andmwdh2o.com/faf pdfs/2_gw_emwd mwd final report pilot scale brine...removed from the desalting...

Final Report including

AquaSel Technology Pilot Scale Demonstration Menifee De-salter Final Test Report

and Economic Analysis: Preliminary Construction and O&M Cost for Full-Scale AquaSel Facility

at the Perris and Menifee Desalters

September, 2016

1

Contents

Executive Summary ....................................................................................................................................... 2

Introduction .................................................................................................................................................. 2

Cost Summary ............................................................................................................................................... 3

Schedule Summary ....................................................................................................................................... 4

Project Results and Analysis ......................................................................................................................... 5

Conclusion ..................................................................................................................................................... 6

Appendices

Appendix A: AquaSel Technology Pilot Scale Demonstration Menifee De-salter Final Test Report

Appendix B: Economic Analysis: Preliminary Construction and O&M Cost for Full-Scale AquaSel Facility at the Perris and Menifee Desalters

2

Pilot Scale Brine Concentrator Final Report

Executive Summary

This Final Report discusses the pilot scale project that took place at the Menifee Desalter facility between November 2014 and July 2015. The technology employed was electrodialysis reversal (EDR) and precipitator, with associated sub-systems designed to further concentrate solids removed from the desalting operation reject stream for solid waste disposal, reducing reliance on the Inland Empire Brine Line. Appendix A, “AquaSel Technology Pilot Scale Demonstration Menifee De-salter Final Test Report” (also referred to as the “Main Report”) describes operational issues that occurred during the pilot test, and the steps taken to resolve these issues. In addition, detailed descriptions of operational parameters such a flow rates, current levels, and comparisons of actual results to computer models are presented. Appendix B, “Economic Analysis: Preliminary Construction and O&M Cost For Full-Scale AquaSel Facility at the Perris and Menifee Desalters” provides preliminary cost analysis of an AquaSel system capable of processing 2.4 million gallons per day (mgd) of brine, sufficient to handle all of the reject flow from both the Perris and Menifee desalters.

Introduction

Overview

The pilot scale brine concentrator was delivered to the site on November 10, 2014, and first made operational on November 27, 2014. It was briefly run December 1 – 3, then shut down until January 1, 2015. The period starting in January was characterized by silica scaling problems, and subsequently became known as Phase 1. This period lasted into March 2015. The first approach to addressing the scaling problem was pre-treatment with ferric chloride and media filtration, but this proved ineffective. Sulfuric acid dosing was also tried, but discontinued when it was discovered that the acid deactivated the antiscalant in the brine water. Eliminating pre-treatment reduced scaling to tolerable levels. This led to Phase 2, in which hydraulic balance was addressed. The cyclic flow of the system was affecting the performance of the precipitator, which was solved by revisions to the piping and addition of a buffer tank. Phase 3 started with the installation of new EDR membranes on April 28. With the brine concentrator functioning at an acceptable level, Desalter recovery was set at 75%, the brine concentrator at 80%, resulting in overall recovery of 95%. 72% of TDS were removed by the system.

3

Findings

Following the revisions made to the system in Phases 1 and 2, Phase 3 demonstrated that the brine concentrator was capable of solids reduction within the parameters set out at the beginning of the project. The study also demonstrated power consumption within expectations. Ultimately, the most important finding of the study is that electrodialysis reversal is potentially applicable to the large-scale treatment of desalter reject water, and a follow-up program involving a 100 gallon per minute system, employing many of the same components as a full-scale system, is warranted by the results.

Roles of supporting entities

Eastern Municipal Water District provided the facility (Menifee Desalter), and arranged for infrastructure to support the study. This included providing electrical power and water (potable and desalter reject water), and dedicating staff to support the study from development of infrastructure through removal of the brine concentrator at the end of the study. EMWD laboratory staff performed analysis of water samples throughout the study. Falcon Engineering handled site construction, including power, water, slab, and placement of the 40 foot container housing the brine concentrator. Carollo Engineers provided technical consulting to review and assess the setup, operation, testing and evaluation of the brine concentrator. Environmental and Chemical Consulting provided the container for solids disposal, and hauled the filled container at the conclusion of the study. GE provided the brine concentrator, and GE staff had primary responsibility for the equipment from installation through removal. In the early period of the study (Phases 1 and 2, November 28 through May 7) GE staff was frequently onsite, resolving operational issues and making improvements to the system. In Phase 3, GE staff was onsite less as the brine concentrator performed more reliably. GE also had primary responsibility for analysis of data generated throughout the study; this analysis is provided in the attached Main Report. The United States Bureau of Reclamation provided grant funding in support of the study, and also provided technical review of the results of the study. Western Municipal Water District, Inland Empire Utilities Agency, and Las Virgenes Municipal Water District were co-applicants and supporters of the project.

Cost Summary

Summary of costs and funds disbursed

Eastern Municipal Water District has recorded $500,785 in costs associated with this project. Included in that figure is approximately $58,223 in staff labor, which is not reimbursable under the terms of the funding agreement with MWD. EMWD has requested $168,121.68 in reimbursement from MWD under the funding agreement. Additionally, EMWD has received $91,766.36 in reimbursement from the Bureau of Reclamation.

4

Planned vs. actual budget

Total project cost in the budget (Exhibit A of the funding agreement) is $422,508. Actual costs of $500,785 exceed that figure by $78,277. Significant variations include extending the rental period of the brine concentrator for six weeks (+$22,000), sales tax on the rental of the brine concentrator (+$13,920), Technical Consultant costs (+$74,378) peer review by USBR (+$77,160), in-house laboratory analysis vs. Babcock Labs (-$89,632), and reduced installation costs (-$29,370). Other cost categories varied by smaller amounts. Please note, USBR review costs are subject to final audit, and possible revision. While individual budgets demonstrated some variation, total costs were in line with expectations, and commensurate with the increased duration and difficulty of the project.

Schedule Summary

Summary of all tasks accomplished

At this time, Tasks 1 – 7 are complete. With the exception of the long-term Post-Project Update, all reports are complete with the delivery of this document.

Planned schedule vs. actual

The following table compares the original schedule to the actual dates experienced during the study. Three factors added considerably to the completion of several tasks. First, delivery of the brine concentrator was delayed nearly 4 ½ months from the originally scheduled delivery on July 1, 2014, arriving on November 10, 2014. The second factor affecting the schedule related to early performance issues with scaling. The Phase 1 period of testing lasted through February 2015. As a result, the pilot evaluation (Task 4), originally scheduled to begin in October 2014, could not begin in earnest until March 2015. The third factor affecting the schedule was the decision to extend the rental period six weeks. This extended the completion of Task 4 from March 31, 2015 to August 5, 2015, but added valuable data and increased confidence in the ability of the brine concentrator to perform reliably.

Task Scheduled Start Actual Start

Scheduled Finish Actual Finish

Task 1 1/1/2014 11/1/2013 3/31/2014 6/30/2014 Task 2 4/1/2014 4/4/2014 6/30/2014 12/31/2014 Task 3 7/1/2014 4/1/2014 9/30/2014 3/31/2015 Task 4 10/1/2014 1/1/2015 3/31/2015 8/14/2015 Task 5 10/1/2014 1/1/2015 3/31/2015 9/30/2015 Task 6 1/1/2015 8/3/2015 3/31/2015 8/14/2015 Task 7 1/1/2015 4/1/2015 6/30/2015 12/31/2015 Task 8 1/1/2014 11/1/2013 4/15/2017 4/15/2017

5

Project Results and Analysis

Analysis of project results

Objectives of the pilot study are listed on Page 15 of the Main Report; a detailed discussion of results can be found starting on Page 84 of the same report. Generally, the brine concentrator demonstrated that it is capable of removing large amounts of dissolved solids from reverse osmosis reject water, lowering TDS levels from above 6,000 mg/L to under 2,000 mg/L at 80% recovery (total recovery of 95%). It was determined that pretreatment of the brine was not necessary. Performance of the brine concentrator was found to correlate well with GE’s computer model, and process variables such as chemical dosing and clean in place requirements were established.

Economic Analysis

Carollo Engineers produced an economic analysis for a full-scale system (Appendix B), based on a facility designed to treat 2.4 mgd of concentrated brine at up to 80% recovery. The analysis included both construction and operation and maintenance costs, with two alternative configurations considered. In both cases, all reject flow from the desalters would be routed to the AquaSel unit for treatment; brine would be sent to the Inland Empire Brine Line, and solid waste would be sent to a landfill. For Alternative 1, AquaSel product water is blended with reverse osmosis permeate upstream from the chlorine contact basin. The AquaSel product water has similar TDS levels to the raw well water, and the volume is more than sufficient to offset the raw water bypass. As a result, this configuration increases overall desalter output by about 618 AFY. For Alternative 2, AquaSel product water is blended with raw well water before entering the reverse osmosis process, resulting in 1,936 AFY of additional water.

Alternative 1 offers lower capital costs, but does not produce enough new water to offset capital and O&M costs. Alternative 2 is more expensive, partly due to a required expansion of RO capacity, but comes closer to breakeven. However, at an assumed interest rate of 5% over a 30 year loan, Alternative 2 is estimated to generate negative cash flow of $743,000 per year. Partial grant funding, combined with a lower interest rate loan, could make the project cost-neutral. Alternative 1 generates negative cash flow even with a 0% interest rate loan.

Were project goals achieved?

Performance of the brine concentrator met performance goals, after early operational issues were addressed. The study illustrated areas where improvements could be made, which will be incorporated into future systems.

Major problems

The initial problem identified during the study was silica fouling of both the cartridge filters and the electrodialysis reversal membranes (described in the Main Report starting on page 27).

6

This was addressed with pretreatment strategies, but they proved ineffective. By eliminating acid dosing, performance of existing antiscalants was restored, and fouling issues were reduced to acceptable levels. The second area of improvement involved refinements to the precipitator to improve stability of the system (see page 33 of the Main Report).

Application to other areas of the region

Brine management is a significant issue in the region for two reasons: first, reject water from the reverse osmosis process represents the loss of a useful resource, which reduces the value of developing independent local water sources. Second, disposal of brine from the Inland Empire relies on the Inland Empire Brine Line, with attendant costs and potential capacity limitations. Any project using reverse osmosis technology to reduce total dissolved solids from local groundwater sources has the potential to benefit from brine concentrator technology with respect to recovering the reject flows, thus reducing reliance on imported water. Coastal agencies have less to gain from a standpoint of brine management, assuming an approved ocean disposal regime is in place. However, inland agencies stand to benefit in both areas, with possible benefits to local aquifers from reduced salt levels.

Conclusion

Lessons learned

The brine concentrator demonstrated the ability to reduce TDS while recovering 80% of the reject water from the desalting operation. However, the demonstration unit was flowing less than 10 gpm, compared to a full scale requirement of 1,600 gpm. More information is needed to determine the cost and potential operational issues with a system roughly 200 times the size of the pilot scale unit.

Next steps

An interim step between the pilot scale unit and a system capable of treating the Menifee and Perris desalter operating at capacity would offer the ability to evaluate a second-generation system at a reasonable cost, while limiting the risk associated with such a new technology. Accordingly, a 100 gpm system, with enhancements to the precipitator and refinements to the filtration systems, is under consideration by Eastern Municipal Water District as a demonstration project.

APPENDIX A

AquaSel Technology Pilot Scale Demonstration Menifee De-salter Final Test Report

April, 2016

GE Water & Process Technologies

AquaSel Technology

Pilot Scale Demonstration Menifee De-salter Final Test Report

Submitted to:

Eastern Municipal Water District 2270 Trumble Rd Perris, CA 92570

Attention: Khos Ghaderi

Submitted by:

GE Water & Process Technologies 3239 Dundas St. W

Oakville, ON, L6M 4B2

April, 2016

Revision 1.3

AquaSel Technology Pilot Demonstration EMWD, CA

2

All patents, trade secrets and other intellectual property in this document shall be the property of GE Water & Process Technologies (GEW&PT). All intellectual property shall be kept confidential in accordance with the Pilot Rental Agreement associated with Pilot Proposal Number “PIL-830667 Rev1”.


3

Acknowledgements GE would like to acknowledge all the efforts of the EMWD staff to help make this

project successful.

John Dotinga was a great coordinator, making sure that the on-site staff and GE

communication was consistent.

Thank you to Phil Lancaster and his staff (notably Jim, Tim, and Matt) who spent

numerous hours on-site helping to troubleshoot issues during operations.

The Carollo team was instrumental in reviewing data and developing action plans

to keep the project moving forward.

Thank you to Khos Ghaderi for organizing the whole project and keeping an

optimistic attitude throughout.


4

Executive Summary

From November 28th, 2014 through July 31st, 2015, a pilot project took place at

the Menifee and Perris de-salter facility at Eastern Municipal Water District

(EMWD). The primary piloting objective was to evaluate the AquaSel system as a

brine reduction treatment option and to increase the de-salter recovery from 75%

to 95%. The pilot system took the reverse osmosis (RO) brine water from the

Menifee de-salter and treated it with an AquaSel system to reduce the salinity.

The AquaSel product water goal was to produce water of similar quality to the

well water feeding the ROs. The AquaSel system product water, with an average

TDS of 1825 mg/L in the final pilot phase, was comparable to the raw well water

that is used for blending with treated water from the de-salters. The pilot results

demonstrated the AquaSel product water quality that can be expected in a full

scale or demonstration scale system, and could be used as a substitute for the

blend water currently used to mix with the RO permeate. The secondary goal was

to demonstrate the ability of the AquaSel system to remove salt, calcium sulfate in

particular, from the RO brine stream and to concentrate it to supersaturated levels

so that it could precipitate as a solid in the precipitation step of the pilot. During

the pilot study solids precipitation did occur as shown by increasing solids

concentration in the precipitator, creating excess solids that were sent to waste.

Electrodialysis Reversal (EDR), a key component of the Aquasel process, operates

by applying a direct current with positive and negative electrodes to a stream of

water flowing between ion exchange membranes to draw ions out of the less

concentrated stream into a more concentrated stream, thus removing impurities

from the feed stream. In EDR, the polarity of the charge is reversed periodically to

prevent ion accumulation on ion exchange membranes. Ion exchange

membranes are arranged in a stack formation, referred to as a membrane stack,

alternating between anion and cation exchange membranes. During initial pilot

operation, scaling of the EDR membrane stack was immediately evident, with inlet

pressures to the stack reaching inoperable levels. Investigation of the stack

indicated that it was silica scale. During the months of December 2014 and

January 2015 lab scale testing was performed to attempt to remove silica from

the RO brine water using ferric chloride. In February 2015, a pretreatment system

involving ferric chloride dosing and media filtration was implemented and tested,

but did not prove to be effective. After a re-evaluation of the feed to the pilot, the

pretreatment system was removed. Sulfuric acid dosing into the RO brine water

feeding the pilot was also stopped. It was learned that the sulfuric acid dosing in

the feed to the EDR stack was deactivating the antiscalant present in the RO brine

water (used in the RO desalter to control silica), allowing silica to come out of

solution and scale the membranes. This proved to be a turning point for pilot


5

operations as the EDR was then able to operate without silica scale accumulation

on the membranes. This key learning of operation without any pretreatment will

be applied to future design and presents a significant savings to a larger scale

system.

The hydraulic balance of the pilot system was a key area of focus during the

months of March through May 2015. The AquaSel process includes flows that are

not constant, such as the solids blowdown flow from the precipitator and the off-

spec product flow from the EDR, where off-spec is product water that does not

meet purity specifications. These inconsistent flows caused issues in the

precipitator (such as solids carry-over) and level issues in other process tanks. By

re-routing some of the pilot piping and including the use of a buffer tank it was

possible to minimize pilot upsets and smooth flows. The piping modifications led

to the final pilot schematic, found in Figure 10 of this report, which will be the

basis of the full-scale design.

After working through the aforementioned challenges, the EDR anion membranes

were replaced on April 28th to allow for enhanced desalting performance from the

EDR stack. Once pilot operation was shown to be stable and reliable, a final phase

of operation occurred where the Menifee RO recovery was set to 75% and the

AquaSel recovery was set to 80%, resulting in an overall recovery of 95%. During

this final phase, the pilot demonstrated the ability to remove 72% of the TDS

present in the RO brine water. A 95% percent recovery would be recommended

for future applications, with the ability to increase if possible.

The precipitation step in the AquaSel process consisted of a clarifier with

concentric rings, where the center rings were continuously mixed. The

concentrate stream from the EDR stack was fed directly to the center rings of the

precipitator. The center rings of the precipitator were called the reaction zone and

consisted of a slurry of calcium sulfate solids that served as precipitation sites for

the supersaturated concentrate solution. It was determined during the pilot study

that a larger precipitator would be preferred for future designs to minimize the

up-flow velocity in the outer clarification ring to <0.4 gpm/ft2 to prevent solids

carry-over. It was also found that using a high velocity centrifugal pump for slurry

recirculation resulted in smaller particle size and increased solids carry-over.

