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Pilot Plant Demonstration of Effluent Desalination by Electrodialysis at the Western Treatment Plant, Werribee Russell J Taylor and Nigel B Goodman Final Report October 2007

Water for a Healthy Country Flagship report series

ISSN: 1835-095X

Copyright and Disclaimer

© 2007 CSIRO To the extent permitted by law, all rights are reserved and no part of this publication covered by copyright may be reproduced or copied in any form or by any means except with the written permission of CSIRO.

Important Disclaimer

CSIRO advises that the information contained in this publication comprises general statements based on scientific research. The reader is advised and needs to be aware that such information may be incomplete or unable to be used in any specific situation. No reliance or actions must therefore be made on that information without seeking prior expert professional, scientific and technical advice. To the extent permitted by law, CSIRO (including its employees and consultants) excludes all liability to any person for any consequences, including but not limited to all losses, damages, costs, expenses and any other compensation, arising directly or indirectly from using this publication (in part or in whole) and any information or material contained in it. Authors: Russell J Taylor and Nigel B Goodman CSIRO Land and Water Private Bag 33, Clayton South Victoria, Australia, 3169 Acknowledgement re funding: The project was partially-funded by the Smart Water Fund that is an initiative of the Victorian water industry in partnership with the Victorian Government. The Smart Water Fund encourages innovation in water recycling, water conservation and bio-solid management to help secure Victoria’s water supplies now and in the future. For further information, visit: http://www.smartwater.com.au

http://www.smartwater.com.au/

CONTENTS

EXECUTIVE SUMMARY ......................................................................................................... 3 INTRODUCTION ..................................................................................................................... 5 HISTORY OF PROJECT ......................................................................................................... 5 DESALINATION OF EFFLUENTS.......................................................................................... 6

Rationale for using EDR over RO................................................................................ 6 Outline of MMF EDR process ...................................................................................... 7

SYSTEM CONFIGURATION................................................................................................... 8 Multimedia filter overview............................................................................................ 8 EDR Overview ............................................................................................................. 10 Communications between MMF and EDR ................................................................ 11

DESALINATION PILOT PLANT TRIAL................................................................................ 11 Installation and Commissioning ............................................................................... 11 MMF Performance....................................................................................................... 13 Electrodialysis optimisation ...................................................................................... 15

Indicators of EDR MMF System Performance ....................................................................... 15 Product conductivity ............................................................................................................... 16 Stack Inlet Pressure ............................................................................................................... 16 pH control of concentrate/brine.............................................................................................. 17 Clean-in-Place (CIP) .............................................................................................................. 18 Mini CIP.................................................................................................................................. 19 Stack Dismantling .................................................................................................................. 19

PERFORMANCE EVALUATION DURING CONTINUOUS OPERATION............................ 20 Logged data ................................................................................................................ 20 Summary of Operations ............................................................................................. 24

Period 1: May 2 2006 – July 31 2006 ................................................................................... 24 Period 2: October 18 2006 – Feb 02 2007............................................................................ 24 Stack Currents ....................................................................................................................... 25 Operational problems............................................................................................................. 26 Brine Management................................................................................................................. 27

Efficiency and Energy calculations with respect to product and waste formation29 Energy usage ......................................................................................................................... 29 Water usage........................................................................................................................... 30 Chemicals usage.................................................................................................................... 31

Chemical analysis of the various streams in the MMF/EDR process. ................... 32 Relation between analyses and time of sampling.................................................................. 32 Composition of various streams from the MMF/ EDR plant................................................... 33 First period of testing.............................................................................................................. 33 Second period of testing ........................................................................................................ 35

Blending to produce 1000 µS/cm product................................................................ 38 Disposal options for brine ......................................................................................... 41

Brine as a medium for nutrient removal via microbiological processes................................. 41 Brine as a source of chemicals .............................................................................................. 42

Performance evaluation of monovalent cation membranes in the CSIRO EDR rig43 Backgound ............................................................................................................................. 43 Results ................................................................................................................................... 43 Conclusions regarding the use of monovalent membranes................................................... 44

CONCLUSIONS .................................................................................................................... 45

Pilot Plant Demonstration of Effluent Desalination by Electrodialysis at the Western Treatment Plant, Werribee 1


ACKNOWLEDGEMENTS ..................................................................................................... 47 APPENDICES........................................................................................................................ 48

EXECUTIVE SUMMARY

A commercial electrodialysis reversal (EDR) pilot plant using pre-treatment with multimedia filtration (MMF) has been operated at the Western Treatment Plant over two three-month periods and demonstrated effectively the removal of salt from treated effluent. Based on an energy cost of $0.10/kWh and using the cost of bulk chemicals, the cost of 1000 µS/cm product blended with filtered feed is $169/ML and $113/ML for single and two-stage EDR respectively and $108/ML and $0.82/ML respectively when blended with unfiltered feed

The typical product after MMF and two-stage EDR has a conductivity of less than 500 µS/cm, considerably less than the 1000 µS/cm criterion established by Melbourne Water as suitable for irrigation. A final conductivity of approximately 1000 µS/cm was achieved by blending the product with raw feed water thus producing a higher yield from the available plant.

Ferric sulphate was found to be an effective coagulant for feed turbidity less than 5 NTU, but it was not sufficient to cope with excursions to values as high as 15 NTU. A contingency to co-dose with a polymeric flocculent was developed to reduce turbidity from higher levels.

The MMF plant operated with a water efficiency of 97% while the two-stage EDR plant ran with a mean water efficiency of approximately 81% with a maximum of 84%. A plant with MMF and two-stage EDR produces a better product at lower operational cost. The sodium adsorption ration is slightly lower (6.1 cf 6.3), the energy usage is 30% less and the chemical usage is 18% less.

A single step of electrodialysis reduces the concentrations of all ionic material by approximately 50%, including sodium chloride which comprises the major component of its salinity, as well as other ions including sulphate, carbonate, potassium and magnesium. Importantly, the ionic nutrients nitrate, nitrite, phosphate and ammonia are also removed by similar amounts. All these salts are removed in concentrated form, so that the reduced volume facilitates any subsequent process for their disposal.

The most efficient operation of the system was observed near the end of the trial when the stack currents and hence salt removal were highest. Based on these data, in order to produce 50 ML/day of 1000 µS/cm product at an overall water efficiency for two-stage EDR of 84%, an MMF plant of 55 ML/day and a two-stage EDR plant of 27 ML/day would be required. Product from the EDR plant producing 27 ML/day would be combined with 23 ML/day of filtered product to produce 1000 µS/cm product at 50 ML/day. The brine blowdown flow would be approximately 5 ML/day.

The capital cost of a two stage plant is greater than a single stage plant although the additional cost of membranes is partially offset by the components that are in common, namely the instrumentation and control system and general infrastructure. A two-stage process involves twice the membrane area, two power supplies and higher pumping capacity.

The energy and chemical components of the operational cost of producing 1000 µS/cm water from EDR of feed water of ~1900 µS/cm depends on the process. After blending with MMF product, the energy cost is 1.18 kWh/m3 after a single stage of EDR and 0.72 kWh/m3 after a two-stage EDR. If blending is performed using unfiltered product, the energy costs are 0.62 and 0.5 kWh/m3 respectively.

In order to establish the overall cost of an installation, the capital cost would need to be provided by an equipment provider. Currently GE is the major global provider, although Eurodia / Ameridia is another potential provider.


There was no apparent deterioration of the MMF system or EDR stack after the completion of these trials. Restoration of original stack performance was always achieved by a conventional hydrochloric acid CIP. A mini-CIP, developed during the trial, was also found to remove the effects of scaling without the necessity to cease production. Staining of the anion membranes, presumably by organics, did not appear to have any significant effect on cell performance.

Disruptions to plant operation occurred largely due to excessively high turbidity from the inlet water supply, as well as electronics and software problems in the sensing and control systems. Based on the experience of this trial, we are convinced that with state-of-the-art EDR plant and improved instrumentation and controls, CIP intervals of greater than 60 days would be feasible. The main factors causing the requirement for CIP in this study related to plant breakdowns that prevented automatic flushing of the EDR system that in turn caused precipitation of scale. More careful control of brine blowdown flow and anti-scalant dosing would be possible on a system that was operating continuously.

The 20ML/day EDR plant in La Jolla, San Diego provides an example of the application of EDR in a situation very similar to the Western Treatment Plant. The plant takes feed from a clarifier in an adjacent municipal waste treatment plant and uses cartridge filtration as the pre-treatment for a two-stage EDR process. This pre-treatment is quite crude in comparison with the coagulation/MMF used in the present study and as a consequence necessitated quite frequent CIP, more than monthly. Its output is pumped to a dam from which it is used for irrigation.

The Western Treatment Plant would be an ideal site for the implementation of an EDR salt-reduction system. An installation would include pre-treatment and EDR facilities with appropriate controls. It could be located in the vicinity of the disinfection facilities at the HORS and direct its output down the pipeline to Werribee South. Although MMF was used as the pre-treatment in this investigation, other pre-treatments that could provide the feed quality required for EDR should be considered e.g. dissolved air flotation or micro- or ultra-filtration.

MMF/EDR could be implemented in a number of configurations:

• Chlorinated effluent could be diverted from the disinfection train to MMF then EDR and the EDR product finally blended with additional disinfected feed prior to dispatch. The turbidity of the blended product would fluctuate according to the turbidity of the feed.

• Turbidity removal, using one of the aforementioned techniques, could be performed prior to disinfection, producing a stream suitable for EDR and blending. This may reduce the chlorine demand in the disinfection process and give a final blended product of more uniform quality with respect to turbidity.

• If microfiltration or ultrafiltration were used, it is expected that the microbiological quality of the product would be improved and there would be a reduced requirement for CIP due to the improved quality of the EDR feed.

This project has demonstrated the robustness of EDR to major fluctuations in the turbidity of the feed. Even when the feed turbidity was 15NTU and the MMF was unable to reduce turbidity sufficiently, the EDR system kept producing in the presence of increasing fouling, albeit with reduced efficiency and increased conductivity.

This study has demonstrated that electrodialysis reversal is a serious candidate technology for the removal of salts from the effluent at the Western Treatment plant or indeed any waste treatment plant or other source of brackish water for which salt reduction is necessary to enable re-use.


INTRODUCTION

The objective of this project has been to develop and demonstrate a process for reducing the salinity of WTP sewage plant effluent, based on electrodialysis. Recent advances in electrodialysis technology, especially electrodialysis reversal (EDR) have resulted in improved performance, lowered cost and extended life of plant and materials. Whereas reverse osmosis (RO) is preferred for desalination of seawater, EDR is generally superior to RO when the levels of dissolved salts in the feed water are low. By using EDR for reduction of the concentrations of dissolved salts and nutrients from sewage effluent, Class A and B products suitable for irrigation, industrial and third-pipe use should be achievable.

The Western Treatment Plant (WTP) produces approximately 150,000 ML of effluent annually at an average rate of 410 ML/day. Archival data obtained from Melbourne Water indicated that this effluent contains an average of 1050 mg/L of dissolved solids including high concentrations of salts but also phosphorus and nitrogen-based nutrients and some heavy metals in addition to smaller quantities of a large number of organic materials. Removal of these substances would expand the opportunities for re-use of this water. The high sodium content can make the water unsuitable for irrigation due to potential damage to soil structure. The high chloride content increases the corrosivity of the water towards metals, even stainless steels, limiting its use in industry.

The advantages of electrodialysis are not widely known. Whereas RO involves high-pressure filtration, forcing water through pico-scale pores, electrodialysis is a low-pressure technique where ions migrate through ion-conducting membranes due to the passage of dc electric current. Recent advances in EDR technology have resulted in improved performance, lowered cost and extended life of plant and materials. Because EDR is not a filtration technique, its feedwater requirements are generally less stringent than that for RO, ultrafiltration and nanofiltration. EDR would be expected to reduce the concentrations of all ionic constituents, including ammonium, nitrate and nitrite ions, by similar factors. Divalent ions (e.g. Ca2+, Mg2+, SO4

2-) experience a greater force in the electric field of an electrolysis cell and tend to be removed at a higher rate than monovalent ions (e.g. NH4

+, Na+, K+, Cl-, NO3

-). A typical reduction in TDS for a single pass through an EDR ‘stack’ is 50%.

EDR also produces a concentrated waste stream, up to 100,000 mg/L. This may, in theory, be used as a resource for salt production, chemical production or metal extraction or be disposed of into the bay.

HISTORY OF PROJECT

In 2003, CSIRO was awarded Smart Water funding to establish a laboratory-scale electrodialysis reversal (EDR) pilot plant to investigate the feasibility of using this method for salt-removal from treated sewage effluent. The reasoning behind the choice of EDR was the claimed economic superiority over reverse osmosis (RO) for applications at low salinities1.

The Smart Water Fund grant was awarded after successful demonstration of salt removal in a laboratory-scale trial at one litre per minute. To scale-up for this project, negotiations to obtain a pilot plant were undertaken with the US-based company Ionics, at that stage the major global supplier of EDR technology. During these negotiations, Ionics acquired Ecolochem, and consequently the agreed configuration was based on an Ecolochem® multimedia filtration (MMF) plant feeding an Ionics EDR pilot plant.

