 m’. The final start-up was initiated on I8 November 1997 and was shut down on 29 November 1997, 5 days of operation (I 20 h) of all 8 trains. SD1 during the operation was between 1 and 2.9. The overall performance was under contractual permeate flow of 294 m3/h. There were only 7 1171 permeators connected (initial set 8 i/8 1). permeate conductivity was less than 400 mg/l. In 200 I after acknowledging a future problem of no replacement of DuPont membranes, Endesa
(8+1
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decided to seek for a solution to support and guarantee the service agreement signed with DuPont. Hydranautics, the technology leader, has demonstrated excellent field data in many SWRO and BRO plants. Some of the plants include: Las Palmas 3 (55,000 m3/d SW open intake); Larnaca (54,000 m’/d SW open intake); Carboneras (123,000 m’/d SW open intake); Cartagena (65,000 m’id SW open intake); Almeria (50,000 m’id SW well) and Santa Cruz de Tenerife (20,000 m’/d SW well) among several others. Endesa decided to use the Hydranautics SW technology. The membranes chosen were Hydranautics’ SWC3 seawater RO membranes. At the end of year 2002, one RO train was refurbished with Hydranautics’ SWC3 seawater RO membrane
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elements installed in Bel Composites pressure vessels. Membrane comparison and standard test conditions are presented in Tables 3 and 4. Table 3 Membrane comparison: train 8 was replaced by 693 SW SWC3 membranes Membrane
Hydranautics
Du Pont
Model
swc3
Surface, ft2 Permeate flow, GPD Rejection, % Diameter, inch
370 5.900 99.6 8
6882TM twin bundle 11.181 19.200 99.35 8.5
Table 4 Membrane standard test conditions Hydranautics
swc3
32,000 ppm NaCl
800 psi applied pressure 77”F, 25°C operating temp 10% permeate recovery 6.5-7.0 pH range
Duoont
6882TM twin bundle
35,000 ppm NaCl
1.OOOpsi applied pressure
Marbella plant
Design summary
Membrane source Membrane model Plant capacity, m’id Membrane surface, ft’ Total membranes surface, ft2 Average brine flow Average product flow Maximum feed press, bar Feed pH guideline Feed SDI, 15 min
DuPont, USA 6882TM 56,400 11,181 28,981,152 0.98 GFDil.68 Imh 0.53 GFD/0.90 Imh 82.7 ~6.9 <3.0
77”F, 2YC, operating temperature 75% permeate recovery 6.5-6.9 pH range
Hydranautics, Japan/USA swc3 56,400 370 2,05 1,280 13.46 GFDi22.88 Imh 7.25 GFDi12.3 Imh 82.7 No guideline ~4.0 recommended
Start-up data Feed salinity, mg/l Feed pressure, bar Permeate salinity, mgil Total recovery, %
38,750 68.5 275 45
38,750 60.8 231 45
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3. Process economics set
1” Polasek et al
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membrane replacement
Direct comparison of the HFF replacement bundles and new plant modification to SW membranes, the ratios of savings is expressed in Table 5 and Table 6 for respectively the KFNB and Marbella plants. Tables 7 and 8 indicate the approximate price/ cost of the replacement membrane elements once the entire rack is converted from HFF to SWC3 membranes.
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Table 5 Comparison for KFNB plant, 3,000 m”/d, 2 trains ..~ _____~ DuPont Hydranautics Ratio savings Quantity PV Totals, $
8Oxtwin bundles Included 800,000
32OxSWCl 64, 5 el/PV 709,800”
0.89
--
*Including engineering cost of plant modification
Table 6 Comparison for Marbella plant, 56,400 m3/d, I out of 8
4. Other plant design considerations Regarding the recovery rate, the conversion has a major impact on the economics of the seawater RO process. The size of all process equipment, which is determined according to feed or concentrate flow, will decrease with increased recovery rate. This applies to the size of the feed water supply system, and power consumption of intake pumps. The sizing of pretreatment equipment including storage tanks, booster pumps, filtration equipment and chemical dosing systems is determined according to the feed flow. The same considerations apply to sizing of concentrate piping and of the outfall facility. The design permeate flux rate affects the number of membrane elements installed, number of pressure vessels, manifold connections and size of membrane skid. The effect ofthe below parameters on investment cost will be examined in an example for a 6 MGD (22,700 m’/d) system operating on Mediterranean seawater from an open intake. The cost estimation of the conventional reference design is based mainly on the data developed by G. Leitner [2], P. Shields and I. Moth [3]. Table 9 contains a comparison of equipment cost (including 35% indirect cost) of the basic design and a system operating at high recovery and high permeate flux, (HRF design). The basic design consists operation at 45% permeate recovery and 8 GFD flux rate. The HRF design operates at 55% recovery and flux rate of 11 GFD.