Future designs should use Variable Frequency Drive (VFD) controlled pumps that

can run continuously at a rate which maintains large enough particle sizes that

easily settle in the clarification zone.

The solids production rate in the precipitator during the final phase of operations

was 0.79 kg/hr of calcium sulfate. The solids produced were sent as a waste


6

stream slurry to a collection bin where they were separated from the water by

means of settling. In a 100 gpm demonstration scale plant, the solids production

would be scaled to 12.3 kg/hr.

Along with the solids slurry waste stream, the precipitator also had a liquid waste

stream from the clarification zone. The liquid concentrate waste stream would be

sent to the Inland Empire Brine Line in a larger scale application. The liquid waste

was supersaturated in calcium sulfate, but also contained residual antiscalant.

The residual antiscalant may have prevented the solution from decreasing in

concentration to the saturation point. Blending of the liquid waste with other

wastes in the brine line would decrease the likelihood of scale in the brine line. A

small dose of antiscalant to the liquid waste would further reduce its scaling

potential in the brine line.

After the precipitation step, the clarified concentrate water passed through a

filtration step to remove any solids that remained in the stream. Initially, a back-

washable membrane candle filter system was used to remove solids. This system

is called a one-pass and consists of candle-shaped filter elements arranged

vertically inside a pressure vessel. However, it was found that backwashing the

filter was not effective and solids were able to accumulate in the filter housings to

the point of inoperability. As an alternative, a set of three parallel 5-micron

cartridge filters was used for the filtration step. The cartridge filters had to be

changed every 3-5 days as solids accumulated in their housings and the pressure

drop increased. While improvements to the precipitation step should help to

reduce the amount of solids carried back through the concentrate loop, a

different filtration system is recommended for use in future designs, such as a

continuously flowing cyclonic separation system. Cartridge filters should still be

used at the concentrate inlet and raw water inlet to the EDR stack as a final

protection step to prevent solids accumulation on the membranes.

The chemicals used for operation of this pilot, which will also be required in future

plants, were sulfuric acid and hydrochloric acid (HCl). Sulfuric acid was dosed

directly into the precipitator to control the concentrate loop pH to 6.3. The

decreased pH in the concentrate loop is to prevent calcium carbonate scale from

occurring. The dose rate of sulfuric acid for the pilot was, on average, 11.9 L/day

(225 mg/L) of 50% sulfuric acid, which would increase based on the precipitator

size. HCl was dosed directly to the electrode streams of the EDR stack. After

several scaling events of the electrode stream, it was determined that a constant

dose of HCl was required to maintain the electrode stream pH at <7. The constant

HCl dose was augmented by an increased HCl dose three times per day, dropping

the electrode pH to <2, to clean any scale that may have accumulated. Based on


7

the pilot consumption, the dose rate of HCl that would be required for a future

design would be 12.9 L/day (956 mg/L) per EDR line. HCl was also used in the

AquaSel process for cleaning the EDR membranes and piping. It was established

that cleaning with a 5% HCl solution every four weeks would be sufficient.

Additionally, it is recommended that future designs incorporate a product flush

step to clear membranes and piping of concentrate water in case of shutdown to

prevent scale build-up.

The DC power consumption of the EDR stack during the final pilot phase was 6.2

kwh/kgal of water treated and the estimated AC power consumption of the

auxiliary equipment was 3 kwh/kgal, for a total power consumption of 9.2

kwh/kgal. This power consumption is lower than the projected power

consumption as calculated by GE’s EDR design software, WATSYS. The WATSYS

software projects 11.9 kwh/kgal for a 100 gpm EDR system and it is expected to

be lower in actual operation.

Based on the data collected during the pilot study, GE recommends proceeding

with a 100 gpm demonstration scale system, where design changes can be

implemented and the technology can be proven at a larger scale. The 100 gpm

system would use four EDR stacks that consist of 600 cell pairs, and would be the

same stack used in full-scale construction.


8

Table of Contents

Glossary of Terms ................................................................................................................................. 12

Introduction ............................................................................................................................................. 14

Pilot Objectives ....................................................................................................................................... 15

Electrodialysis Reversal Technology ............................................................................................. 16

Pilot Study Process Design Parameters ...................................................................................... 21

Water Quality ...................................................................................................................................... 21

Pilot Equipment Description ............................................................................................................. 23

Operations ................................................................................................................................................ 27

Problem Solving Step 1: Pretreatment Review – November 28, 2014 – March 20, 2015 ........................................................................................................................................................ 27

Problem Solving Step 2: Process Stabilization – January 28th, 2015 – April 13th, 2015 ........................................................................................................................................................ 33

Problem Solving Step 3: New Membrane Installation – April 28th, 2015 ................... 36

System Design: Recirculation Loop ........................................................................................... 39

Precipitator Performance .............................................................................................................. 43

Chemical Consumption .................................................................................................................. 46

Discussion ................................................................................................................................................. 48

Pilot Operation Phases .................................................................................................................... 48

Operational Results – Phase 1 – November 28, 2014 to December 3, 2014 .......... 48

Process Pressures, Stream Conductivities, Voltage and Current, DC Power Consumption – Phase 1 .................................................................................................................. 49

Operational Results – Phase 2 – January 28, 2015 to May 7, 2015 ............................ 58


Operational Results – Phase 3 – May 7, 2015 to July 31, 2015 .................................... 68

Process Flows and Pressures – Phase 3 ................................................................................. 70

DC Voltage and Current – Membrane Stack – Phase 3 .................................................... 70

Conductivity and Constituent Removal – Phase 3 ............................................................. 72


EDR Concentrate Loop and Precipitator Performance .................................................... 82

Conclusions .............................................................................................................................................. 84

Future Design Recommendations ................................................................................................. 89

APPENDIX A.1: EMWD Phase 3 EDR Conductivity Charts .................................................... 90

APPENDIX A.2: EMWD Phase 3 EDR Flow Charts ..................................................................... 95

APPENDIX A.3: EMWD Phase 3 EDR Voltage and Current Charts ................................. 100

List of Tables Table 1: Raw Water Quality .............................................................................................................. 21

Table 2: Product Quality Targets .................................................................................................... 22

Table 3: Recovery Phases .................................................................................................................. 48

Table 4: WATSYS Design - Phase 1 ................................................................................................ 49

Table 5: Operational Data - Phase 1 ............................................................................................. 50

Table 6: Mass Balance For Phase 1 Operations ...................................................................... 56

Table 7: WATSYS Design - Phase 2 ................................................................................................ 59


9

Table 8: Analytical Results, April 17, 2015 .................................................................................. 60

Table 9: AquaSel Process Mass Balance for Phase 2 Operations ................................... 66

Table 10: WATSYS Design For Phase 3 Operations ................................................................ 69

Table 11: Average Analytical Results From Phase 3 .............................................................. 73

Table 12: AquaSel Mass Balance for Phase 3 Operations .................................................. 80

Table 13: Precipitator Summary for Phase 3 Operations .................................................... 83

Table 14: WATSYS Design For 100 GPM EDR System ............................................................ 86

Table 15: Sulfuric Acid Dose Rate During Pilot Study ............................................................ 86

Table 16: Power Consumption of Pumps and Motors in Pilot ........................................... 87

Table 17: Hydrochloric Acid Consumption for Pilot ............................................................... 87

List of Figures Figure 1: Electrodialysis Process..................................................................................................... 16 Figure 2: EDR Membrane Cell Pair and Full EDR Membrane Stack ................................. 17 Figure 3: EDR Process Flow Streams ............................................................................................ 18 Figure 4: EDR Polarity Reversal Effect on Colloidal matter ................................................. 18 Figure 5: Simplified EDR Process Flow Diagram ...................................................................... 19 Figure 6: Simplified AquaSel Process Flow Diagram ............................................................. 20 Figure 7: Containerized EDR Pilot Unit ......................................................................................... 23 Figure 8: Pilot Precipitator ................................................................................................................. 23 Figure 9: Pilot EDR Stack .................................................................................................................... 24 Figure 10: AquaSel Process Flow Diagram ............................................................................ 26 Figure 11: EDR Pressure Trend in Phase 1 ................................................................................. 28 Figure 12: Silica Coated Filter and Membrane ......................................................................... 29 Figure 13: Pretreatment System Process Flow Diagram ..................................................... 30 Figure 14: EDR Pressures Between March 1st and April 26th ........................................... 32 Figure 15: Calcium Sulfate Deposition in Precipitator and Candle Filter ...................... 34 Figure 16: EDR Pressures Between March 1st and May 8th .............................................. 35 Figure 17: Product Conductivity Between March 1st and July 31st ............................... 37 Figure 18: Calcium Sulfate Buildup in Precipitator Piping ................................................... 39 Figure 19: Stokes Law Settling Velocity Versus Particle Size ............................................. 40 Figure 20: One-Pass Filter and Cartridge Filter Process Flow ........................................... 41 Figure 21: Calcium Sulfate Build-up in Candle-Filters .......................................................... 42 Figure 22: Calcium Sulfate Build-up on Cartridge Filters .................................................... 43 Figure 23: Precipitator Solids Settling Test ................................................................................. 45 Figure 24: Precipitator Solids Percentage by Volume ........................................................... 46 Figure 25: EDR Pressures During Phase 1 Operations .......................................................... 51 Figure 26: EDR Conductivities During Phase 1 Operations................................................. 52 Figure 27: EDR Voltage and Amperage During Phase 1 Operations.............................. 53 Figure 28: Product Conductivity During Phase 1 Operations ............................................ 54 Figure 29: EDR Stack Power Consumption During Phase 1 Operations ....................... 55 Figure 30: Simplified Process Flow Diagram For Phase 1 Operations ........................... 57 Figure 31: EDR Pressures During Phase 2 Operations .......................................................... 61 Figure 32: EDR Conductivities During Phase 2 Operations................................................. 62 Figure 33: EDR Voltage and Amperage During Phase 2 Operations.............................. 63 Figure 34: EDR Product Conductivity During Phase 2 Operations .................................. 64 Figure 35: EDR Stack Power Consumption During Phase 2 Operations ....................... 65


10

Figure 36: Simplified Process Flow Diagram For Phase 2 Operations ........................... 67 Figure 37: AquaSel Flows June 19, 2015 .................................................................................... 70 Figure 38: EDR Voltage and Amperage, June 19, 2015 ....................................................... 71 Figure 39: AquaSel Conductivity and Calcium, June 19, 2015 .......................................... 72 Figure 40: EDR Pressures During Phase 3 Operations .......................................................... 75 Figure 41: EDR Conductivities During Phase 3 Operations................................................. 76 Figure 42: EDR Voltage and Amperage During Phase 3 Operations.............................. 77 Figure 43: EDR Product Conductivity During Phase 3 Operations .................................. 78 Figure 44 Stack Power Consumption During Phase 3 Operations ................................. 79 Figure 45: AquaSel Simplified Process Flow Diagram for Phase 3 Operations ......... 81 Figure 46: Calcium, Sulfate, % Solids in Precipitator Effluent ............................................ 82 Figure A.47: Conductivity and Ca 6/19/2015 Figure A.48: Conductivity and Ca 6/22/2015 Figure A.49: Conductivity and Ca 6/24/2015 Figure A.50: Conductivity and Ca 6/24/2015 ......................................................................... 91 Figure A.51: Conductivity and Ca 6/29/2015 Figure A.52: Conductivity and Ca 7/2/2015 Figure A.53: Conductivity and Ca 7/6/2015 Figure A.54: Conductivity and Ca 7/8/2015 ........................................................................... 92 Figure A.55: Conductivity and Ca 7/10/2015 Figure A.56: Conductivity and Ca 7/14/2015 Figure A.57: Conductivity and Ca 7/15/2015 Figure A.58: Conductivity and Ca 7/20/2015 ......................................................................... 93 Figure A.59: Conductivity and Ca 7/22/2015 Figure A.60: Conductivity and Ca 7/27/2015 Figure A.61: Conductivity and Ca 7/30/2015 Figure A.62: Conductivity and Ca 7/31/2015 ......................................................................... 94 Figure A.63: Flow Snapshot 6/19/2015 Figure A.64: Flow Snapshot 6/22/2015 Figure A.65: Flow Snapshot 6/24/2015 Figure A.66: Flow Snapshot 6/26/2015 .................................................................................... 96 Figure A.67: Flow Snapshot 6/29/2015 Figure A.68: Flow Snapshot 7/2/2015 Figure A.69: Flow Snapshot 7/6/2015 Figure A.70: Flow Snapshot 7/8/2015 ....................................................................................... 97 Figure A.71: Flow Snapshot 7/10/2015 Figure A.72: Flow Snapshot 7/14/2015 Figure A.73: Flow Snapshot 7/15/2015 Figure A.74: Flow Snapshot 7/20/2015 .................................................................................... 98 Figure A.75: Flow Snapshot 7/22/2015 Figure A.76: Flow Snapshot 7/27/2015 Figure A.77: Flow Snapshot 7/30/2015 Figure A.78: Flow Snapshot 7/31/2015 ................................................................................ 99 Figure A.79: Voltage and Amperage 6/19/2015 Figure A.80: Voltage and Amperage 6/22/2015 Figure A.81: Voltage and Amperage 6/24/2015 Figure A.82: Voltage and Amperage 6/26/2015 ................................................................. 101


11

Figure A.83: Voltage and Amperage 6/29/2015 Figure A.84: Voltage and Amperage 7/2/2015 Figure A.85: Voltage and Amperage 7/6/2015 Figure A.86: Voltage and Amperage 7/8/2015 ................................................................... 102 Figure A.87: Voltage and Amperage 7/10/2015 Figure A.88: Voltage and Amperage 7/14/2015 Figure A.89: Voltage and Amperage 7/15/2015 Figure A.90: Voltage and Amperage 7/20/2015 ................................................................. 103 Figure A.91: Voltage and Amperage 7/22/2015 Figure A.92: Voltage and Amperage 7/27/2015 Figure A.93: Voltage and Amperage 7/30/2015 Figure A.94: Voltage and Amperage 7/31/2015 ........................................................... 104


12

Glossary of Terms

1. Anion: a negatively charged ion resulting from the dissociation of salts,

minerals, or acids in water

2. Anion membrane (anion transfer membrane): a membrane through

which only anions will transfer 3. Anode (positive electrode): the electrode that attracts negatively charged

anions 4. Calcium Sulfate (CaSO4) Saturation: the point beyond which any further

addition of CaSO4 in a given solution will cause precipitation 5. Candle filter: a back washable, batch operated filter with candle-shaped

filter elements arranged vertically inside a pressure vessel, operated as

outside-in filtration 6. Cathode (negative electrode): the electrode that attracts positively

charged cations 7. Cation: a positively charged ion resulting from the dissociation of salts,

minerals or acids in water 8. Cell pair: repetitive section of a membrane stack consisting of a cation

membrane, a demineralized water-flow spacer, an anion membrane and a

concentrate water-flow spacer 9. Concentrate stream: the stream in the membrane stack into which ions

are transferred and concentrated 10. Conductivity: the ability of a solution to conduct electrical current,

commonly expressed in microsiemens/centimeter (micromhos/cm) 11. Cross leakage: refers to the water leakage between demineralized and

concentrate streams in the membrane stack 12. Demineralize (de-salt): to reduce the quantity of minerals or salts in an

aqueous solution 13. Product (demineralized, dilute) stream: the stream in the membrane

stack from which ions are removed 14. Electrodialysis (ED): a process in which ions are transferred through

membranes from a less concentrated to a more concentrated solution as

a result of the passage of a direct electric current 15. Electrodialysis reversal (EDR): an electrodialysis process in which the

polarity of the electrodes is reversed on a prescribed time cycle thus

reversing the direction of ion movement in a membrane stack 16. Electrical staging: the addition of electrode pairs in ED/EDR systems to

optimize the DC electrical system within a membrane stack 17. Electrode: a thin metal plate which carries electric current in and out of a

membrane stack, normally constructed of platinum coated titanium alloys


13

18. Electrode compartment (stream): the water flow compartment containing

the metal electrode where oxidation/reduction reactions occur 19. Fouling: a phenomenon in which organic or other materials are deposited

on the membrane surface causing inefficiencies 20. Heavy cation membrane: a cation membrane made twice normal

thickness (1.0 mm) to withstand greater differential pressures 21. Hydraulic staging: multiple passes of a water between electrodes used in

ED/EDR systems to achieve further demineralization 22. Membrane (ED): an ion exchange resin cast in a sheet form which is

essentially water-tight and electrically conductive 23. Off-spec product (OSP): product water which does not meet purity

specifications 24. One-pass: see candle filter 25. Percent recovery: the percentage of feed water which becomes product

water (the amount of product water produced divided by the total amount

of feed water multiplied by 100) 26. Precipitation: the action or process of precipitating a substance from a

solution 27. Scaling: the formation of a precipitate on a surface in contact with water

as the result of a physical or chemical change 28. Water flow spacer: a die-cut sheet of plastic which forms discreet flow

paths for the demineralized and concentrate streams within an ED

membrane stack 29. Water transfer: phenomenon in which water molecules are transferred

through a membrane along with an ion 30. WATSYS: GE design software used to calculate EDR process parameters.