1 Passanini, J, Persechino, J and Reynolds, TK, EDR, NF and RO at a brackish water reclamation facility’, Proceedings 2000 AWWA Annual Conference.


Meanwhile, Melbourne Water (MW) called for expressions of interest (EOI) in a technology demonstration trial for desalination of recycled water at the Western Treatment Plant (WTP). A consortium comprising Ionics and United KG made an EOI based on two configurations incorporating ultra-filtration (UF): UF-RO and UF-EDR. The EDR configuration offered differed from the MMF-EDR that was to be used in the Smart Water project. However the Ionics-UKG EOI was unsuccessful and in the intervening period Ionics was taken over by GE Water.

CSIRO proceeded to assemble equipment in two shipping containers, one for MMF pre-treatment and the other for EDR. At about the same time, contact was made with the new local GE management to try and revive the possibility of a acquiring a larger pilot plant for lease by CSIRO. GE initially proposed purchasing a new Aquamite III unit (1 m3/hr) that included MMF pre-treatment in a single container. However, because of likely delays in supply, GE then suggested leasing to CSIRO an existing EDR Aquamite V pilot plant and a purpose-built MMF unit. Ironically this was the same EDR plant that had been the subject of earlier negotiations with Ionics. An agreement was then made with GE to provide MMF and EDR equipment for the trial, and a further revision of the SWF work plan was made.

Before the GE equipment arrived, the two containers of CSIRO equipment that had been configured for the amended program were installed on the test site at WTP. Considerable time was expended refining the coagulation/MMF until a fault in the filter was discovered. The GE equipment arrived on 13th February 2006 and was installed and commissioned by GE personnel.

DESALINATION OF EFFLUENTS

Rationale for using EDR over RO Unlike RO, EDR does not involve filtration. The ionic material is removed by transport through cationic and anionic conducting membranes, depleting the feed water as it passes through the stack. As the schematic in Figure 1 shows, the membrane stack consists of alternating cationic and anionic conducting membranes with electrodes, typically Pt/Ti or carbon, at each end.

Figure 1 Schematic of EDR stack

Electrolysis results in production of hydrogen at the cathode and a mixture of chlorine and oxygen at the anode. The consequential ion flow in the electrolyte causes the ion transport, which results in desalination. Cations (e.g. sodium, potassium) migrate through the cation membrane, and anions (e.g. phosphate, chloride) migrate through the anion membrane into the concentrate or brine stream leaving the feed/product stream depleted in salts.


Because it does not involve filtration, EDR is tolerant of feed quality with respect to suspended material as indicated by SDI, turbidity and the other parameters shown in Table 1. On the down side, there is no removal of pathogens, although the capacity to operate continuously with a free chlorine residual of 0.5mg/L partially addresses this issue.

Table 1 EDR feed requirements; units mg/L unless otherwise indicated

pH range 1-9 pH (cleaning) 1-11

Fe (dissolved) 0.3 Mn (dissolved) 0.1

Aluminium 0.1 H2S 0.1

Free Cl2 (continuous) 0.5 Free Cl2 (shock) 30

Oil (Infrared method) 2 COD 50

TOC 15 MBAS 1

Turbidity (continuous) 0.5 NTU Turbidity (intermittent) 2 NTU

SDI 15 (continuous) 10-12 SDI 15 (intermittent) 15

Outline of MMF EDR process

Class C treated effluent was delivered from the high level carrier from the head of the road storage facility at WTP to a four-port manifold on the site. A number of measures were used to facilitate reliable and stable operation of the process:

• Pre-filtration though an AMIAD 500 µm automatic cleaning filter was used to remove any large particles and living material including blood worms.

• Coagulation/filtration pre-treatment was used to remove suspended material that could deposit in the stack. Although a feed turbidity of zero is ideal, a practical objective was 0.5 NTU, with intermittent increases to 1 NTU.

• The feed water was dosed with sodium hypochlorite to reduce microbiological fouling of the multimedia filters and the EDR stack. Unlike RO, the EDR membranes were tolerant of free chlorine 0.5 mg/L (continuous) and 5 mg/L (intermittent). Initially there were problems with the hypochlorite pump becoming un-primed. This was rectified by placing the feed pump on top of the hypochlorite container, thus reducing the length of the feed tube.

• Precipitation of insoluble material in the brine stream was controlled by three means: o Its concentration, indicated by conductivity, was regulated by constant dilution

by a stream of feed with overflow (brine blow down – BBD) to waste. o Its pH was controlled by constant addition of hydrochloric acid to raise the

solubility of potential scaling compounds. o A chemical anti-scalant was also added to the brine stream as an additional

measure. • The degree of salt removal was directly determined by the electrical current in the

two stages of the stack. As the temperature decreased during winter, it was necessary to apply a higher voltage to maintain the current to overcome the increased electrical resistance of the stack due to the temperature dependence of electrolyte conductivity.


SYSTEM CONFIGURATION

Multimedia filter overview

The Plumbing and Instrumentation Diagrams (P&ID) for the equipment leased from GE are included as Appendix 1.

A multimedia filtration (MMF) system, custom built by GE Water, was leased to CSIRO for the purpose of this trial. This unit consisted of two fibreglass multimedia filters that were fed by two centrifugal pumps under suction and managed with pneumatically operated valves. The unit also stored filtered product in a large (1500 L) backwash tank with a separate back-flush pump. The sand filter pumps operated at a flow rate of approximately 7.5 m3/hr and had a maximum flow rate of 12 m3/hr. Operation of the filtration system was controlled by a programmable logic controller (PLC) in a parallel, duty/standby configuration which allowed for continuous operation of the units. When a changeover occurred, the standby filter came on duty and the duty filter was backwashed and rinsed and placed on standby.

Approximately 6.5 m3/hr passed through the MMF plant with 0.3 m3/hr initially filling up the backwash tank. The remainder of the 6.2 m3/hr of filtered water then went to the EDR. Of the 6.2 m3/hr approximately 5.2 m3/hr is product and 1.0 m3/hr was brine blowdown (BBD). Turbidity and conductivity of the filtered product were monitored continuously. The conductivity range of the product was between 450 and 550 microsiemens/cm.

The user interface for control of the MMF system was via a touch-sensitive panel-mounted screen, which allowed each of the two multimedia filters to be operated manually or automatically in several modes: service, backwash and rinse. There were several bugs in the software which caused the MMF system to lock up and prevent a changeover of the filters. This was one factor which led to overshooting the operational recommendation of 0.5 NTU for the turbidity of the MMF product (EDR feed). This, along with periods of excessively high turbidity water being supplied from the main inlet carrier, further complicated the pre-treatment using MMF, raising the question of whether using ultrafiltration (UF) or nanofiltration (NF) may have been more reliable.

Feed entering the MMF was first passed through an AMIAD 500 µm pre-filter which removed debris and heavy contaminants from the feed. To minimise biological fouling in the system and possible health risks of working with tertiary effluent, sodium hypochlorite (NaOCl) disinfectant was injected downstream of the pre-filter before multi-media filtration and providing a contact time prior to MMF of only a few seconds. Free chlorine levels of approximately 0.5 mg/L were maintained throughout the system.

Injection of hypochlorite took place after this filtration stage, whereas addition of the Fe2(SO4)3 coagulant was done about twenty metres upstream of the pre-filter, to allow greater reaction time for floc formation. The dose rate of ferric sulphate was approximately 3 mg/L as Fe, this produced an iron residual in the filtrate of <0.1 mg/L; minimising ferric was essential as the maximum allowable concentration of Fe3+ in the solution exposed to the EDR membranes was 0.3 mg/L.

Injection of NaOCl and Fe2(SO4)3 took place using Iwaki dosing pumps connected to the main feed supply line with 6 mm nylon hosing and a non-return dosing valve. The dosing pumps were controlled by the PLC and were switched on only during filtration. There were initially approximately 15 seconds of contact time before the ferric dosed water entered the pump and filter. Although the freeboard volume of the filter provided for additional contact time, it was found that the silt density index (SDI) of the MMF feed was high resulting in deposition of scale in the EDR system. The ferric dosing point was moved to allow a one minute contact time, resulting in much lowered SDI and improved performance.


Under conditions where the feed water was approximately 2 NTU the filters required backwashing approximately every 8-12 hours. These conditions were adjusted several times during the trial to achieve a stable turbidity of the order of 0.5 NTU.

The chemicals used in the project were obtained from Omega Chemicals, with the exception of Hypersperse MDC150 that was provided by GE Water. Table 2 lists the chemicals used and describes their purpose.

Table 2 Chemicals used in the MMF EDR pilot plant

Chemical Application Average Dosage Total Number of Dosages/CIP

Sodium hypochlorite 12.5%

Disinfection of pre-treatment equipment and EDR membranes

Free chlorine residual 0.1 to 0.5 mg/L

continuous

Hydrochloric acid 33%

Clean-in-place 5% active solution as required

Hydrochloric acid 26%

Anti-scaling To lower pH to pH 6-7 continuous

Ferric sulphate 12.5% as Fe

Coagulation 2.5-3.0 ppm as Fe continuous

Hypersperse MDC150 150 g/L

Anti-scalant 12-15 ppm continuous since 23 June 2006


EDR Overview

In the preliminary stage of this project, a 10 membrane-pair laboratory-scale EDR stack was obtained from Ionics (now GE) and incorporated as the central element following the general schematic in Figure 2.

Figure 2 Schematic of electrodialysis equipment

After commissioning, secondary treated effluent from the Western Treatment Plant (WTP), Werribee, Victoria was transported to the CSIRO laboratory at Clayton in 1000 L batches. The laboratory program involved developing procedures for coagulation and clarification of this effluent feed water to achieve the minimum feed specification suggested in the commercial literature.

After success at the laboratory scale, additional support was obtained to undertake a pilot plant trial at WTP. Equipment was leased from GE and operated independently but in cooperation with GE Infrastructure. The equipment consisted of a multimedia filtration (MMF) plant directly connected to an EDR pilot plant. The MMF consisted of two parallel systems (pump, filter) to maintain continuous operation during back wash. The system was sized to provide feed for the 5.25 m3/hour EDR plant, and 1.5 m3 of the filtered product was stored for use as backwash water.

The EDR container included a two-stage electrodialysis stack, each stage comprising 85 cation and anion membrane pairs with Type 4 separators. There were electrodes at each end, and a central pair separating the two stages. Figure 3 is a view of partially dismantled stack. The holes at the end are flow ports for electrode, feed and brine streams, the perforated component on the top layer is the Type 4 separator that separates the cation and anion membranes and the bolts for tensioning the upper electrode against the stack, sealing it against leakage. The upper electrode had been removed to gain access to the stack, but the central and lower electrodes are visible in Figure 3.


The stack was fed by feed and brine pumps and powered by a dc power supply. There was extensive monitoring of the various streams for flow, pressure, pH and conductivity. Sampling points were provided for the MMF feed and product, EDR product and brine. An additional sampling point was provided to intercept the product from the first stage of EDR as it flowed between the two stages in the stack.

Figure 3 View of partially dismantled EDR stack

Electrodes

Separator

Ports

Communications between MMF and EDR

The MMF and EDR systems were controlled by separate programmable logic controllers (PLCs) each of which originated from a different manufacturer. The control module on the MMF was a current model Fanuc PLC. The control unit on the EDR was an Allen Bradley unit and was approximately 15 years old.

Communication between the systems was controlled through the MMF PLC via a single link to the EDR PLC. The MMF PLC functioned in automatic mode (master) and the EDR in remote mode (slave). This meant that the EDR would not operate unless flow was detected, and if the flow from the MMF stopped, an alarm signal caused the EDR to shut down. The system could also be shut down manually, and in the case of an orderly shut down, a 2 minute rinse of the EDR stack was implemented before pump operations ceased.

DESALINATION PILOT PLANT TRIAL

Installation and Commissioning

The equipment housed in two 20 m containers was delivered on Monday 13 February (see Figure 4). The multimedia filters were in the container in the foreground and the EDR equipment was in the rear container. The other equipment on the site were three containers that were part of a salt removal technology trial organised by Melbourne Water involving Veolia and United Utilities, two containers operated by CSIRO containing small-scale MMF



and EDR equipment, a site office and toilet facilities. Plumbing and electrical interconnections were made over the following few days.

Figure 4 Containers on site at WTP

.

The two photographs in Figure 5 show overview and close-up plan views of the location of the equipment at the test site at Werribee. The two containers are indicated by the arrow in Figure 5b.

Figure 5 a and b Views of test site at WTP taken from Google Earth

Initial commissioning took place during the week commencing 20 February. Initially there was no data logging and no automatic backflush control of the filters, so the equipment could only be operated in manual mode. This restricted operational time of the plant to regular working hours when personnel would be on site. The plant continued to be run using manual control from February 24 until the end of March 2006.

GE staff returned on March 28 and installed a notebook computer and modem with software that collected data from the two programmable logic controllers (PLCs). The communications were programmed to dial out every 2 hours to a server that transmitted the data to the Watereye web site. A username and password enabled access to the current data from the Watereye site; this could be viewed and analysed using a basic plotting capability. The data were periodically downloaded for storage and evaluation by CSIRO and GE.