239--217
Quantity PV Totals, $
DuPont
Hydranautics
324xhvin bundles included I ,134,ooo
693 SWC3 99, 7 904,000*
Ratio savings
0.80
*Including engineering cost of plant modification
Table 7 Marbella plant, 56,400 mj/d, total membrane replacement price ______~_...~~~~. _~___~~__ Membrane model Ratio
Quantity in 8 RO trains Totals, $
B-10,6882TM 2.592
SWC3 5.544
12,960,OOO
3,149,200
0.24
Table 8 KFNB plant, 3,000 m’id, total membrane replacement price ____Quantity in 5 RO trains Totals, $
Membrane model B- 10,688OT 320
SWCI 800
2,000,000
495,000
Ratio
_~
0.25
The combined effect of higher recovery and higher flux rate result in significant reduction of the investment cost components. These savings are directly transferred to the total water cost.
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’Desalination
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Table 9 Effect of recovery rate and permeate flux on cost of 6 MGD, 22,700 m’id, RO seawater system, in 1000’s US$ Investment cost component
Reference design: 45% recovery, 8 GFD flux
HRF design: 55% recovery, 11 GFD flux
Equipment cost change, %
Intake and outfall Pretreatment Membranes Process equipment Product water treatment Site development Total investment Soecific investment. Wm’id
940 5,000 2,000 16,050 400 670 25,060 1.104
830 4,390 1,450 13,650 400 640 21,360 940
-11.7 -12.2 -27.5 -15.0 0 -4.5 -14.8 -14.8
However, the higher recovery rate and higher permeate flux require a higher feed pressure. Therefore, specific power consumption will be higher. The increase in feed pressure with recovery rate at a given permeate flux is due to the increase in the average feed salinity and osmotic pressure. For pressure calculations, a 20% flux decline has been assumed due to fouling and compaction. Fig. 1 contains values of specific power consumption for increasing recovery rate. As shown in Fig. 1, for a given flux rate, the minimum specific power consumption corresponds to a recovery rate of about 50%. Changing the design parameters from a 45% recovery rate and 8 GFD flux to a 55% recovery rate and 11 GFD results in increase of specific power consumption of the highpressure pump from 4.2 kWh/m’ to 4.6 kWh/m3. This difference of 0.4 kWh/m3 at 18°C will de-
crease slightly at a higher feed water temperature. At a feed water temperature of 28°C (the high temperature limit for this evaluation) this difference will be about 0.3 kWh/m3. Table 10 contains results of comparison of water cost components for operation at standard conditions with projected costs for operation at high flux and a high recovery rate. For the calculations of water cost, the following economic parameters were used: interest rate: 8%, plant life: 20 years, replacement cost of 8 incn seawater membrane element: $800; maintenance: 3% of capital; power cost: $O.O6/kWh; plant load factor 95%. The cost of chemicals is based on the use of chlorine, inorganic flocculents, polymer, sodium bisulfate, and scale inhibitor in the pretreatment system. Operation at the high recovery rate and increased permeate flux improves process economics. However, maintaining stable membrane
Fig. I. Prqjected power consumption for Mediterranean, Marbella, feed, 18°C.