Inputs include feed water analysis, recovery, and product rate. Outputs

include stack voltages and currents, stream flows and compositions,

scaling potential, pumping and electrical power requirements.


14

Introduction

Eastern Municipal Water District (EMWD) services a wide area in southern

California including the cities of Temecula, Perris, Hemet, Murrieta, Menifee, and

San Jacinto. Currently the EMWD has two operating membrane de-salter plants.

The Menifee and Perris 1 de-salters have a combined capacity of 8,500 acre-feet

per year, or 8.5 mgd. These plants operate at approximately 75% water recovery.

Therefore, the remaining 25% of the water pumped from the ground is a non-

usable product that must be disposed. This brine stream is sent to the Inland

Empire Brine Line (IEBL). EMWD must pay for this disposal, so finding methods of

decreasing the brine flows has become an objective of EMWD to reduce cost. The

water shortage in California is also a motivator to find ways to save water and

reduce waste.

Multiple technologies have been evaluated to help increase the overall recovery of

the de-salter system. Some of these technologies were piloted previously, but

none of these tests have yielded significant results. Therefore, EMWD approached

GE with the hopes of gaining a better understanding of the EDR process. As part

of GE’s product development, a precipitation step has been added to the brine

loop of the EDR process, resulting in the AquaSel technology. This precipitation

step focuses on removing calcium sulfate from the brine stream in a controlled

environment, preventing precipitation in the membrane stack. GE teamed with

EMWD to perform a pilot scale test of the AquaSel technology, and the results of

that test are contained herein.

This report first describes the AquaSel process and the pilot used at EMWD. The

Operations section of this report discusses challenges encountered during the

pilot process, and the solutions to those challenges. The Discussion section of this

report details the operational specifics of pilot operation, such as recovery, power,

and de-salting performance. Finally, the results are summarized in the

Conclusions and Future Design Recommendations sections.


15

Pilot Objectives

The following were the specific objectives of the EMWD pilot study:

1. Establish the performance capacity of the EDR process to remove dissolved salts.

2. Determine if pretreatment is needed to run the EDR process without issue.

3. Demonstrate precipitator’s ability to remove solids from the system.

4. Establish treatment process water recovery.

5. Establish EDR process criteria including velocity and amperage.

6. Determine chemical dosages.

7. Determine electric power requirements.

8. Determine electrode cleaning chemical dosage.

9. Establish Clean In Place requirements


16

Electrodialysis Reversal Technology

Electrodialysis (ED) is an electrochemical separation process that removes ions

and other charged species from water and other fluids. ED uses small quantities

of DC voltage electricity to transport these species through membranes

composed of ion exchange material to create a separate purified and

concentrated stream.

When a DC voltage is applied across a pair of electrodes positive ions, such as

sodium, move towards the negatively charged cathode and these are called

cations. Negative ions, such as chloride, move towards the positively charged

anode, and these are called anions. If membranes are placed between the

electrodes, different flow paths are made.

A membrane permeable to cations only is placed nearest to each of the

electrodes. Cations move through the cation-transfer membranes, while anions

move through the anion-transfer membrane. Flow spacers are placed between

the membranes to support the membranes and create a turbulent flow path.

Water flows tangentially across the membranes, not through them. The ions

travel through the membranes so that one stream is demineralized as product

while the other is concentrated. Since the water does not need to be forced

through the membranes, the cost of electrodialysis treatment is only in removing

the salt.

Cation-Transfer Membrane

Desalinated Product

Anion-Transfer Membrane

Concentrate

Cation-Transfer Membrane

Anode (+)

Cathode (-)

Figure 1: Electrodialysis Process


17

Electrodialysis Reversal (EDR) is a continuous self-cleaning electrodialysis process

by means of periodic reversal of the DC polarity, thereby switching the

concentrating and diluting flow streams. A membrane stack is assembled by

compiling multiple EDR membrane cell pairs between two identical electrodes

which act as both cathode and anode during the reversal cycle. An EDR cell pair

contains an Anion Exchange Membrane, a concentrating spacer, a Cation

Exchange Membrane and a diluting spacer. An EDR stack can contain between

500 and 600 cell pairs.

In the EDR treatment process there are 3 main streams: feed water, concentrate

and product. The feed water is pumped through a feed pump and then through

the membrane stacks to make desalinated product. The concentrate is pumped

through the concentrate pump to make ion saturated concentrate. The flow of

feed water and concentrate through the stack is essentially equal. Most of the

concentrate leaving the stack is recycled through the concentrate pump so that a

high recovery can be achieved. However, in order to prevent the concentrate

from becoming too concentrated, which could result in salts precipitating or

forming a scale on the membranes, a small amount of concentrate is wasted.

This quantity is made up with fresh water from the feed stream. Figure 3 depicts

the flow streams for the EDR process.

Figure 2: EDR Membrane Cell Pair and Full EDR Membrane Stack


18

The electrical polarity, and thus the demineralized and concentrate flow passages

are automatically reversed two to four times every hour. This reverses the

direction of ion movement, which provides “electrical flushing” of scale forming

ions and colloidal matter from the membrane surfaces. This “electrical flushing”

controls scaling and fouling of membranes and can eliminate the need for

extensive pretreatment of the feed water and also reduces the need of chemical

cleaning.

Figure 3: EDR Process Flow Streams

Figure 4: EDR Polarity Reversal Effect on Colloidal matter


19

The alternating exposure of membrane surfaces to the product and concentrate

streams provides a self-cleaning capability that enables desalting of scaling or

fouling waters. At reversal, automatically operated valves switch the two inlet and

outlet streams so that the incoming feed water flows into the new demineralizing

compartments and the recycled concentrate stream flows into the new

concentrating compartments. The effect of this reversal is that the concentrate

stream remaining in the stack whose salinity is higher than the feed water, must

now be desalted. This creates a brief period of time in which the demineralized

stream (product water) salinity is higher than the specified level. This slug of water

is known as off-spec product. Conductivity controlled valves shunt the product to

waste until specifications are met.

The manner in which the membrane stack array is arranged is called staging. The

purpose of staging is to provide sufficient membrane area and retention time to

remove a specified fraction of salt from the demineralized stream. Staging is the

process of adding additional passes through an EDR stack for each increment of

water processed. The ultimate goal of staging is to increase product water purity.

In larger systems, additional stages are created by simply adding more stacks in

series to achieve the desired water purity.

Figure 5: Simplified EDR Process Flow Diagram

As an addition to the standard EDR process, a precipitation step has been added

to the concentrate recycle loop in order to remove calcium sulfate from the EDR

process. Instead of recycling directly back to the concentrate side of the EDR

stack, the concentrate stream is fed into a precipitator. The addition of this

precipitation step is what differentiates the standard EDR process from the

AquaSel, which was piloted at EMWD.


20

The precipitator has two main sections. The interior column of the precipitator is

the active mixing zone. The concentrate water is fed here, and the mixer

continuously recirculates the high salinity water to control precipitation and

particle size. The other section is the outside clarification section. This section

works like a typical clarifier. The water flows slowly and gently upwards, allowing

any particles of calcium sulfate to stay at the bottom of the tank. The clarified

water at the top of the precipitator is then fed back to the suction of the

concentrate pump.

Between the precipitator and the concentrate pump, there is a filtration step to

ensure that gypsum (calcium sulfate) particles are not recycled back into the EDR

system. This filtration steps consists of two options: candle filters or cartridge

filters. The candle filters are designed as a back-washable option that

automatically cleans any particles off the interior surface on a regular timed basis.

The backwash is intended to keep these filters clean and free from gypsum

deposits over time. If this filtration method is not successful, there are cartridge

filters as a backup option. By adjusting the valves on the system, the operator can

choose which set of filters will be used.

Figure 6: Simplified AquaSel Process Flow Diagram


21

Pilot Study Process Design Parameters

Water Quality

This pilot study was conducted at the Menifee De-salter, treating reverse osmosis

brine water. Table 1 below reflects raw water analyses used in process

calculations for each operational phase of the pilot. Phase 1 was based on water

analysis provided before the pilot commenced (RO plant recovery = 75%). In

Phase 2, de-salter RO recovery was reduced to approximately 70% in order to

manage scaling issues that developed during Phase 1. In Phase 3, RO plant

recovery was restored to 75%. The latest water analysis was used as it became

available to model EDR operation using GE’s WATSYS software, which is used to

provide mass balance and operational models for EDR processes.

Table 1: Raw Water Quality

Parameter Units Phase 1* Phase 2 Phase 3

RO Recovery 75% 70% 75%

Calcium mg/l 1150 976 1079

Magnesium mg/l 293 211.7 260.6

Sodium mg/l 1159 627 774

Potassium mg/l 31.4 21.5 25.7

Strontium mg/l 6.1 6.4 6.1

Barium mg/l 1.05 0.73 0.75

Ammonia mg/l < 0.5 < 0.5 < 0.5

Bicarbonate mg/l 936 736 809

Sulfate mg/l 666 422 489

Chloride mg/l 3581 2350 2953

Fluoride mg/l 0.63 < 1 0.61

Nitrate mg/l 151 119 97

Total PO4 mg/l 0.8 4.1 4.1

Silica (total) mg/l 166 148 154

Silica (reactive) mg/l Not Tested Not Tested 129

CO2 (calculated) mg/l 22 22 48

Total Hardness CaCO3 4075 3312 3772

TDS mg/l 7965 5623 6652

Conductivity µS/cm 10600 8760 10500

pH 7.8 7.7 7.4

* Phase 1 Raw Water Quality data was provided by EMWD for the proposal prior to pilot operations; Phases 2 and 3 were measured during operation.


22

A key focus of the pilot was total dissolved solids (TDS) removal. The targets are

shown below in Table 2 for each phase. Quality targets were improved with each

phase as it was demonstrated that fouling tendency of RO brine chemistry toward

EDR membranes and EDR electrode processes could be managed. Since TDS is

difficult to measure in the field, a conductivity target was used to verify that the

proper ion removal was taking place.

Table 2: Product Quality Targets

Parameter Phase 1 Phase 2 Phase 3

TDS (mg/l) < 2900 < 2500 < 2200

Conductivity (µS/cm) 4650 3620 3300

Throughout the course of this study, water samples were taken for: feed water,

product water, precipitator feed water, precipitator effluent water, and

concentrate blowdown. During Phases 1 and 2, these streams were tested on site

with some samples sent to EMWD’s lab for analysis. During Phase 3 samples were

sent to EMWD’s lab for analysis three times per week. Some simple tests and

measurements were performed daily on-site such as pH, turbidity, and

conductivity. Results of these analyses are discussed in the section titled

“Discussion”, and specifically sub-section “Operational Results – Phase 3 – May 7,

2015 to July 31, 2015”.


23

Pilot Equipment Description

GE W&PT supplied the AquaSel pilot unit to demonstrate treated water quality and

to collect operational data for full-scale design. During the study GE W&PT used

one membrane stack with two electrical stages and 50 cell pairs per phase. The

AquaSel pilot was designed to demonstrate similar process flow and operation to

the system that would be used in full-scale operation. That is, the water would be

treated using the same process but with scaled up equipment. Photos of the pilot

system at EMWD can be seen below in Figures 7 – 9.

Figure 7: Containerized EDR Pilot Unit

Figure 8: Pilot Precipitator


24

Figure 9: Pilot EDR Stack

The AquaSel pilot system consisted of a standard EDR pilot modified to

incorporate a precipitation and filtration step. The EDR pilot was automated with a

PLC and SCADA control to maintain process set-points and to collect online flow,

conductivity, pH, and temperature data. The EDR pilot is shown above in Figure 9.

The AquaSel system also included a mixed precipitator as shown above in Figure

8. The system had various tanks and pumps to facilitate flow throughout the

system. These components of the AquaSel system are similar to what would be

found in a demonstration or full-scale system. Please see Future Design

Recommendations for modifications from pilot operation to larger scale

operation.

A complete process flow diagram of the AquaSel system piloted at EMWD, in its

final iteration can be found in Figure 10, and a description of the process is as

follows. RO brine water was fed into a Raw Feed Tank, essentially an equalization

tank, and then was pumped through a 5-micron cartridge filter before entering

the dilute stream of the EDR stack. The dilute stream was de-salted in the EDR

stack and was then sent to discharge as product water. A small stream of RO

brine water was diverted to flow through the electrode streams in the EDR stack

and hydrochloric acid (HCl) was added to that stream. The electrode stream was

sent to waste. Another stream of RO brine water was fed into the concentrate

stream to function as a make-up stream. The concentrate stream was a

continously recirculated stream that went through various parts of the AquaSel

process flow. Starting at the EDR, the concentrate stream was pumped through

the EDR stack and was further concentrated with ions removed from the dilute

stream. From the EDR stack the concentrate stream was fed directly to the


25

precipitator. The precipitator consisted of a clarifier with three concentric rings.

The inner most ring was continously mixed with two impeller blades and

contained a concentrated slurry of calcium sulfate particles. The concentrate

stream entered the precipitator in the inner most ring and flowed out through

piping in the outer most ring after being clarified. Sulfuric acid was added to the

precipitator to maintain a pH of 6.3. The precipitator was intially seeded with

calcium sulfate dihydrate (gypsum) to achieve 10% by volume solids in the centre

ring of the precipitator. The precipitator also had continous recirculation that

pulled slurry from the bottom of the tank and returned it to the centre ring. Solids

were wasted from the precipitator using the recirculation stream by sending the

slurry to a waste bin. Some clarified concentrate water was also overflowed out of

the precipitator and sent to waste. The remaining clarified concentrate water from

the precipitator was then pumped using the Filter Feed Pump through a set of

three parallel 5-micron cartridge filters to remove any remaining solids. After the

filtration step, the concentrate was collected in a tank called the BW Feed Tank

and blended with off-spec product water generated by polarity reversal of the

EDR stack. From that tank, concentrate was pumped back through the EDR stack,

completing the concentrate loop. Excess water in the BW Feed Tank was pumped

to the BW Buffer Tank and returned to the precipitator using the LC Pump.

Figure 10: AquaSel Process Flow Diagram 26


27

Operations

After the startup of the system, the intention was to run the system at steady

state parameters for the duration of the test. The whole test was intended to be

one phase where the operations and performance were measured and evaluated.

As there were some issues along the way, the test was broken out into steps

where specific issues were resolved. The following sections outline the major time

periods for problem solving issues with the pilot as well as discussion around

specific components of the pilot flow sheet. Operational phases, related to plant

recoveries, are addressed in the Discussion section of this report.

Problem Solving Step 1: Pretreatment Review – November 28, 2014 – March 20, 2015

Problem solving step 1 was predominantly focused on the silica situation and

understanding the impact of fouling on the system. During the time period of this

first problem solving step, the de-salters were initially operated at 75% recovery

and the AquaSel system was operated at 90% recovery. After a short period of run

time, it was quickly evident that silica was depositing on both the cartridge filters

and EDR membranes. The concentration of silica in the RO brine water feeding the

AquaSel unit was greater than 140 ppm, which was higher than the design

recommendation of 125 ppm. Fouling was indicated by an increase in EDR stack

inlet pressure. The EDR stack inlet pressure was constantly monitored by pressure

transmitters. From the data below in Figure 11, the pressure increased

significantly after only a few days of operation.


28

Figure 11: EDR Pressure Trend in Phase 1

Shutdown over holidays, replace anion membranes

Acid CIP of EDR stack



Step changes in Cartridge Filter Diff Pressure represent cartridge filter replacement


29

Early in the pilot operation, this pressure increase was noticed. The baseline

pressure for the system was approximately 30 psi. During operation, the pressure

was observed going higher than 45 psi, or a 50% increase in operating pressure.

An acid clean in place (CIP) was performed using a 5% solution of hydrochloric

acid to remove scale and return the stack to original operating pressures.

However, the pressure would soon increase again as seen in Figure 11, above.

Inspection was necessary to understand this problem further. The cartridge filter

assembly was taken apart so the filter could be examined. A coating of gelatinous

material was covering the cartridge filter as shown in Figure 12. It had the look

and texture of silica gel, and it was sent to the lab to be verified. The lab results

confirmed that it was in fact silica. A similar inspection was done on the EDR

membrane stack as well. The membrane surface was also covered in silica. The

pictures below are from the inspection. Note the rough surface of the membrane

representative of silica fouling in the right hand side of Figure 12.

Figure 12: Silica Coated Filter and Membrane

A pre-treatment system was added to the process flow in an attempt to remove

the silica before the feed water was fed to the EDR cartridge filter. The purpose

was to reduce and hopefully eliminate the silica fouling within the EDR system. A

multimedia filtration (MMF) system with ferric chloride addition was selected. The

MMF also included a cartridge filter as a second barrier before the feed tank, as

shown in Figure 13, below.