Before continuous operation commenced in early May 2006, the system was stopping intermittently for a number of reasons. Initially this related to software issues and modem failures in particular. On April 10, 2006 and again about 1 week later, a local programmer was engaged by GE to upload a revised version of the control software and undertake some on-line modifications. Attempts to get the modem operating so that the GE personnel could dial in from Perth were unsuccessful.

A major process issue which also lead to intermittent stopping of the system was the filtrate turbidity, which during April regularly exceeded 1 NTU, well above the recommended operating level of 0.5 NTU. A significant amount of work was required to solve this problem. The only coagulant used during the trial was ferric sulphate at a dose rate of approximately 2-3 mg/L as Fe. This produced a Fe residual in the filtrate of <0.1 mg/L that complied with the requirement for the EDR membranes of <0.3 mg/L. The coagulant dose rate was adjusted and optimised and the pH of the feed was varied in an attempt to lower the turbidity, but with no success. Because of the disruptions to continuous flow, conditions may have been favourable for microbial growth within the system and in particular in the MMF. It is thus possible that biofouling was contributing to the high turbidity.

Once continuous operation was established, then the regular backflushing of the MMF filters along with effective chlorination meant that turbidity began to reduce and stabilise. From that point on, increases in filtrate turbidity only occurred as a result of the filter becoming overloaded or when gross increases in feed turbidity occurred. As a result, the EDR product flow that was intermittent during the commissioning period became more reliable towards the end of April, when continuous operation of the system began.

MMF Performance In Figure 6, turbidity variations in feed and MMF product over the total duration of the trial, May 2, 2006 – Feb 4, 2007, are shown. The plant was not operated between August 1 and October 18, 2006. The EDR system required a feed water turbidity of less than 0.5 NTU with occasional excursions to 2 NTU being tolerated. This was achieved generally for the first period shown, where the average turbidity was approximately 0.42 NTU. As explained later in this report and highlighted in, large increases in feed turbidity attributed by Melbourne Water staff2 to wind events, made it difficult at times to maintain turbidity with a constant coagulant dosing system. A full scale system would require automatic feedback control of coagulant dosing in response to feed quality.

2 Trevor Gulovsen, Melbourne Water, private Communication


Figure 6 Turbidity of MMF feed and product for duration of trial

Turbidity of MMF Feed (HORS) and MMF Product for duration of trial

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No operation in this period

The variations in MMF product turbidity were due to a number of factors:

• The variable turbidity of the feed. • As the filter became conditioned, the turbidity declined • As the filter capacity approached saturation the turbidity increased. • The turbidity peaked after changeover during backflushing because coagulant dosing

was halved during the rinse cycle when the flow rate was approximately doubled for 10 minutes. The purpose of the rinse cycle was to condition the filter with coagulated feed, in preparation for its use while the second filter was in backwash mode.

Figure 7 Flow-time plots for operation of the dual multimedia filters

Filter 1&2 Flow Graph

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Date/Time

m3 /h

r

The operation of the two MMF filters that supplied feed to the EDR stack is summarised in Figure 7. The narrow peak after changeover was due to the rinse cycle that followed the


backwash (which consisted of 10 minutes backflushing and a rest period of 5 minutes). During the rinse cycle, an increase in turbidity was observed. This was due to the flow rate of feed increasing (approximately double) and a subsequent halving of the coagulant dose rate. As a result there was insufficient coagulant in the feed water to effectively remove suspended material.

A comparison of the raw feed and the MMF product is shown in the concentrations of selected parameters in Table 3. Because the purpose of filtration is the removal of suspended solids, it is not surprising that there was no great chemical difference for most of the entries, apart from sulphate from the ferric sulphate coagulant. The main measurable effect of MMF performance was the turbidity of the raw feed that was typically 2-3 NTU during the first test period but became more erratic towards the end of the first and during the second test period – see Figure 6, the turbidity was typically reduced to approximately <0.5 NTU after MMF filtration in the first period, but during the second period it became harder to control.

Table 3 Average composition of selected parameters in the MMF feed and product during the first period of continuous testing.

Species, mg/L Raw Feed MMF Product Conductivity, µS/cm 1906 1926 TDS 1135 1146 Calcium 38.5 37.4 Magnesium 28.5 28.4 Sodium 325 329 SAR 9.69 9.86 Sulphate 108 113 Alkalinity 137 134 Total Phosphorus 35.9 32.8 TOC 8.6 8.8 Iron 0.12 0.06 Manganese 0.09 0.01 Silicon 13.9 13.2

Heavy metals are of concern in EDR systems, especially Fe3+ because of its use as the coagulant. Other cations that bind strongly to sites in the cation membrane (e.g. Al3+ and Mn2+) are also potential problem species that must be controlled. From Table 3 it is clear that Fe3+ concentrations in the MMF product were less than 0.3mg/L, as required by the process (Table 1).

Electrodialysis optimisation

Indicators of EDR MMF System Performance

There were a number of indicators - physical, electrical, and chemical - of performance of the MMF and EDR systems.

Physical indicators:

• Turbidity of the MMF product. This is required to be as low as possible, certainly less than 0.5 NTU, to minimise deposition of suspended material in the EDR separators.

• EDR stack inlet pressure (SIP). For a given feed flow rate, the SIP was 1.9 kg/cm2. If there was physical fouling due to deposition of suspended material or scale, this pressure increased and eventually required removal by clean-in-place.


Electrical indicators:

• Stack Current. The best electrical indicators of EDR performance were the currents and the conductivity of the product, both for each stage of the stack. Fouling effectively increased the stack resistance with a consequent drop in current because the voltages applied to stack, although process variables, were constant over extended periods and independent of fouling.

• Product conductivity. The electrical conductivity of the product was the ultimate indicator of system performance and was the consequence of most process problems that inevitably resulted in decreased current and decreased conductivity.

Product conductivity

Figure 8 EDR product conductivity over a short period showing spikes due to off spec product

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Figure 8 shows EDR product conductivity over a short period. The spikes in product conductivity were occasionally captured during current reversal when some brine flowed past the product conductivity sensor on the way to disposal as off-spec product (OSP). Current reversal occurred every 15 minutes and although this can be varied, it was not altered during the trial. The period of “off-spec product” is approximately 30-40 seconds or 3.3-5.0 % of flow. If the reversal time were extended, this percentage would decrease with a consequent increase in production efficiency. The conductivity at which product collection recommenced was normally set at 600 µS/cm. If this was increased, the OSP period would decrease but the mean product salinity would increase. The rapid drop in product conductivity after reversal is highlighted by the arrow in Figure 8.

Stack Inlet Pressure

Figure 9 shows the stack inlet pressure (SIP) during the first 6 weeks of continuous operation. For the first days of May, the stack inlet pressure was relatively constant, after which there was a period when the EDR stack was shut down. The rises in SIP indicated fouling of the separators between membranes. The system was designed to operate with a SIP of approximately 2 to 2.5 kg/cm2. Fouling of the system caused the product flow rate to drop, as well as the product conductivity to increase. An overall decrease in stack current was also observed. The inlet pressure typically increased over a 4-6 day period, shown as ‘SIP increase’ in Figure 9. During these excursions, the pressure exceeded 2.6 kg/cm2 and was as high as 3 kg/cm2.


Figure 9 EDR stack Inlet Pressure during the first 6 weeks of continuous operation

EDR Stack Inlet Pressure: May 2 - June 14

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Managing the SIP by preventing fouling is a major issue in achieving reliable and continuous operation. If the fouling that causes the increase in SIP is chemically removable, the major remedial measure is the clean in place, while if it cannot be removed chemically, dismantling of the stack may be required.

pH control of concentrate/brine

It was originally anticipated that the electrodes would be fed using feed, product or dilute sulphuric acid. In the GE EDR plant provided, water from the feed pump was directed through the electrode compartments and then into the circulating brine stream. Hydrochloric acid was pumped via the Electrode Clean-in-Place (ECIP) pump into the electrode feed to acidify the stream to avoid scaling of the electrodes, and in turn acidified the brine stream. The ECIP pump rate was adjusted to maintain the brine pH at a value required to avoid precipitation of insoluble compounds in the stack using a program obtained from Ionics (now GE). This software allowed calculation of the concentrations of Ca2+, Mg2+ and PO4

3- in the brine stream which would allow up to 400% over-saturation by the major limiting supersaturated species (Ca HPO4) by varying the pH and the temperature.

Figure 10 shows that fluctuations in the brine blow-down (BBD) pH occurred because, on reversal, the brine was diluted by a volume of feed (pH ~7.5 – 8.0) whereupon the pH rose and fell gradually due to continued ECIP acid dosing. The BBD pH used during the continuous operational stages of the project was significantly lower than those shown in Figure 9 because of fouling that occurred at these higher initial pH values of BBD.


Figure 10 Blowdown pH variations during successive 15 minute reversal cycles in the commissioning stage

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Clean-in-Place (CIP)

An increase in stack inlet pressure is an indication of physical fouling. Acid soluble fouling may be removed by acidic clean-in-place (CIP). Several CIPs were performed during the trial when their necessity was indicated by an increase in stack inlet pressure and/or declining stack currents. During the CIP, the feed/product and brine streams were isolated and circulated while the CIP chemical was inducted into the stream. A typical CIP took 60 minutes followed by a 60 minute system flush to remove all cleaning chemicals and residues from the system.

A decrease in stack current was an indication of increased stack resistance caused either by physical fouling or fouling by organic or other adsorbed material. Organic fouling may be reversed by soaking in brine, especially at elevated pH (~10) to accelerate the process. This would normally be followed by an acid CIP. Several brine CIPs were carried out by inducting sufficient sodium chloride concentrate and sodium hydroxide to achieve 5% NaCl and pH 10. However, there was no apparent benefit from this procedure, the major fouling being removed by acid CIP alone.

An analysis of a typical CIP is included in Table 4. It is clear that the major species removed were Fe, Ca, Mg, and P. The iron represents the coagulant that is responsible for the flocculation process. Iron, along with the calcium and magnesium were most likely deposited as phosphates, sulphates or carbonates, despite efforts to prevent scaling by the constant addition of hydrochloric acid to the brine stream via the electrode stream, and by the addition of anti-scalant to the brine stream. It was also clear from the data over time in the bottom three rows of Table 4 that these materials were removed quite quickly and a CIP shorter than the standard hour may suffice with less loss of production time.

Table 4 Composition in mg/L of key components in feed water and solution after CIP

Solution Ca Fe K Mg Na P SBrine Blowdown 193.9 0.0 178.1 110.9 625.5 26.6 118.4MMF Product 38.7 0.1 25.0 21.9 223.7 8.1 27.1EDR Product 9.4 0.0 7.5 5.6 92.2 3.7 12.3CIP after 5 minutes 1420.0 109.9 196.2 581.6 1183.0 189.7 332.6CIP after 30 minutes 1408.3 90.6 168.5 576.5 1066.4 184.8 352.1CIP after 60 minutes 1476.8 92.5 199.5 631.9 1165.6 191.3 376.0


In a full-scale system, measures to neutralise acidic CIP waste would be required and would constitute an additional cost. Likewise the use of alkaline NaCl brine would require acid neutralisation and would constitute the discharge of additional salt in the waste stream.

Mini CIP

During the trial, we initiated what we termed a ‘MINI CIP’. This involved increasing to the maximum the rate of HCl dosing into the electrode stream. Typically it was found that normal operation was restored after about 30 minutes of treatment. This had the advantage of not requiring the system to be taken off-line to conduct a CIP, and also used substantially less acid. In an operational EDR plant, the quantity of HCl and the time required for conventional CIP could be established over time and incorporated into the Standard Operating Procedure.

Stack Dismantling

The ease with which the stack may be dismantled is another advantage of EDR over RO and other filtration membranes systems. Once opened up, the membranes and separators may be cleaned or replaced if necessary, and re-assembled. Although time consuming, it is a major difference from standard RO systems in which the membrane surfaces are inaccessible.

On Thursday 8 June 2006, the EDR stack was dismantled, cleaned and re-assembled to enable examination of the stack at the midpoint of the 3-month trial and to look for the cause of the inlet pressure increase. Figure 11 shows the accumulation of a gelatinous material in the separator between the membrane pairs. There were also beige-coloured ‘scale deposits’ around some of the main channels (the holes through the membranes and separators on the left side of Figure 11. The rinse water from stack cleaning was an accumulation of all material removed during the stack cleaning.

Figure 11 View inside stack during dismantling on June 7 2006

Analysis of the material removed from the dismantled stack, summarised in Table 5 , indicated that the major species present were iron and phosphate, followed by calcium and magnesium. It is likely that precipitation of ferric, calcium and magnesium phosphates is occurring in the brine circuit and that their rates of formation may be increasing as feed temperature decreases. This was confirmed using calculations based on the solubility product of FePO4 (pKs = 26.4).

In order to reduce the concentration of free Fe3+ to avoid FePO4 formation, an anti-scalant to preferentially complex Fe3+ in the presence of Ca2+ and Mg2+ was sourced and implemented in mid-June. The use of alternative coagulants, e.g. alum, was considered, however an even lower pH would be required to keep AlPO4 in solution, and cation membranes are also sensitive to Al3+>0.1 mg/L. Given that the concentration of Fe3+ is ~0.1 mg/L, and the flow rate is 5.25 m3/hour, then 525 mg of Fe is available to form FePO4 each hour or about 40 g of ferric phosphate per day or 1.2 kg per month.