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Table IO Effect of recovery rate and permeate flux on total water cost in RO seawater system, in US$ Water cost components
Ref design: 45% recovery,
8 GFD flux Capital cost Membrane replacement Maintenance Power consumption Chemicals and cartridge filters Labor Total water cost, $/In’
0.320 0.050 0.095 0.252 0.060 0.050 0.827
performance at the higher flux rate requires a significant improvement in feed water quality compared to the quality obtained from conventional pretreatment applied to surface seawater. The combined effect of lower investment and improved operating conditions results in an 8% decrease of total water cost. An increase in recovery rate and permeate flux in seawater systems can improve the economics of the desalting process. Implementation of high recovery, high flux operation requires better quality of the feed water. New capillary membrane technology used as a pretreatment step has the potential to improve feed water quality, which will enable seawater membranes to operate at a higher flux rate. The new technology has demonstrated reliable operation at a variety of operating conditions. It is cost competitive with conventional pretreatment technology, and will result in higher reliability and better overall economics of RO seawater desalting. The combined savings due to lower investment and lower operating cost coupled with the ability to optimize system-operating conditions due to better feed water quality should result in about a 10% reduction in total water cost. 5. Conclusions Based on the above study of the 2 SWRO plants, Hydranautics’ SWC technology contributes greatly not only to the significant plant investment cost reduction but also to the improvement
HRF design: 55% recovery, 11 GFD flux
Water cost change, %
0.275 0.037
-4.0 -26.0 -12.0
0.084 0.276 0.050 0.040 0.762
+9.5 -17.0 -20.0 -7.9
of the water production cost. This is achieved by the unique design of the SW elements, which provide the client with excellent performance at lower feed pressure. The cleaning efficiency of the SW elements is significantly reduced. This is quite obvious taking into consideration that membrane surface of the SW elements is 30 times less than HFF modules. Based on this, the overall plant utility factor is significantly improved when using the SW elements. The Hydranautics’ SW elements do not require costly and time-consuming membrane post treatment in order to maintain the salt rejection performance during the membrane lifetime. Thanks to vast experience and the know-how of liydranautics, the plant builders as well as the end users benefit from trouble free and economic plant operation. The pricing ofthe Hydranautics’ SWC membrane elements is approximately one-quarter that of the HFF replacement membranes. This leads to a very significant improvement of the plant economics. Thanks to Hydranautics experience and expertise in large plant membrane substitution, the plant builders can convert from an HFF plant to Hydranautics SWC membranes at significantly lowet investment cost. Acknowledgments This paper has been prepared with kind agreement of Wesco in Saudi Arabia and Endesa-
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Decosol in Spain. We would like to extend our thanks to these companies for providing all information to complete this paper.
References
[‘I Endesa-Decosol plant operational data. PI Wesco plant operational data. c31 M. Wilf and K. Klinko, Performance of commercial seawater membranes. Desalination, 96 (1994) 4% 478. L41 G Leitner, Cost of seawater desalination in real terms, 1979 through 1989, and projections for 1999. Desalination 76 (1989) 201-213. VI X. Zhu and M. Elimelech, Fouling of reverse osmosis membranes by aluminum oxide colloids. J. Environ. Eng., Dec. ( 1995) 884-892. S. Hong and M. Elimelech, Fouling of nanofiltration [61 membranes by natural organic matter. Proc. 1996 ADA Conf., Monterey, 1996, pp. 717-727. [71 E. Ebrahim, M. Abdel-Jawad and M. Safar. Convent ional pretreatment for the Doha reverse osmosis plant:
247
technical and economical assessment. Proc. ADA 1994 Biennial Conf., Pal Beach, Florida, 1 (1994) 149-163. Y. Ayash, H. Imai, T. Yamada, T. Fakuda, I’. Yanaga PI and T. Taniyama, Performance of reverse osmosis membranes in Jeddah Phase I plant. Desalination, 96 (1994) 215-224. [91 S. Kremen, M. Wilf andP. Lange, Operating results and economics of single stage and two stage large size sea water RO systems. Desalination, 82 (1991) 15-30. DOI G. Tanny, R. D’Agostino and M. Wilf, Membrane microfiltration as a pretreatment for seawater reverse osmosis. Proc. 7th International Symposium on Fresh Water from the Sea, Amsterdam, 2 ( 1980) 307-3 17. B. [I 11 Ericson and B. Hallmans, Membrane filtration as a pretreatment method - a cost comparison, Proc. IDA World Conf. on Desalination and Water Reuse, Washington, Vol 2, 1991. [ 121 M. Miller, M. Silbemagel, T. Kuepper and W. Vamava, Testing prefiltration for the military ROWPU. Proc. ADA Conf., Monterey, 1996, pp. 467-485.