30

Figure 13: Pretreatment System Process Flow Diagram

After a few weeks of operation with the pre-treatment in-line, there was not a

noticeable improvement to the system. The pressure data shown in Figure 11

was taken from this second part of problem solving step 1 (or operational Phase 2)

with pre-treatment in place. The RO recovery upstream was even decreased from

75% to 70% to decrease the concentration of silica feeding the EDR system, but

the problem still persisted. At this point, a new solution needed to be devised.

Previous to this pilot study, an EDR pilot had been run treating similar RO

concentrate and did not have the same silica fouling issues. The data from the

previous pilot and the EMWD pilot were similar. Throughout step 1, sulfuric acid

was continuously added to the feed water of the AquaSel pilot. The sulfuric acid

dosing is shown in the process flow diagram in Figure 13. However, the other pilot

did not reduce EDR feed pH. Up until this pilot, the AquaSel process called for

reduction in feed pH to 5.5 in order to control calcium carbonate scaling. The RO

system uses a non-GE (King Lee) antiscalant. The GE process team contacted the

manufacturer to better understand the phenomenon.

GE learned that, at the pH of 5.5, the antiscalant became inactive. The antiscalant

was being added to the RO system to ensure that silica stayed in solution. As the

chemical became less effective due to the low pH, the silica was able to drop out

and deposit on the filters and membrane. Since another pilot had successfully run

at a pH range of 7-8, a test was run to see if eliminating acid addition would also

eliminate silica fouling. This test was run and the silica fouling immediately

stopped. It was decided to run without acid addition to the feed for the duration

of the pilot, meaning that the EDR was treating water between pH 7 and 8.

However, substantial removal of alkalinity was still required. This acidification

load was transferred to the concentrating side of the process. The typical AquaSel

process had operated at pH of 5.5 – 6.0 on the concentrating side of the process.

Maintaining this pH without acidification of the feed required additional acid

dosing to the concentrate such that CO2 solubility limits were exceeded. Initially


31

this led to CO2 gas generation in the precipitator and flotation of solids. An

optimal pH was arrived at where CO2 solubility was not exceeded and calcium

carbonate scaling was under control (pH 6.3 – 6.5). Figure 14, below, shows the

pressure stabilization after the acid dosing was removed from the feed and the

media filter pre-treatment system by-passed.


32

Figure 14: EDR Pressures Between March 1st and April 26th

Stack inspection and cleaning

Pretreatment removed, acid addition stopped


33

Problem Solving Step 2: Process Stabilization – January 28th, 2015 – April 13th, 2015

Once the silica fouling problem was solved, the team was able to review other

issues that were occurring in the rest of the process. The second problem solving

step focused on getting the system to operate in a steady state. During this

period the de-salter was operated at 70% recovery, with the AquaSel pilot at 90%

recovery. The work in this step focused on balancing flows within the AquaSel

pilot.

The precipitator is essentially an up-flow clarifier. For this clarifier to work, a

consistent slow flow velocity was crucial. Flow upsets could cause high levels of

solids to float to the top of the precipitator and get recycled back to the EDR

system.

Within the process, there are two flow rates that are intermittent: off-spec

product and candle filter backwash flow. Typically, off-spec product (OSP) occurs

for about 36 seconds every 15 minutes during the polarity reversal. Due to

membrane fouling, off spec times had become extended and recycle of this extra

flow contributed to flow imbalances in the system. Candle filter backwash flow

was being recycled at too high a rate to the precipitator, causing an additional

temporary influx of water to the precipitator.

The effect of these varying recycle flows was periodic overloading of the

precipitator. This caused unintended overflows to waste and a loss of solids from

high carryover rate. Due to the high up-flow in the precipitator, solids losses

overloaded the candle filters. The pictures in Figure 15 show the candle filters and

the accumulation of solids on the upper ring of the precipitator.


34

Figure 15: Calcium Sulfate Deposition in Precipitator and Candle Filter

As part of a GE process visit to the site, the flows were altered to minimize the

impact of these intermittent flows. A tank was repurposed to collect these

streams. Within that tank, there was a pump that would reintroduce these wastes

to the precipitator more slowly and over a longer period of time. Rather than

flowing to the precipitator at 7 gpm for one minute, the change allowed for

approximately 1 gpm over 7-8 minutes.


35

Figure 16: EDR Pressures Between March 1st and May 8th

Pretreatment removed, acid addition to feed stopped

Replace EDR anion membranes


36

With the flow changes, the process stabilized. There were less frequent

shutdowns and the system was able to run continuously for a few days without

alarm. This was an important step as up to this point the operation would

typically shut down overnight and need to be restarted in the morning. The above

operational chart, Figure 16, shows the pressure stabilization from March through

to late April. With the confidence in multi-day operation, a decision was made to

replace the anion membranes in the EDR stack. Now that there was stabilization

in the process, it was important to show how a new unit would be expected to

operate. These new membranes take us into problem solving step 3.

Problem Solving Step 3: New Membrane Installation – April 28th, 2015

With the lessons learned in Steps 1 and 2, the system was ready for a longer-term

steady-state operation. To enhance the performance of the system, the anion

membranes in the EDR stack were replaced. This replacement gave a fresh start

to the overall operation of the plant. The goal was to show better product water

quality with the new membranes. Since the previous membranes had been fouled

with silica in Step 1, the product water quality was not meeting expectations until

Step 3.

Process stabilization and new membranes defined a new phase of operation. It

was in this phase that the operation performed as expected against most of the

parameters. Pressure stabilized and product quality was within the expected

parameters. The data in Figure 17 shows the improvement throughout the study.

On the left side of the graph (representing the performance before the membrane

change), the product quality is quite varied and inconsistent. With the membrane

replacement, an immediate change was noted. The product conductivity band is

tighter and more consistent.


37

Figure 17: Product Conductivity Between March 1st and July 31st


Menifee shut down for maintenance, inspection and cleaning of EDR stack

Reversal noise: high conductivity recorded briefly during reversal/off-spec step


38

With the improvement in flow stability and product quality, the team felt confident

to return to the original RO recovery of 75%, but the AquaSel recovery was

decreased to 80% to achieve an overall recovery of 95%. The parameters of the

EDR system were adjusted slightly to accommodate this change. From this point,

the product quality became even better, achieving the goals for blending back to

the RO plant.

The scattering of data points in Figure 17 is due to reversal noise in the EDR. As

the EDR reverses, the polarity of the stack changes and for this short time, the

process sends product to the off-spec stream until it returns to the expected

values. While this high conductivity product does not mix with the actual product,

the conductivity sensor does not know the difference. Therefore, the sensor reads

during the entire operating time and measures both product and off spec. In the

pilot, the off-spec stream goes directly into the concentrate loop.


39

System Design: Recirculation Loop

This section will discuss observations and events related to operation of the brine

recirculation loop. As discussed, the brine stream flowed from the EDR stack,

through the precipitator, and then back to the EDR stack, where the goal of the

precipitator was to allow the calcium sulfate to drop out of solution and collect at

the bottom of the precipitator tank. The outside chamber of the precipitator is an

up-flow clarifier. As the concentrate flows up through this chamber, the solids

that form settle back down to the bottom of the tank. Ultimately, the concentrate

leaving the precipitator should have a negligible amount of solids in the stream.

During operation, solids were flowing through the precipitator. There was

noticeable buildup around the collection ring where the concentrate water was

pumped out back toward the EDR via the filtration step. This buildup was noticed

through the operation of the pilot and the solids were identified as gypsum. The

pictures below in Figure 18 show the collection ring with the buildup. Solids could

have deposited during upsets in the precipitator flow experienced early in pilot

operation before flow stabilization. However, it is more probable that the up flow

velocity of the precipitator was too high. With a high velocity, the water would be

able to carry the particles up through the tank and onto the collector ring.

Figure 18: Calcium Sulfate Buildup in Precipitator Piping

As previously mentioned, in order to control solids intrusion back into the

recirculation loop to the EDR stack, the precipitator requires a moderate up-flow

rate. Based on the flows through the EDR, the pilot precipitator achieved

approximately 0.46 gpm/ft2. Since the size of the precipitator could not be

changed for the pilot, the system continued to run at this rate. As GE looks to the

next level of design for this project, this value will be reduced to ~0.4 gpm/ft2. By

reducing the velocity of the flow through the precipitator, the system will be able

to regulate and reduce the amount of solids flowing back to the EDR stack.


40

Lowering the loading rate of the precipitator below 0.4 gpm/ft2 is possible, but will

require significant increases in tank volumes. A rate of 0.4 gpm/ft2 will result in

conservative loading of the precipitator while maintaining reasonable tank

volume. The value of 0.4 gpm/ft2 is based on an analysis of particle rise rate

versus settling rate for varying particle size. For a 20 µm particle the rise rate at

0.46 gpm/ft2 is 1.12 m/h and the settling rate is 1.03 m/h, so the 20 µm particle

will have a tendency to rise. At 0.4 gpm/ft2 the rise rate for a 20 µm particle is 0.98

m/h versus the settling rate of 1.03 m/h, so the particle will have a tendency to

settle. The ideal particle size for precipitator operation ranges between 30 and 100

µm, but the 0.4 gpm/ft2 flow rate will accommodate the small particles (less than

30 µm. Figure 19 shows the settling velocity at 20°C and the rise rates of 0.46 and

0.4 gpm/ft2 graphed against particle size.

Figure 19: Stokes Law Settling Velocity Versus Particle Size

In addition to the up-flow velocity in the precipitator, the particle sizes of the solids

are important as well. During initial pilot operation, a centrifugal pump was used

for internal recirculation of the slurry of solids in the central reaction zone of the

precipitator. The nature of operation of the centrifugal pump was intermittent and

at a high velocity, which was difficult to control and caused shearing of the

particles creating smaller particle sizes. The smaller particles settled more slowly

than larger particles, allowing them to be carried out of the precipitator more

easily. In order to maintain a consistent recirculation rate and low particle shear, a


41

peristaltic pump was used in place of the centrifugal pump. This allowed the

particle size to increase, as measured by the short period of time the particles

took to settle, and decreased particle carryover. It also allowed for more

consistent recirculation at ~4 gpm, ensuring that the precipitator was well mixed

and that solids blowdown to waste could occur easily.

As a second line of defense to prevent solids from entering the EDR stack, a filter

skid was placed in the process. As the water flowed out from the precipitator, it

was pumped through the filter skid before re-entering the EDR stack. The

intended method of filtration on this skid was to utilize candle filters in the one-

pass system. These filters are back-washable filters that will go into a backwash

sequence based on a timer. With two filters in parallel, one filter is always in

operation while the other is in backwash or standby. The benefit of these filters is

that the backwash sequence should clean off the scale that has been filtered out

of the concentrate stream. The solids filtration option on this skid is cartridge

filters. There are valves on the skid that allow the operator to choose which filter

style to operate. The diagram in Figure 12 shows the setup for this filter skid. If

the one-pass filter is down, then the cartridge filters could be placed into

operation. The issue with running cartridge filters is that they cannot be

backwashed, so they need to be replaced when used.

Figure 20: One-Pass Filter and Cartridge Filter Process Flow

During the pilot, problems developed in the one-pass filter during operation. The

filters were backwashed as part of normal operation, but over time, the backwash

procedure did not return the filter to a clean state. The filters were disassembled,

and there was a high level of solids in the filter housings as shown below in Figure

21, so the filter housing was cleaned out, the filter socks were replaced, and were

put back into service. Following this clean it was evident that there was buildup

accumulating in the housing again as noted by the pressure drop across the

filters. The buildup in the filter housing could be attributed to several factors. First


42

the solids passing through the precipitator (as mentioned above) and second, the

backwash sequence (frequency and flow). As the candle filters became fouled

again, the one-pass filter was taken off-line, and the cartridge filters were put into

service.

Figure 21: Calcium Sulfate Build-up in Candle-Filters

Operation continued with the cartridge filters. During this operation scheme,

there was still solids breakthrough from the precipitator. Therefore, solids would

build up on the cartridge filters quickly. It was observed that the cartridge filters

required replacement every 3-5 days in order to prevent them from becoming

completely clogged. For the pilot study, this frequency was manageable. The

system could be shut down and restarted easily. In a full scale system, downtime

represents money. Frequent filter changes would be costly and lead to extended

periods of downtime. The picture below in Figure 22 shows a filter in operation for

five days as compared to a brand new filter (shown in the background).

Several cartridge filters were collected and dried to determine the approximate

loading rate of solids in the five-micron filters. In general, cartridge filters had to

be changed every 3-5 days, depending on how quickly the pressure drop

increased across them. Based on the filters collected and weighed by GE, on

average, the filters had to be changed when they reached between 15-25 pounds

between the three filters, or 5-8 pounds per filter. The calculated average loading

rate based on the pilot run time while those filters were in operation was 0.3 – 0.6

lbs/hr.


43

Figure 22: Calcium Sulfate Build-up on Cartridge Filters

Once the process was stabilized and operating for many weeks at good steady

state flows, GE decided to further evaluate the one-pass filters. The hope of this

evaluation was to develop a good cleaning procedure. If an effective cleaning

strategy could be devised, then the one-pass filters could be put back in service

for the remainder of the study. Several cleaning chemicals were identified, and a

few were selected for trials. While on-site, a GE representative attempted

cleanings on these filters. Unfortunately, the cleaning was not successful. The

filter could have been too heavily loaded with solids or the chemicals were

ineffective. Since these tests were conducted in late June, there was not enough

time left in the study to continue to test alternative cleaning chemicals.

In the next size system, the design of the intermediate filter skid needs to be

reviewed. Improved precipitator performance is critical to the success of the

filtration step. Even with an updated filter system, the level of solids should be

reduced coming into this filter skid. Reduced influent solids will greatly improve

the functionality of any filter used at this part of the process. Further design

evaluation will be part of the next phase.

Precipitator Performance

This section will discuss the operation of the precipitator throughout the pilot

study. A key aspect of the pilot study was to confirm ability to adequately

precipitate calcium sulfate from the recirculating EDR concentrate solution in a

CaSO4 seeded slurry process unit. Calcium and sulfate were added to the

concentrating side of the EDR process and subsequently removed in the form of

CaSO4 as precipitated solids, and as dissolved Ca and SO4 ions in a liquid

blowdown stream of the concentrate. The precipitation and blowdown allowed


44

excess (supersaturated) calcium sulfate to come out of solution during the

precipitation and filtration steps as opposed to in the EDR stack, which would

have caused significant scale and inoperability.

As such, the precipitator is an essential part of the process. In order for the

AquaSel process to work, a supersaturated solution containing calcium sulfate

solids must be maintained in the central column of the precipitator. If the

percentage of solids in the precipitator gets too high, the large number of solids

can scale or clog the precipitator and piping. On the other hand, little or no

precipitation will take place if the solids percentage is too low and if the

concentrate stream drops down below the saturation point for calcium sulfate.

The presence of the right amount of solids is important in the precipitator as it

allows for precipitation sites for calcium sulfate. Therefore, maintaining optimal

operation through the precipitator needs to be monitored throughout the testing.

With the exception of two mechanical failure events, the precipitator was in

operation for the entire pilot study. The mixer effectively kept the calcium sulfate

crystals from settling out and hardening on the bottom of the tank. There was a

constant slurry of supersaturated calcium sulfate solution in the central column of

the precipitator, allowing precipitation to occur when the EDR was in operation.

The precipitator levels were continuously monitored during the pilot study to

ensure the correct loading. The recirculation stream was sampled and measured

daily to understand the amount of solids and particle size. A graduated cylinder

was used for this testing. In the picture on the left of Figure 23 the water was

cloudy. The particles were suspended in solution and slowly settled out as seen in

the picture on the right of Figure 23. Both the volume of solids and the time to

settle were recorded on the log sheet. This data was used to calculate solids

percentage and particle size in the precipitator.

As discussed previously, solids recirculation in the precipitator played an

important role in operation of the pilot. The internal recirculation in the

precipitator pulled the slurry of solids from the bottom cone of the precipitator

and fed it back to the top. During the pilot, the most consistent recirculation was

achieved using a peristaltic pump that allowed for continuous solids recirculation

at a rate of ~4 gpm. The peristaltic pump did not shear the solids as it was

recirculating, allowing the particle size to grow and improving the settling ability of

the solids.


45

Figure 23: Precipitator Solids Settling Test

Figure 24, below, shows the solids percentage by volume in the precipitator during

the months of June and July, 2015. During the month of June and July there were

several events that affected the solids percentage in the precipitator. In general,

calcium sulfate did prove to precipitate in the precipitator, which is represented on

the graph by periods with increasing solids percentage. Periods of decreasing

solids percentage were either a result of solids blowdown to the waste bin, or due

to an upset event such as a pump failure. For example, during the period of July

14th to 16th solids blowdown occured at a rate of 1.25 gpm on average. However,

on July 17th, there was a flow imbalance in the precipitator causing it to overflow,

losing a large amount of solids. Solids were also lost during mechanical failures of

the solids recircuation pump, when the peristaltic pump tubing ruptured.