Table 5 Samples from stack dismantle June 7 2006

Element Atomic Mass

Deposit in separators mg/L

Deposit in separators mM

Scale deposit mg/L

Scale deposit mM

Rinse water from stack cleaning, mg/L

Rinse water from stack cleaning, mM

Ca 40 70 1.75 35260 881 845 21.1 Cu 63.5 20 0.31 75 1.18 25 0.39 Fe 55.85 460 8.24 120 2.15 2995 53.6 K 39 14 0.36 275 7.05 250 6.4

Mg 24.3 13 0.53 9260 381 95 3.9 Na 23 128 5.57 2145 93 575 25 P 31 235 7.58 14870 479 1650 53.2 S 32 30 0.94 635 19.8 115 3.59

The material removed during dismantling was found to be soluble in hydrochloric acid. It is thus not surprising that, as required over the remainder of the project, acid CIP effectively removed deposited material and restored the low SIP and typical stack currents and product conductivity without the need for dismantling.

PERFORMANCE EVALUATION DURING CONTINUOUS OPERATION

Logged data

Data from the MMF and EDR processes were monitored by the respective PLCs and logged by a laptop computer via serial interface to the controllers. Data were transmitted to a central data storage facility, Watereye, from which it was possible to recall data for further processing. Lists of the logged parameters are included as Appendices 4 and 5.

There were two periods of continuous operation:

• May 2, 2006 – July 31, 2006 • October 18, 2006 – Feb 4, 2007

Plots of the key logged parameters are presented in Figures 11 – 22, and are presented to facilitate observation of the relationship between the various indicators.

In the first period, the system was running virtually continuously, as shown in the following graphs of EDR product flow (Figure 13 and Figure 19 ) and conductivity (Figure 14 and Figure 20). The causes of interruptions to flow are described in the log in Appendix 2, and were due to:

• Occasional power failure. • Requirements for clean-in-place. • Stoppages for maintenance and calibration. • Unavailability of feed water from the main inlet carrier.

The product flow set point of 5.25 m3/hr was generally achieved as indicated, with normal fluctuations. The periods during which the indicated flow was approximately 4.5 m3/hr (e.g. 5-6 June, 14-16 January) occurred when the EDR system had shut down while the MMF kept running and pumping though the EDR stack under the residual pressure of the MMF



pump. This represented a shortcoming of the control system that would not be present in a full-scale system.

Reasonably constant conductivity was achieved up to mid June (once the problems with fouling were remedied by frequent CIP), after which there were no further CIPs for the remaining 6 weeks. This was also achieved despite variations in feed pH and feed turbidity after MMF. In the plots that follow, spikes (e.g. see Figure 8) were removed for clarity.

For the first period some of the key features were:

• The relatively ideal EDR feed turbidity (approx. 0.5 NTU) achieved after coagulation and filtration

• The frequent interruptions to product flow, especially in the early stages, which contributed to fouling of the EDR stacks by precipitation.

• The inverse relation between stage 1 stack current and stack inlet pressure, especially obvious in the first half of the period.

• The clear increase in conductivity with fouling as indicated by SIP. • The relatively minor dependence of stage 2 stack current on fouling – the arrows in

Figure 16 indicate a region where the stage 2 current actually increased as the stage 1 current decreased. The reason for this is the feed to stage 2 increased as the stage 1 current decreased and current is directly related to feed concentration. In most other cases, this inverse effect was less pronounced, resulting in a relatively constant stage 2 current

For the second period some of the key features were:

• The inability to achieve turbidity less than 0.5 NTU due to poorer quality raw feed from the HORS at WTP – see also Figure 6 and consequent interruptions to EDR operation and product flow.

• By comparison with the first period, the higher stack currents and lower conductivity due to the use of higher stack voltages – see also Figure 27 and Figure 29.

• The continuation of operation, as indicated by product flow, despite periods of elevated SIP, but the recovery of the system after CIP, e.g. 4 January, but the gradual increase in conductivity and decrease in stack current as fouling increased as indicated by SIP.

Figure 12 MMF Filtration Turbidity May 2 – July 31

Filtration Turbidity May 2 - Jul 31

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Figure 13 EDR Product Flow May 2 – July31

EDR Product Flow May 2 - July 31

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Figure 14 EDR Product Conductivity May 2 – July 31

EDR Product Conductivity 2 May - 31July 2006

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Figure 15 Stack Inlet Pressure May 2 – July 31

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Figure 16 Stack currents during periods of EDR operation May 2 – July 4

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Figure 17 EDR Brine Blowdown Flow May 2 – July 31

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Figure 18 Filtration turbidity Oct 18 – Feb 4

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Figure 19 EDR Product Flow Oct 18 – Feb 4

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Figure 20 EDR Product Conductivity Oct 18 – Feb 4

EDR Product Conductivity Oct 18 2006 - Feb 4 2007

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Figure 21 Stack Inlet Pressure October 18 – Feb 4

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Figure 22 Stack currents during periods of EDR operation Oct 18 – Feb 4

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Figure 23 EDR Brine Blowdown pH October 18 – Feb 4

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Summary of Operations

Details of the process during the two periods are tabulated in Appendices 2 and 3 and described below.

Period 1: May 2 2006 – July 31 2006

During the first period of operation, the EDR produced water with conductivity ranging from 450-550 μS/cm with a recovery of around 85%. The ferric sulphate coagulant was dosed prior to the MMF at a rate between 80 to 100 mL per hour as product. This equates to between 18 to 23 ppm as product or 2.2 ppm to 2.9 ppm as Fe. Some refinement of the dosing was necessary to make sure the turbidity of the filtrate was kept below 0.5 NTU. Sodium hypochlorite was dosed prior to the MMF to maintain a free chlorine residual of approximately 0.5 mg/L. This was measured colorimetrically using an on-line Hach Chlorine analyser and also indicated by between 450 to 520 mV on the ORP monitor. This equates to 0.1 to 0.4 ppm of available free chlorine.

A hydrochloric acid solution was dosed continuously into the electrode stream to prevent deposition of acid-soluble compounds on the cathode. The electrode stream discharged into the brine recycle stream thereby lowering its pH to prevent fouling by deposition of insoluble salts on the membranes and in the separators between the membranes. For the first 6 weeks, the dose rate was set to maintain the brine pH at around 6.7. Later the acid dose rate was increased to lower the brine pH to 6.0 (see BBD Figure 17) to cope with increased feed turbidity. However, despite lowering the brine pH, fouling still occurred so dosing a cationic anti-scalant Hypersperse MDC150 into the brine stream at 15 ppm commenced on 23 June. The acid dosing was decreased and maintained pH at 6.5 - 6.7 throughout most of the remainder of the first period.

CIPs were carried out on 1 May, 22 May, 30 May, 8 June, 19 June and 23 June as indicated in the SIP fluctuations in Stack Inlet Pressure. When a CIP was carried out, a 5% hydrochloric acid solution was achieved by inducting approximately 20 litres of 33% HCl into the recirculating stream through the EDR stack during the CIP.

Figure 9 shows the stack inlet pressure (SIP) over the first 6 weeks of continuous operation. For the first few days of May, the stack inlet pressure was relatively constant, after which there was a period when the EDR stack was shut down. The rises in SIP indicate fouling of the separators between membranes. The system is designed to operate with SIP of approximately 2 to 2.5 kg/cm2. Fouling of the system causes the product flow rate to drop, the product conductivity to increase, and an overall decrease in stack current. The inlet pressure typically increased over a 4-6 day period, shown as ‘SIP increase’ in Figure 9. During these excursions, the pressure exceeded 2.6 kg/cm2 and was as high as 3 kg/cm2.

Power and water outages occurred on a number of occasions. Some of the power outages coincided with a MMF backwash that prevented the multimedia filters changing over. This meant the system had to be reset, which required the plant to be shut down.

For the first 6 weeks operation, both the EDR and MMF were operating satisfactory, but improvement was needed to extend times between CIP cleans. The MMF was producing water consistently below 0.7 NTU, and for most of the time below 0.5 NTU. After the alterations in mid June no further CIP were required.

Period 2: October 18 2006 – Feb 02 2007

Plant operation recommenced on Monday October 18th and a CIP was carried out close to this date. As shown in


Figure 6, the feed turbidity also rose from 2.0 NTU to 5 - 6 NTU in November and then in December came back down to ~2.0 NTU. The plant was shut down on 21 December due to the Christmas break. Only 51 days were available for operation during these months. After the initial CIP there were no further CIP in November or December. The pH of the brine was adjusted to around 7.0 to assist in preventing scaling, and the EDR was run for 24 days without a CIP. Due to the high turbidity in the feed there was evidence of fouling as the product conductivity increased and current decreased in the first stage of EDR. Despite this less than ideal feed, the EDR product conductivity still remained between 475 to 670 µS/cm, which demonstrated the robustness and flexibility of the EDR system. Under similar circumstances, an RO system would require shut down to prevent blockage by particulate material. During mid-November when the feed turbidity was 5.0 - 6.0 NTU, trials were conducted using 2 to 3 ppm of Klariad PC1190 (a cationic polymer) in conjunction with the ferric sulphate coagulant to help reduce turbidity of the MMF filtrate. Without Klaraid PC1190 turbidity was averaging around 1.30 NTU, and with Klaraid PC1190 it decreased to an average of 0.9 NTU, still greater than the desirable 0.5 NTU. Over the entire trial there were very few problems with the EDR but it is ultimately dependent on the pre-treatment and the scaling control strategies. With more sophisticated pre-treatment giving turbidity readings continually below 0.5 NTU, the EDR would be expected to run at least 60 days plus between CIP. During the trial it was realised that it would have been better to have a break tank between the EDR and MMF. This would have provided a buffer between the systems and also provided a source of filtered water with which the EDR system could have been flushed in the event of shutdown of the MMF. There were many outages caused when the MMF filters changed over, since this resulted in a low flow and a low differential pressure for about 10 seconds that was sufficient to initiate an automatic immediate critical shut down of the EDR. Plant operation ceased over the Christmas break and recommenced on Tuesday 2 January 2007. A CIP was carried out on Thursday 4 January. The feed turbidity remained below 2.0 NTU for most of the month but rose dramatically in early February so that the whole plant had to be shut down on 4 February. Only 33 days were available for EDR during these months.

After the CIP in January there were no further CIP cleans. The pH of the brine was adjusted to around 7.0 (Figure 17 and Figure 23) to assist in preventing scaling from the brine stream. EDR ran successfully for 25 days without needing a CIP with EDR product conductivity between 398 and 655 µS/cm. On 19 January the recovery rate was increased to 89% and the EDR continued at this level until 4 February. During this time the quality of the EDR product was still within acceptable levels. It was evident during the higher recovery that scaling was occurring as both conductivity and current levels were changing.

The SIP increased during November and stabilised in December – see Figure 21. The product conductivity (Figure 20) was high, approximately 600 µS/cm at this stage coinciding with the high SIP and low stack current, however the system kept operating despite less than optimum performance. The CIP on 4 January restored optimum performance until ~26 Jan when BBD flow rate was decreased causing an increase in conductivity and an increase in SIP.

Stack Currents

From the plots of currents in the two stacks in Figure 16 and Figure 22 it is clear that by mid-June, the fouling issues were solved. This meant that the requirements for CIP declined and the currents and hence the other parameters, especially the product conductivity (Figure 14), stabilised.


To highlight the direct link between stack current and conductivity, Table 6 contains conductivity and current data for the two stages at the time of sample collection of the various streams for analysis. The changes in conductivity, delta K1 and delta K2 after each stage of EDR were calculated and plotted in Figure 24. These results show the direct relationship between current and salt removal from each stage - in EDR, salt removal is controlled by current. Current in turn is maintained by minimising scaling and increasing voltage to compensate for decreases in temperature.

Table 6 Dependence of conductivity changes with stack current (* shows best performance)

Sampling Date

*

MMF Prod

µS/cm

EDR PROD 1 µS/cm

Current 1 Ampere

Delta K1 per amp

EDR PROD 2 µS/cm

Current 2 Ampere

Delta K2 per amp

31-Oct-06 1880 942 16 58.63 571 6.4 58.0 20-Nov-06* 1880 874 18 55.89 440 7.2 60.3 23-Nov-06 1900 978 15.8 58.35 553 7 60.7

8-Dec-06 1930 1040 14.5 61.38 512 9.2 57.4 15-Dec-06 1930 1200 13.5 54.07 656 9.8 55.5 21-Dec-06 1940 1160 14.5 53.79 588 9.9 57.8 12-Jan-07* 1950 975 18 54.17 445 9.2 57.6 17-Jan-07* 1940 956 18 54.67 441 9.5 54.2 29-Jan-07 1960 1130 16 51.88 529 10.4 57.8

Figure 24 Dependence of conductivity changes with stack current for both stages of EDR

Conductivity Dependence on Stack Current

0

200

400

600

800

1000

1200

0 5 10 15 20

Current, Ampere

Cond

uctiv

ity D

ecre

ase

µS/c

m

Delta K Stage 1Delta K Stage 2

Operational problems

The need for clean-in-place (CIP) was indicated by an increase in stack inlet pressure (see Figure 9), which occurred due to physical obstruction in the separators. This also caused increased electrical resistance and a rise in product conductivity (see Figure 14 and Figure 20). Other fouling such as the adsorption of organic species also contributed to reducing the current flow through the stack. Acid CIP’s took place on May 1, 22 and 30 and June 8, 19 and 23, as indicated in Figure 9. These resulted in a decrease in product conductivity and stack inlet pressure, and also an increase in stack current, especially for the first stage. The current in the second stage was not as affected by fouling because the conductivity of its feed usually increased because it was the product from the first stage.