Desalination 156 (2003) 239-247 www.elsevier.com/locate/desal
Conversion from hollow fiber to spiral technology in large seawater RO systems - process design and economics Vashek Polasek”*, Santi Talob, Tamer ShariP “Hydranautics Europe, Zeppelinstrasse 4, 85399 Munchen-Hallbergmoos, Germany Tel. +49 (811) 99 8790; Fax +49 (811) 998 7979 “Hydranautics (Spain), CKonstitucio, 3er 5a, edtjkio Diagonal 2, 08960 Sant Just Desvern, Barcelona, Spain hTel. +34 (93) 473 1722; Fax +34 (93) 473 1485; email: [email protected] ‘Hydranautics Middle East, POB 5728, Jeddah 21431, Saudi Arabia Tel. + 966 (2) 699 2280; Fax f 966 (2) 660 6508; email: [email protected]
Received 14 February 2003; accepted 19 February 2003
Abstract
In the fast growing desalination market, Hydranautics has succeeded as the technology leader, meeting the continuous challenge of providing innovative products to the desalination industry. One of the recent challenges was to develop products, which allow recovering reject water streams from varying applications. Using Hydranautics’ products, water is reused by converting used “waste” water resources into high standard quality water to be used in various applications. For example, Hydranautics’ membranes successfully produce high quality water from municipal wastewater. Other challenges include the successful conversion of aging plant membrane populations such as hollow tine fiber (HFF) into the latest state of art spiral wound (SW) technology for the production of high quality potable water from seawater. This paper presents two case studies of membrane conversion from HFF aramid membrane to Hydranautics’ SW membrane technology. The first plant site discussed is in the Kingdom of Saudi Arabia and the second is located in Spain. Interchangeable reverse osmosis (RO) systems using SW membranes have emerged as a short/medium term objective for many RO plant owners and operators throughout the world. Conversion to SW RO technology is necessitated by the lack of availability of direct HFF membrane replacements from DuPont. In addition, the conversion from HFF membranes to SW technology helps reduce the operating costs of the plant, reduce of membrane replacements costs, and the total water cost. Keywords:
Hollow fiber; Spiral wound; Conversion; Seawater reverse osmosis; Design; Economics
*Corresponding
author.
Pre.sented at the European Conference on Desalination and the Environment: Fresh Waterfor All, Malta, 4-8 May 2003. European Desalination Societ;v International Water Association. 00 I I-91 64/031$- See front matter 0 2003 Elsevier Science B.V. All PTI: so01
l-9164(03)00346-l
rights reserved
340
I: Polasek
ef ul.
Desalination
I.56 (2003)
239.-2-17
1. Introduction
2. Case studies
of water desalting using The economics reverse osmosis (RO) technology have been continuously improving. Due to the improvement of RO membrane technology, the present cost of desalting low and medium salinity brackish water using RO is competitive with conventional water supply. This is especially true if pumping over a long distance is required or if extensive water treatment has to be applied. Even though the prices of seawater membrane elements are only slightly higher than the prices of brackish membrane elements, the cost of RO desalting of seawater is significantly higher than the cost of the RO process applied to treat the brackish feed. This is primarily because seawater systems require significantly more expensive pumping equipment, and use more corrosion resistant piping materials made of expensive alloy steel. Also, due to the nature of an open-intake seawater feed and a relatively low recovery rate, the pretreatment system is significantly more expensive than that which is required in brackish RO systems. At the open-intake design, the cost of spiral wound (SW) membrane elements contributes to only 6%-8% of the total water cost produced in RO seawater system. The major water cost contribution results from the cost of process equipment and power consumption. The RO process parameter, which has the largest effect on investment and operating cost, is the permeate recovery rate. The feed flow is inversely proportional to the design recovery rate. Therefore, the recovery rate affects directly the size and cost of all process equipment and power consumption. However, in seawater RO systems, the recovery rate cannot be increased at will, as higher recovery results in higher average feed salinity, which results in higher osmotic pressure and increased, permeate salinity. In order to overcome the high osmotic pressure generated by high recovery, it is necessary to increase the feed pressure generating higher power requirements.