46

Figure 24: Precipitator Solids Percentage by Volume

Chemical Consumption

This section will discuss the chemicals used during pilot operation. As mentioned,

some adjustment was made to the chemical dosing program during the pilot

study. Initially there were two injection points for sulfuric acid, as well as ferric

chloride addition. However, during the final operating phase two chemicals were

used in normal operation. Hydrochloric acid (HCl) was dosed directly into the

electrode stream of the EDR stack to prevent calcium carbonate scale and to

prevent deactivation of residual antiscalant (present in the RO brine). Deactivation

occurs at higher pH (typically generated at cathodes during EDR operation)

according to the manufacturer. The antiscalant controls silica scaling and activity

must be maintained through the EDR process. Sulfuric acid was dosed directly to

the precipitator to lower the pH and prevent calcium carbonate scale in the

concentrate loop.

HCl was dosed at a constant rate of 300 mL/h (534 mg/L) to the electrode stream.

The electrode stream flowing next to the cathode generated hydroxide alkalinity,

which increased the pH and the potential for calcium carbonate and silica scale.

By dosing 300 mL/h (534 mg/L) it was possible to maintain the electrode stream


47

cathode outlet below a pH of 7, thus decreasing the risk of scale and antiscalant

deactivation. Every seven hours, the HCl flow rate was increased to 1120 mL/h

(1993 mg/L) for 60 minutes. The periods of increased HCl flow were to perform an

electrode clean in place (ECIP), which is essentially acidifying both electrode

streams to a pH of less than 2 to remove any calcium carbonate scale that may

have built up. The average of these two flows over a 24 hour period resulted in a

flow rate of 402.5 mL/h (715 mg/L) of 32% HCl dosed into the system.

As discussed, the initial operating condition had sulfuric acid being dosed into the

EDR feed stream as well as into the precipitator. However, the sulfuric acid dose to

the EDR feed stream was removed, so the only dosing location for sulfuric acid

was directly into the center reaction zone of the precipitator. Sulfuric acid dosing

to the precipitator was controlled using a pH probe and an SC200 controller to

dose acid based on a pH set point of 6.3. Due to the nature of this dosing system,

and the interruptions in the process, the sulfuric acid consumption rate was

calculated by correlating the day by day consumption of acid from the dosing

tank and the hours of run time on each respective date. Based on the above

method, the calculated average dose rate of sulfuric acid was 402 mL/h of 50%

sulfuric acid, which corresponds to approximately 182 mg/L. The calculated dose

rate ranged between 211 mL/h (95 mg/L) at the lowest and 655 mL/h (296 mg/L)

at the highest.


48

Discussion

Pilot Operation Phases

Pilot operation is reported in three phases as related to adjustments made in

recovery of the Menifee RO de-salter and pilot system. The basic conditions are

described below.

Table 3: Recovery Phases

Phase Menifee RO Recovery Pilot System Recovery

1 75% 90%

2 70% 90%

3 75% 80%

Operational Results – Phase 1 – November 28, 2014 to December 3, 2014

Pilot experiences early during Phase 1 led to modifications of the system. An

expected mass balance has been developed for Phase 1, based on the modified

process arrangement. A key modification was to recycle EDR off spec product to

the concentrate loop rather than to EDR feed.

Conditions targeted for Phase 1 included:

Menifee de-salter recovery 75%

Pilot feed flow rate 7.1 gpm

Pilot feed acidified to pH 5.5

EDR brine precipitator process recovery 90%

Brine precipitator acidified to pH 5.5

Solids blowdown rate 6 gallons per hour

Precipitator solids concentration 10 – 16% by volume, 5 – 8% by weight

Initial estimation of system performance for phase 1 was obtained from GE’s

WATSYS (EDR design software) projection in Table 4 below. Compositional

information for development of mass balance was taken from the WATSYS

projection below. Feed water analysis indicated in the pilot proposal and

projected product quality from the projection below was used.


49

Table 4: WATSYS Design - Phase 1

Process Pressures, Stream Conductivities, Voltage and Current, DC Power Consumption – Phase 1

Process modifications were made during early operation to manage scaling in the

front end of the system. The following charts in Figures 25 - 29 illustrate EDR feed

pressures, process stream conductivities, DC voltage and amperages for

membrane stacks and estimated DC power consumption. Table 6 and Figure 30

show a mass balance table and process flow diagram (PFD) for Phase 1

operations.

EDR feed pressures were rising due to fouling from scaling (previously described)

as the feed pumping maintains constant flow through the diluting side of the

membrane stack. Product conductivity was better than expected (< 3,000 µS/cm)

Project Name EMWD

Company

User KJI

WATSYS Run Date Thursday, September 03, 2015 Recycled Feed Product Conc. BD Waste Raw Feed

Calcium mg/l 1173.9 299.4 8805.4 8805.3 1150.00

Number of Lines 1 Magnesium mg/l 298.6 89.1 2128.3 2128.3 293.00

EDR System AQ3-1-4 with 1 Line(s) 4 Stage(s) Sodium mg/l 1173.1 536.9 6757.7 6757.6 1159.00

Anion Membrane AR204 Potassium mg/l 31.9 11.5 210.9 210.9 31.40

Cation Membrane CR67 Strontium mg/l 0.0 0.2 0.0 0.0 0.00

Spacer Mark IV-2 Barium mg/l 1.1 0.3 8.1 8.1 1.05

Ammonia mg/l 0.0 0.0 0.0 0.0 0.00

Bicarbonate mg/l 121.1 83.9 442.4 442.3 119.70

EDR Product 6.4 USGPM Sulphate mg/l 1337.7 352.2 9940.4 9940.3 1311.00

Dilute In 7.1 USGPM Chloride mg/l 3640.0 1285.2 24243.7 24243.5 3581.00

Dilute Flow Losses 0.5 USGPM Fluoride mg/l 0.6 0.4 2.5 2.5 0.63

Dilute Out 6.6 USGPM Nitrate mg/l 153.3 58.7 982.0 982.0 151.00

Off-Spec Product 0.2 USGPM 3% OSP Total PO4 mg/l 0.8 0.3 8.0 8.0 0.80

Feed Pump 8.2 USGPM HPO4 mg/l 0.0 0.0 0.4 0.4 0.00

Concentrate Pump 6.4 USGPM H2PO4 mg/l 0.8 0.3 7.6 7.6 0.80

Electrode Waste 0.9 USGPM Silica mgl 166.0 166.0 166.0 166.0 166.00

Concentrate Makeup Flow 0.2 USGPM CO2 mgl 671.90 671.90 671.95 671.95 679.19

Net System Feed into EDR 7.1 USGPM Carbonate mgl 0.00 0.00 0.02 0.02 0.00

Total System Waste 0.7 USGPM 10% Waste w/o Bypass Total Hardness CaCO3 4157.4 1113.8 30726.6 30726.4 4074.80

Concentrate Blowdown 0.7 USGPM TDS mg/l 8098.1 2884.0 53695.2 53694.8 7964.60

System Feed w/ Bypass 7.1 USGPM Conductivity uS/cm 12211.3 4651.3 63663.6 63663.3 12029.50

Bypass Feed to Product 0.0 USGPM pH 5.51 5.35 6.07 6.07 5.50

Minimum Velocity 7.11 cm/s WATSYS % Saturation

First Stage Inlet Pressure 29.41 psig CaSO4 78.0 17.7 608.4 608.4 76.27

Last Stage Outlet Pressure 5 psig BaSO4 806.5 337.5 2452.0 2452.0 796.65

SrSO4 0.6 8.0 0.0 0.0 0.00

Temperature 18 C CaF2 93.8 45.2 456.4 456.4 92.40

Pumping Power 3.48 kWh/kgal CaHPO4 20.6 5.5 326.4 326.4 19.76

DC Power 15.93 kWh/kgal Ca3(PO4)2 7.6 1.7 173.2 173.2 7.22

Total Power 19.41 kWh/kgal

Total DC KVA 6.87 KVA Langelier Index (LI) -1.14 -2.04 0.91 0.91 -1.17

Feed Pump Power 0.8 hp Stiff-Davis Index (SDI) -1.58 -2.26 -0.23 -0.23 -1.60

Concentrate Pump Power 0.72 hp SAR 7.90 6.99 16.75 16.75 7.89

Flow Rate USGPM 7.1 6.4 0.7 0.7 8.0

See table on PFD tab for additional chemical dosage information

Electrical Stages 1 2

Voltage (V) 76 63

Current (Amps) 40.7 28.6

Surge (Amps) 123.3 142.0

Hyd Stages / Elect Stage 2 2

Hydraulic Stages 1 2 3 4

% Polarization 12.86 16.99 14.58 20.09

Cut Fraction 0.20 0.25 0.22 0.29

Current Efficiency 0.73 0.72 0.68 0.68

% Manifold Shorting 29.61 31.28 24.15 26.17

Cell Pairs 40 40 40 40


50

and concentrate conductivity was effectively driven up to where expected (>

60,000 µS/cm).

Table 5 shows data collected at site for period of November 30 to December 9,

2014, using onsite HACH test kits (for calcium and sulfate) and a handheld meter

(conductivity and pH). It should be noted that these values differ from the Phase 1

values in Table 1, likely due to water quality at the time of testing and the test

method.

Table 5: Operational Data - Phase 1

Site Data pH Conductivity Sulfate Calcium

µS/cm mg/l mg/l

Raw water 7.51 10780 390 1032

Acidified EDR feed 5.6 10730 590 1056

Product 5.5 2154 48 48

EDR Concentrate out

6.06 67650 1200 9880

Precipitator effluent

6.56 56740 - -

One-pass effluent - - 900 7140

Above data supports the conductivity expectations for product and concentrate

streams, and removal of salts in the precipitator (CaSO4), however the relative

removals of calcium and sulfate through the precipitator were not consistent.

This may have been from periodic instabilities where internal recycle of process

streams contributed to wide variation in analytical results, even though samples

were taken at the same time.


51

Figure 25: EDR Pressures During Phase 1 Operations

Localized scale found in concentrate piping at acid injection point, scale cleaned.


52

Figure 26: EDR Conductivities During Phase 1 Operations

Test of precipitator solids blowdown


53

Figure 27: EDR Voltage and Amperage During Phase 1 Operations

Stage 2 amperage increases as membranes acclimatize and alter efficiency


54

Figure 28: Product Conductivity During Phase 1 Operations

Product quality decreases as membranes scale with silica


55

Figure 29: EDR Stack Power Consumption During Phase 1 Operations

Power increases, due to stage 2 amperage increase, as membranes acclimatize during initial operation


56

Table 6: Mass Balance For Phase 1 Operations

Stream ID 1 2 3 4 5 6 6A 7 7A 8 9 10 11

Description Raw Feed EDR Feed

Pump

EDR Feed

(dilute in)

Electrode

Feed/Recycle

Concentrate

MakeupProduct

Off Spec

Product

EDR

Concentrate

Out, pH

adjusted

Precipitator

Out

EDR

Concentrate

In

Concentrate

Blowdown

Precipitator

Blowdown

CaSO4

removed

(100%

basis)

Flow Rate (GPM) 7.10 8.00 7.1 0.9 0.00 6.4 0.2 6.9 6.20 6.40 0.60 0.10 Balance Total

Pressure (PSIG) 10 38 33 30 33 5 5 4 4 32 3 3 Check Liquid Solid

Temperature (°C) 18 18 18 18 18 18 18 18 18 18 18 18 Feed Product Blowdown Blowdown

kg/hr kg/hr kg/hr kg/hrCalcium mg/l 1150.00 1150.00 1150.00 1150.00 1150.00 297.34 1150.00 5929.2 5579.2 5360.1 5495.9 5495.9 350.0 1.8543 1.8543 0.4322 0.8737 0.548447Magnesium mg/l 293.00 293.00 293.00 293.00 293.00 88.60 293.00 2182.8 2182.8 2103.2 2161.6 2161.6 0.4724 0.4724 0.1288 0.3436 0Sodium mg/l 1159.00 1159.00 1159.00 1159.00 1159.00 540.36 1159.00 6899.2 6899.2 6638.5 6815.2 6815.2 1.8688 1.8688 0.7854 1.0834 0Potassium mg/l 31.40 31.40 31.40 31.40 31.40 11.57 31.40 214.6 214.6 206.6 212.2 212.2 0.0506 0.0506 0.0168 0.0337 0Strontium mg/l 6.10 6.10 6.10 6.10 6.10 0.23 6.10 60.0 60.0 57.9 59.5 59.5 0.0098 0.0098 0.0003 0.0095 0Barium mg/l 1.05 1.05 1.05 1.05 1.05 0.27 1.05 8.1 8.1 7.8 8.0 8.0 0.0017 0.0017 0.0004 0.0013 0Bicarbonate mg/l 119.70 119.70 119.70 119.70 119.70 84.30 119.70 47.2 47.2 41.1 38.5 38.5 0.1930 0.1930 0.1225 0.0061 0Sulfate mg/l 1311.00 1311.00 1311.00 1311.00 1311.00 349.60 1311.00 3074.4 2234.4 2113.5 2139.4 2139.4 840 2.1645 2.1645 0.5081 0.3401 1.316272Chloride mg/l 3581.00 3581.00 3581.00 3581.00 3581.00 1286.80 3581.00 24815.8 24815.8 23900.8 24556.3 24556.3 5.7740 5.7740 1.8703 3.9037 0Fluoride mg/l 0.63 0.63 0.63 0.63 0.63 0.43 0.63 2.41 2.41 2.31 2.36 2.36 0.0010 0.0010 0.0006 0.0004 0Nitrate mg/l 151.00 151.00 151.00 151.00 151.00 58.77 151.00 1005.3 1005.3 968.1 994.4 994.4 0.2435 0.2435 0.0854 0.1581 0Total PO4 mg/l 0.81 0.81 0.81 0.81 0.81 0.33 0.81 5.22 5.22 5.02 5.16 5.16 0.0013 0.0013 0.0005 0.0008 0Silica mgl 166.00 166.00 166.00 166.00 166.00 166.00 166.00 178.3 178.3 166.2 166.3 166.3 0.2677 0.2677 0.2413 0.0264 0CO2 mgl 619.86 616.8 616.8 616.8 616.8 620.01 616.8 243.3 243.3 211.6 198.6 198.6

Total Hardness CaCO3 4075.00 4082.2 4082.2 4082.2 4082.2 1108.4 4082.2 23816.4 22941.4 22065.4 22645.5 22645.5

TDS mg/l 7965.00 7970.7 7970.7 7970.7 7970.7 2884.6 7970.7 44422.6 43232.6 41571.1 42655.0 42655.0

Conductivity uS/cm 12029.00 12029.0 12029.0 12029.0 12029.0 4654.00 12029.0 63630.0 61925.5 59545.5 61098.0 61098.0pH 5.5 5.5 5.5 5.5 5.5 5.4 5.5 5.5 5.5 5.5 5.5 5.5

acid added acid added %CaSO4 CaSO4

H2SO4 (100%) 658.44 mg/l H2SO4 33 mg/l 133 kg/hr

HCO3 0.0643 kg/hr 1.8647

50% H2SO4 1522 ml/hr 50% H2SO4 50% H2SO4 74 ml/hr

gpm 0.1 8.211 wt%


57

Figure 30: Simplified Process Flow Diagram For Phase 1 Operations

H2SO4 Stream ID Description Flow Rate

HCl 4 (gpm)

1 Raw Feed 7.1

4 2 EDR Feed 8.0

1 3 Dilute In 7.1

2 3 6 4 Electrode Feed/Recycle 0.9

5 Concentrate Makeup 0.0

6 Product 6.4

5 6A 6A Off Spec Product 0.2

7 Concentrate Out 6.9

8 7 7A Precipitator Effluent 6.2

8 Concentrate In 6.4

9 Concentrate Blowdown 0.6

10 Precipitator Blowdown 0.1

H2SO4

BW Waste

9

7A 10

BW Feed

Feed Tank

EDR

Precip-itator

BW Tank

Filter


58

Operational Results – Phase 2 – January 28, 2015 to May 7, 2015

Experiences during Phase 1 led to revisions in the piloting process. To manage

fouling concerns, EMWD agreed to lower the RO plant recovery. This would

provide an RO concentrate for treatment in the pilot with lower levels of saturated

salts (particularly silica). In addition, pH adjustment of feed to the pilot was

removed. A mass balance was developed for Phase 2, based on the modified

process arrangement.