As discussed in the pre-treatment section, there were significant turbidity fluctuations during this second period of extended operation.


Figure 6 shows turbidity variations of feed from HORS over the period of the trial. It is evident that up until mid December, there were large fluctuations in raw water turbidity that made the pre-treatment and EDR operation challenging to say the least. Nonetheless, the system continued to operate despite difficulty in producing low turbidity feed from the MMF.

In all, 55.6 days of additional operation were achieved. The longest continuous runs were 10 days 18 hours, 12 days 18 hours and 13 days 4 hours.

During this time there were many shutdowns, occurring predominantly during the changeover between MMF filters. The hours of operation of the MMF EDR pilot over the period are summarised in Appendices 2 and 3

Brine Management

The objectives of the second stage were

• to achieve a longer period until a CIP is required (greater than 1 month) and • further optimisation of major operational parameters (current, flow, chemical dosing)

to improve efficiency with respect to product yield and consumption of energy and chemicals.

The system was restarted on 2 January and an alkaline-brine CIP and an acid CIP were carried out on 4 January after which there was a drop in product conductivity. At that stage there had been approximately 33 days operation since the previous CIP.

The intermittent stoppages highlighted in Appendix 3 made it very difficult to carry out the desired optimisation. The main parameters that were varied were stack voltage and brine blowdown flow rate.

The trial started using both stages of the EDR stack on the lowest stack voltage available. During winter when the conductivity dropped due to the lower temperatures, the stack voltages were increased to the second lowest values available. During the second period, the first stage was operated predominantly on the lowest voltage, with approximately 50% reduction in conductivity. By operating the second stage at the third lowest voltage, the salt removal in the second stage was increased to approximately 50% also, with an increase in production efficiency.

Provided the pre-treatment is under control, management of the brine is the key element to keeping the stack free of fouling and maximising production efficiency. As the brine blowdown flow decreases, the production efficiency increases, as does the brine concentration. Precipitation in the brine was controlled by the addition of hydrochloric acid to lower the pH, and by the addition of anti-scalant, forming complexes.

The brine blow-down (BBD) flow rate determines production efficiency, since the feed to the EDR from the MMF is shared between the product and the BBD:

Production efficiency = {Feed flow rate –BBD}/Feed flow rate.

As BBD is increased, the brine becomes more prone to scaling by calcium phosphate in particular and to a lesser extent by calcium fluoride.

An algorithm obtained from Ionics (now GE) was used to calculate the minimum pH required to avoid scaling. It required the collection of brine samples over a range of BBD flow rates and analysed for calcium, magnesium and phosphate. The results of the analysis are included as Table 7. This suggested that a BBD of 0.7 m3/hr could be used with a BBD pH of 6.7. This was implemented during the last 10 days of the trial commencing on January 11


and continuing after a stoppage between January 12 - 17. However, as indicated in Figure 20, the product conductivity gradually increased, indicating gradual accumulation of scale in the system. Clearly the brine pH was set too high and stable operation would have required maintenance of the BBD pH at a slightly lower value, consuming more hydrochloric acid and/or anti-scalant.

Table 7 Brine analyses for a range of brine blowdown flow rates

BBD kL/hr Conductivity, mS/cm Ca Mg Na P pH min, calc 0.5 13.05 390 230 1850 75 6.63 0.6 11.74 405 230 1535 65 6.69 0.6 12.13 350 210 1695 65 6.77 0.7 10.80 280 175 1520 60 6.97 0.8 10.06 280 170 1337 55 7.02

0.89 8.89 220 140 1205 52 7.3 Units mg/L unless stated otherwise


Efficiency and Energy calculations with respect to product and waste formation

Energy usage

Total electrical energy usage for the combined MMF EDR system was displayed on an energy meter in the EDR container. This was recorded manually during visits to the plant and is plotted for the two periods of continuous operation in Figure 25 and Figure 26. There was no separate logging of electrical energy for MMF and EDR. The energy used by the MMF was measured with the EDR turned off but still in-line with the MMF product. It was estimated to be 0.35+0.05 kWh/m3. It was also calculated from extended periods that included joint MMF-EDR operation and MMF operation alone.

Figure 25 Electrical energy meter readings for MMF+EDR from May 2 to July 31

May-July 2006 MMF EDR Energy Meter

71000

72000

73000

74000

75000

76000

77000

78000

79000

13-May-06 27-May-06 10-Jun-06 24-Jun-06 8-Jul-06 22-Jul-06

Date

Ener

gy, k

Wh

Figure 26 Electrical energy meter readings for MMF+EDR from Oct 16 – Feb 4

Oct-Feb 2007 MMF EDR Energy Meter

82000

83000

84000

85000

86000

87000

88000

89000

90000

27-Oct-07

10-Nov-07

24-Nov-07

08-Dec-07

22-Dec-07

05-Jan-08

19-Jan-08

02-Feb-08

16-Feb-08

Date

Ener

gy, k

Wh


It is clear from data in Figure 25 over the period 27 May to 20 July that the process was continuous and energy usage was basically constant apart from a few stoppages. It is also clear from the predominantly non-linear data in Figure 26 that the second period of operation from mid-October to early February was intermittent with only brief periods of continuous operation.

Energy and water efficiency calculations were based on three periods of extended continuous operation:

• 26 May 2006 – 23 June 2006 • 27 June 2006 – 14 July 2006 • 24 January 2007 – 30 January 2007.

Flow volumes and energy usage for the three selected periods of operation are summarised in Table 8 and Table 9 along with efficiency calculations. The energy usage per unit volume was reasonably consistent over the three test periods at 0.99+0.05 kWh/m3 product, after correction for off-spec product (OSP). After subtracting the component estimated to be consumed by the MMF (approximately 0.4 kWh/m3), the electrical energy usage by pumping and electrolysis in the EDR was 0.59+0.10 kWh/m3 product. This is consistent with estimates provided by Ionics at the start of the project using feed water quality data for WTP treated effluent.

Water usage

Efficiencies of water usage were calculated by taking into account the losses by backwashing in the case of the MMF, and in losses as brine blowdown and off spec product in the case of EDR. Thus the volume of filtered product from the MMF was 97.2% of the volume of Class C feed and the volume of EDR product was on average 80% of the volume of MMF filtrate. The average overall efficiency of the combined MMF/EDR process was thus approximately 78%.

There is scope for improving the efficiency of the MMF by further optimisation of coagulation to allow for an increase in the backflush time. In the case of EDR, the volume loss as brine blowdown could be further decreased by increased use of the acidic and/or complexing anti-scalants that have been used; however there would be an increase in operation cost that would need to be considered. Likewise if the set point for OSP was adjusted to collect product at a higher conductivity after reversal, more product could be collected but the overall conductivity of the product would increase slightly. All of these variables could be altered in a fully-operational plant to achieve optimum performance.

Table 8 Water usage and energy calculations for three selected regions of continuous operation

Period Total of two Filters m3

Backflush m3

EDR Product m3

EDR less OSP m3

BBD m3

Total Energy kWh

Energy per m3 kWh/m3

EDR+MMF Power, kW

26 May 2006 – 23 June 2006

5022 114.4 3231 3125 667 2965 0.95 4.4

27 June 2006 – 14 July 2006

3706 102.3 2730 2640 602 2656 1.00 5.0

24 Jan 2007 – 30 Jan 2007

979 24.3 751 725 95 751 1.03 5.0


Table 9 Water efficiency calculations for three selected regions of continuous operation

Period % MMF Product

% EDR Product

Overall efficiency

Av EDR m3/h

Av MMF m3/h

26 May 2006 – 23 June 2006

97.3 78.7 76.5 4.59 6.10

27 June 2006 – 14 July 2006

97.0 77.9 75.5 4.95 6.4

24 Jan 2007 – 30 Jan 2007

97.5 84.3 82 4.95 6.6

Average 4.83

Chemicals usage

Consumption rates of chemicals were measured throughout the trials. These are summarised in Table 10, along with indicative pricing of bulk chemicals on the scale appropriate for a 50 ML/day plant. Using those prices, the cost for each chemical is indicated. Although the main use for hydrochloric acid was for controlling the pH of the brine, if there was a requirement for an acid CIP every 60 days using 20L hydrochloric acid, this would constitute an additional daily HCl requirement of 0.3 L/day, an additional 3% or 0.6 cents per m3.

Thus, the cost of chemicals per 1000 L of as-produced EDR product at approximately 500 µS/cm is 5.91 + 0.06 cents per m3. The MMF uses 2.12 cents per m3 and the EDR 3.79 cents per m3. The cost for a full scale plant are calculated later (see Table 15).

Table 10 Chemical usage for three selected regions of continuous operation

Period EDR Product m3/day

Sodium Hypochlorite L/day

Ferric Sulphate L/day

Hydrochloric Acid L/day

Anti-scalant, L/day

Total cost

Mean consumption rate

118 6.7 2.0 8.7 .334 n/a

Cost bulk chemical n/a $300/ m3 $250/tonne $320/tonne $7.56/kg n/a Cost bulk chemical, cents/L

n/a 0.30 0.17 0.28 6.69 n/a

Chemical cost, cents/day

n/a 2.01 0.33 2.40 2.23

Cost 2-stage EDR product, cents/m3

n/a 1.70 0.28 2.03 1.90 5.91


Chemical analysis of the various streams in the MMF/EDR process.

Relation between analyses and time of sampling

Samples of feed water, pre-treated feed, product and concentrate were taken at regular intervals for analysis by Melbourne Water as described in the Smart Water testing program. Flows, pressures, conductivity, temperature, pH, current and voltage were monitored continuously and collected and stored from the Watereye website.

Samples of MMF feed, MMF product, EDR Product 1 (from 1 June 2006), EDR Product 2, Brine were taken weekly, and MMF backflush waste was sampled occasionally. This was undertaken by Melbourne Water contractors during the first test period and by CSIRO staff during the second test period. Representative results are included in Appendices 3 and 4, and these results are snapshots that are complementary to the data that was logged continuously. The interpretation of these analyses is complicated by the fact that they are representative of the conditions at the instant of sampling.

Figure 27 Stack voltages and currents, and stage 2 conductivity at the instants of water sampling during the first period of operation.

Stack voltages and currents, and conductivity after stage 2 EDR

40

50

60

70

80

90

100

2-May-06 12-May-06 22-May-06 1-Jun-06 11-Jun-06 21-Jun-06 1-Jul-06 11-Jul-06 21-Jul-06 31-Jul-06

Date

Stac

k vo

ltage

(V) o

r Con

duct

ivity

/10

(µS/

cm)

0

4

8

12

16

20

24

Stac

k cu

rren

t (A

)

Voltage stage 1 Voltage stage 2 EDR product conductivity Current stage 1 Current stage 2 Total Curent

Voltage 1

Voltage 2

Current 1

Current 2

Total Current

Stage 2 Product conductivity /10

Figure 27 and Figure 29 show stack currents and voltages during the two test periods. The removal of salt by EDR is directly related to the direct current passing though the stacks. In a well-run system the operational parameters, i.e. stack voltage and current, would be relatively constant, only changing with temperature, feed composition and stack fouling. As shown in Figure 27, the applied voltages were held constant until 15 June whereupon they were increased to compensate for the decrease in temperature, with the consequence that the currents in both stages increased.

These plots are included because they show the stack conditions at the time of water sampling. The plots in Figure 27 and Figure 29 show the variations of the measured parameters over time for the various streams. In response to the variations in currents at the sampling times, there are corresponding variations in removal of species by EDR. In order to


calculate mean optimum removal efficiencies from the analytical data, averages were taken for sampling dates when the system was running at its optimum. During the first test period, optimum stable operation was observed for sampling dates over the period 6-27 July.

When viewed along with Figure 27, Figure 28 shows the dependence on stack current of electrical conductivity and hence salt content in the various streams during the first test period. As the current decreases the conductivity increases because there is less salt being removed. Conversely, as the current increases the conductivity decreases because there is more salt being removed. The relatively constant conductivity of products from both stages over the preferred sampling period 6-27 July is clear from Figure 28.

Figure 28 Conductivities of EDR feed and two product streams May 2- July 31

Electrical Conductivity

0

200

400

600

800

1000

1200

1400

1600

1800

2000

04-May-06

11-May-06

18-May-06

25-May-06

01-Jun-06 15-Jun-06 06-Jul-06 13-Jul-06 20-Jul-06 27-Jul-06

Date

µS/c

m Electr. Cond. - EDR PROD 2Electr. Cond. - EDR PROD 1Electr. Cond. - EDR FEED

Towards the end of the trial in the summer of 2006/7, there was a gradual increase in feed conductivity presumably due to the drought and the impact of water restrictions and reduction in the diluting effect of low salinity waters like shower and other grey water sources that were progressively diverted to domestic applications like gardens and other uses.