2. I. C’ase .vtuc+ 1: King Fuisal 5x I.515 rd/dSWROplant
Nuvul
Base
The plant was built and commissioned by Metito Riyadh in 1995 and put into operation in 1996. It is located about 20 km south of Jeddah, Saudi Arabia. The naval base chose to use RO technology after using multi-stage flush (MSF) plants (2x0.67 MGD or 2,536 m3/d) for more than 15 years. The two MSF plants were built in 1978 by the US Corps of Engineers (subcontractor is Enregenecs). This plant produces potable water for the consumption at the naval base. Both the MSF and RO plants are presently in operation. The seawater source is an open water intake from approximately I .5 km inside the Red Sea and 6 m deep. Salinity is 40,000 ppm - 41,000 mg/l, and the feed water temperature range is between 28°C and 32°C. The pretreatment consists offilter feed pumps (7 on-line with 2 on standby, the capacity is 2 IO m’/h each at 4.2 bar). There are 6 epoxy coated horizontal media filters (5 on-line and I on standby), with a capacity of 210 m’/h each. Air scouring is available with 3 blowers (each with 360 n-+/h at 0.55 bar). Coagulant used is FeCI: with a dosing rate of -3 ppm. Acid is also dosed to reduce RO feed pH down to 6.9. There were 2 12 cartridge filters/vessel. There is no continuous chlorination in place, however a provision is made for shock dosing with I-1.5 mg/l residual chorine for about IO h (once every 30 d). High-pressure pumps are lngersol Rand (2 10 m3/h at 74.8 bar each). There is no energy recovery system installed. The plant consists of 5 RO trains with a 400,000 GPD (I ,5 15 m3/d) capacity each. The plant has been built with Du Pont B- 10, 68807 Twin Permeators. Each permeator has a surface area of 9,3 18 ft2. The design was to produce permeate quality of <500 mg/l at 35% overall plant recovery. In the year 2000, two RO trains were refurbished with Hydranautics’ SWCI seawater RO
I! Polasek et al. / Desalination
membrane elements installed in Codeline pressure vessels. Water and Environmental Services Co. (W ESCO) in Jeddah refurbished the 2 RO trains. WESCO has also provided extensive engineering and pretreatment optimization support to the end user to ensure successful performance ofthe plant after the refurbishment. These two Hydranautics trains have been operating successfully since November 2000. In the two years of operation, and despite poor raw water feed quality, Hydranautics’ Train I has only been cleaned every 6 months and Train 2 every 9 months. The cleaning cycle consisted of Step 1: Sod TPP + DDBS (pH=lO); Step 2: Citric acid 2% at pH= 4 solution. ??
??
DuPont trains were operated for 5 years. A third train was partially replaced (65% of the elements). Hydranautics’ RO trains, after more than 2 years of operation, did not require any membrane replacement. After replacing the hollow fiber membranes with Hydranautics’ SW elements, the pretreatment has been optimized. The coagulant was replaced (cationic polyelectrolyte to ferric chloride) and the SD1 improved, contributing significantly to the reduction of the cleaning frequency and improvement of the overall plant utility factor. Membrane comparison and test conditions are presented in Tables I and 2. 2.2. Case stuc$2: Marbella, Spain 8 x 7,050 mi/d SWROplant The desalination plant of Marbella, built between 1995 and 1996 (Inima-Endesa), was conceived initially to produce potable water starting from seawater beach wells. It was considered in the design that the salinity of the seawater extracted from the wells would be 36.747 ppm. The design capacity of this plant is 56,400 m3/d, produced by 8 trains or 162 (144 connected) permeators, containing the Permasep model BIO, 6882TM manufactured by DuPont. Each train’s contractual production is 293,75 mj/h at conversion of 45%.
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Table 1 Membrane comparison Membrane
Hydranautics
DuPont
Model
SWCl
Surface, f? Permeate flow, GPD Rejection, % Diameter, inch
315 5,000 99.6 8
6880T twin bundle 9.318 16,000 99.55 8.5
After the initial start-up in September and November of 1996 (only train No. 1) and April 1997 (for all the trains) it was observed that all the trains worked under the design conditions, fulfilling the contractual flow. The feed conductivity was 30.000 and 35.000 pS. Since the flow ofthe wells was not enough to feed the whole plant, Decosol decided to build a seawater open intake. The salinity of the open intake is 38.750 mg/l. This resulted in use of all the permeators installed to meet the contractual flow as per original design. The change from RO feed from sea wells to open-intake followed a series of recommendations. Namely, the chlorine dosing point must be located near the open-intake of seawater (not used at present), installation of a detection system for free chlorine to eliminate potential hazard of permeator chlorination, thus destruction due to oxidation. All necessary design modifications have been made to obtain RO feed SD1 of less or equal to 3 .O. Dosing of polyelectrolyte and coagulant has been inserted into the pretreatment section of the plant. The configuration of the desalination plant after the plant update has resulted in the following design: the reception of seawater is some 360 m off the coast, and the water arrives by a collector piping of 2 m diameter until a break tank. From there, the seawater is pumped with 4 pumps (3-t 1 for reserve) at a flow velocity of 1956 m3/h. In the aspiration of these pumps there is dosing of an oxidant as sterilizer. After 2.5 km of piping, the seawater arrives at an intermediate tank. From this tank, 9 pumps with a capacity of 652 m3/h