Conditions targeted for Phase 2 included:

Menifee de-salter recovery 70%


Pilot feed pH 7.7 (not acidified)


Brine precipitator acidified to pH 6.3

Solids blowdown rate 6 gallons per hour

Precipitator solids concentration 10 – 16% by volume, 5 – 8% by weight

Initial estimation of system performance for Phase 2 was obtained from the

WATSYS projection below. Analytical results obtained during Phase 2 were used

to develop a representative mass balance, as shown in Table 7, below.


59

Table 7: WATSYS Design - Phase 2


During Phase 2, hydraulic balance of the pilot system was improved (with regard

to management of internal recycle streams). This reduced volatility of precipitator

operation. Filtration and recycle of precipitator effluent back to the EDR

membrane stack had been challenged by higher solids than intended and

pressure drop in the membrane stack was rising quickly. The improvement can

be seen in the EDR pressure trend below. The corresponding buildup in EDR

concentrate conductivity followed from this more stabilized period of operation.

Project Name EMWD

Company

User KJI

WATSYS Run Date Thursday, September 03, 2015 Recycled Feed Product Conc. BD Waste Raw Feed

Calcium mg/l 968.0 316.9 6828.4 6828.3 968.00






Ammonia mg/l 0.0 0.0 0.0 0.0 0.00

Bicarbonate mg/l 783.2 515.4 3193.4 3193.4 783.20











Total System Waste 0.7 USGPM 10% Waste w/o Bypass Total Hardness CaCO3 3343.3 1141.2 23163 23162.9 3343.30







SrSO4 34.6 8.4 117.0 117.0 34.59





Total DC KVA 2.64 KVA Langelier Index (LI) 1.68 0.85 3.68 3.60 1.68

Feed Pump Power 0.79 hp Stiff-Davis Index (SDI) 1.34 0.64 2.65 2.57 1.34


Flow Rate USGPM 7.1 6.4 0.7 0.7 7.0



Voltage (V) 47 50


Surge (Amps) 47.4 66.4



% Polarization 9.82 11.99 14.10 18.99

Cut Fraction 0.16 0.18 0.21 0.27





60

An analytical data point was obtained on April 17, 2015 to gauge the process

conditions during Phase 2. The results from GE Woodlands analytical laboratory

are listed below in Table 8.

Table 8: Analytical Results, April 17, 2015

Parameter Units Feed EDR Product

EDR Concentrate

Precipitator Inlet

Precipitator Outlet

RO Recovery 70%

Calcium mg/l 976 124.8 8040 7360

Magnesium mg/l 211.7 40.5 2043.7 1912.6

Sodium mg/l 627 435 2530 2400

Potassium mg/l 21.5 10.9 125 118

Strontium mg/l 6.4 0.8 62.3 58.5

Barium mg/l 0.73 0.07 4.99 4.51

Ammonia mg/l

Bicarbonate mg/l 735.7 466.0 856.4 671

Sulfate mg/l 422 47.9 1750 1570

Chloride mg/l 2350 949 21500 22300

Fluoride mg/l 0.6

Nitrate mg/l 119 29.1 1100 1100

Total PO4 mg/l 2.5

Silica (total) mg/l 148 149 107 107

Aluminum mg/l 0.6 0.1 1.6 1.6

CO2 (calculated) mg/l

Total Hardness CaCO3 3320 480 28600 26300

TDS mg/l 5618 2256 38120 37602

Conductivity µS/cm 8760 4130 56600 54500

pH 7.7 7.5 6.6 6.5

Operational charts from Phase 2 operations can be found in Figures 31 – 35 and a

mass balance can be found in Table 9 accompanied by a simplified PFD in Figure

36. As discussed previously, key changes during Phase 2 operation included the

removal of the pretreatment system, especially sulfuric acid dose to the feed, and

replacement of EDR anion membranes. These and other significant events are

highlighted on Figures 31 – 35, below.


61


Acid CIPs of EDR stack

Pretreatment removed, acid addition to feed stopped



62



Precipitator mixer falls off; precipitator is drained, then refilled and re-seeded

Off-spec flow increased, diluting concentrate stream

Off-spec flow adjusted to achieve 90% EDR recovery


63


Voltage adjusted; Stage 1 @ 48 V, Stage 2 @ 46 V

Voltage adjusted; Stage 1 @ 40 V, Stage 2 @ 40 V Voltage

adjusted; Stage 1 @ 38 V, Stage 2 @ 38 V


64

Figure 34: EDR Product Conductivity During Phase 2 Operations


Voltage adjusted; stage 1 @ 40V, stage 2 @ 40V

Voltage Adjusted

Product quality decreases to due scale in EDR stack


65

Figure 35: EDR Stack Power Consumption During Phase 2 Operations





66

Table 9: AquaSel Process Mass Balance for Phase 2 Operations

Stream ID 1 2 3 4 5 6 6A 7 7A 8 9 10 11


Pump

EDR Feed

(dilute in)

Electrode

Feed/Recycle

Concentrate

MakeupProduct

Off Spec

Product

EDR

Concentrate

Out, pH

adjusted

Precipitator

Out

EDR

Concentrate

In

Concentrate

Blowdown

Precipitator

Blowdown

CaSO4

removed

(100%

basis)






TDS mg/l 5622.9 5622.9 5622.9 5622.9 5622.9 2255.2 5622.9 30927.9 30298.9 29097.6 29880.1 29880.1


acid added %CaSO4 CaSO4H2SO4 (100%) H2SO4 205 mg/l 138.7 kg/hr

HCO3 0.3882 kg/hr 0.9571

50% H2SO4 50% H2SO4 447 ml/hr

gpm 0.14.2143

wt%


67

Figure 36: Simplified Process Flow Diagram For Phase 2 Operations

Stream ID Description Flow Rate

HCl 4 (gpm)

1 Raw Feed 7.1

4 2 EDR Feed 8.0

1 3 Dilute In 7.1



6 Product 6.4







H2SO4

BW Waste

9

7A 10

BW Feed

Feed Tank

EDR

Precip-itator

BW Tank

Filter


68

Operational Results – Phase 3 – May 7, 2015 to July 31, 2015

Results for Phase 3 are discussed in terms of consistency and stability of throughput

(pressures and flows), electrical performance of the EDR process (membrane stack DC

voltage and stability of current), desalting performance (salt removal and product quality)

and precipitator performance (calcium removal from precipitator influent to effluent and

calcium sulfate solids generation).

Conditions targeted for Phase 3 testing included:

Menifee RO de-salter recovery 75%



EDR product conductivity < 3,200 µS/cm

Solids blowdown rate 6 gallons/hour

Precipitator solids concentration 10 - 16% by volume, 5 – 8% by weight

Initial estimation of system performance for Phase 3 was obtained from the WATSYS

projection below in Table 10.


69

Table 10: WATSYS Design For Phase 3 Operations

The Phase 3 test ran virtually continuously for 962 operating hours from the final

recovery adjustment until the end of the study. During this time 16 sets of samples were

collected and analyzed in the laboratory. Key results of these analyses were viewed in

association with operating data at the times of sampling. Analytical results obtained

during Phase 3 were used to develop a representative mass balance.

Project Name EMWD

Company

User

WATSYS Run Date Wednesday, May 06, 2015 Recycled Feed Product Conc. BD Waste Raw Feed

Calcium mg/l 1150.0 200.6 4947.4 4947.4 1150.00






Ammonia mg/l 0.0 0.0 0.0 0.0 0.00

Bicarbonate mg/l 935.7 467.6 1546.4 1546.4 935.74


















SrSO4 0.0 3.9 0.0 0.0 0.00








Flow Rate USGPM 8.0 6.4 1.6 1.6 8.0



Voltage (V) 75 62


Surge (Amps) 70.1 82.4



% Polarization 15.58 21.98 19.29 29.98

Cut Fraction 0.23 0.30 0.28 0.38





70

Process Flows and Pressures – Phase 3

A sample chart of flows and pressures for the EDR process is shown below in Figure 37.

Figure 37: AquaSel Flows June 19, 2015

This is for an operating period when the first samples were collected for analysis in Phase

3. Flows read on the left axis show the product flow rate stable at 6.6 gpm and

concentrate makeup stable at 0.9 gpm. The electrode flow rate varied between 0.9 and

1.8 gpm. This illustrates the “bumping” or periodic flowing of the anode stream was

working correctly.

The dilute inlet and concentrate inlet streams to the EDR membrane stack are read on

the right axis. These data show a slightly higher inlet pressure for the dilute versus

concentrate inlet, which is by design as the EDR process is intended to operate with a

slightly higher inlet pressure on the diluting side of the membranes in the stack. The

higher pressure on the dilute side is to prevent the concentrate stream from

contaminating the dilute product water. Any cross-leakage between streams will flow

from the higher pressure (dilute) stream to the lower pressure stream (concentrate). The

data point where inlet and outlet pressure become inverted occurs when DC polarity is

reversed on the membrane stack every 15 minutes. This illustrates the control system is

adjusting the inlet pressure differential as required when each polarity reversal occurs.

Fifteen more similar charts for flow and pressure can be found in Appendix A.1. They are

all very similar demonstrating consistent flow and pressure during Phase 3.

DC Voltage and Current – Membrane Stack – Phase 3

A sample chart of DC voltage and current for each of the two stages of the EDR

membrane stack is shown below in Figure 38.


71

Figure 38: EDR Voltage and Amperage, June 19, 2015

Individual data points are shown for DC voltages applied to electrical stage 1 and

electrical stage 2. Solid lines illustrate current flow in each stage during this period of

operation. The steps up and down illustrate reversal of polarity from positive to negative.

During this time stage 1 voltage was stable at 48 volts and stage 2 at 36 volts on

average. Stage 1 current was stable at 35 amps and stage 2 at about 20 amps. The

current spikes seen in stage 2 are from current surges at polarity reversal as the

concentrating and diluting streams displace one another. This is a normal phenomenon

in EDR process and did not have an impact on the pilot test. Recent EDR technology

upgrades utilize DC drives which minimize current surges during reversal.

Fifteen more similar charts for DC voltage and current can be found in Appendix A.2. The

flatness of the current profile during each polarity as well as the product conductivity

profile during each operating polarity indicates no detectable increase in electrical

resistance of the membranes was seen during the final phase of the pilot. The charts in

Appendix A.2 are all very similar demonstrating that the feed water (Menifee de-salter RO

brine) did not foul the EDR membranes during the final phase of the pilot.


72

Conductivity and Constituent Removal – Phase 3

A sample chart of conductivity and calcium in each of three main process streams is

shown below in Figure 39. These are the feed, product and concentrate blowdown

(overflow from precipitator sent to waste).

Figure 39: AquaSel Conductivity and Calcium, June 19, 2015

This is for same operating period as the previous chart. As expected, feed conductivity

was stable at 10,500 µs/cm throughout (RO brine from Menifee de-salter operating at

75% recovery). Product water from the EDR was consistent, averaging about 3,130

µS/cm. Note the spurious readings were not included in the average. The slightly higher

product conductivities observed in Figure 39 at >5000 µs/cm were associated with EDR

polarity reversal, where the positive and negative poles on the EDR membrane stack

switch along with the streams in the EDR stack, during which off spec product is

produced as a normal part of EDR operation. In the pilot process, off spec product was

sent to the recirculating concentrate loop (at inlet to precipitator) where it served as a

small portion of the makeup water to the brine loop. The concentrate conductivity at the

outlet of the EDR membrane stack was stable at 36,000 µS/cm.

Samples for laboratory analysis were taken at the time of this data. Analytical points for

calcium are displayed in Table 11. Feed calcium was 1,079 mg/l and product calcium

was 60 mg/l. Calcium in the effluent from the precipitator (after the precipitation process)

was 4,399 mg/l. Fifteen more similar charts for conductivity and Ca can be found in

Appendix A.3. They are all very similar demonstrating consistent salt removal and

concentration of dissolved salts in EDR effluent.

Averaged analytical results from 16 samples between June 19th and July 31st from the

final phase of testing are listed in Table 11 below.


73

Table 11: Average Analytical Results From Phase 3

Parameter Units Raw Water

Product Waste

Calcium mg/l 1079 60 4399

Magnesium mg/l 261 81 957

Sodium mg/l 774 405 2141

Potassium mg/l 25.7 8.4 90.4

Strontium mg/l 6.1 0.34 22.5

Barium mg/l 0.75 0.03 1.59

Ammonia mg/l < 0.5 < 0.4 < 0.9

Bicarbonate mg/l 809 416 1460

Sulfate mg/l 489 33 1930

Chloride mg/l 2953 629 12544

Fluoride mg/l 0.61 0.48 1.39

Nitrate mg/l 96.9 15.5 427

Total PO4 mg/l 4.1 2.8 5.3

Silica (total) mg/l 154 153 140

Silica (reactive) mg/l 129 124 128

CO2 mg/l 53 53 365

Total Hardness CaCO3 3773 484 14940

TDS mg/l 6653 1825 24119

Conductivity µS/cm 10500 2744 35600

pH 7.4 7.3 6.8

Flow gpm 8.0 6.4 1.6

The above analytical results were fitted with pilot operating data to develop a

representative mass balance for Phase 3 of pilot testing, shown in Table 12 and is

accompanied by Figure 45, which shows a simplified PFD of the process.


Figures 40 – 44 highlight the key operational data points during Phase 3 operation. At the

beginning of Phase 3 the Menifee RO recovery was set to 75% and the AquaSel recovery

was set to 80%, for a total recovery of 95%. The recovery change occurred at the y-axis

on the charts below. During this operational phase, there were fewer changes and upsets

than the previous two phases. It can be seen in Figure 42 that the voltages of EDR stage

1 and stage 2 were adjusted several times. The adjustment to voltage is to maintain the

target product conductivity and amperage as per the WATSYS design projection.


74

During this phase there were two instances where a CIP was required to clean the EDR

stack, as indicated on Figure 40, below. After the first CIP of the phase, the stack was also

disassembled for inspection and in-depth cleaning, removing and replacing any

membranes that looked particularly fouled. Although CIPs were performed during this

phase, significant pressure increase in the EDR stack was not observed, indicating that

fouling was minimal and localized, not widespread as observed previously.

Phase 3 also experienced mechanical and electrical problems which contributed to small

fluctuations in the data. An impact on run-time was an electrical issue caused by loose

wiring which sent false alarms to the PLC and interrupted operation. Mechanically, the

HCl pump dosing to the electrode stream had a worn seal at one point, causing the pump

to frequently lose prime and stop dosing. The lack of HCl dosing was the cause of the

isolated fouling indicated in Figure 40. Finally, the recirculation and pump on the

precipitator had several failures resulting in loss of solids and concentration in the

precipitator, as can be seen in the concentrate conductivity trend in Figure 41.


75


Start of final test period


Acid CIP Acid CIP

Isolated fouling in one polarity


76



Precipitator upset causes loss of solids and concentrate

After recovery change to 80%, start concentrate make-up


77



Stage 1 voltage @ 48 V

Stage 1 @ 47 V, Stage 2 @ 36 V


Stage 1 @ 45 V, Stage 2 @ 35 V, stack inspection and cleaning




78

Figure 43: EDR Product Conductivity During Phase 3 Operations

Reversal noise: high conductivity recorded briefly during reversal/off-spec step, product diverted during this time

Stage 1 @ 50 V, Stage 2 @ 41 V, decreases product conductivity





79

Figure 44 Stack Power Consumption During Phase 3 Operations

Power decrease due to stack fouling and resulting amperage decrease



80

Table 12: AquaSel Mass Balance for Phase 3 Operations

Stream ID 1 2 3 4 5 6 6A 7 7A 8 9 10 11


Pump

EDR Feed

(dilute in)

Electrode

Feed/Recycle

Concentrate

MakeupProduct

Off Spec

Product

EDR

Concentrate

Out, pH

adjusted

Precipitator

Out

EDR

Concentrate

In

Concentrate

Blowdown

Precipitator

Blowdown

CaSO4

removed

(100%

basis)






TDS mg/l 6652.3 6652.3 6652.3 6652.3 6652.3 1805.6 6652.3 24519.8 24009.8 20437.5 23618.7 23618.7


acid added % CaSO4 CaSO4ppm H2SO4 H2SO4 225 mg/l 138 kg/hr

HCO3 0.4324 kg/hr 0.7876

50% H2SO4 498 ml/hr

gpm 0.1 3.468 wt%


81

Figure 45: AquaSel Simplified Process Flow Diagram for Phase 3 Operations

Stream ID Description Flow Rate

HCl 4 (gpm)

1 Raw Feed 8.0

4 2 EDR Feed 8.9

1 3 Dilute In 7.1



6 Product 6.4







H2SO4

BW Waste

9

7A 10

BW Feed

Feed Tank

EDR

Precip-itator

BW Tank

Filter


82

EDR Concentrate Loop and Precipitator Performance

As described previously, in the AquaSel process the concentrate is recirculated through

the EDR membrane stack where dissolved ions are transported from the diluting stream

into the concentrate stream, along with a small amount of water. This increases the

concentration of Ca and SO4 as well as other dissolved ions as the stream moves from

concentrate inlet to concentrate outlet of the membrane stack. From the outlet the

concentrate goes through a seeded slurry (CaSO4) precipitator where calcium and sulfate

are precipitated as calcium sulfate di-hydrate (gypsum). A small amount of barium

sulfate precipitation is also expected.