Composition of various streams from the MMF/ EDR plant

The simplest indicator of concentration is conductivity since most of the dissolved materials in treated sewage are salts. Conductivity of the product after two-stage EDR was measured continuously during this study as displayed in Figure 14. The fluctuations in conductivity are largely due to fouling and to a lesser extent the applied voltage. Because samples for analysis were collected at discrete moments, they do not reflect optimum performance but rather the performance at that moment.

First period of testing

In Figure 27 there was a period of stable operation between July 6 and July 27 so it is reasonable to expect that the analyses during that period were a good indicator of best performance for the MMF EDR system. The analytical results for this period are listed in


Table 11 and displayed in decreasing order with respect to % removal after the first stage and after the combined two stages of EDR.

MMF feed and product

As mentioned earlier, the main influence of the coagulation process was the removal of suspended material (turbidity) and a slight increase in the sulphate content of the feed – see Table 3.

Product stream

Table 11 Analytical data for various streams over the period July 6 - July 27 2006

Species EDR Feed

EDR Prod1

EDR Prod2

EDR Brine

% removed Stage 1

% removed Stages 1 & 2

Calcium 33 12 5.4 177 63 84 Strontium 0.18 0.08 0.035 0.98 57 81 Magnesium 21 9.75 4.45 130 54 79 NO3

- as nitrate 41 18 9.2 211 56 77 Chloride 365 167 88 2500 54 76 TDS 998 462 302 4000 53 70 Conductivity, µS/cm 1700 897 542 7050 47 68 Sodium 245 130 88 1075 47 64 Fluoride 1.85 1.15 0.85 7 38 54 Alkalinity 132 86 61 360 35 54 Total P as phosphate 26.4 18.8 12.8 94.5 29 52 TSS 2.5 1.75 1.25 6.25 30 50 Manganese 0.00175 0.00125 0.001 0.0055 29 43 Total Kjeldahl N 1.8 1.7 1.2 3.4 5.6 33.3 Iron 0.077 0.062 0.06 0.18 19 23 TOC 12.5 11.5 10.3 24 8 18 S as sulphate 98.5 78 88 342.5 21 11 SAR 7.57 6.72 6.8 8.17 n/a n/a Boron 0.28 0.26 0.26 0.28 6.3 6.3 Ammonia 0.15 0.15 0.15 0.125 0.0 0.0 Barium 0.002 0.002 0.002 0.0053 0.0 0.0 Silicon 9.6 9.65 9.85 9.07 ~0 ~0 COD 12.7 14 14.7 66.5 ~0 ~0 Nitrite 0.0077 0.0075 0.009 0.011 ~0 ~0

TOC - Total organic carbon; SAR Sodium adsorption ratio; COD Chemical oxygen demand; TDS Total dissolved solids; TSS Total suspended solids Units mg/L unless stated otherwise

EDR Product

• It is clear that the divalent cations Ca, Mg and Sr are removed most effectively, followed by the monovalent anions nitrate and chloride and then the monovalent species sodium, fluoride and bicarbonate (alkalinity).

• The polyvalent anions sulphate and phosphate were less effectively removed. In the case of phosphate, nitrate and ammonia this may be advantageous for the intended use of the desalinated effluent in irrigation.

• There was some removal of organic carbon (TOC), presumably the anionic components, but negligible change in COD.

• As expected, silicon was not removed by EDR, which is in contrast with RO where it can present a substantial fouling problem.


• Boron was also not removed by EDR because it exists in neutral form at the feed pH. • SAR, the sodium adsorption ratio, decreased in the first stage of EDR but the

expected desirable impact of removal of calcium and magnesium was partially offset by the lower removal of sodium.

Brine stream

Table 11 also includes analysis of the brine stream and represents the accumulation of ionic material removed from the product stream. Of particular relevance are the concentrations of ions at levels at which precipitation of insoluble materials, especially calcium phosphates and calcium fluoride is likely. This is controlled by the use of anti-scalants that reduce pH or complex calcium, and by limiting the concentrations of the individual ions (Ca2+, PO4

3- and F-

by dilution with overflow as brine blowdown.

Second period of testing

Figure 29 Stack voltages and currents and stage 2 conductivity at the instants of water sampling during the second period of operation

Stack voltages and currents, and conductivity after stage 2 EDR

30

40

50

60

70

80

90

100

31-Oct-06 20-Nov-06 23-Nov-06 08-Dec-06 15-Dec-06 21-Dec-06 12-Jan-07 17-Jan-07 29-Jan-07

Date

Stac

k vo

ltage

(V) o

r Con

duct

ivity

/10

(µS/

cm)

6

10

14

18

22

26

30

34

Stac

k cu

rren

t (A

)

Voltage Stage 1 Voltage Stage 2 EDR product conductivity Current Stage 1 Current Stage 2 Total current

Current 1

Voltage 2

Current 2

Voltage 1

Stage 2 product conductivity/10

Total Current

Table 12 shows the compositions of streams during the second test period and Figure 29 shows stack voltages and currents for the second stage of EDR at the instants of water sampling. The sample dates chosen for the data in Table 12 were 20 Nov, 12 Jan and 17 Jan when the total currents were highest and the conductivity was lowest (see Figure 29). Comparison of the percentage removed columns of Table 11 and Table 12 indicates greater removal in the second test period. This is best indicated from the conductivity data that shows 51.5 % removal in the first stage and a further 25.6% removal (equivalent to 50% of the material remaining after the first stage). The reason for the greater removal of salts during the second period of testing was the use of increased voltages on the stack with consequential higher currents in periods of low stack fouling (see Figure 29). Had higher currents been used in the first period, similar performance would have been expected.

Disposal or further processing of the brine stream is a key aspect of a desalination installation. The brine contains those components that have been removed by EDR and also any organic anti-scalant that may have been used. It also contains all of the components of


treated effluent that are not affected by EDR as does the EDR product – e.g. dissolved silica and boron species as well as uncharged organic and biological material.

Table 12 Analytical data for various streams over the period October 18 – Feb 4

Species Raw Feed

EDR Feed

EDR Prod 1

EDR Prod 2

EDR Brine

% removed Stage 1

% removed Stages 1 & 2

Strontium 0.23 0.21 0.08 0.02 1.50 61 90 Calcium 38.5 37.4 14.8 4.00 254 60 89 Magnesium 28.5 28.4 11.4 3.32 189 60 88 Potassium 39.1 39.3 15.17 4.94 176 61 87 Chloride 429 4450 1780 66.3 2915 60 85 NO3

- as nitrate 55 54 23 8.77 282 57 84 TDS 1135 1146 563 260 7088 51 77 Conductivity, µS/cm 1906 1926 935 442 8690 52 77 Sodium 325 329 188 84.6 1515 43 74 S as sulphate 108 113 72.33 40 546 36 65 Alkalinity 137 134 81.5 48.1 469 39 64 Total P as phosphate 35.9 32.8 23.5 13.2 130 28 60 Fluoride 2.3 2.3 1.4 0.93 9.9 39 59 SAR, ratio 9.69 9.86 8.92 7.57 17.5 n/a n/a TOC 8.63 8.83 8 7.37 17.7 9.4 17 Iron 0.12 0.06 0.05 0.05 0.05 11 12 Manganese 0.09 0.01 0.01 0.01 0.07 0.0 0.0 Silicon 13.9 13.1 14.6 13.5 11.5 ~0 ~0

TOC - Total organic carbon; SAR Sodium Adsorption Ratio; TDS Total dissolved solids Units mg/L unless stated otherwise

MMF Waste stream.

Very few analyses were performed on the MMF waste stream. In fact the sample collected on 11 May 2006 showed the highest concentrations of species expected to be removed by coagulation/filtration, as shown for total suspended solids in Figure 30. The period of the backflush cycle at the time was nine hours.

The major components in the MMF backflush were Fe, COD and TSS. Apart from those species, there was little difference between other components in the MMF feed and product.

Figure 30 Total suspended solids (TSS) in MMF waste stream over the first test period May2 – July 31

Total Suspended Solids

0

50

100

150

200

250

300

350

400

13/0

4/20

06

20/0

4/20

06

4/05

/200

6

11/0

5/20

06

18/0

5/20

06

25/0

5/20

06

1/06

/200

6

15/0

6/20

06

6/07

/200

6

13/0

7/20

06

20/0

7/20

06

27/0

7/20

06

3/08

/200

6

17/0

8/20

06

Date

mg/

L

TSS - EDR PROD 2TSS - EDR BRINETSS - MMF WASTETSS - EDR PROD 1TSS - EDR FEED


The bar chart in Figure 31 shows the concentrations in the MMF waste stream along with the corresponding concentrations in the feed. A logarithmic scale is used to highlight the differences where one of the concentrations was very small. Given that the backflush time was 10 minutes and the backflush flow rate was ~10.4 m3/hour, the removed species were actually dissolved in 1.73 m3. Thus the actual amount of material removed in the backflush was 1730 times the concentration of the diluted sample. It should be noted that, although alkalinity is expressed as mg CaCO3/L, a substantial proportion of the alkalinity is phosphate because the acid titration does not distinguish between the various basic anions.

Figure 31 Concentrations of major species in the MMF waste stream cf feed, 11 May 2006

Major components in MMF waste compared with MMF feed

0.001

0.01

0.1

1

10

100

1000

Cal

cium Iro

n

TSS

CO

D

Stro

ntiu

m

Man

gane

se

Pho

spha

te

Nitr

ite

Alk

alin

ity

TKN

TOC

Bar

ium

Component

Con

cent

ratio

n, m

g/L

MMF feed MMF Waste

The masses of key components removed by MMF per cycle are shown in Table 13. Given that the backflush frequency was 9 hours, and that during that time the volume of water that passed though the filter was 61 m3, the mass of material removed by MMF per megalitre of EDR product at approximately 500 µS/cm can be calculated and is indicated in the right hand column of Table 13. As expected, suspended material was the major component of the waste stream. It clearly contains most of the added ferric coagulant and a large amount of chemically oxidisable organic material as well as precipitated calcium and ferric phosphates. As indicated earlier, phosphate would have been a significant component of the measured alkalinity. The presence of these components would need to be taken into account in the disposal of this waste stream. Separation as sludge would be the preferred option.

Table 13 Masses of key components removed by MMF per cycle

Component MMF feed mg/L

MMF Waste mg/L

Mass removed in 1730m3 backflush cycle, g

Mass removed per ML EDR product, g

Calcium 36 64 111 2343 Iron .05 87 151 3185 TSS 5 380 657 13913 COD 19 230 398 8421 Strontium .19 0.77 1.3 28 Manganese .003 1.5 2.6 55 Phosphorus 29 95 164 3478 Nitrite .01 0.57 1.0 21 Alkalinity 96 170 294 6224 TKN 2.2 15 26 549 TOC 13 24 41.5 879 Barium .002 0.16 0.3 6


Blending to produce 1000 µS/cm product

Given the formula for calculating the sodium adsorption ratio (SAR)

SAR = [Na+]/√([Ca2+] + [Mg2+])

it was anticipated that, with EDR, there may have been selective removal of divalent ions calcium and magnesium with an increase in the SAR or at best maintenance of SAR. However the reality has been the opposite due to the square root factor in the denominator.

It is unlikely that any adjustment to the composition of the product steam after two stage EDR will be necessary, with a typical pH of 7.0-7.5, TDS of 250-300 mg/L and sodium adsorption ratio of approximately 7. Addition of lime, possibly in conjunction with CO2 to control pH, could be used to lower the SAR if required.

Table 14 Properties of the solutions before and after blending to produce 1000µS/cm mixtures

Sample Description Parameter

Units

EDR Feed

Product Stage 1

Product Stage 2

1000µS/cm Mixture from

Stage 1

1000µS/cm Mixture from

Stage 2

Conductivity μS/cm 1653 828 483 1000 1000

Volume mL 56.5 200 256.5

Volume mL 85 100 185 m3 1000µS/cm

product

per m3 EDR product

1.28 1.85

Calcium mg/L 43 21 9.5 25 26

Potassium mg/L 26 10 4.8 15 14

Magnesium mg/L 29 15 7.3 18 18

Sodium mg/L 275 137.5 87.5 170 165

Phosphorus mg/L 8.3 6 4.4 6.2 6.5

Sulphate mg/L 31 20 14 22 23

SAR n/a 7.94 5.60 5.19 6.33 6.09

The only treatment post-EDR would be any blending with feed to raise the salinity within acceptable limits to increase yield, followed by disinfection.

The conductivity records throughout the trial indicated that EDR through a single stage was sufficient to produce salt-reduced product of conductivity less than 1000 µS/cm and that two stage EDR resulted in a product of much lower conductivity than the Melbourne Water specification for its salt-reduction trial. This means that by blending with treated feed, the yield from single or two-stage EDR could be increased.

Table 14 shows the levels of some major components of solutions that were prepared by blending the product streams from each stage of EDR with EDR feed (MMF product) to produce two 1000 µS/cm mixtures. For this particular set of samples, the SAR after stage 2 was lower than either the feed or that after stage 1. The SAR increased after blending with feed because of the higher sodium content of the feed. However, the SAR of the blended products are still less than the feed and thus would be expected to be more suitable for use as irrigation water because of the lower SAR and lower salt content over all, while retaining some of the nutrients.