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Table 2 Test conditions Hydranautics
SWCl
32,000 mg/l NaCl
800 psi, 56.3 bar, applied pressure 77”F, 25”C, operating temperature 10% permeate recovery 6.5-7.0 pH range
DuPont
6880T
Twin bundle
35,000 mgil NaCl
I .OOOpsi, 70.4 bar, applied pressure
77’F, 25°C operating temperature 75% permeate recovery 6.556.9 pH range
KFNB plant design
Technical summary
Membrane source Membrane model Plant capacity, m’/d Membrane surface, ft’ Total membranes surface, t? Average brine flow Average product flow Maximum feed pressure, bar Feed pH guideline Feed SDI. 15 min
DuPont, USA 6880TR 7,575 9,318 3,727,200 1 GFDil.7 Imh 0.54 GFDi0.92 Imh 82.7 ~6.9 ~4.0 recommended
Hydranautics, Japan/USA SWCl 7,575 320 256,000 14.67 GFD/24.9 Imh 7.1 GFD02.7 Imh 82.7 no guideline ~4.0 recommended
_--.
Start-up data Feed salinity, ppm Feed pressure, bar Permeate salinity, mg/l
42.295 62 415
42.295 61.5 275
Total recovery, %
35
35
(8+1 for reserve) transfer the seawater to the 16 horizontal closed sand filters. In the aspiration of these pumps the sodium metabisulphite is dosed to eliminate any oxidizer traces. The filtering speed is 11 m3/m2/h in operation and I 5 mj/m2/h during backwash cycle of one filter. The water goes out of the sand filters with a SDI of 2. After acid dosage to pH 6.9, the seawater goes to a common collector directly to the 12 vertical cartridge filters. In each cartridge filter are 69 cartridges of Pall’s High Flow type. The antiscalant dosage is injected before the cartridge filters. From the cartridge filters the seawater enters in another common collector that feeds the 9 high-pressure pumps with capacity of 652 m’/h
for reserve). Starting from this point the plant can work in a common collector or with independent lines. Each high-pressure pump feeds one RO train. The trains of RO present in the superior part a deposit of osmotic balance of 14 m’. The final start-up was initiated on I8 November 1997 and was shut down on 29 November 1997, 5 days of operation (I 20 h) of all 8 trains. SD1 during the operation was between 1 and 2.9. The overall performance was under contractual permeate flow of 294 m3/h. There were only 7 1171 permeators connected (initial set 8 i/8 1). permeate conductivity was less than 400 mg/l. In 200 I after acknowledging a future problem of no replacement of DuPont membranes, Endesa
(8+1
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decided to seek for a solution to support and guarantee the service agreement signed with DuPont. Hydranautics, the technology leader, has demonstrated excellent field data in many SWRO and BRO plants. Some of the plants include: Las Palmas 3 (55,000 m3/d SW open intake); Larnaca (54,000 m’/d SW open intake); Carboneras (123,000 m’/d SW open intake); Cartagena (65,000 m’id SW open intake); Almeria (50,000 m’id SW well) and Santa Cruz de Tenerife (20,000 m’/d SW well) among several others. Endesa decided to use the Hydranautics SW technology. The membranes chosen were Hydranautics’ SWC3 seawater RO membranes. At the end of year 2002, one RO train was refurbished with Hydranautics’ SWC3 seawater RO membrane
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elements installed in Bel Composites pressure vessels. Membrane comparison and standard test conditions are presented in Tables 3 and 4. Table 3 Membrane comparison: train 8 was replaced by 693 SW SWC3 membranes Membrane
Hydranautics
Du Pont
Model
swc3
Surface, ft2 Permeate flow, GPD Rejection, % Diameter, inch
370 5.900 99.6 8
6882TM twin bundle 11.181 19.200 99.35 8.5
Table 4 Membrane standard test conditions Hydranautics
swc3
32,000 ppm NaCl
800 psi applied pressure 77”F, 25°C operating temp 10% permeate recovery 6.5-7.0 pH range
Duoont
6882TM twin bundle
35,000 ppm NaCl
1.OOOpsi applied pressure
Marbella plant
Design summary
Membrane source Membrane model Plant capacity, m’id Membrane surface, ft’ Total membranes surface, ft2 Average brine flow Average product flow Maximum feed press, bar Feed pH guideline Feed SDI, 15 min
DuPont, USA 6882TM 56,400 11,181 28,981,152 0.98 GFDil.68 Imh 0.53 GFD/0.90 Imh 82.7 ~6.9 <3.0
77”F, 2YC, operating temperature 75% permeate recovery 6.5-6.9 pH range
Hydranautics, Japan/USA swc3 56,400 370 2,05 1,280 13.46 GFDi22.88 Imh 7.25 GFDi12.3 Imh 82.7 No guideline ~4.0 recommended
Start-up data Feed salinity, mg/l Feed pressure, bar Permeate salinity, mgil Total recovery, %
38,750 68.5 275 45
38,750 60.8 231 45
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3. Process economics set
1” Polasek et al
Desalination
membrane replacement
Direct comparison of the HFF replacement bundles and new plant modification to SW membranes, the ratios of savings is expressed in Table 5 and Table 6 for respectively the KFNB and Marbella plants. Tables 7 and 8 indicate the approximate price/ cost of the replacement membrane elements once the entire rack is converted from HFF to SWC3 membranes.