Below in Figure 46 are the analytical data for calcium, sulfate and % CaSO4 solids in

concentrated waste and sludge blowdown from the precipitator in the final phase of

testing.

Figure 46: Calcium, Sulfate, % Solids in Precipitator Effluent

On average, the effluent from the precipitator contained 4400 mg/l Ca and 1930 mg/l

SO4. However, there is fluctuation in Ca and SO4 concentrations. This is due to recycling

of streams with different compositions and periodic hydraulic imbalances in the pilot

system. The average percent solids from the pilot varied, with values ranging from ~2 to

22% by weight. The high data points of solids percentage are not considered optimal

operation. Averaging all of the recorded data points results in an average of ~11.5% by

weight solids in the precipitator. Removing the high outliers averaged 10% CaSO4 by

volume (5% by weight), which is roughly the preferred operating range. A mass balance


83

developed to reflect approximate operation of the final phase of testing is shown above

in Table 12. A simplified process flow diagram is shown above in Figure 45.

Key points to note in the mass balance are the sulfuric acid dosage required to control

pH in the precipitator, the level of CaSO4 super saturation in the precipitator effluent and

CaSO4 solids generated and removed with sludge blowdown. Note that the weight

percent of CaSO4 sludge is variable depending on the actual flow and duration of periodic

sludge blowdowns. For the purpose of mass balance calculation, an average 0.1 gpm

precipitator sludge blowdown flow rate was assumed and resulted in about 3.5 weight

percent solids (7% by volume). The relative amounts of concentrate blowdown (stream 9)

and precipitator blowdown (stream 10) may be varied to optimize percent solids in the

sludge blowdown. I.e. Mass balance indicates 1.5 gpm concentrate blowdown and 0.1

gpm precipitator blowdown (for 3.5 weight percent solids in 0.1 gpm precipitator

blowdown stream). A 1.55 gpm concentrate blowdown stream and 0.05 gpm precipitator

blowdown would result in 7.0 weight percent solids in the 0.05 gpm precipitator

blowdown stream. The dissolved ions in the combined liquids of these streams remained

the same. A summary of the precipitator performance is shown below in Table 13. As

shown in Table 13, the average solids production during Phase 3 of the pilot was 0.79

kg/hr as 100% CaSO4. Scaling this solids production linearly to a 100 gpm system would

result in a solids production rate of 12.3 kg/hr.

Table 13: Precipitator Summary for Phase 3 Operations

Precipitator Unit Summary AquaSel 80% recovery

Pilot system feed 8.0 gpm

Pilot system product 6.4 gpm

H2SO4 acid consumption 11.95 lpd as 50% H2SO4 3.52 kg/day as 100% H2SO4

Effluent saturation ~ 140% CaSO4

CaSO4 solids 0.1 gpm as 3.5 wt% sludge

0.79 kg/hr as 100% CaSO4

Rise rate 0.46 gpm/ft2

As previously mentioned, there was antiscalant present in the RO brine fed to the pilot.

This antiscalant was also present in the water initially used to fill the precipitator and was

present in the concentrate stream. Based on the results of this pilot, the effect of the

antiscalant on the precipitation process cannot be definitively described, but it may have

prevented the concentrate solution from reaching saturation levels. The precipitator

effluent was measured as a supersaturated solution (140% saturation), which would also

be sent to the Inland Empire Brine Line for ultimate disposal. While a supersaturated

solution may cause scale in the brine line, it is less likely to scale due to the trace

antiscalant present and the blending with other waters which will occur in the brine line.

If scaling does become an issue in future applications, a small dose of antiscalant could

be added to the precipitator liquid effluent to prevent scale.


84

Conclusions

This section highlights the conclusions that can be drawn from the EMWD AquaSel pilot

study. In the context of previously stated objectives the following is concluded.

Establish the performance capacity of the EDR process to remove dissolved salts

For brine from Menifee RO de-salter (operating at 75% recovery), the EDR process

demonstrated it was capable of reducing TDS by 72%. Given RO brine of

approximately 6653 mg/L TDS as feed, the product from a four stage EDR process

can provide quality of approximately 1,825 mg/L TDS at 80% recovery. For the

diluting side of the EDR process, the raw water did not demonstrate any fouling

tendencies toward EDR membranes during Phase 3 testing. This means membrane

performance properties were not affected by the chemistry of the RO brine.

Increased electrical resistance of membranes was not detected (reduction in stack

current).

To ensure necessary flow through electrode streams in the EDR stack, continuous pH

adjustment was required to maintain cathode stream outlet pH at 7.0. This is

required to prevent silica and CaCO3 from coming out of solution, coating the cathode

and leading to flow blockage. (Antiscalant used in RO de-salter for control of silica is

deactivated at low pH). The reduced pH also helps controls calcium carbonate

scaling since the typical higher pH cathode environment is avoided.

In terms of hydraulic throughput, some particulate fouling of the EDR membrane

stack occurred from the concentrating side leading to an increase in stack pressure

drop during the first two phases of testing. These events were remedied by use of

cartridge filtration on both dilute feed and concentrate feed streams to EDR

membrane stack. Clean in Place (CIP) needs to be done every four weeks to prevent

scale from taking firm hold thus increasing difficulty to clean.

Determine if pretreatment is needed to run the EDR process without issue

No pre-treatment of raw water is necessary. Holdup of any RO brine should be for a

minimum amount of time to prevent silica from coming out of the solution in feed

tank.

In a larger scale plant, provision for flushing with product water should be included to

avoid scaling of the system during shutdowns.

Demonstrate precipitator’s ability to remove solids from the system


85

Illustrated in Figure 23 and Figure 45, although measured sludge solid percentages

varied during pilot, ability of precipitation process to remove solids from the

recirculating EDR concentrate stream was demonstrated. Some solids were also

removed from the precipitator effluent, (discussed previously) by cartridge filters or

candle filters.

Membrane filtration is not suitable for filtering of effluent from precipitator. With the

tendency for precipitator effluent to range from 130% – 150% calcium sulfate super-

saturation, the candle filters struggled to maintain adequate recovery of permeability

after backwashing. The supersaturated environment led to scaling of the membrane

that could not be removed by backwashing and could not be cleaned with cleaning

solutions. Much of the final period of operation used only cartridge filters, which also

experienced fouling by precipitation.

Use of a solid impermeable media such as discs or micro screens or microfibers

would allow physical displacement of solids by backwashing with a jet of water

(AMIAD) or physical separation of components (Arkal discs). However, based on pilot

learnings, less disturbance from recycle of backwash effluents will improve stability of

both hydraulic balance and stream compositions in the process. For this reason a

continuous flowing cyclonic process would fit best for primary removal of solids

carryover in precipitator effluent. Continuous flow with no moving parts with lowest

cost and minimal maintenance will maximize process performance. After the cyclonic

process, a break tank and a cartridge filter would be required to stabilize flow to the

EDR and prevent any residual fine particulate from entering the EDR stack.

Establish treatment process water recovery.

Minimum of 80% recovery was demonstrated during Phase 3. 80% recovery resulted

in the most stable operation of the pilot unit.

Establish EDR process criteria including velocity and amperage.

EDR process details are described for a one line AQ15 EDR 100 gpm EDR (production)

system in the projection below, including flow velocity and electrical characteristics.


86

Table 14: WATSYS Design For 100 GPM EDR System

Determine chemical dosages.

Maximum, minimum and average chemical dosing (50% sulfuric acid) observed in the

pilot for pH adjustment of the precipitator to approximately 6.3 is described below.

Production basis is 6.4 gpm at 80% recovery.

Table 15: Sulfuric Acid Dose Rate During Pilot Study

Chemical Dosing Pilot Dosing rate, liters/hr (mg/L) 50% Sulfuric acid

Pilot Maximum 0.655 (182 mg/L)

Pilot Average 0.402 (296 mg/L)

Pilot Minimum 0.211 (95 mg/L)

Conclusion based on pilot results and mass balance modeling

0.498 – 0.565 (225 mg/L – 255 mg/L)

Determine electric power requirements.

Project Name EMWD

Company

User

WATSYS Run Date Tuesday, September 15, 2015 Recycled Feed Product Conc. BD Waste Raw Feed

Calcium mg/l 1079.0 194.9 4615.6 4615.6 1079.00


EDR System AQ15-1-3-600 with 1 Line(s) 4 Stage(s) Sodium mg/l 774.0 276.1 2765.7 2765.7 774.00




Ammonia mg/l 0.0 0.0 0.0 0.0 0.00

Bicarbonate mg/l 809.0 400.2 1095.6 1095.6 809.00

















Last Stage Outlet Pressure 6.59 psig BaSO4 451.6 133.4 1235.4 1027.7 451.64

SrSO4 31.4 6.2 85.7 71.1 31.40








Flow Rate USGPM 125.0 100.0 25.0 25.0 125.0


Electrical Stages 1 2 3 4

Voltage (V) 412 410 407 399 Details of Acid Dosing to Concentrate Blowdown.

Current (Amps) 31.6 29.6 26.7 22.3 Acid Name 98% H2SO4

Surge (Amps) 45.6 51.5 59.8 70.4 Amount of Acid 325.86 lb/day 148.12 kg/day

Hyd Stages / Elect Stage 1 1 1 1 pH after Acid Dosing 6.28

The amount of acid used to reduce LI to 1.20 is as follows

Hydraulic Stages 1 2 3 4 68.50 USGPD of 36% HCl to reduce LI to 1.20

% Polarization 14.79 18.48 23.88 31.07 21.70 USGPD of 98% H2SO4 to reduce LI to 1.20

Cut Fraction 0.22 0.26 0.32 0.39 259.27 l/day of 36% HCl to reduce LI to 1.20

Current Efficiency 0.85 0.84 0.83 0.82 82.13 l/day of 98% H2SO4 to reduce LI to 1.20


Cell Pairs 600 600 600 600


87

Very stable operation and consistent desalting performance over the last 30 days of

Phase 3 resulted in 6.2 kwh/kgal net production.

Power consumption estimates for pumps and motors in the pilot are listed in Table 16

below along with DC power.

Table 16: Power Consumption of Pumps and Motors in Pilot

Item Flow Pressure Pump efficiency

Motor efficiency

VFD efficiency

kW kWh/kgal

gpm psig

EDR feed pump

8.9 65 0.7 0.95 0.97 0.39

EDR concentrate pump

6.4 65 0.7 0.95 0.97 0.28

Filter feed pump

7 25 0.7 0.95 0.97 0.12

Backwash pump

8 25 0.7 0.95 0.14

LC pump 1 10 0.6 0.95 0.008

Bleed pump 1 10 0.6 0.95 0.008

Slurry pump 3 5 0.6 0.95 0.011

HCl dosing pump

0.1

H2SO4 dosing pump

0.1

Mixer motor

AC total 1.16 3.0

DC total 2.33 6.2

Determine electrode cleaning chemical dosage.

Chemical dosing utilized in pilot test for electrode cleaning included continuous pH

adjustment and periodic Electrode Clean-in-Place (ECIP) process. Maximum

productivity for one line of EDR in scale up is 100 gpm of product water.

Table 17: Hydrochloric Acid Consumption for Pilot

Chemical Frequency Injection rate (per line of EDR)

Daily consumption (per line of EDR)

Liters/hour liters/day

32% HCl Continuous 0.3 7.2

32% HCl Periodic (3 x 60 min) per day

1.9

5.7

Total 32% HCl 12.9


88

Establish Clean In Place requirements

EDR CIP frequency of 2 - 4 weeks was demonstrated in pilot test. With optimization of

hydraulic balance, improved precipitator and effluent filtration design, CIP frequency

in larger scale is expected to be closer to 4 weeks. This CIP is standard 5% HCl acid

cleaning used in EDR equipment.

Learnings from the pilot will benefit larger scale operation for improving CIP

frequency. The learnings include the removal of periodic high recycle streams back to

precipitator feed and improved control of precipitator process operation, by

controlling pH and stability of sludge recycle and blowdown.

Further consideration of CIP requirements for CaSO4 precipitation system is

necessary. This includes a CIP process for cleaning piping from the EDR stack outlets,

through precipitator, sludge recycle system, precipitator effluent piping and solids

removal equipment, effluent collection tanks and EDR concentrate inlet pump.

Heating of cleaning solution (to 80°C) is important to ensure ability to remove CaSO4

scale from the brine precipitation system/loop as recommended CaSO4 cleaning

chemicals are most effective at 80°C. Frequency of such cleanings may be two times

per year.


89

Future Design Recommendations

From the pilot, there were many lessons learned to be implemented on the next level

design for EMWD. The pilot was an effective demonstration project, showing that the

technology could achieve the removals and recovery desired. However, there are many

potential improvements to the process that should be implemented in the next

demonstration level project.

For the full scale design of the system, the final desired flowrate is 1600 gpm. Since the

pilot was only ~7 gpm, it is a large scale up from the pilot to the full scale

design. Because of this, GE and EMWD have discussed an intermediate step between

these two system sizes. A 100-gpm demonstration unit would provide a good step in

validating the design with a full scale EDR stack, allowing further optimization of the

precipitation process. Many of these improvements have been discussed in this

report. They will be summarized here.

Pretreatment: Future design case will not include the acid reduction of the feed. The

system will continue to have a cartridge filter before the EDR system. HCl will be dosed

into the electrode streams of the EDR stack.

EDR Stack: The stack in the next level design will be the GE Standard 600 cell pair EDR

stack. This stack is the same size stack that would be used in a full scale design system.

Precipitator: The design flow of the precipitator will be slower than the current pilot

design. Pilot rise rate was 0.46 gpm/ft2, but in the next level of design, it should be

reduced to 0.4-0.45 gpm/ft2. Precipitator internal recirculation should be continuous and

controlled to maintain particle size adequate for good settling.

Precipitator Effluent Filtration: GE has been evaluating other options to the existing

setup. At the time this report was written, the evaluation is still underway. A Cyclonic

treatment process will be evaluated. A cartridge filter after the cyclonic process will be

required as a protection for the EDR stack in case of upset.

Cleaning: The system will be designed to clean not just the EDR stack, but the precipitator

piping as well. This change enables full cleaning of the system piping during CIP

procedure. The cleaning system will be capable of heating to 80°C to enhance removal of

calcium sulfate scale formed.

Flush: When the system shuts down, there will be ample supply of water to flush the EDR

stack so it does not sit with saturated brine on the membrane surface.


90

APPENDIX A.1: EMWD Phase 3 EDR Conductivity Charts


91

Figure A.47: Conductivity and Ca 6/19/2015 Figure A.48: Conductivity and Ca 6/22/2015



92




93




94




95

APPENDIX A.2: EMWD Phase 3 EDR Flow Charts


96

Figure A.63: Flow Snapshot 6/19/2015 Figure A.64: Flow Snapshot 6/22/2015



97




98




99




100

APPENDIX A.3: EMWD Phase 3 EDR Voltage and Current Charts


101

Figure A.79: Voltage and Amperage 6/19/2015 Figure A.80: Voltage and Amperage 6/22/2015



102




103




104



APPENDIX B

Economic Analysis: Preliminary Construction and O&M Cost For Full-Scale AquaSel Facility at the Perris and

Menifee Desalters

June, 2016

Carollo Engineers

i

EASTERN MUNICIPAL WATER DISTRICT

GE AQUASEL PILOT TESTING AND EVALUATION

ECONOMIC ANALYSIS

TABLE OF CONTENTS Page No.

INTRODUCTION ........................................................................................................................ 1 COST ESTIMATE ASSUMPTIONS ............................................................................................... 1 Operation and Maintenance Cost Assumptions .................................................................... 1 Capital Cost Assumptions ....................................................................................................... 2 PROJECT COST ESTIMATES ....................................................................................................... 3 Capital Cost Estimate ............................................................................................................. 3 Operating Cost Estimate ........................................................................................................ 5 IMPLEMENTATION ................................................................................................................... 5 Existing Desalter Streams and Flows ..................................................................................... 6 Alternative 1 - AquaSel Product Directly to Disinfection ....................................................... 7 Alternative 2 - AquaSel Product Into Raw Feed Water Stream ............................................. 8 OVERALL WATER COST ANALYSIS ............................................................................................ 9 CONCLUSIONS ........................................................................................................................ 11

LIST OF TABLES Table 1 Operation and Maintenance Cost Assumptions ................................................ 2 Table 2 Capital Cost Assumptions ................................................................................... 3 Table 3 Capital Cost Estimate - 2.4 mgd AquaSel System ............................................... 4 Table 4 Operation and Maintenance Cost Estimate - 2.4 mgd AquaSel System ............ 5 Table 5 Annual Cost of Treatment ................................................................................ 10

LIST OF FIGURES Figure 1 Schematic of Existing Desalter Arrangement ..................................................... 6 Figure 2 Schematic of Desalter Arrangement Including AquaSel System – Alt. 1 ........... 7 Figure 3 Schematic of Desalter Arrangement Including AquaSel System – Alt. 2 ........... 9

1

PRELIMINARY COST ESTIMATE

INTRODUCTION

The purpose of this report is to provide preliminary construction and operation and maintenance (O&M) cost estimates (order-of-magnitude) for a concentrate treatment facility designed to recover additional water from the Menifee and Perris I desalters. The full-scale facility would be designed to treat 2.4 mgd of concentrated brine and recover up to 80 percent of the influent flow. All of the brine from the desalters would be treated by the AquaSel unit, with reject flow being discharged to the Inland Empire Brine Line. The conceptual facility consists of the following major components.