Using this information, the volume of EDR product water available at 1000 µS/cm may be calculated. Thus each cubic meter of stage 1 product results in 1.85 cubic meters at 1000 µS/cm, and each cubic meter of stage 2 product results in 1.28 cubic meters at 1000 µS/cm.

The energy cost per cubic metre of product at 1000 µS/cm may be calculated using the earlier findings that the energy for MMF/EDR and MMF alone were 1.0 kWh/m3 and 0.4 kWh/m3 respectively. Since twice the amount of salt is removed in stage 1 than in stage 2, it is reasonable to assume that the electrical energy for stage 1 is double that for stage 2. Consequently, if the energy for EDR stage 1 and 2 are 0.4 and 0.2 kWh/m3 respectively, then the energy for MMF +EDR stages 1 and 2 are 0.8 and 0.6 kWh/m3 respectively.

The following relationship can be used to calculate the energy requirements of EDR:

Energy blended product/m3 = volume MMF Prod * energy EDR-MMF Prod/m3 + volume MMF Prod * energy MMF Prod/m3

Then:

E1 = 1.0 * 0.8 + 0.28 * 0.4 = 1.52 kWh for 1.28 m3 or 1.18 kWh/m3 of product at 1000 µS/cm,

E2 = 1.0 * 1.0 + 0.85 * 0.4 = 1.34 kWh for 1.85 m3 or 0.72 kWh/m3 of product at 1000 µS/cm,

where E1 and E2 are the energy required per cubic metre of 1000 µS/cm product from Stage 1 and 2, respectively.

Alternatively, if the blending was achieved using raw feed, the respective energies would be:

E1 (Blending with raw feed) = 1.0 * 0.8 /1.28 = 0.625 kWh/m3 of product at 1000 µS/cm

E2 (Blending with raw feed) = 1.0 * 1.0 /1.85 = 0.5 kWh/m3 of product at 1000 µS/cm

The cost of chemicals for the different scenarios may also be estimated using the figures calculated earlier, namely $.0198/m3 for MMF alone and $.0591/m3 for MMF/EDR. The chemical cost for single or dual stage EDR would be similar because the purpose of anti-scalants dosing is to control scaling in the brine and that is independent of the whether the EDR process is single or double (or even triple) stage. For a given feed, the brine composition is likely to be similar regardless of the number of stages of EDR, although some fine tuning could be necessary.

A similar approach may be taken to determine the cost of chemicals, C1 and C2, for 1000 µS/cm product from EDR stage 1 and 2 respectively using the equation:

Chemicals/m3 = volume MMF Prod * chemicals for EDR-MMF Prod/m3 + volume MMF Prod * chemicals for MMF Prod/m3.

Thus:

C1 (Blending with filtered feed) = 1.0 * 0.0591 + 0.28 * 0.0198 or $0.0.0646 for 1.28 m3 or $0.051/m3 of product at 1000 µS/cm

C2 (Blending with filtered feed) = 1.0 * 0.0591 + 0.85 * 0.0198 or $0.0759 for 1.85 m3 or $0.041/m3 of product at 1000 µS/cm.


Of course if the blending were achieved using raw feed, the component of coagulant and anti-scalant would be reduced so that:

C1 (Blending with raw feed) = $0.046/m3 of product at 1000 µS/cm

C2 (Blending with raw feed) = $0.032/m3 of product at 1000 µS/cm

The results of these calculations are summarised in Table 15

Table 15 Summary energy and chemical data for product at 1000µS/cm using different scenarios.

Conditions Parameter Single Stage EDR Dual Stage EDR Electrical Energy kWh/ML 1180 720 Electrical Energy /ML @ $0.10/kWh

118 72

Chemical Cost $/ML 51 41

Blended to1000 µS/cm with filtered feed

Cost $/ML 169 113 Electrical Energy kWh/ML 625 500 Electrical Energy $/ML @ $0.10/kWh

62 50

Chemical Cost $/ML 46 32

Blended to1000 µS/cm with unfiltered feed

Cost $/ML 108 82

In summary, a plant with MMF and two-stage EDR produces a better product at lower operational cost than one with a single stage. The sodium adsorption ration is slightly lower (6.1 cf 6.3), the energy usage is 30% less and the chemical usage is 18% less.

The most efficient operation of the system was observed near the end of the trial when the stack currents and hence salt removal were highest. Based on these data, in order to produce 50 ML/day of 1000 µS/cm product at an overall water efficiency for two-stage EDR of 84%, an MMF plant of 55 ML/day would be required. Product from an EDR plant producing 27 ML/day would be combined with 23 ML/day filtered product to produce 50 ML/day. The brine blowdown flow would be approximately 5 ML/day.

However the capital cost of a two stage plant is greater than a single stage plant although the additional cost of membranes is partially offset by the components that are common, namely the instrumentation and control system and general infrastructure. A two-stage process involves twice the membrane area, two power supplies and higher pumping capacity.

Based on the experience of this trial, we are convinced that with state-of-the-art EDR plant and improved instrumentation and controls, CIP intervals of greater than 60 days would be feasible. The main factors causing the requirement for CIP in this study related to plant breakdowns that prevented automatic flushing of the EDR system that in turn caused precipitation of scale. More careful control of brine blowdown flow and anti-scalant dosing would be possible on a system that was operating continuously.

Figure 32 is a distant view of the stacks at a 20ML/day EDR plant in La Jolla, San Diego and a close-up full-height view of a stack with the side panel removed and showing the membranes. The fittings on the grass in the foreground are provision for expansion of the plant in the form of 4 ML/day containerised EDR units. The plant takes feed from a clarifier in an adjacent municipal waste treatment plant. It uses cartridge filtration only as the pre-treatment for a two-stage EDR process. Its output is pumped to a dam from which it is used for irrigation. This pre-treatment is quite crude in comparison with the coagulation/MMF used in this study and as a consequence it necessitated quite frequent CIP, more than monthly.


Figure 32 EDR stacks at 20ML/day wastewater desalination plant at La Jolla, San Diego, USA

In order to establish the overall cost of an installation, the capital cost would need to be provided by an equipment provider. Currently GE is the major global provider, although Eurodia / Ameridia is another potential source of EDR equipment and expertise.

Disposal options for brine

The composition of the reject stream from EDR of effluent at WTP will be closely monitored and options for its use as a resource are now considered.

The simplest method for disposal of EDR concentrate is by discharge into Port Philip Bay. The salinity of the EDR concentrate is closer to that of seawater and should be expected to make less environmental impact than the present outfall, apart from elevated levels of nutrients.

By blending the concentrate with the present outfall, there would be a slight increase in concentration, but an overall decrease in the total quantity of solids discharged since some of the dissolved solids previously discharged into the bay would be included in the desalinated product.

Take, for example, the case of a 50 ML/day EDR plant producing a concentrate stream with a conductivity of 8 mS/cm at 10 ML/day. For a treatment plant outfall of 500 ML/day with conductivity 2 mS/cm, this would represent a 1/50 dilution. Thus, the concentration of salts in the outfall would increase by ~approximately 6.5%, but the actual amount of material in the discharge would decrease by ~10%.

Brine as a medium for nutrient removal via microbiological processes.

The ‘waste’ concentrate stream from the EDR process is enriched in ionic species that have been transferred across the electrodialysis membranes. As well as the major components sodium and chloride ions, the minor constituents include phosphate, nitrate and ammonia, essential nutrients for photosynthesis and potential causes of environmental problems in the Bay.

Reduction of the concentrations of these components are key objectives of the Environment Improvement Plan for the Western Treatment Plant (WTP) at Werribee.


Table 163, describes current and future limits for the Western Treatment Plant according to accreditation granted under Section 26 of the Environment Protection Act 1970.

Table 16 Current and future limits for the Western Treatment Plant Waste Indicator* Median 75th%ile** 90th %ile Maximum Carbonaceous Biochemical Oxygen Demand 10 - 20 - Suspended Solids 30 - 60 - Total Phosphorus (as P) - - 15 - Colour (Pt-Co Units) 200 - 400 - Ammonia (as N) † 10 - - 30 Anionic Surfactants - - 0.80 1.0 Cadmium - 0.0011 - 0.005 Chromium - 0.040 - 0.10 Copper - 0.028 - 0.10 Lead - 0.015 - 0.06 Mercury - 0.0004 - 0.001 Nickel - 0.029 - 0.06 Zinc - 0.060 - 0.20 Benzene - 0.005 - 0.01 Toluene - 0.005 - 0.01 Phenols - 0.04 - 0.08 Total PAHs‡ - 0.0004 - 0.0008 Acute Toxicity None - - -

* Unless otherwise specified all units expressed in milligrams per litre. **%ile is the percentile value in the 12 months monitoring data obtained in any financial year. † Ammonia performance limits may be lowered if trials predict a lower ammonia limit is achievable ‡ Includes all alkyl derivatives

Sewage treatment plants operators are increasingly contemplating methods for nutrient removal from effluents, including wetlands, grasses, algal ponds and membrane processes. These methods are generally applied to the effluent after treatment, with the nutrients at a relatively low concentration. After EDR, the nutrients in the waste stream along with the major ionic species are typically concentrated by a factor 4-5.

One advantage of the EDR process over RO is that the current reversal process allows the process to be operated with over-saturation with respect to substances prone to cause scaling (e.g, calcium phosphate by 400% and calcium carbonate by 200%). This means that the allowable concentration of all substances including nutrients is even higher than might be expected and the volume smaller. Thus the economics of any process for nutrient removal should be better, since a smaller volume can be handled at higher concentration.

Brine as a source of chemicals

Considerable energy has been expended in producing the salt-reduced product but also the brine. As seen in Table 11 and Table 12, the concentrations of all ionic species are substantially increased compared to those in the feed, the values being typically higher than those achieved by RO because of the advantages of current reversal in allowing supersaturation with respect to calcium phosphate and calcium fluoride. Further processing of the brine could produce chemicals of value, in particular calcium phosphate. It has been observed during this project that carbon dioxide could be used as an alternative to hydrochloric acid to control the brine pH to avoid scaling. If the CO2 is stripped from the brine blowdown, calcium phosphate is precipitated. If this were implemented in either an RO or an EDR system, not only would a potentially valuable product be made, but the environment impact of the brine disposal would be reduced. 3 Melbourne Water Accreditation discharge waste to the environment, granted under Section 26 of the Environment Protection Act, State of Victoria, Australia, 1970


Performance evaluation of monovalent cation membranes in the CSIRO EDR rig

Backgound

Monovalent-selective membranes are claimed to offer the possibility of selectively removing a larger proportion of monovalent ions while allowing divalent ions to pass through in the product stream of an EDR process. Monovalent-selective cation CMS membranes were obtained from Neosepta and installed on the small EDR stack and commissioned using as feed the product from the MMF in the GE pilot plant. The stack was assembled using alternating monovalent cation membranes, conventional anion membranes and standard Type 1 separators, with conventional heavy cation membranes immediately adjacent to the electrodes.

The stack current was adjusted to generate product with approximately the same conductivity as the first stage of the GE rig. The second stage product was collected and then used as feed with the stack current adjusted to generate product with approximately the same conductivity as the second stage of the GE rig. The GE pilot plant and the CSIRO EDR Stack were operated in parallel and samples of feed and product taken for analysis.

Results

From the results summarised in Table 17 it is clear that the proportion of sodium, indicated by the ratios Ca/Na and Mg/Na, was smaller for the monovalent-selective membranes. For a given product conductivity, the concentration of sodium in each product stream is less when the monovalent membrane is used, and there is proportionally more calcium and magnesium remaining in EDR products from both stages using the monovalent membranes. This resulted in a significant lowering of the sodium adsorption ratio from 9.8 to 7.0 for the first stage product and from 8.1 to 4.5 for the second stage product. In fact the conductivity of the second stage product was higher in the case of the monovalent membrane stack, so it would be expected that the separation of sodium from calcium and magnesium would have been even greater had the conductivities been the same.

Table 17 Electrical parameters and composition* of feed and streams for GE pilot and CSIRO EDR plants

Sample Description

Voltage Current Ampere

Conductivity µS/cm Ca Mg Na Na/Ca Na/Mg SAR

MMF Product n/a n/a 2130# 50 27 250 5 9 10.2

GE Product 1 124 18.4 1045 18 11 155 8.3 14 9.8 Mono. Product 1 24.7 2.0 1039 24 17 135 5.5 7.7 7.0 GE Product 2 99 8.4 534 5.7 4.1 76.9 14 20 8.1 Mono. Product 2 19.1 1.06 602 14 11 70.3 5 6.2 4.5

* Units mg/L unless otherwise stated

The rig was operated for several weeks continuously. As the period of operation progressed, an increasing stack voltage was required for maintenance of the current and hence product conductivity. This indicated increased fouling to a degree not observed previously on the stack with conventional membranes. The image in Figure 33 clearly indicates the deposition of calcium phosphate in the electrode flow channel in the double-thickness outer separator between the electrode and the heavy cation membrane. A hydrochloric acid CIP was

# The feed conductivity at this stage was substantially higher than earlier in the trial, due to the consequences of water saving measures during the drought.


performed on the stack to establish whether the fouling was reversible. The stack inlet pressure and stack voltage returned to values more typical of a freshly cleaned stack.