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Table 5 Comparison for KFNB plant, 3,000 m”/d, 2 trains ..~ _____~ DuPont Hydranautics Ratio savings Quantity PV Totals, $
8Oxtwin bundles Included 800,000
32OxSWCl 64, 5 el/PV 709,800”
0.89
--
*Including engineering cost of plant modification
Table 6 Comparison for Marbella plant, 56,400 m3/d, I out of 8
4. Other plant design considerations Regarding the recovery rate, the conversion has a major impact on the economics of the seawater RO process. The size of all process equipment, which is determined according to feed or concentrate flow, will decrease with increased recovery rate. This applies to the size of the feed water supply system, and power consumption of intake pumps. The sizing of pretreatment equipment including storage tanks, booster pumps, filtration equipment and chemical dosing systems is determined according to the feed flow. The same considerations apply to sizing of concentrate piping and of the outfall facility. The design permeate flux rate affects the number of membrane elements installed, number of pressure vessels, manifold connections and size of membrane skid. The effect ofthe below parameters on investment cost will be examined in an example for a 6 MGD (22,700 m’/d) system operating on Mediterranean seawater from an open intake. The cost estimation of the conventional reference design is based mainly on the data developed by G. Leitner [2], P. Shields and I. Moth [3]. Table 9 contains a comparison of equipment cost (including 35% indirect cost) of the basic design and a system operating at high recovery and high permeate flux, (HRF design). The basic design consists operation at 45% permeate recovery and 8 GFD flux rate. The HRF design operates at 55% recovery and flux rate of 11 GFD.
239--217
Quantity PV Totals, $
DuPont
Hydranautics
324xhvin bundles included I ,134,ooo
693 SWC3 99, 7 904,000*
Ratio savings
0.80
*Including engineering cost of plant modification
Table 7 Marbella plant, 56,400 mj/d, total membrane replacement price ______~_...~~~~. _~___~~__ Membrane model Ratio
Quantity in 8 RO trains Totals, $
B-10,6882TM 2.592
SWC3 5.544
12,960,OOO
3,149,200
0.24
Table 8 KFNB plant, 3,000 m’id, total membrane replacement price ____Quantity in 5 RO trains Totals, $
Membrane model B- 10,688OT 320
SWCI 800
2,000,000
495,000
Ratio
_~
0.25
The combined effect of higher recovery and higher flux rate result in significant reduction of the investment cost components. These savings are directly transferred to the total water cost.