• GE's AquaSel Process

• Solids Handling Facility

• Chemical Dosing Systems

This chapter also discusses how a full scale AquaSel system could be integrated into the existing desalter facilities to provide the most economical approach.

COST ESTIMATE ASSUMPTIONS

Several baseline cost assumptions were required to complete the cost estimate for the concentrate treatment system. These assumptions include O&M factors, such as the cost for power, chemicals, etc. and estimates of process capital cost based on past project and vendor quotes. At the time of this cost estimate (October 2015), the Los Angeles-based Engineering News-Record (ENR) construction cost index was 10981.

Operation and Maintenance Cost Assumptions

A list of the pertinent O&M cost assumptions used in the process cost estimate is presented in Table 1.

2

Table 1 Operation and Maintenance Cost Assumptions Technical Feasibility Report Eastern Municipal Water District

Parameter Value Units Estimated Consumption

Sulfuric Acid (50%) 3.3 $/gal 317 gal/d

Hydrochloric Acid (32%) 4.5 $/gal 55 gal/d (1)

Cartridge Filters 300 $/filter monthly changes

EDR Membrane Replacement 0.2 $/1,000 gal -

Electrical Power 0.13 $/kWh 9.2 kWh/kgal

Sludge Disposal 50 $/ton Based on biosolids

disposal cost

Notes: (1) Excludes acid used for clean-in-place process - about 4,000 gal/month

Capital Cost Assumptions

Several planning level cost assumptions were made based on both vendor quotes for equipment and established rule-of-thumb parameters for membrane system treatment costs. A summary of capital cost assumptions is shown in Table 2.

3

Table 2 Capital Cost Assumptions Technical Feasibility Report Eastern Municipal Water District

Parameter Value Units

Yard Piping(1) 3 %

Interconnecting Pipework(1) 14 %

Electrical and Instrumentation(1) 18 %

Building Costs 150 - 200 $/sq ft

Contingency 25 %

Contractor General Conditions 5 %

Contractor Overhead and Profit 10 %

Bid Market Allowance (2) 0 %

Sales Tax (3) 8 %

Engineering 10 %

Construction Management 10 %

Administration, Environmental and Permits 8 %

Notes: (1) As a percentage of direct equipment costs (2) Biding conditions were not taking into consideration. (3) Assumed that this applies to 50% of the equipment costs and 25% of the yard piping costs

PROJECT COST ESTIMATES

Using the assumptions and approach stated above, an overall project cost was developed.

Capital Cost Estimate

The planning level capital cost estimate for a treatment train to treat 2.4-mgd (1,667 gpm) of Menifee and Perris I desalter concentrate at a recovery of 80% is summarized in Table 3. The project cost estimate is $40.1 million. This estimate includes chemical systems, process equipment, solids handling, pumping, and other ancillary facilities. The treatment system would be made up of four 400 gpm modules; each module would include 16 electrodialysis reversal (EDR) stacks. The modules would be arranged within a building and would also include precipitator tanks, one or two per 400 gpm module.

4

Table 3 Capital Cost Estimate - 2.4 mgd AquaSel System Technical Feasibility Report Eastern Municipal Water District

Component Description Cost(2) (3)

Site Civil and Yard Work 2,095,459 AquaSel System (1) 8,966,000 Dewatering / Loading Facility 2,310,000 Chemicals, Product Storage and Transfer 1,050,231 Equipment Installation 4,034,700 Buildings 1,250,000 Electrical & Instrumentation and Controls 2,218,721

Contractor General Conditions, 5% 1,096,256 Contractor Overhead and Profit, 10% 2,192,511 Bid market Allowance, 0% 0 Sales Tax @ 8.0%(4) 666,312 Contingency, 25% 5,481,278

Construction Cost 31,361,467 Engineering, 10% 3,136,147 Construction Management, 10% 3,136,147 Administration, Environmental and Permits, 8% 2,508,917

Total Estimated Project Cost 40,142,678

Notes: (1) Estimate provided by GE (2) ENR CCI Index for Los Angeles (October 2015, 10,981). (3) Sub-total of direct costs which includes all equipment, buildings and installation is

$21,925,111 (4) Calculated assuming 50 percent of equipment costs and 25 percent of yard piping are

taxable. The cost estimate herein is based on our perception of current conditions at the project location. This estimate reflects our professional opinion of accurate costs at this time and is subject to change as the project design matures. Carollo Engineers, Inc. have no control over variances in the cost of labor, materials, equipment; nor services provided by others, contractor's means and methods of executing the work or of determining prices, competitive bidding or market conditions, practices, or bidding strategies. Carollo Engineers, Inc. cannot and does not warrant or guarantee that proposals, bids, or actual construction costs will not vary from the costs presented as shown.

5

Operating Cost Estimate

The planning level O&M costs for the treatment facilities are summarized in Table 4. The total annual O&M cost is estimated to be approximately $3.1 million. This amount includes electrical costs, chemical costs, membrane replacement, brine disposal, solids disposal and an allowance for providing one additional operator. No credit has been taken for the reduction in final brine volume for disposal.

Table 4 Operation and Maintenance Cost Estimate - 2.4 mgd AquaSel System Technical Feasibility Report Eastern Municipal Water District

Parameter Value Units Source

Equipment O&M Costs (1) 123,262 $/yr See note

Chemical Costs 256,686 $/yr Pilot testing

Solids Disposal Costs 415,917 $/yr $50/ton(2)

Brine Disposal Costs 493,880 $/yr Based on IEBL

EDR Membrane Replacement Costs 175,347 $/yr Estimate

Cartridge Filters 43,200 $/yr Pilot/Estimate

Power Costs 1,433,889 $/yr Pilot testing(3)

Labor and Staffing Costs Allowance (4) 115,000 $/yr Estimate

Total Cost 3,057,182 $/yr

Notes: (1) Assume 1% of the Equipment Cost; compared with about $60,000 for existing desalter

O&M Costs (2) Based on 17 wet t/d of solids produced (3) Based on 9.2 kWh/kgal estimate provided by GE (4) Allowance for one additional operator

IMPLEMENTATION

The product water from the AquaSel system (1,333 gpm or 2,154 AFY) would be returned to the desalters for production of potable water. There are two ways in which the AquaSel water could be utilized. In one case, Alternative 1, the water could be combined with the permeate flow from the Menifee and Perris I desalters, as mentioned above, then disinfected in the existing chlorine contact tank and discharged to the system as potable

6

water. In the other case, Alternative 2, the AquaSel product would be returned to the feed side of the desalters and used to increase the raw feed flowrate. This would provide additional feed water for desalting and would allow the desalters to be expanded by an amount equal to the AquaSel product stream. Each approach is discussed briefly below after a description of the existing system.

Existing Desalter Streams and Flows

Figure 1 presents a schematic of the arrangement of the existing desalters and how their treated and bypass flow is combined, prior to final disinfection and distribution. The Menifee desalter comprises two RO trains, each with a maximum permeate production capacity of 930 gpm. The Perris I desalter also has two trains, but these are larger, with a maximum permeate production capacity of 1,680 gpm each. As shown in Figure 1, in the case of the Menifee Desalter about 250 gpm of raw water bypasses the RO trains and combines with the RO permeate downstream.

A similar arrangement with bypass flow takes place with the Perris I desalter. The total bypass flow for the two desalters is around 950 gpm.

Figure 1: Schematic of Existing Desalter Arrangement

7

Alternative 1 - AquaSel Product Directly to Disinfection

For the first alternative, all of the brine from the desalters would be treated by the AquaSel unit. The proposed AquaSel product stream would be combined with the product streams from the desalters just upstream of the chlorine contact basin, and concentrate from the AquaSel process would be discharged to the Inland Empire Brine Line. This is illustrated in Figure 2.

Figure 2: Schematic of Desalter Arrangement Including AquaSel System - Alternative 1

Due to the quality of the AquaSel product water, which is expected to have a TDS concentration of around 2,000 mg/L, which is almost the same as the raw well water TDS, the bypass flow around the desalters would need to be turned off in order to maintain the finished water TDS concentration. With the AquaSel product flowrate around 1,333 gpm, and the bypass flow turned off completely, combining the AquaSel product flow with the

8

permeate from the desalters would result in an overall increase in the product water production of about 383 gpm (618 AFY), and an expected TDS in the combined finished potable water of around 750 mg/L.

This finished water TDS is significantly higher than the TDS of around 470 mg/L reported in the finished water between July and September, 2015. However, based on mass balance calculations, if the full 950 gpm bypass flow is used, the delivered water TDS in the existing configuration should be around 650 mg/L; not too different from the calculated value of 750 mg/L mentioned above.

Presumably, EMWD can adjust the amount of bypass flow to get the desired finished water quality. However, once the AquaSel system is in operation, the system would have to accept all the product water. A full mass balance blend of the finished water still needs to be done to check the concentration of other constituents such as calcium, alkalinity, hardness etc. to confirm that the finished water stability is acceptable. This will require additional analytical tests on the desalter permeate streams to obtain the required data to perform such a balance.

Alternative 2 - AquaSel Product Into Raw Feed Water Stream

The other alternative for the AquaSel product water would be to blend it with the raw well water being pumped to the desalters. This approach would allow either:

• one or more wells to be taken off line to reduce the raw well production rate by the AquaSel product water flow rate, and maintain the current hydraulics through the desalters, or

• increase the production of the desalters by an estimated amount of 1,280 gpm (2,060 AFY), including about 280 gpm of bypass water for blending.

For the purposes of this evaluation, the second approach was selected, as shown in Figure 3 below. A new 1.4 mgd (1,000 gpm) RO train would be installed as part of the Perris I desalter. This unit would allow more bypass water to be used. The total brine produced by the RO trains would be collected and transferred to the AquaSel unit for treatment; concentrate from the AquaSel unit would be discharged to the Inland Empire Brine Line. Because there would be additional RO production, the flow of brine to the AquaSel unit could be greater than the original 1,600 gpm which would mean a slightly larger AquaSel unit may be required. However, based on current production rates of the desalters, the total brine produced by all RO units would be close to the rated capacity of the AquaSel system. A third Alternative was considered in which the AquaSel product water would be fed directly to a separate RO train. This would have worked well except that the silica

9

concentration in the AquaSel product is at a concentration of 130 mg/L, too high to be fed directly to an RO train.

Figure 3: Schematic of Desalter Arrangement Including AquaSel System - Alternative 2

OVERALL WATER COST ANALYSIS

Alternatives 1 and 2, presented above were compared on a cost basis to identify the most economically attractive alternative for implementing a full-scale AquaSel system for EMWD.

Detailed cost estimates for Alternative 1 were presented earlier in Tables 3 and 4. These costs are summarized in Table 5, together with estimates for Alternative 2. Annualized capital and operating costs are presented. The difference in the capital cost estimate between Alternatives 1 and 2 is a result of the 1.4-mgd RO train that is included in Alternative 2. The O&M cost estimate for Alternative 2 is also slightly higher than that for Alternative 1, due to the additional RO train and additional chemicals for disinfection of the higher product flow. The capital costs for both Alternatives were amortized using a 5% annual interest rate and a loan period of 30-years.

10

Also included in the table are total water production estimates for both Alternatives, as well as the benefit cost associated with the production of additional "new" water compared with the existing desalter production of approximately 9,600 AFY. The benefit was calculated using a selling price of $1,552/AF as well as accounting for savings for water which would no longer need to be purchased from MWD, at a rate of $942/AF.

As indicated in Table 5 the annual cost of Alternative 2 is about $400,000 higher than Alternative 1. However, Alternative 2 would result in 1,318 AFY more water production and therefore significantly higher savings, totaling more than $5 million per year.

Table 5 Annual Cost of Treatment Technical Feasibility Report Eastern Municipal Water District

Parameter Alternative 1(1) Alternative 2(2)

Capital Cost Estimate, $ million 40.1 41.9

Annual Amortized Capital Cost (3) $ 2,611,339 $ 2,728,093

Annual Operating Cost (4) $ 3,057,182 3,329,733

Total Estimated Annual Cost $ 5,668,521 6,057,826

Total Water Production, AFY 10,216 11,534

New Water Production, AFY(5) 618 1,936

Benefit from Sale of new water (6) $ 1,542,194 $ 4,827,737

Benefit from Less Brine for Disposal $ 486,463 $ 486,463

Total Estimated Annual Benefit $ 2,028,657 $ 5,314,200

Total Cost Minus Total Benefit $ 3,639,864 $ 743,627

Notes: (1) Alternative 1 includes returning the AquaSel product stream to blend with the permeate from

the existing desalters. The current practice of blending the RO permeate with bypass water would stop.

(2) Alternative 2 includes returning the AquaSel product to the feed side of the existing desalters and mixing it with the incoming raw well water. A new 1.4-mgd RO train would be provided to increase overall production and bypass blending would continue.

(3) Based on a 5% interest rate and a 30-year loan period. (4) Includes electrical power, chemicals, solids disposal, brine disposal, membrane replacement,

equipment O&M, membrane and cartridge filter replacement, and additional labor costs. (5) New water over and above the existing system capacity of 9,598 AFY (6) This is based on selling water at $1,552/AF and the added benefit of avoiding imported water

costs of $942/AF

11

Despite the significant savings associated with Alternative 2, the project costs are still higher than the anticipated savings, by about $743,000 per year. However, if the interest rate on the capital sum could be reduced to around 2.5% by, for example, having a portion of the loan paid for using grant money, the costs for Alternative 2 would equal the savings and the project would therefore pay for itself. The cost difference for Alternative 1 ($3.6 M) is too high to enable that project to pay for itself.

CONCLUSIONS

This document presents capital and operating cost estimates for a 2.4-mgd AquaSel system, large enough to treat the RO concentrate from the existing Menifee and Perris I desalters. An estimate of the preliminary cost of the AquaSel system and how it could be integrated into the existing Menifee and Perris I desalters has shown the following:

1. The construction cost for an AquaSel system to treat 1,667 gpm of brine from the

Menifee and Perris I desalters is estimated to be $31.4 million. The total project cost is expected to be around $40 million.

2. Annual operating costs are estimated to total about $3.0 million, with over 55% of that cost associated with electrical power costs and chemicals. The annual operating costs include almost $500,000 for disposal of the final concentrate stream produced by the AquaSel system. Before disposal, this stream would be diluted with excess recycled water to prevent scale formation in EMWD's brine pipeline. Another $416,000 would be required for the annual disposal of solids produced by the system.

3. The AquaSel product water could be either combined with the desalter permeate streams to replace the current use of bypass flow (Alternative 1), or returned to the raw well water feed stream to increase the available flow to the desalters (Alternative 2). A comparison was made between these two alternatives. Alternative 2 would require the addition of a 1.4-mgd RO train to treat the additional flow. This would increase both the capital and operating costs of this Alternative, but it would also increase the drinking water production by over 1,300 AFY more than Alternative 1. A cost analysis showed that despite the higher capital and operating costs of Alternative 2, it would be the most economically attractive project.

4. Neither Alternative 1 nor Alternative 2 would pay for themselves based on the assumptions used in the analysis. However, if the average annual interest rate for the capital sum of Alternative 2 could be reduced from 5% over 30-years to 2.5%, the costs and benefits would balance. Even with a 0% interest rate, Alternative 1 would not pay for itself, because its "new" water production is relatively small at only 618 AFY.

12

5. A full mass balance blend calculation of the AquaSel product water with the raw well water still needs to be done to check the concentrations of constituents such as calcium, alkalinity, hardness, silica etc. that would result, to confirm no impact on the existing desalters current production and recovery rates This will require additional analytical tests on the desalter feed and permeate streams to obtain the required data to perform such a calculation.

6. If EMWD is considering implementation of the AquaSel technology then the best approach would be to use Alternative 2. In addition, further testing of a larger scale AquaSel system (around 400 gpm) should be undertaken to confirm that consistent water quality can be produced, and to gain additional operational experience on a larger scale system, and how to integrate it into the existing facilities.

andmwdh2o.com/faf pdfs/2_gw_emwd mwd final report pilot scale brine...removed from the desalting...

Documents