Figure 33: Blocked electrode flow channel in the separator adjacent to the outer electrode

Conclusions regarding the use of monovalent membranes

This brief trial indicated the benefit of monovalent cation membranes in achieving sodium adsorption ratio that was much lower than when the conventional cation membranes were used. The calcium phosphate fouling was attributed to poor brine management (BBD flow and pH) and was not related to the monovalent membranes. There would be great advantages if these membranes could be successfully incorporated in an EDR plant. The sodium adsorption ratio would be substantially lower than that from conventional membranes and the brine stream could be operated at a higher concentration, with benefits in productivity and/or anti-scalant requirements.


CONCLUSIONS

A commercial electrodialysis reversal (EDR) pilot plant using pre-treatment with multimedia filtration (MMF) has been operated at the Western Treatment Plant over two three-month periods and demonstrated effectively the removal of salt from treated effluent.

The typical product after MMF and two-stage EDR has a conductivity of less than 500 µS/cm, considerably less than the 1000 µS/cm criterion established by Melbourne Water as suitable for irrigation. A final conductivity of approximately 1000 µS/cm was achieved by blending the product with raw feed water thus producing a higher yield from the available plant.

Ferric sulphate was found to be an effective coagulant for feed turbidity less than 5 NTU, but it was not sufficient to cope with the excursions to values as high as 15 NTU. A contingency to co-dose with a polymeric flocculent was developed to reduce turbidity from higher levels.

The MMF plant operated with a water efficiency of 97% while the two-stage EDR plant ran with a mean water efficiency of approximately 81% with a maximum of 84%. A plant with MMF and two-stage EDR produces a better product at lower operational cost. The sodium adsorption ration is slightly lower (6.1 cf 6.3), the energy usage is 30% less and the chemical usage is 18% less.

A single step of electrodialysis reduces the concentrations of all ionic material by approximately 50%, including sodium chloride which comprises the major component of its salinity, as well as other ions including sulphate and carbonate, potassium and magnesium. Importantly, the ionic nutrients nitrate, nitrite, phosphate and ammonia are also removed by similar amounts. All these salts are removed in concentrated form, so that the reduced volume facilitates any subsequent process for their removal.

The most efficient operation of the system was observed near the end of the trial when the stack currents and hence salt removal were highest. Based on these data, in order to produce 50 ML/day of 1000 µS/cm product at an overall water efficiency for two-stage EDR of 84%, an MMF plant of 55 ML/day and a two-stage EDR plant of 27 ML/day would be required. Product from the EDR plant producing 27 ML/day would be combined with 23 ML/day of filtered product to produce 1000 µS/cm product at 50 ML/day. The brine blowdown flow would be approximately 5 ML/day.

The capital cost of a two stage plant is greater than a single stage plant although the additional cost of membranes is partially offset by the components that are common, namely the instrumentation and control system and general infrastructure. A two-stage process involves twice the membrane area and two power supplies and higher pumping capacity.

The energy and chemical components of the operational cost of producing 1000 µS/cm water from EDR of feed water of ~1900 µS/cm depends on the process. After blending with MMF product, the energy cost is 1.18 kWh/m3 after a single stage of EDR and 0.72 kWh/m3 after a two-stage EDR. If blending is performed using unfiltered product, the energy costs are 0.62 and 0.5 kWh/m3 respectively.

In order to establish the overall cost of an installation, the capital cost would need to be provided by an equipment provider. Currently GE is the major global provider, and Eurodia / Ameridia is another potential provider.

Based on an energy cost of $0.10/kWh and using the cost of bulk chemicals, the cost per ML or 1000 µS/cm product blended with filtered feed is $169/ML and $113/ML for single and two-stage EDR respectively and $108/ML and $0.82/ML respectively when blended with unfiltered feed.


There was no apparent deterioration of the MMF system or EDR stack after the testing program. Restoration of performance of the stack to its original was always achieved by a conventional hydrochloric acid CIP. A mini-CIP, developed during the trial, was also found to remove the effects of scaling without the necessity to cease production. Staining of the anion membranes, presumably by organics, did not appear to have any significant effect on cell performance.

Disruptions to plant operation occurred largely due to excessively high turbidity from the inlet water supply as well as electronics and software problems in the sensing and control systems. Based on the experience of this trial, we are convinced that with state-of-the-art EDR plant and improved instrumentation and controls, CIP intervals of greater than 60 days would be feasible. The main factors causing the requirement for CIP in this study related to plant breakdowns that prevented automatic flushing of the EDR system that in turn caused precipitation of scale. More careful control of brine blowdown flow and anti-scalant dosing would be possible on a system that was operating continuously.

The 20 ML/day EDR plant in La Jolla, San Diego provides an example of the application of EDR in a situation very similar to the Western Treatment Plant. The plant takes feed from a clarifier in an adjacent municipal waste treatment plant and uses cartridge filtration as the pre-treatment for a two-stage EDR process. This pre-treatment is quite crude in comparison with the coagulation/MMF used in the present study and as a consequence necessitated quite frequent CIP, more than monthly. Its output is pumped to a dam from which it is used for irrigation.

The Western Treatment Plant would be an ideal site for the implementation of an EDR salt-reduction system. An installation would include pre-treatment and EDR facilities with appropriate controls. It could be located in the vicinity of the disinfection facilities at the HORS and direct its output down the pipeline to Werribee South. Although MMF was used as the pre-treatment in this investigation, other pre-treatments that could provide the feed quality required for EDR should be considered e.g. dissolved air flotation or micro- or ultra-filtration.

MMF/EDR could be implemented in a number of configurations:

• Chlorinated effluent could be diverted from the disinfection train to MMF then EDR and the EDR product finally blended with additional disinfected feed prior to dispatch. The turbidity of the blended product would fluctuate according to the turbidity of the feed.

• Turbidity removal, using one of the aforementioned techniques, could be performed prior to disinfection, producing a stream suitable for EDR and blending. This may reduce the chlorine demand in the disinfection process and a final blended product of more uniform quality with respect to turbidity.

• If microfiltration or ultrafiltration were used, it is expected that the microbiological quality of the product would be improved and there would be a reduced requirement for CIP due to the improved quality of the EDR feed.

This project has demonstrated the robustness of EDR to major fluctuations in the turbidity of the feed. Even when the feed turbidity was 15 NTU and the MMF was unable to reduce turbidity sufficiently, the EDR system kept producing in the presence of increasing fouling, albeit with reduced efficiency and increased conductivity.

This study has demonstrated that electrodialysis reversal is a serious candidate technology for the removal of salts from the effluent at the Western Treatment plant or indeed any waste treatment plant or other source of brackish water for which salt reduction is necessary to enable re-use.


ACKNOWLEDGEMENTS

A large number of people have contributed to the success of this project. The authors wish to acknowledge the assistance of staff in the mechanical and electronics workshops of CSIRO at Highett and Clayton in the preparation of the two CSIRO containers used to accommodate the small-scale component of the trial. We also thank CSIRO materials characterisation staff, especially Peter Curtis and Yeşim Gözükara for analysis of a large number of samples throughout the trial. Peter also assisted in the planning stages of the project.

Personnel from GE Water and Infrastructure have been extremely cooperative throughout, especially Nan-Marie Schoerie who made it all happen, and Allan Clements who was very generous with his time and allowed us to draw on his extensive experience in water treatment as well as making many out-of-hours visits as required.

We have appreciated the encouragement and assistance of staff of Melbourne Water and their contractor Connell Wagner. On the ground Bob Clifton was always willing to go beyond the call of duty as the need arose.

Personnel from the other salt-removal trial operators were also helpful, especially Marrack Payne of United Utilities who with the assistance of Imants Didrichsons provided project management for the installation and ultimate removal of equipment at the site, along with various informal scientific/technical discussions.

Finally, thanks go to Simon Lees of the Smart Water Fund and David Gregory of Melbourne Water for their thoughtful oversight of this project.


APPENDICES

Appendix 1: P&ID for GE-Betz MMF and EDR equipment

GE-Betz MMF P&ID


GE-Betz MMF P&IDMMF Layout


GE-Betz EDR P&ID


GE-Betz EDR container layout


GE-Betz EDR Flow diagram


GE-Betz EDR Flow diagram (continued)


GE-Betz EDR Equipment layout


Appendix 2: Pilot Plant log for the first trial period 2 May and 31 July 2006

Date; Time

Response Time

Duration of Down-time

Action/s Reason for Down-time

1/5/06 8.15am

8 hrs CIP Shut down for CIP

3/5/06 8.24am

5 hrs 6hrs Power Power outage

5/5/06 8.26pm

58hrs 62 hrs Adjust DP set-point

Pressure gauge faulty giving high DP across MMF 1.

9/5/06 12.15pm

1hr 2hrs Power Power outage

13/5/06 10.50am

1hr 3.5 hrs

19/5/06 11.08am

1hr 2hrs

22/5/06 10.53am

1hr 3.5hrs Reset MMF / CIP

MMF 2 ran for 18hrs as didn’t change over to F1. Reset program. CIP carried out

30/5/06 2.45am

9hrs 13hrs CIP High inlet carrier pressure to EDR CIP carried out

6/6/06 1.20pm

1 hr 1 hr Mini CIP Inlet pressure high on EDR. Hypochlorite not dosing. Increase acid on brine pH for a short time.

8/6/06 10.00am

1 hr 5hrs Dismantle stack

Looking to take samples of deposit in stack. No CIP. Just gave membranes a wash with a hose.

9/6/06 11.00am

1 hr 1.5 hrs Brine CIP Brine CIP carried out to try and increase current. No significant benefit. Changed Backwash frequency to 9 hours on MMF.

14/6/06 12.15pm

1 hr 0 No action MMF1 high diff pressure alarm

19/6/06 4.08am

8hrs 2 hrs Acid CIP High inlet pressure on EDR and drop off on product conductivity

21/6/06 12.11am

10hr 11hrs Cleans Amiad filter

Filter 2 stopped on low flow.

23/6/06 9.42am

1hr 2.5hrs Acid CIP CIP carried out as membranes still fouling. Started dosing Hypersperse anti-scalant. Increased voltage as feed water temperature down at 10’C.

27/6/06 2.45am

1hrs 0 Primed coagulant pump

High turbidity as pump lost prime.

19/7/06 10.20am

1 hr 5 hrs Restarted Trip on low-low flow Was in the process of converting over from F2 to F1 .Manually backwash Filter 2

21/7/06 9.50am

1 hr 3hrs Restarted Power outage. Restarted both MMF and EDR.

23/7/06 1.40am

11 hrs 11 hrs Restarted Trip on low-low flow. Automatic flush on Amiad filter lost power. Caused blockage to filter and low flow too MMF pumps.

23/7/06 2.25pm

18 hrs 18hrs Restarted Trip on low-low flow. Same problem as above not picked up as on the weekend.

25/7/06 4.53pm

1hrs 3 hrs Acid CIP High inlet pressure on EDR and drop off on product conductivity

26/7/06 3.40am

6hrs 6hrs Clean Amiad filter

Filter 2 stopped on low flow.

28/7/06 12.50pm

1hr 2hrs Trip on High- turbidity

Reset plant and restarted

28/7/06 5.50pm

22hrs 22hrs Power trip Not picked up until weekend.


Appendix 3: Pilot Plant log for the second trial period: August 18 2006 – Feb 04 2007

Date; Time

Response Time

Duration of Down-time

Action/s Reason for Down-time

2/11/06 9.40am

1 hr 1 hr Restarted Trip on high-high turbidity on MMF Was in the process of converting over from F1 to F2 .

7/11/06 4.20am

6 hrs 13 days Kept system down

High-high Turbidity. Did not retart system as feed turbidity above 6 NTU

20/11/06 3.00pm

Restarted Plant restarted .

21/11/06 2.40am

6 hrs 6hrs Restarted Trip on low differential pressure on EDR between Filter change over.

21/11/06 9.30pm

36hrs 36 hrs Restarted Trip on low differential pressure on EDR between Filter change over.

24/11/06 2.30pm

24hrs 24hrs Restarted Trip on low differential pressure on EDR between Filter change over.

26/11/06 10.30pm

4 days 4 days Restarted Trip on low differential pressure on EDR between Filter change over.

30/11/06 1.00pm

7 days Restarted 6/12 Only ran system for 3 hours as Turbidity of feed was high above 6.0 NTU.

7/12/06 6.30am


11/12/06 10.30am

21 hrs 21 hrs Restarted Trip on low differential pressure on EDR between Filter change over.

18/12/06 3.30pm

17hrs 17 hrs Restarted Trip on low differential pressure on EDR between Filter change over.

21/12/06 Shutdown Shutdown for Christmas. 2/1/07 2.14pm

3 hrs 3 hrs Restarted Trip on low differential pressure on EDR between Filter change over.

4/1/07 12.22am


9/1/07 8.46am


15/1/07 10.40am

1.5 days 1.5days Restarted Trip on low differential pressure on EDR between Filter change over.

18/1/07 2.15pm

3hrs 3hrs Restarted Trip on low differential pressure on EDR between Filter change over.

19/1/07 7.25am

1 hr 1 hr Restarted Altered PLC on MMF as duration of settling step was far too long. Originally 10 min. and then changed to 2hours moved back to 10min.

19/1/07 9.30am

Restarted Initiated a backwash on F1 as this would reset program.

4/2/07 11.00am



Appendix 4: Quick View of data for MMF filter



Appendix 5: Quick View of data for EDR

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