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’Desalination
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1.56 (2003) 239-247
Table 9 Effect of recovery rate and permeate flux on cost of 6 MGD, 22,700 m’id, RO seawater system, in 1000’s US$ Investment cost component
Reference design: 45% recovery, 8 GFD flux
HRF design: 55% recovery, 11 GFD flux
Equipment cost change, %
Intake and outfall Pretreatment Membranes Process equipment Product water treatment Site development Total investment Soecific investment. Wm’id
940 5,000 2,000 16,050 400 670 25,060 1.104
830 4,390 1,450 13,650 400 640 21,360 940
-11.7 -12.2 -27.5 -15.0 0 -4.5 -14.8 -14.8
However, the higher recovery rate and higher permeate flux require a higher feed pressure. Therefore, specific power consumption will be higher. The increase in feed pressure with recovery rate at a given permeate flux is due to the increase in the average feed salinity and osmotic pressure. For pressure calculations, a 20% flux decline has been assumed due to fouling and compaction. Fig. 1 contains values of specific power consumption for increasing recovery rate. As shown in Fig. 1, for a given flux rate, the minimum specific power consumption corresponds to a recovery rate of about 50%. Changing the design parameters from a 45% recovery rate and 8 GFD flux to a 55% recovery rate and 11 GFD results in increase of specific power consumption of the highpressure pump from 4.2 kWh/m’ to 4.6 kWh/m3. This difference of 0.4 kWh/m3 at 18°C will de-
crease slightly at a higher feed water temperature. At a feed water temperature of 28°C (the high temperature limit for this evaluation) this difference will be about 0.3 kWh/m3. Table 10 contains results of comparison of water cost components for operation at standard conditions with projected costs for operation at high flux and a high recovery rate. For the calculations of water cost, the following economic parameters were used: interest rate: 8%, plant life: 20 years, replacement cost of 8 incn seawater membrane element: $800; maintenance: 3% of capital; power cost: $O.O6/kWh; plant load factor 95%. The cost of chemicals is based on the use of chlorine, inorganic flocculents, polymer, sodium bisulfate, and scale inhibitor in the pretreatment system. Operation at the high recovery rate and increased permeate flux improves process economics. However, maintaining stable membrane
Fig. I. Prqjected power consumption for Mediterranean, Marbella, feed, 18°C.
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Table IO Effect of recovery rate and permeate flux on total water cost in RO seawater system, in US$ Water cost components
Ref design: 45% recovery,
8 GFD flux Capital cost Membrane replacement Maintenance Power consumption Chemicals and cartridge filters Labor Total water cost, $/In’
0.320 0.050 0.095 0.252 0.060 0.050 0.827
performance at the higher flux rate requires a significant improvement in feed water quality compared to the quality obtained from conventional pretreatment applied to surface seawater. The combined effect of lower investment and improved operating conditions results in an 8% decrease of total water cost. An increase in recovery rate and permeate flux in seawater systems can improve the economics of the desalting process. Implementation of high recovery, high flux operation requires better quality of the feed water. New capillary membrane technology used as a pretreatment step has the potential to improve feed water quality, which will enable seawater membranes to operate at a higher flux rate. The new technology has demonstrated reliable operation at a variety of operating conditions. It is cost competitive with conventional pretreatment technology, and will result in higher reliability and better overall economics of RO seawater desalting. The combined savings due to lower investment and lower operating cost coupled with the ability to optimize system-operating conditions due to better feed water quality should result in about a 10% reduction in total water cost. 5. Conclusions Based on the above study of the 2 SWRO plants, Hydranautics’ SWC technology contributes greatly not only to the significant plant investment cost reduction but also to the improvement
HRF design: 55% recovery, 11 GFD flux
Water cost change, %
0.275 0.037
-4.0 -26.0 -12.0
0.084 0.276 0.050 0.040 0.762
+9.5 -17.0 -20.0 -7.9
of the water production cost. This is achieved by the unique design of the SW elements, which provide the client with excellent performance at lower feed pressure. The cleaning efficiency of the SW elements is significantly reduced. This is quite obvious taking into consideration that membrane surface of the SW elements is 30 times less than HFF modules. Based on this, the overall plant utility factor is significantly improved when using the SW elements. The Hydranautics’ SW elements do not require costly and time-consuming membrane post treatment in order to maintain the salt rejection performance during the membrane lifetime. Thanks to vast experience and the know-how of liydranautics, the plant builders as well as the end users benefit from trouble free and economic plant operation. The pricing ofthe Hydranautics’ SWC membrane elements is approximately one-quarter that of the HFF replacement membranes. This leads to a very significant improvement of the plant economics. Thanks to Hydranautics experience and expertise in large plant membrane substitution, the plant builders can convert from an HFF plant to Hydranautics SWC membranes at significantly lowet investment cost. Acknowledgments This paper has been prepared with kind agreement of Wesco in Saudi Arabia and Endesa-
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Decosol in Spain. We would like to extend our thanks to these companies for providing all information to complete this paper.
References
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